Surface Process, Transportation, and Storage 0128238917, 9780128238912

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Surface Process, Transportation, and Storage

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Oil and Gas Chemistry Management Series

Surface Process, Transportation, and Storage Volume IV

Edited by

Qiwei Wang Saudi Aramco, Dhahran, Saudi Arabia

Gulf Professional Publishing is an imprint of Elsevier 50 Hampshire Street, 5th Floor, Cambridge, MA 02139, United States The Boulevard, Langford Lane, Kidlington, Oxford, OX5 1GB, United Kingdom Copyright © 2023 Elsevier Inc. All rights reserved. No part of this publication may be reproduced or transmitted in any form or by any means, electronic or mechanical, including photocopying, recording, or any information storage and retrieval system, without permission in writing from the publisher. Details on how to seek permission, further information about the Publisher’s permissions policies and our arrangements with organizations such as the Copyright Clearance Center and the Copyright Licensing Agency, can be found at our website: www.elsevier.com/permissions. This book and the individual contributions contained in it are protected under copyright by the Publisher (other than as may be noted herein). MATLABs is a trademark of The MathWorks, Inc. and is used with permission. The MathWorks does not warrant the accuracy of the text or exercises in this book. This book’s use or discussion of MATLABs software or related products does not constitute endorsement or sponsorship by The MathWorks of a particular pedagogical approach or particular use of the MATLABs software. Notices Knowledge and best practice in this field are constantly changing. As new research and experience broaden our understanding, changes in research methods, professional practices, or medical treatment may become necessary. Practitioners and researchers must always rely on their own experience and knowledge in evaluating and using any information, methods, compounds, or experiments described herein. In using such information or methods they should be mindful of their own safety and the safety of others, including parties for whom they have a professional responsibility. To the fullest extent of the law, neither the Publisher nor the authors, contributors, or editors, assume any liability for any injury and/or damage to persons or property as a matter of products liability, negligence or otherwise, or from any use or operation of any methods, products, instructions, or ideas contained in the material herein. ISBN: 978-0-12-823891-2 For Information on all Gulf Professional Publishing publications visit our website at https://www.elsevier.com/books-and-journals

Publisher: Charlotte Cockle Senior Acquisitions Editor: Katie Hammon Editorial Project Manager: Aleksandra Packowska Production Project Manager: Prasanna Kalyanaraman Cover Designer: Miles Hitchen Typeset by MPS Limited, Chennai, India

Contents List of contributors ................................................................................................ xiii

CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans.................................................................... 1 1.1 1.2 1.3 1.4

1.5

1.6

1.7

Sunder Ramachandran Introduction ....................................................................................2 Hydrogen sulfide and mercaptan measurement ............................4 Hydrogen sulfide and mercaptan partitioning in oil, water, and gas............................................................................................5 Chemical scavengers ......................................................................5 1.4.1 Solid scavengers.................................................................. 6 1.4.2 Oxidizing chemicals............................................................ 7 1.4.3 Aldehydes............................................................................ 7 1.4.4 Formaldehyde reaction products......................................... 8 1.4.5 High valence metal compounds........................................ 10 1.4.6 Aqueous alkaline solutions ............................................... 10 1.4.7 Hydrogen fluoride ............................................................. 11 1.4.8 Novel hydrogen sulfide scavengers from biological sources ............................................................................... 11 Physical chemistry of scavengers ................................................11 1.5.1 Scavenging kinetics .......................................................... 12 1.5.2 Scavenger thermodynamics .............................................. 12 Laboratory testing of hydrogen sulfide and mercaptan scavengers.....................................................................................13 1.6.1 ASTM D5705 test methodology with modifications..................................................................... 13 1.6.2 Laboratory assessment of hydrogen sulfide and mercaptan scavengers in towers ....................................... 13 1.6.3 Continuous gas flow apparatus......................................... 14 Hydrogen sulfide and mercaptan scavenging process.................15 1.7.1 In-line injection ................................................................. 15 1.7.2 Gas lift injection................................................................ 16 1.7.3 Capillary injection............................................................. 17 1.7.4 Contact towers................................................................... 17 1.7.5 Storage tanks ..................................................................... 17 1.7.6 Rail cars............................................................................. 19 1.7.7 Scavengers in acidizing treatment .................................... 19

v

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1.8 Case studies ..................................................................................20 1.8.1 Optimization of South Texas system.............................. 20 1.8.2 Optimization of scavenging cost from joint industry program ........................................................................... 20 1.8.3 Scavenger automation at sour gas processing facility ... 20 1.8.4 Development of sour reservoir in giant field in Kuwait ............................................................................. 21 1.8.5 Development of sour gas field in the Netherlands using scavenger with scale inhibitor .............................. 21 1.8.6 Scavenging dry gas pipeline in Western Oklahoma ...... 21 1.8.7 Scavenging in coiled tubing drilling operations in Saudi Arabia.................................................................... 22 1.8.8 Reduction of sulfur oxide content of flare gas............... 22 1.8.9 Capillary string downhole injection in South Texas...... 22 1.8.10 Fixed bed hydrogen sulfide removal in the North Sea ...... 23 1.9 Challenges associated with scavenging treatment.......................23 1.9.1 Reaction products.............................................................. 23 1.9.2 Induced scaling problems ................................................. 24 1.9.3 Corrosion issues ................................................................ 24 1.9.4 Formation damage ............................................................ 25 1.9.5 Emulsion problems in oil and water separation............... 25 1.9.6 Overconsumption of scavenger ........................................ 25 1.10 New developments .......................................................................26 1.10.1 Safe operation ................................................................. 26 1.10.2 Digital transformation ..................................................... 26 1.10.3 Environmentally friendly products ................................. 26 1.11 Summary and conclusions ...........................................................27 Nomenclature............................................................................... 27 References.................................................................................... 28

CHAPTER 2 Natural gas sweetening .............................................. 37 Sebastien Duval Introduction ..................................................................................38 Gas conditioning to satisfy sales gas quality...............................39 Natural gas sweetening methods..................................................42 Chemical absorption.....................................................................43 2.4.1 Chemical reactions between H2S and CO2 and amine .... 44 2.4.2 Amine process overview................................................... 47 2.4.3 Design best practices ........................................................ 49 2.5 Physical absorption ......................................................................51 2.5.1 Propylene carbonate process............................................. 52

2.1 2.2 2.3 2.4

Contents

2.6 2.7

2.8

2.9

2.10

2.11

2.5.2 Dimethyl ether of polyethylene glycol (DEPG or DMEPEG) solvents ......................................... 53 2.5.3 N-Methyl-2-pyrrolidone .................................................... 53 2.5.4 Refrigerated methyl alcohol (methanol)........................... 53 2.5.5 Combined physical and chemical absorption................... 54 Adsorption ....................................................................................54 Permeation or membrane based technologies .............................55 2.7.1 Principle ............................................................................ 57 2.7.2 Polymeric membrane type ................................................ 58 2.7.3 Membrane module types................................................... 59 2.7.4 Gas pretreatment ............................................................... 59 Sulfur recovery .............................................................................60 2.8.1 Thermal section................................................................. 61 2.8.2 Catalytic section................................................................ 62 2.8.3 Major equipment ............................................................... 63 2.8.4 Quality of the acid gas...................................................... 64 2.8.5 Reduction absorption tail gas treatment ........................... 64 Emerging approaches for treating highly sour gas......................65 2.9.1 Cryogenic distillation........................................................ 66 2.9.2 Membranes for high H2S .................................................. 66 CO2 capture technology at gas plant ...........................................67 2.10.1 CO2 capture from flue gas.............................................. 69 2.10.2 CO2 captured from the acid gas stream ......................... 70 Final remarks................................................................................76 Nomenclature............................................................................... 76 References.................................................................................... 76

CHAPTER 3 Emulsion separation ................................................... 79 Thomas Krebs and Mohamed Reda Akdim Introduction ..................................................................................80 Emulsion formation......................................................................82 Emulsion stabilization ..................................................................88 Theory of emulsion separation ....................................................93 3.4.1 Settling velocity of droplets.............................................. 95 3.4.2 Coalescence rates ............................................................ 100 3.4.3 Semi-empirical approaches ............................................. 105 3.5 Emulsion separation techniques.................................................107 3.6 Thermal demulsification ............................................................110 3.6.1 Effect of heating on emulsion properties ....................... 110 3.6.2 Heater technology ........................................................... 112 3.6.3 Case studies..................................................................... 113 3.1 3.2 3.3 3.4

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3.7 Mechanical internals ..................................................................114 3.7.1 Separator vessels ............................................................. 114 3.7.2 Perforated baffles ............................................................ 115 3.7.3 Plate packs....................................................................... 116 3.7.4 Pipe separators ................................................................ 119 3.7.5 Case studies..................................................................... 120 3.8 Chemical demulsification...........................................................121 3.8.1 Effect of demulsifier on separation rates ....................... 122 3.8.2 Mechanisms of demulsifier action.................................. 123 3.8.3 Demulsifier formulation.................................................. 125 3.8.4 Case studies..................................................................... 128 3.9 Electrostatic demulsification......................................................129 3.9.1 Droplet migration in electric fields ................................ 130 3.9.2 Droplet collisions in electric fields................................. 131 3.9.3 Effect of electric field properties on droplet coalescence...................................................................... 134 3.9.4 Electrocoalescer technology ........................................... 136 3.9.5 Case studies..................................................................... 138 3.10 Concluding remarks ...................................................................139 Nomenclature............................................................................. 139 References.................................................................................. 140

CHAPTER 4 Foam control ............................................................. 153 Kalman Koczo, Mark D. Leatherman and Jonathan J. Wylde 4.1 Introduction and overview .........................................................154 4.1.1 Foam basics..................................................................... 155 4.1.2 Oil-based versus water-based foams .............................. 159 4.1.3 Antifoaming versus defoaming....................................... 159 4.1.4 Antifoaming versus deaeration ....................................... 160 4.1.5 Solid-stabilized foams..................................................... 161 4.1.6 Overview of foam stabilizer and antifoam chemistries...... 163 4.2 Oil-based foams..........................................................................165 4.2.1 Defoaming versus demulsification ................................. 165 4.2.2 Nonaqueous foaming ...................................................... 165 4.2.3 Nonaqueous foams of crude oil...................................... 168 4.2.4 Chemistry of antifoams for oil-based foams .................. 175 4.3 Water-based foams .....................................................................185 4.3.1 Chemistry of antifoams for water-based foams ............. 185 4.3.2 Water-based applications ................................................ 189 4.4 Mechanical defoaming ...............................................................198 4.5 Defoaming by chemical reaction ...............................................199

Contents

4.6 Mechanisms of antifoaming action............................................199 4.6.1 Antifoaming of nonaqueous foams................................. 200 4.6.2 Antifoaming of aqueous foams....................................... 204 4.6.3 Breaking solid stabilized foams...................................... 211 4.7 Concluding remarks ...................................................................212 Nomenclature............................................................................. 213 References.................................................................................. 213

CHAPTER 5 Polymeric drag reduction in pipelines .................... 227 Yung N. Lee and Ray L. Johnston 5.1 Drag-reducing agent history.......................................................228 5.2 Basic pipeline hydraulics tutorial ..............................................229 5.2.1 Reynolds number .......................................................... 230 5.2.2 Laminar flow................................................................. 231 5.2.3 Turbulent flow............................................................... 231 5.2.4 Pressure drop................................................................. 231 5.2.5 Static head ..................................................................... 232 5.2.6 Friction pressure............................................................ 233 5.2.7 Gradient......................................................................... 235 5.2.8 Profile ............................................................................ 236 5.2.9 Pipeline pumps.............................................................. 237 5.2.10 Operating point ............................................................. 237 5.2.11 Calculating drag reduction performance in a pipeline system.............................................................. 238 5.3 Drag-reducing agent chemistry..................................................240 5.4 Drag reduction mechanism ........................................................240 5.4.1 Misconceptions................................................................ 242 5.5 Application to the pipeline—drag-reducing agent theory.........243 5.5.1 Applications in oil/water or multiphase pipelines.......... 247 5.6 Utilization of drag-reducing agent in pipeline operations ........249 5.6.1 Example cases for utilization in pipelines...................... 251 5.7 Conclusion ..................................................................................257 Nomenclature............................................................................. 257 References.................................................................................. 258

CHAPTER 6 Natural gas storage by adsorption........................... 261 Yuguo Wang and Rashid Othman 6.1 Introduction ................................................................................262 6.2 Fundamentals of adsorption .......................................................265 6.2.1 Definition ........................................................................ 265 6.2.2 Adsorption forces ............................................................ 265 6.2.3 Adsorption separation and storage mechanism .............. 266

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6.2.4 Adsorption processes ...................................................... 267 6.3 Industrial adsorbents ..................................................................268 6.3.1 Adsorbent selection......................................................... 268 6.3.2 Silica gel.......................................................................... 268 6.3.3 Activated alumina ........................................................... 269 6.3.4 Zeolites ............................................................................ 270 6.3.5 Activated carbons............................................................ 272 6.3.6 Potential novel industrial adsorbents.............................. 275 6.3.7 Summary of natural gas storage adsorbents................... 275 6.4 Case study: screening activated carbon for natural gas storage ...... 276 6.4.1 Experimental ................................................................... 276 6.4.2 Method of determining the amount of methane adsorbed ... 278 6.4.3 Experimental results........................................................ 280 6.4.4 Empirical modeling with adsorption potential theory.... 282 6.4.5 Isosteric heat of adsorption modeling ............................ 284 6.5 Heat management modeling ......................................................286 6.5.1 Mathematical modeling .................................................. 286 6.5.2 Performance analysis through thermal simulation ......... 288 6.6 Summary.....................................................................................294 Nomenclature............................................................................. 294 References.................................................................................. 295

CHAPTER 7 Crude oil storage ...................................................... 299 7.1 7.2

7.3

7.4

James Alexander McRae, Kennith Van Ness and Chandrashekhar Khandekar Introduction ................................................................................299 Types of storage .........................................................................300 7.2.1 Storage tank..................................................................... 300 7.2.2 Concrete gravity-based structures................................... 301 7.2.3 Floating tanks.................................................................. 301 7.2.4 Underground caverns ...................................................... 302 Chemistry-related issues and solutions......................................302 7.3.1 Corrosion ......................................................................... 303 7.3.2 Bacteria ........................................................................... 308 7.3.3 Emulsion.......................................................................... 312 7.3.4 Carboxylate soaps ........................................................... 314 7.3.5 Paraffin and asphaltene................................................... 315 7.3.6 Inorganic solids ............................................................... 319 Summary.....................................................................................321 Nomenclature............................................................................. 321 References.................................................................................. 322

Contents

CHAPTER 8 Geologic carbon storage: key components ............. 325 8.1 8.2

8.3 8.4

8.5

8.6 8.7 8.8 8.9 8.10

Hakan Alkan, Oleksandr Burachok and Patrick Kowollik Introduction ................................................................................326 Geologic carbon storage classifications, definitions, types.......330 8.2.1 Definitions....................................................................... 331 8.2.2 Geologic carbon storage types........................................ 333 Key components of geologic carbon storage projects...............333 Surface components: capture, conditioning, and transport .......335 8.4.1 Capture ............................................................................ 335 8.4.2 Conditioning.................................................................... 342 8.4.3 Transport ......................................................................... 347 Subsurface components: exploration and reservoir...................353 8.5.1 Exploration and screening .............................................. 353 8.5.2 Storage capacity .............................................................. 356 8.5.3 Injectivity ........................................................................ 373 8.5.4 Containment .................................................................... 378 Risk assessment, monitoring, and validation.............................383 Monitoring and validation..........................................................389 Regulations and certification .....................................................391 Economics ..................................................................................396 Outlook .......................................................................................401 Nomenclature............................................................................. 405 References.................................................................................. 408

CHAPTER 9 Carbonate geochemistry and its role in geologic carbon storage .......................................... 423 Sylvain Delerce, Chiara Marieni and Eric H. Oelkers 9.1 Introduction ................................................................................424 9.2 Review of subsurface carbon dioxide trapping mechanisms ....425 9.3 Thermodynamic considerations .................................................429 9.3.1 The carbon dioxide system ............................................. 429 9.3.2 The CO2-H2O-NaCl system............................................ 431 9.3.3 The CO2-H2O-MeCO3 system........................................ 435 9.4 Kinetic considerations ................................................................436 9.4.1 Rates of CO2 dissolution into water ............................... 436 9.4.2 Rates of mineral dissolution reactions with CO2-charged water.......................................................... 437 9.4.3 Mineral precipitation rates .............................................. 437 9.5 Kinetic modeling ........................................................................440 9.6 Case studies ................................................................................445

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9.6.1 The Weyburn EOR and CO2 storage project ................. 445 9.6.2 The fate of the injected CO2 at Weyburn ...................... 447 9.6.3 The CarbFix mineral storage project.............................. 450 9.7 Carbonate chemistry and wellbore integrity..............................453 9.7.1 Carbonation and cement alteration ................................. 455 9.7.2 Effects of the fluid composition on cement alteration... 458 9.7.3 Impacts of the cement composition................................ 460 9.7.4 Calcite precipitation and self-sealing effects ................. 461 9.7.5 Corrosion and cement degradation of the Weyburn wells ................................................................ 462 9.8 Conclusions ................................................................................464 Nomenclature............................................................................. 465 References.................................................................................. 466

CHAPTER 10 Carbon conversion: opportunities in chemical productions ............................................................... 479

10.1

10.2

10.3 10.4

10.5

Peter Richard Ellis, Martin John Hayes, Norman Macleod, Stephen J. Schuyten, Cathy L. Tway and Christopher Mark Zalitis Introduction ................................................................................480 10.1.1 The carbon dioxide question......................................... 480 10.1.2 Carbon dioxide thermodynamics, reactivity, and catalysis ......................................................................... 480 10.1.3 Current and potential uses of carbon dioxide .............. 481 10.1.4 Routes for carbon dioxide oxidation state reduction ... 482 Supporting concerns for technology selection...........................484 10.2.1 Sources and costs of carbon dioxide ............................ 484 10.2.2 Purification of carbon dioxide ...................................... 486 Examples of commercialized technologies addressing the criteria and concerns ............................................................487 Promising technology areas for the future ................................489 10.4.1 Examples from chemical transformations .................... 489 10.4.2 The use of electrochemical systems for carbon dioxide conversion ............................................ 499 10.4.3 Continuous bio-catalysis for carbon dioxide conversion ..................................................................... 505 Final remarks..............................................................................511 Acknowledgments ..................................................................... 513 Nomenclature............................................................................. 513 References.................................................................................. 515

Index ......................................................................................................................525

List of contributors Mohamed Reda Akdim TechnipFMC, Arnhem, The Netherlands Hakan Alkan TU Bergakademie Freiberg, Freiberg, Saxony, Germany Oleksandr Burachok Wintershall Dea AG, Kassel, Hessen, Germany Sylvain Delerce Ge´osciences Environnement Toulouse, CNRS – CNES – IRD – OMP – Universite´ de Toulouse, Toulouse, France Sebastien Duval Research & Development Center, Saudi Aramco, Dhahran, Saudi Arabia Peter Richard Ellis Johnson Matthey, Reading, United Kingdom Martin John Hayes Johnson Matthey, Cambridge, United Kingdom Ray L. Johnston Liquid Power Specialty Products, Inc., Ponca City, OK, United States Chandrashekhar Khandekar Schlumberger, Houston, TX, United States Kalman Koczo Momentive Performance Materials, Tarrytown, NY, United States Patrick Kowollik Wintershall Dea AG, Kassel, Hessen, Germany Thomas Krebs TechnipFMC, Arnhem, The Netherlands Mark D. Leatherman Momentive Performance Materials, Tarrytown, NY, United States Yung N. Lee Liquid Power Specialty Products, Inc., Houston, TX, United States Norman Macleod Johnson Matthey, Billingham, United Kingdom Chiara Marieni Ge´osciences Environnement Toulouse, CNRS – CNES – IRD – OMP – Universite´ de Toulouse, Toulouse, France James Alexander McRae Schlumberger, Houston, TX, United States

xiii

xiv

List of contributors

Eric H. Oelkers Ge´osciences Environnement Toulouse, CNRS – CNES – IRD – OMP – Universite´ de Toulouse, Toulouse, France Rashid Othman Research & Development Center, Saudi Aramco, Dhahran, Saudi Arabia Sunder Ramachandran Baker Hughes, Sugar Land, TX, United States Stephen J. Schuyten Johnson Matthey, Savannah, GA, United States Cathy L. Tway Johnson Matthey, Oakbrook Terrace, IL, United States Kennith Van Ness Schlumberger, Houston, TX, United States Yuguo Wang Research & Development Center, Saudi Aramco, Dhahran, Saudi Arabia Jonathan J. Wylde Heriot-Watt University, Edinburgh, Scotland, United Kingdom; Clariant Oil Services, Clariant Corporation, Houston, TX, United States Christopher Mark Zalitis Johnson Matthey, Reading, United Kingdom

CHAPTER

1

Chemical scavenging of hydrogen sulfide and mercaptans

Sunder Ramachandran Baker Hughes, Sugar Land, TX, United States

Chapter Outline 1.1 1.2 1.3 1.4

Introduction ................................................................................................... 2 Hydrogen sulfide and mercaptan measurement ................................................. 4 Hydrogen sulfide and mercaptan partitioning in oil, water, and gas ................... 5 Chemical scavengers ...................................................................................... 5 1.4.1 Solid scavengers ...........................................................................6 1.4.2 Oxidizing chemicals ......................................................................7 1.4.3 Aldehydes .....................................................................................7 1.4.4 Formaldehyde reaction products .....................................................8 1.4.5 High valence metal compounds ....................................................10 1.4.6 Aqueous alkaline solutions ...........................................................10 1.4.7 Hydrogen fluoride ........................................................................11 1.4.8 Novel hydrogen sulfide scavengers from biological sources .............11 1.5 Physical chemistry of scavengers .................................................................. 11 1.5.1 Scavenging kinetics .....................................................................12 1.5.2 Scavenger thermodynamics ..........................................................12 1.6 Laboratory testing of hydrogen sulfide and mercaptan scavengers ................... 13 1.6.1 ASTM D5705 test methodology with modifications ........................13 1.6.2 Laboratory assessment of hydrogen sulfide and mercaptan scavengers in towers ....................................................................13 1.6.3 Continuous gas flow apparatus .....................................................14 1.7 Hydrogen sulfide and mercaptan scavenging process ..................................... 15 1.7.1 In-line injection ..........................................................................15 1.7.2 Gas lift injection .........................................................................16 1.7.3 Capillary injection .......................................................................17 1.7.4 Contact towers ............................................................................17 1.7.5 Storage tanks ..............................................................................17 1.7.6 Rail cars .....................................................................................19 1.7.7 Scavengers in acidizing treatment ................................................19

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00004-1 © 2023 Elsevier Inc. All rights reserved.

1

2

CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

1.8 Case studies ................................................................................................ 20 1.8.1 Optimization of South Texas system ............................................20 1.8.2 Optimization of scavenging cost from joint industry program .........20 1.8.3 Scavenger automation at sour gas processing facility ....................20 1.8.4 Development of sour reservoir in giant field in Kuwait ..................21 1.8.5 Development of sour gas field in the Netherlands using scavenger with scale inhibitor ....................................................................21 1.8.6 Scavenging dry gas pipeline in Western Oklahoma .......................21 1.8.7 Scavenging in coiled tubing drilling operations in Saudi Arabia .....22 1.8.8 Reduction of sulfur oxide content of flare gas ..............................22 1.8.9 Capillary string downhole injection in South Texas .......................22 1.8.10 Fixed bed hydrogen sulfide removal in the North Sea ...................23 1.9 Challenges associated with scavenging treatment .......................................... 23 1.9.1 Reaction products .......................................................................23 1.9.2 Induced scaling problems ............................................................24 1.9.3 Corrosion issues ..........................................................................24 1.9.4 Formation damage .......................................................................25 1.9.5 Emulsion problems in oil and water separation ..............................25 1.9.6 Overconsumption of scavenger .....................................................25 1.10 New developments ....................................................................................... 26 1.10.1 Safe operation ..........................................................................26 1.10.2 Digital transformation ................................................................26 1.10.3 Environmentally friendly products ...............................................26 1.11 Summary and conclusions ............................................................................. 27 Nomenclature ........................................................................................................ 27 References ............................................................................................................ 28

1.1 Introduction Hydrogen sulfide (H2S) and mercaptans are toxic and corrosive substances that are naturally prevalent in natural gas and crude oil. The substances occur naturally and are abundant in natural gas and oil. The substances are toxic and hazardous. They are corrosive to many materials. It is important to remove these substances and not introduce them to the environment. H2S is formed in many regions of the world [1,2]. H2S in reservoirs originates from bacterial sulfate reduction (BSR), thermochemical sulfate reduction (TSR), thermochemical cracking of sulfur-rich oil, and fluid migration from other reservoirs [3]. BSR occurs in low-temperature reservoir environments with temperatures ranging between 0
C and 80
C while TSR occurs at higher temperature reservoir environments between 80
C and 200
C [4]. BSR can introduce H2S in reservoirs that previously did not have H2S. It has done so in some North Sea reservoirs [5,6]. Sulfur is among the thirteen most abundant elements on the crust of the earth [5].

1.1 Introduction

Mercaptans are organic thiol compounds. Some common mercaptans are methyl mercaptan (CH3SH) and ethyl mercaptan (C2H5SH). Mercaptans are found in crude oil in Arab crude [7], Russian crudes [8], North Sea crudes [9], and Canadian condensates [10]. They are found in sour gas in Southeast Asia [11] and the Middle East [12]. The Claus process creates mercaptans causing a reduction in sulfur recovery efficiency [13]. H2S and mercaptan removal are an important part of total sulfur emission reduction [12]. H2S and methyl mercaptan are hazardous at low concentrations and have a nuisance odor at extremely low concentrations [14]. The Threshold Limit Value (TLV) concentrations for hazardous effects and odor detection limit concentrations are, respectively, 10 ppm and 4.7 ppb for H2S and 0.5 ppm and 2.1 ppb for methyl mercaptan. The low concentrations at which hazardous and nuisance effects occur means that it is important to eliminate even the smallest trace of H2S or mercaptans from any escaping gas. Higher pressures and concentrations of these gases create large unsafe areas in cases of leakage. The effect of H2S partial pressure can cause large changes in the length of the fatal concentration zone [15]. The impact of changes in H2S partial pressure on the length of the fatal H2S concentration zone in case of leakage is shown in Table 1.1. For large H2S partial pressures, one has to be able to design systems extremely carefully. Ethyl mercaptan can cause potential underground water contamination [16]. Mercaptans and H2S can also cause severe corrosion problems. Mercaptans in crude oil contribute to the total corrosion of carbon steel in the 235
C300
C temperature range [9]. Vehicle fuel level sensors contain electrical contacts that are made of silver alloys that can corrode when the gasoline contains trace values of mercaptans [17]. H2S partial pressures cause material failures of materials due to sulfide cracking corrosion (SSC) [18]. When the partial pressure of H2S is below 0.0035 bar, it is not necessary to use special precautions for material qualification and welding that are otherwise needed to prevent SSC [18]. This encourages the use of H2S scavengers to bring the partial pressure of H2S below 0.0035 bar. It has been recently suggested that fugacity of H2S [19] or solution concentration of H2S in the aqueous phase [20,21] may be better measures of the limits of safe operation of a system without SSC. H2S scavengers are used to keep H2S concentrations below the limits where SSC occurs. Table 1.1 Effect of H2S partial pressure on length of fatal H2S concentration zone in case of leakage. H2S partial pressure (bar)

Length of fatal H2S concentration zone in case of leakage (m)

2 40 128

35 250 750

3

4

CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

1.2 Hydrogen sulfide and mercaptan measurement As H2S and mercaptans are toxic at low concentrations, it is important to be able to measure the concentration accurately and reliably at low concentrations. There are several methods available to measure H2S concentrations [22]. Methyl mercaptan and H2S concentrations in the gas phase have been measured using Ultraviolet (UV)/Visible light detection systems [14]. Multigas monitors use electrochemical sensors to measure the concentration of H2S. The method has interferences with mercaptans [23]. Typical electrochemical methods utilize the oxidization of H2S [24]. The electrochemical reaction for the oxidation of H2S is shown in Eq. (1.1). H2 S2S 1 2H1 1 2e2

(1.1)

This reaction has the standard reduction potential of 0.141 V [25]. Voltage changes will be proportional to concentration. Potentiometric titration is an important quantitative method to measure H2S and mercaptans in liquid hydrocarbons [26]. In this method, a liquid hydrocarbon sample is weighed into a 2-propanol solution containing a small amount of ammonium hydroxide. The solution is titrated potentiometrically with alcoholic silver nitrate with a glass standard electrode and a silver-silver sulfide electrode indicating system. To obtain reliability for mercaptans in hydrocarbons using potentiometric titration, one should use samples that are H2S free. Meter reading for mercaptans occurs in the 250 mV to 350 mV level [27]. One standard method to measure H2S concentrations in the gas phase is colorimetric tubes [28]. These methods are also known as the length of stain detector tubes [29]. Length of Stain detector tubes most commonly use lead [30]. The color is formed by the chemical reaction provided in Eq. (1.2). H2 S 1 PbðCH3 COOÞ2 -PbS 1 2CH3 COOH

(1.2)

The length of stain detector tubes is easy to use and inexpensive. They are often used in field situations. Gas chromatographic (GC) methods to detect H2S are reliable and can have lower detection limits of less than 0.5 ppm when a sulfur chemiluminescence detector (SCD) is used [30]. H2S can be measured in crude oil and the crude oil head space using multidimensional gas chromatography [31]. The method is not normally used in field operations. Volatile alkyl mercaptans in gasoline were measured by precipitating the mercaptans using silver nitrate to silver mercaptides [17]. The silver mercaptides are converted to pentafluorobenzyl derivatives and measured using gas chromatography with mass spectrometric detection [17]. Catalytic adsorption sensors use changes in metal oxide resistance with the adsorption of H2S to measure H2S concentrations. The method was found to be not accurate at low concentrations of H2S (020 ppm) [32].

1.4 Chemical scavengers

A method that uses continuous photometric analysis has been found to be a reliable means of detecting H2S concentration at low values [32]. A roll of lead acetate standardized paper is exposed to a continuous flow of gas. A light-sensing photometer reads the color density of the exposed paper and sends a signal to a microprocessor that converts it to a digital signal. In-line H2S measurement in different phases was accomplished using a multiphase flow meter that utilized gamma-ray fraction measurements and the measurement of the mass attenuation of sulfur in comparison with water and hydrocarbon [33]. The measurement technique allows continuous measurement of H2S in all three phases. Wireless sensor networks have been used to expand the measurement area of H2S detection [34]. This improves the safe and reliable operation of the oil and gas production area where H2S naturally occurs.

1.3 Hydrogen sulfide and mercaptan partitioning in oil, water, and gas The prediction of H2S concentrations in high pressure and high-temperature wells is extremely important for proper material selection [3]. The adsorption of H2S in brine oil mixtures at different water cuts and high pressures at room temperature has been measured experimentally [35]. Modeling and simulation of H2S removal by chemical scavengers created correlations to estimate the equilibrium constant of H2S between oil and water phases from temperature, pressure, and the gas-oil ratio [36]. H2S is rapidly absorbed in alkaline solutions [37]. The thermodynamics of gas condensate water systems enables the assessment of the possibility of corrosion or scale formation in high-temperature sour gas wells [38]. The information can be used to assess the concentrations of H2S in different phases in the good bore. The solubility of mercaptans and their distribution between isooctane and water have been measured at 20
C [39]. Lower molecular weight mercaptans have a greater solubility than higher molecular weight mercaptans do. The solubility of mercaptans in water at 20
C is shown in Fig. 1.1. It is difficult to remove high molecular weight mercaptans using aqueous alkaline solutions.

1.4 Chemical scavengers Ultra sour gas fields contain large amounts of H2S. Different processes for sulfur removal are effective at different removal rates [40,41]. Sweetening processes to remove H2S involve the use of chemical solvents, chemical and hybrid solvents, physical solvents, and hybrid solvents [2,42]. Some chemical solvents that are used in the processes are methanol, N-Methyl-2pyrrolidone, poly (ethylene glycol)

5

CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

0.12

Solubility in Water (Mols/litre)

6

0.1 0.08 0.06 0.04 0.02 0

FIGURE 1.1 Aqueous solubility of different mercaptans at 20
C.

dimethyl ether, sulfolane, and disopropanol amine [2]. The use of refrigeration and membranes have been involved in some processes for acid gas sweetening [2,43,44]. Refrigeration allows for the recovery of propane, butane, and natural gas liquids and allows acid gas reinjection for CO2 and H2S [45]. Amines have limited mercaptan removal ability so new hybrid solvents have been developed to allow better mercaptan removal in amine units [4547]. These processes will be discussed in detail in Chapter 2  Natural gas treatment. Chemical scavengers are substances that react irreversibly with H2S or mercaptan. Chemical scavengers are restricted to the treatment of sulfur compounds below one ton per day for economic reasons, except in those cases where one needs to produce natural gas or oil immediately for other reasons. Both solids and liquids are used to scavenge H2S and mercaptans.

1.4.1 Solid scavengers Solids have been used for many years to remove H2S from sour gas. Among the solid chemicals used are iron oxide, zinc oxide (ZnO), and solid nano-particles. Iron oxide has been used many years ago to remove H2S [48,49]. Iron oxide is often used in alkaline drilling muds [48]. It has been used in over 60 sour gas wells in Canada [48]. Fixed beds of iron oxide can often allow the removal of H2S to very low concentrations [49]. The process is easy to operate. The process has been limited to small streams. If one wishes to regenerate the iron oxide, the bed can become plugged with solids [49]. The disposal of spent products from iron sponge products has been known to be a fire hazard [50].

1.4 Chemical scavengers

ZnO has also been used for several years [50]. Liquid H2S scavengers proved to be more economic than ZnO for an offshore North Sea application [50]. ZnO scavengers were effective in removing H2S and mercaptans from NGL liquids in Canada [51]. High porosity ZnO granules allow enhanced lattice diffusion resulting in more complete utilization of the solid ZnO [51]. Solid nano-particles can be made from ZnO and magnesium oxide [52]. The nano particles are below 10 μm [52]. The particles have been used to remove H2S from gas streams by using these particles in a fixed bed [52].

1.4.2 Oxidizing chemicals H2S can be oxidized to elemental sulfur and further on to different thiosulfates and sulfate environments. Hydrogen peroxide, sodium nitrite, and other chemicals have been used to accomplish this. Hydrogen peroxide has been used to reduce H2S in industrial, wastewater, and geothermal applications [53]. In drilling applications for geothermal systems, a combination of hydrogen peroxide with sodium hydroxide (NaOH) has been sprayed in steam. Hydrogen peroxide is also used periodically to treat steam condensate in the geothermal power plant [53]. It has also been used in offshore applications [54]. One advantage of hydrogen peroxide is that it is an environmentally friendly H2S scavenger [54]. Sodium nitrite has been introduced downhole to remove H2S by oxidizing it [55]. Sodium nitrate was used in a gas well in New Mexico and an oil well in the Netherlands [55]. In some instances, the H2S reduction lasted for some time [55]. This effect may be due to the ability of sodium nitrite to inhibit the activity of sulfate-reducing bacteria [55]. Other oxidizing chemicals are sodium chlorite, peracetic acid, and several other persalts [56]. Mercaptans have been removed by oxidation using strong oxidant reagents such as sodium hypochlorite, oxygen, and hydrogen peroxide [57]. Oilfield brines are kept in oxygen-free conditions to prevent oxygen corrosion and precipitation. The introduction of oxidizers can cause corrosion and precipitation. This is a disadvantage of using oxidizing chemicals.

1.4.3 Aldehydes H2S reacts with aldehydes by adding on to the carbon-oxygen double bond [58]. There are several aldehydes used as H2S scavengers, such as formaldehyde, glutaraldehyde, glyoxal, and acrolein. The molecular structures of these aldehydes are shown in Fig. 1.2. Formaldehyde was one of the first H2S scavengers to be discovered [59]. When H2S reacts with formaldehyde it forms a complex mixture of mercapto derivatives [60]. Formaldehyde is known to be a human carcinogen [61]. This greatly reduces its use around the world.

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

FIGURE 1.2 Chemical structure of different popular aldehyde hydrogen sulfide scavengers.

Glutaraldehyde is normally used as a biocide [62]. It also acts as an H2S scavenger [63]. Glyoxal is a frequently used H2S scavenger in cases where a nonnitrogen-containing H2S scavenger is asked for. One reaction product of glyoxal with H2S has been the crystalline adduct formed from three glyoxal and two H2S molecules: (C2H2O2)3(H2S)2 [64]. Acrolein has been used for many years as an H2S scavenger in oilfield waterflood systems [65]. It has also been used to remove mercaptans in storage terminals. The reaction product of mercaptan with acrolein will result in an aldehyde with a thioether [10].

1.4.4 Formaldehyde reaction products Several reaction products of amines with formaldehyde are commercial H2S and mercaptan scavengers. Some of them have been economically used in applications where even 50 kg/day needed to be removed at the application [56]. Among the most popular is 1,3,5 tris (2hydroxyethyl) hexahydro-s-triazine [66]. This is the reaction product of monoethanol amine with formaldehyde to form the cyclic product popularly described as MEA triazine (Fig. 1.3). One of the first references for use of this compound to reduce the levels of H2S and organic sulfides from gaseous and liquid hydrocarbons is a 1989 patent [67]. The generally accepted pathway to react with H2S is shown in Fig. 1.4. Carbon dioxide has little effect on the H2S scavenging ability of MEA triazine [68]. Another slightly less popular amine formaldehyde reaction product H2S scavenger is 1,3,5 trimethyl hexahydro-s-triazine. This product is the reaction product of methyl amine with formaldehyde [69]. This is popularly known as MMA triazine and is shown in Fig. 1.5. The reaction product of 3 methoxy propyl amine with formaldehyde is also used in some circumstances [70]. The reaction products of alkanolamines with paraformaldehyde form bisoxazoladines [71]. One bisoxazolidine that is formed is the reaction product of mono isopropanol amine with paraformaldehyde. Some popular H2S scavengers that are reaction products of amines with formaldehyde are listed in Table 1.2.

1.4 Chemical scavengers

FIGURE 1.3 Chemical structure of MEA triazine (C9H21N3O3).

FIGURE 1.4 The reaction pathway of hexahydrotriazine reacts with hydrogen sulfide.

FIGURE 1.5 Chemical structure of MMA triazine (C6H15N3).

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

Table 1.2 Hydrogen sulfide scavengers resulted from reactions of amines and formaldehyde. H2S scavenger

Reaction

1,3,5 tris (2hydroxy ethyl) hexahydro-s-triazine N,N’ methylenebis-oxazolidine

Monoethanol amine with formaldehyde Mono isopropanol amine with formaldehyde Methyl amine with formaldehyde 3 methoxy propyl amine with formaldehyde Ammonia with formaldehyde

1,3,5 trimethyl hexahydro-s-triazine 1,3,5 tris (3 methoxy propyl) hexahydro-s triazine Hexamethylenetriamine

Formaldehyde reaction products with amine-free alcohols or urea have also been found to be used as H2S scavengers [72]. Ethylene dioxy (dimethanol) has been used as an H2S scavenger on the Norwegian continental shelf [73].

1.4.5 High valence metal compounds Metal chelates have been used as H2S Scavengers [56]. Iron and zinc are commonly used in this category of H2S scavengers. Ferrous iron gluconate has been used as a H2S scavenger for drilling muds [74]. The product is highly water-soluble and does not precipitate at pH 12 [74]. The product is environmentally friendly and has been used in the drilling of reservoirs with 36% H2S. Non-triazine organic acid-based metal complex H2S scavengers have worked well in downhole mixed production applications [75]. The product was suitable for capillary injection downhole. The product has a very fast reaction time that allows it to be used in systems with low residence times. The product lowers H2S concentrations rapidly. It is cost-effective in many applications where H2S concentrations need to be brought down to low levels. The product has been used in gas pipeline systems [76].

1.4.6 Aqueous alkaline solutions Aqueous alkaline solutions have been used to remove H2S from gas streams [77]. Aqueous alkaline solutions also react with CO2. The solutions are used in situations where H2S concentrations are high and CO2 concentrations are lower. The reaction of NaOH is faster with H2S than with CO2 so contact times are kept low to achieve selective adsorption [37]. Aqueous alkaline solutions have been used to extract mercaptans from an oil phase [39,56,77]. Lower molecular weight mercaptans are more easily removed by these solutions than higher molecular weight mercaptans [39]. It has also been observed that increasing the amount of NaOH increases the removal of

1.5 Physical chemistry of scavengers

mercaptans. Mercaptans are more easily removed at lower temperatures in aqueous alkaline solutions [39].

1.4.7 Hydrogen fluoride Anhydrous liquid hydrogen fluoride (HF) was found to extract organic sulfur compounds such as low molecular weight mercaptans from high sulfur petroleum stock [78]. The reaction is an acid-base reaction. The reaction product is a sulfur HF complex [78]. It is easier to extract lower molecular weight mercaptans than higher molecular weight mercaptans. The HF can be extracted and reused. The process was studied in a 100 gallons per day pilot plant [78]. It is not clear if this process is currently used commercially. HF is an extremely hazardous chemical that would make its use in any process requires a large number of safety precautions for its use.

1.4.8 Novel hydrogen sulfide scavengers from biological sources H2S scavengers have been obtained from biological sources. In some cases, enzymes are used. In other cases, sugar is oxidized to form novel compounds that act as H2S scavengers. Cloning the cDNA sequence from thermophilic bacteria can create an H2S scavenger that is derived from a biological source [79,80]. This scavenger is a Sulfide Quinone Reductase (SQR) enzyme system [79]. The enzyme system was created in large quantities from a yeast expression system [80]. This scavenger has been used to treat an oil well in the Baaken shale formation [80]. Novel oxidized sugar compounds have also been examined as H2S scavengers [81]. The products performed better than MEA triazine when H2S was bubbled through a formulation of tapwater with 0.5% H2S Scavenger [81]. The comparison is not a fair evaluation of MEA triazine as in practice one uses much higher concentrations of triazine in tower application so the pH of the solution in the tower is much higher so the MEA triazine does not hydrolyze.

1.5 Physical chemistry of scavengers In many applications of H2S scavengers, the scavenger is used in amounts substantially greater than stoichiometric amounts. This often occurs due to the fact that the choice of the scavenger does not take into account the way the scavenger contacts H2S in the given oilfield application. In different applications, different phases are present. The residence times vary. All of these considerations impact the economic choice of the scavenger.

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

1.5.1 Scavenging kinetics The kinetics of H2S scavenging by different 1,3,5 triazinanes has been investigated using NMR spectroscopy [69]. The rate of decomposition of the triazinane is provided by Eq. (1.1). 2


 ⅆ½T  5 k½T  H 1 ½HS ⅆt

Where [T] is the concentration of triazinanes, t is time, [H1] is the concentration of hydrogen ions, [HS2] is the concentration of bisulfite ions and k is a kinetic rate constant. Determined by Bakke and Buhang [69], the kinetic rate constants (in M22s21) is 9.1 6 0.2 3 1027 for 1,3,5 tris (2hydroxy ethyl)-1,3,5-triazinane, 5.6 6 0.4 3 1027 for 1,3,5-triethyl-1,3,5 triazinane, and 0.29 6 0.03 3 1027 for 1,3,5 trimethyl -1,3,5 triazinane, respectively, at 22
C. 1,3,5 tris (2-hydroxy ethyl)-1,3,5 triazinane also known as 1,3,5 tris (2hydroxy ethyl) hexahydro-s-triazine or MEA triazine has the highest reaction rate of the triazinanes studied. The reaction rate of this compound is substantially greater than 1,3,5 trimethyl-1,3,5 triazinine which is also known as MMA triazine. Bakke and Buhang also studied the hydrolysis of the triazinanes [69]. They found that the triazinanes hydrolyze rapidly in acidic solutions. This may explain the lack of performance of triazines in water systems that are in equilibrium with acid gases. Kinetic studies were made on the removal of gaseous H2S using Fenton’s reagent in a spraying tower [82]. Experiments in a minitower allowed a comparison of the kinetics of an MEA triazine and a neutral pH hydroxyl derivative scavenger [83]. The rate constants obtained in this study showed the MEA triazine formulation reacts much faster than the neutral pH Hydroxyl derivative. The reaction rate in the tower is 0.275 (hour21) with MEA triazine formulation but only 0.016 (hour21) with neutral pH Hydroxyl derivative at 22
C. In the experiment, the MEA triazine formulation works faster. The rate of reaction may not be best modeled using first-order kinetics for this type of process. In some work, the rate of removal of H2S by a triazine scavenger was modeled using a model that combined mass transport with reaction kinetics [36]. The kinetics of the adsorption of H2S in a NaOH and sodium carbonate (Na2CO3) were investigated by Cassinis and Farone [77]. The enhanced kinetics allow caustic soda scrubbing of H2S in sour gas that contains CO2.

1.5.2 Scavenger thermodynamics H2S Scavengers exhibit different partitioning in oil and water. It is important to understand the partitioning to design treatments with the appropriate amount of scavenger. The partitioning of glutaraldehyde and acrolein between oil and water with different salt concentrations was studied [63]. The information was used to find appropriate binary coefficients for the Elliott-Suresh-Donohue equation of state [63].

1.6 Laboratory testing of hydrogen sulfide and mercaptan scavengers

Both MEA triazine (1,3,5 tris (2hydroxy ethyl)-1,3,5-triazinane) and MMA triazine (1,3,5 trimethyl-1,3,5 triazinane) are water-soluble [84]. Glyoxal exists as hydrated oligomers in an aqueous solution [84]. The hydrated oligomers are in equilibrium with each other. The concentration of larger molecular weight oligomers is higher in concentrated glyoxal solutions whereas dilute solutions of glyoxal have large concentrations of hydrated glyoxal monomer [85]. The chemical equilibrium and liquid-liquid equilibrium of aqueous solutions of formaldehyde and n-butanol have been studied [86]. Formaldehyde reacts with water and alcohol to form oligomers. This work provides values of chemical reaction constants of the oligiomeration reaction, vapor pressure equation coefficients, and liquid phase activity coefficients for the formaldehyde water and n-butanol system [86].

1.6 Laboratory testing of hydrogen sulfide and mercaptan scavengers There are a number of different means to test H2S and Mercaptan Scavenger performance in the laboratory. The purpose of the laboratory tests could be for basic screening, evaluation of performance in a tower, or evaluation of performance in mixed production in inline injection. Evaluation of performance in systems where the reduction of H2S is desired in the vapor phase is different than when an impurity such as mercaptan is desired to be removed from crude oil.

1.6.1 ASTM D5705 test methodology with modifications The ASTM D5705 test apparatus [87] is a testing method originally intended to determine H2S in the residual space above a residual fuel oil. In the method, an H2S-free 1 L container is filled with fuel oil to 500 mL. The vapor space is sparged with nitrogen to displace air. The test container is heated with a sample in an oven to 60
C and agitated in an orbital shaker at 220 rpm for 3 minutes. The H2S concentration is measured using a length of stain detector [28,29]. This test procedure has been modified to allow the assessment of scavengers in multiphase fluids and crude oils at different temperatures (20
C50
C), different shaking rates, and time of exposure (124 hours) to an H2S scavenger [83]. These modifications create an easy-to-use procedure with the facility to investigate the effects of different water to oil ratios, and length of exposure with different H2S scavengers.

1.6.2 Laboratory assessment of hydrogen sulfide and mercaptan scavengers in towers A miniature gas tower apparatus [83], and glass sinter tower apparatus [88] are methods to evaluate H2S scavengers in tower apparatus. H2S containing gas is

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

sparged through a column containing liquid H2S Scavenger. In different examples, the H2S concentration in effluent gas was measured with an online electrochemical monitor was used [83] or a gas chromatograph [88]. The H2S concentration of effluent gas will be low for a while when a breakthrough time is reached. At this time, the H2S concentration will rapidly climb to high levels. The breakthrough time determines the specific consumption of the scavenger. The miniature gas tower apparatus has been used to evaluate the effect of different mercaptan scavengers [14]. In this evaluation of an H2S and mercaptan scavenger, the inlet gas contained 2500 ppm H2S and 500 ppm methyl mercaptan. A UV-Visible online detector system was used to measure the concentration of H2S and methyl mercaptan [14]. The breakthrough time of the scavenger for methyl mercaptan was much faster than for H2S [14].

1.6.3 Continuous gas flow apparatus The continuous gas flow apparatus is a test procedure to evaluate H2S scavenger performance [75,76,89,90]. Comprehensive details of the apparatus are provided in the papers by Petrobras [89,90]. A picture of one such apparatus is shown in Fig. 1.6. In the continuous gas flow apparatus, the gas containing H2S is sparged through a reaction vessel. The reaction vessel will contain a mixture of brine and

Glass reaction vessel

FIGURE 1.6 Continuous gas flow test setup.

Gas analyzer

Mass flow meter

1.7 Hydrogen sulfide and mercaptan scavenging process

hydrocarbon mimicking the actual system. Gas is sparged through the liquid until the H2S concentration reaches a steady-state level. This level should be commensurate to the concentration of the inlet gas. H2S scavenger. With the injection, the H2S concentration will immediately fall. The fall in H2S concentration is an indication of the fast-acting ability of the H2S scavenger. As more gas is sparged through the liquid. The H2S concentration will gradually reach the steady-state concentration of H2S that was obtained before the injection of the H2S scavenger. The total number of moles consumed by the scavenger can be computed from the difference between steady-state H2S concentration with actual H2S concentration when the H2S scavenger is injected. This can allow the determination of the specific consumption of H2S by the scavenger.

1.7 Hydrogen sulfide and mercaptan scavenging process Dependent on the process conditions and chemical scavengers used there are a variety of different methods to introduce chemical scavengers in a process stream and remove H2S and mercaptans from the process. These vary from in-line injection, gas lift injection, injection using a capillary string, contact tower applications, and crude oil storage tanks.

1.7.1 In-line injection In-line injection of H2S scavenger is an inexpensive method to introduce chemicals in a process stream. Velocities to achieve optimal mixing have been mentioned to be in the range of 2545 ft/s [90]. Retention times should be greater than 20 seconds [91]. Pulsation dampeners (Fig. 1.7) and the use of appropriately sized atomizers can aid the efficient use of scavengers in these systems [91,92]. Atomizing nozzles deliver the scavenger more effectively into the gas phase than the fine mists which increase the relative chemical surface area. Improved measurement of H2S in the export gas allows better control and optimization of H2S Scavenger consumption. Automation of the H2S scavenging process can result in paying off the investment in less than one year. There have been several examples of successful automation of in-line injection [32,92,93]. An innovative hydrocarbon-based H2S scavenger that contained zinc was used in one system [93]. Computational Fluid Dynamics has been used to improve the efficiency of in-line injection systems [94]. Direct injection of triazine scavengers has allowed the development of low sour gas reserves [95]. Depending on the application scavenging efficiencies can vary. In one instance a combination H2S scavenger /scale inhibitor minimized calcium carbonate formation. Direct injection of H2S scavengers in mixed production systems with high calcium content can result in calcite scaling [96]. Computer modeling indicates the reduction of triazine consumption with an increase in residence time. The optimum temperature for

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

FIGURE 1.7 Pulsation dampeners used to stabilize the scavenger delivery.

minimum triazine scavenger consumption was found to be between 60
C and 75
C [97]. Acrolein has been used in mixed production systems in Alaska [98]. Acrolein has also been used in water injection systems to remove H2S and iron sulfide (FeS) scale solids [99]. In-line injection of a fast-acting H2S scavenger in a dry gas line allowed the operator to bring H2S concentrations to near-zero levels [76]. Static mixers have been recently examined in laboratory studies to increase the efficiency of in-line injection [100]. Static mixers have been used in field applications to improve efficiency [91]. In coiled tubing drilling operations, triazine H2S scavengers have been injected in a separated water stream after caustic soda injection as an alternative to the use of iron gluconate use to ensure that the H2S is not regenerated [101]. H2S scavengers have also been injected in line to minimize sulfur oxide (SOx) emissions in refinery flare gas systems [102].

1.7.2 Gas lift injection Gas lift is one means of artificial lift. In some instances, gas-lift injection of H2S scavengers has been used for several years to mitigate the effects of H2S in wells, flow lines, risers, and topside facilities [89]. This is especially important when materials that do not resist the effects of sour cracking corrosion are used in the facilities. It is important to assess the thermal stability of H2S Scavengers in this application [89]. Evaluation of heat stability, the effect of evaporation, and assessment of the viscosity of the H2S scavenger after evaporation can aid the assessment of the H2S Scavenger in gas lift applications [103]. An offshore production

1.7 Hydrogen sulfide and mercaptan scavenging process

system that utilized H2S scavengers in gas lift operations was simulated using multiphase flow simulation [36]. In these instances, the H2S scavenger scavenges the gas phase and removes the H2S in this phase. The simulation indicated the need for better atomization in the gas phase [36].

1.7.3 Capillary injection H2S scavengers have been introduced downhole using a capillary string [75,104]. In order to qualify the H2S scavenger for capillary injection, the chemical must undergo appropriate thermal stability, corrosion, and other compatibility tests before the scavenger are used in the capillary string [75,104]. In downhole applications, the H2S Scavenger will contact a very acidic environment. When water is present some H2S scavengers such as MEA and MMA triazines can hydrolyze and decompose [69]. For this reason, these scavengers are not suitable for downhole injection. New H2S Scavengers have been developed for multiphase systems [75,104,105]. Placing the H2S scavenger downhole has many advantages in allowing for inexpensive materials on the surface and sufficient residence time for the H2S scavenger. The disadvantage with downhole applications is that one has to remove the H2S in all phases which may increase the amount of H2S scavenger needed for the application.

1.7.4 Contact towers Contact towers have been used for both solid and liquid H2S scavengers (Fig. 1.8). Fig. 1.9 shows its key components. In one instance natural gas with H2S contacted a tower containing ZnO granules [50]. This system was cumbersome to use as the tower had to be changed out on average every 11 days [50]. Fixed bed contact towers have been used to treat offshore fuel gas to keep the fuel gas at very low H2S concentrations in California [106]. Two beds were used in this application to prevent sulfur dioxide emissions [106]. Mechanical and chemical alterations resolved operational problems with the iron oxide slurry process of H2S removal [107]. Contact bubble towers are often used to remove H2S from a natural gas stream [91]. The use of these systems can often result in the close stoichiometric operation of the tower [108]. Liquid-based H2S Scavengers are typically used in bubble towers [108]. In the field studies, it was found in this study that the maximum superficial velocities that could be sustained without foaming and carryover were in the 0.0400.055 ft/seconds range [108]. Continuous contact towers have been a popular way to remove H2S due to their flexibility and lower capital costs [109].

1.7.5 Storage tanks When H2S or mercaptans are present in large volume storage tanks, communities that are close to the storage tanks often complain about the odors. If the H2S or

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

FIGURE 1.8 A typical contact tower for hydrogen sulfide removal.

mercaptan concentrations are high, the conditions are hazardous for operators and can pose a safety risk to the community. Sour crude is often treated with an H2S scavenger in large volume crude oil storage tanks to remove mercaptans [10] and H2S [110]. Acrolein was used to remove mercaptans from Canadian sour crude [10]. Water-based H2S scavengers can remove H2S from crude oil while keeping the Nitrogen content of the oil is low as the water-based formulation separates from the oil [110]. Efficient removal of H2S using an H2S scavenger requires sufficient mixing and contact times [110]. In some instances, it is necessary to recirculate oil in the storage tank to create sufficient mixing [110]. Mercaptans have been successfully removed from mixed fluids collection tanks using an innovative mercaptan scavenger in Southern Italy [111].

1.7 Hydrogen sulfide and mercaptan scavenging process

FIGURE 1.9 Illustration of the key components in a contact tower.

1.7.6 Rail cars When sour crude is transported in rail cars, H2S concentrations in the vapor space can be high and hazardous. Sour crude has been treated with non-nitrogen non-corrosive H2S scavengers to reduce H2S concentration in crude oil from 5000 ppm to less than 10 ppm H2S when the rail car arrived at the refinery [83,103]. Non-nitrogen-containing scavengers are used in these applications to avoid amine salt corrosion issues in the refinery. The H2S scavenger used in this application should also not contain water.

1.7.7 Scavengers in acidizing treatment Often high H2S concentrations are detected in the flow back after an acidizing treatment [112]. To prevent this, H2S Scavengers have been used in acidizing treatments [112,113]. H2S scavengers can affect the reaction kinetics of acid dissolution and acid dissolution capacity. It is important to know the effect of the H2S scavenger on the reaction kinetics of acid dissolution, the acid dissolution capacity, and scavenging capacity. It is important to know the potential of the spent scavenger to cause formation damage [112]. This is assessed in core flood

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

tests [112]. Triazine should be removed from acid recipes, as it does not work under these conditions [113].

1.8 Case studies In this section case studies will be discussed in greater detail different instances where different approaches were used to solve different H2S treatment challenges.

1.8.1 Optimization of South Texas system A South Texas system was removing H2S from a 35 MMscfd gas stream that had a dewpoint of 20 lb/MMscfd [91]. 1215 ppm of H2S were removed in the operation. The system was treated at a single point using the continuous injection of a triazine scavenger using an atomizer. Mixing was poor due to low velocities between 8 and 10 ft./seconds and low residence times between 12 and 15 seconds. Installation of a treating loop that increased velocities to 2025 ft./seconds and retention times to 1620 seconds increased the reaction rate by 50%. Doubling the length of the loop further, introducing static mixers, introducing a pump pulsation dampening bottle, and introducing water to the gas upstream of the chemical injection point increased the reaction rate further. The final stage of optimization involved adding a polishing bubble tower. The improvements in the system decrease specific consumption by half.

1.8.2 Optimization of scavenging cost from joint industry program A major European multinational oil and gas company optimized its H2S Scavenger costs in a field by utilizing a computer model and experimental insights from a joint industry program [92]. The most significant improvement obtained in this work was to relocate injection points for the scavenger at their most optimal location. Other improvements were obtained by improved monitoring and measurement of H2S in the export gas, improved delivery of chemicals by priming the pulsation dampeners and using an appropriately sized atomizer, and using an additional H2S scavenger injection location at the high-pressure separator discharge. Further savings were obtained from using suitable pumps, fitting the system with a retrievable insertion probe, and automation.

1.8.3 Scavenger automation at sour gas processing facility A natural gas field was producing 101160 MMscfd of sour natural gas [32]. The sales natural gas pipeline required the concentration of H2S to be below 4 ppm. The concentration of H2S was brought to this concentration using an H2S scavenger. The H2S scavenger was injected upstream using an appropriately sized

1.8 Case studies

atomizer. Prior to automation, the facility used on average 50 gallons per day at the facility. A photometric inline H2S analyzer recorded H2S concentrations at the natural gas sales point. A 420 mA signal was created from the sensor that was transmitted using a radio transmitter to an actuator that controlled a chemical feed pump. The chemical pump is designed to produce 60 strokes per minute. The stroke length controls the volume of chemicals injected. With the automation method scavenger consumption was reduced to 33 gallons per day.

1.8.4 Development of sour reservoir in giant field in Kuwait A sour reservoir in a giant field in Kuwait had H2S concentrations between 1% and 3% H2S [93]. The high concentration of H2S prevented the development of this field. The wells had a water cut of 1%. A pre-engineered flow loop with proprietary packing was used with sampling and chemical injection points in the system. Three H2S analyzers were in the system. One analyzer was before chemical injection. One was after and a third acted as quality control. A Zinc-based H2S scavenger was used for the system. The system was installed at the well manifold. The system reduced H2S concentration from 13,000 ppm to less than 5 ppm with an oil flow rate of 3400 bbl/d and the use of 450 gallons per day of H2S scavenger. At oil prices between $23 and 34 bbl, the project generated an internal rate of return (IRR) of 13.3% for the pilot trial. The new skid with the zinc-based H2S scavenger allows the operator to produce from the sour reservoir without increasing the complexity of well design and production operations.

1.8.5 Development of sour gas field in the Netherlands using scavenger with scale inhibitor A new H2S Scavenger was developed to minimize calcium carbonate scale formation that would normally arise from using a highly alkaline triazine scavenger with high concentrations of calcium and bicarbonate [95]. The product significantly reduced calcium carbonate with no adverse effect on H2S scavenging or fluid foaming. In the given application it was important to use direct pressure atomization nozzles that allow high-pressure gas to boost the velocity in the atomizer nozzle. The use of these nozzles with a dual injection system allowed greater than 99% H2S removal at a specific consumption rate of 9 L/kg with effective control of calcium carbonate scaling. In some instances, specific consumption is as low as 7 L/kg in this application.

1.8.6 Scavenging dry gas pipeline in Western Oklahoma A 16-inch 30-mile dry gas pipeline in Western Oklahoma transported 100120 MMscfd [76]. The inlet gas H2S concentrations varied between 2 and 2.8 ppm. It was desired to bring the concentrations of H2S below 1 ppm with

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

scavengers that did not contain water. This was previously being done with an oil-soluble at specific consumption rates that varied between 3.58 and 6.67 gallons per pound of H2S. A new non-triazine-based scavenger with high-temperature stability and fast kinetics was used in the line. H2S concentrations were brought down below 1 ppm at specific consumption rates that ranged between 1.23 and 1.89 gallons per pound of H2S.

1.8.7 Scavenging in coiled tubing drilling operations in Saudi Arabia Chemically treated water is used in coiled tubing drilling operations of lateral wells [101]. In some lateral wells in Saudi Arabia, high levels of H2S were found. It was important to scavenge the H2S in the drilling fluids with scavengers that would not regenerate before disposal. The drilling fluid passed through a highpressure gas separator, a low-pressure gas separator, and a condensate and water separator before being treated with a scavenger. Caustic soda was added to the water stream prior to treatment with an H2S scavenger. A Triazine-based H2S scavenger replaced ferrous gluconate in this process. The process was able to scavenge the H2S in the water phase ranging between 1363 and 1611 mg/L to values between 69 and 83 mg/L with efficiencies between 91% and 99.9%.

1.8.8 Reduction of sulfur oxide content of flare gas SOx is an environmental pollutant that causes acid rain and is regulated in the United States [102]. An emission of greater than 500 lbs/day requires a formal root cause analysis by the Environmental Protection Agency. The customized injection of H2S scavengers in the gas stream is one means to prevent this. In one case of a US refinery, exceeding the SOx regulation limit would force refinery operations to cease resulting in losses of over $2 million dollars per day. A custom H2S treatment facility was designed for this refinery to limit SOx emissions to below 200 lbs/day. A four-nozzle injection system in an existing gas flare system was used. Field trials were conducted to meet SOx specifications. SOx emissions were reduced from 500 lb SOx/day to 70 lb SOx/day. With the H2S scavenger, the refinery avoided the installation of a multi-million dollar processing unit and prevented any decrease in production or any downtime. Avoiding downtime saves the refinery millions of dollars annually.

1.8.9 Capillary string downhole injection in South Texas A new automated injection system was developed to deliver an H2S scavenger downhole using a capillary string [75]. The new system combined a new intelligent pump controller with H2S monitoring and feedback control to deliver the H2S Scavenger downhole. This system was used to keep H2S concentrations

1.9 Challenges associated with scavenging treatment

below 4 ppm with a well that produced between 110 and 170 barrels of oil per day with water production between 30 and 40 barrels per day and gas production between 70 and 250 Mscfd [75]. The H2S concentration in the gas varied between 500 and 1000 ppm. Downhole injection of a new non-triazine organic acid metal complex-based H2S scavengers kept H2S concentrations below 4 ppm.

1.8.10 Fixed bed hydrogen sulfide removal in the North Sea A fixed bed system was used to remove H2S from a North Sea platform [50]. The field has had production rates of 220,000 barrels of oil per day and 120 MMscfd. It was important to keep the H2S content of the gas low. The fixed bed system was 8 m3 in volume [50]. When ZnO was used, the tower had to be changed out every 11 days. The change would take 16 hours to complete. A triazine-based H2S scavenger was selected to replace the ZnO. Laboratory tests were done to assess H2S scavenger performance and compatibility of the scavenger with triethylene glycol, ZnO, and Nylon 11 compatibility. Two products were fieldtested. Both products delivered better results than ZnO. One triazine product had better performance but an undesirable odor which lead to the operator choosing the triazine that did not have this odor. The cost of removing H2S in this facility was reduced by 70% by changing to the triazine product from the ZnO product.

1.9 Challenges associated with scavenging treatment There are several challenges associated with the use of H2S scavenging. Some of these are associated with hard-to-remove reaction byproducts, scaling induced by the basicity of the scavenger, corrosivity of the H2S scavenger, formation damage, emulsion problems with oil and water separation, and the inability of the H2S scavenger to remove H2S to a low enough concentration at economical levels.

1.9.1 Reaction products The reaction product of hexahydro 1,3,5-(2-hydroxyethyl)s-triazine with H2S can form byproducts with large amounts of sulfur [114]. The deposit in some instances can only be removed physically (Fig. 1.10). This deposit is found in applications involving pipelines utilizing in-line injection and low-velocity pipelines [109]. The problem also occurs in contact towers when the scavenger is overspent. A mechanism that explains the formation of high amounts of sulfur in the reaction product has been proposed [115]. When operating with hexahydro 1,3,5-(2-hydroxyethyl) s-triazine it is very important not to overspend the product. Automation of the H2S scavenging process can prevent this. Other nonnitrogen-containing H2S scavengers can avoid the problem of the formation of hard-to-remove reaction byproducts [83].

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

FIGURE 1.10 Amorphous polymeric dithiazine deposit formed by overspending with MEA triazine scavenger.

1.9.2 Induced scaling problems In upstream production, carbon dioxide is often present in relatively high concentrations. This allows the process to carry and transport brines with calcium, magnesium, and iron. Aqueous alkaline solutions and amine formaldehyde reaction products are basic. When these chemicals are water-soluble, they can induce precipitation in calcium-containing brines. It is important in brines such as the Haynesville brine, to use MEA Triazine products formulated with scale inhibitors [95] or non-triazine H2S scavengers that cause substantially less scaling than MEA triazine [96,105].

1.9.3 Corrosion issues Glyoxal is an acidic H2S scavenger with significant corrosivity to many metals [83,103]. It can cause significant corrosion in stagnant pipelines containing oil and water. Fig. 1.11 shows the corrosion rates of different metals placed in MEA triazine, a neutral pH hydroxyl derivative H2S scavenger, and glyoxal at 60
C for 14 days. The corrosion rate of aluminum placed in glyoxal for 14 days at 60
C is 16.6 mpy. The corrosion rate of mild steel placed in glyoxal for 14 days at 60
C is 750 mpy. The high corrosivity of glyoxal with mild steel makes it very important to allow significant concentrations of glyoxal in water to contact mild steel for any significant amount of time. This is one of the disadvantages of using glyoxal as an H2S scavenger. MEA Triazine and the neutral pH hydroxyl H2S Scavenger do not have this drawback.

1.9 Challenges associated with scavenging treatment

1000

Corrosion Rate (mpy)

100

10

1

0.1

0.01

0.001 Admirality Aluminium Brass MEA Triazine

Copper

Mild Steel

Neutral pH Hydroxy Derivative

SS 304

SS 316

Glyoxal

FIGURE 1.11 Corrosion rates of different metals placed in different hydrogen sulfide scavengers for 14 days at 60
C.

1.9.4 Formation damage During acid stimulations, it is important to choose an H2S scavenger that does not cause a loss in permeability. Some spent scavenger products can cause a reduction in permeability of 30%40% in Berea sandstone cores [112]. This reduces the value of using these products in acid stimulation. It was also found in a particular study that a commercial H2S scavenger caused a loss of 8% in permeability [112].

1.9.5 Emulsion problems in oil and water separation Organometallic H2S scavengers can often deliver excellent H2S scavenging performance in mixed production lines but create undesirable emulsion problems for oil and water separation [105]. The undesirable emulsion tendency often precludes their use in many mixed production systems. New non-triazine scavenger chemistry delivers good performance in mixed production without these emulsion problems [100]. This allows the use of an H2S scavenger in a mixed production system.

1.9.6 Overconsumption of scavenger In order to prevent overconsumption of H2S scavenger during in-line injection it is important to operate at optimal velocities to achieve requisite mixing [91].

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

One needs to have sufficient residence time and use pulsation dampeners and appropriately sized atomizers [91,92]. Measurement of H2S in the export gas allows better control and optimization of H2S Scavenger consumption [92]. Automation can result in paying off the investment in less than one year [92]. Computational Fluid Dynamics can help decrease H2S scavenger consumption [94]. If one is in a situation where the inline injection is used with low residence times, one may consider the use of fast-acting H2S scavengers [76]. Contact bubble towers should be operated in optimum velocity ranges to prevent inefficient use of H2S scavengers [108]. Automation of equipment will result in more efficient operation.

1.10 New developments Safe operation, digitization and use of environmentally friendly chemicals are growing trends in the oil and gas industry. These trends are reflected in new developments of application of new H2S scavengers to mixed production, automation of scavenging applications, and novel environmentally friendly H2S scavengers.

1.10.1 Safe operation Keeping oilfield fluids at the surface at low concentrations at the wellhead can increase the safety of people who work in an oilfield. This can be accomplished using H2S scavengers that are injected using capillary tubing [75,104] or H2S scavengers that have been introduced using gas lift [36,89,103]. The H2S Scavenger used in these applications needs to be compatible for use in capillary strings or gaslift systems. Upstream product conditions are high temperature and have a low pH. These conditions cause difficulties for H2S scavengers such as MEA triazine. Several new H2S scavengers have been developed for a mixed fluid environment [75,76,104,105].

1.10.2 Digital transformation Digital transformation of oilfield operations results in safer operations, greater reliability, and increased production [116]. Automation of H2S Scavenger applications results in additional safety when the automation keeps downstream H2S concentrations low. Advances in this field should decrease the possible exposure of personnel to high concentrations of H2S. Automation projects have also provided a quick return on the additional investment involved during the automation [32,92,93].

1.10.3 Environmentally friendly products Ideally one would like to use safe and environmentally friendly H2S scavengers. Many of the older H2S scavengers are not safe or environmentally friendly. New

Nomenclature

developments to create environmentally friendly products have been based on biological sources and fermentation. H2S scavengers used to treat an oil well have been created from a biological source [79,80]. Oxidized sugar compounds from fermentation processes have also been examined as H2S scavengers [81]. These products are environmentally friendly but so far are not as cost-effective as other H2S scavengers.

1.11 Summary and conclusions In this chapter, H2S and mercaptan scavenging using substances that irreversibly react with H2S or mercaptans were reviewed. H2S and mercaptans are prevalent in many sources of natural gas and crude oil. The substances are toxic and cause corrosion problems. This motivates their removal. The use of substances that irreversibly react with H2S and mercaptans is one means of their removal. There are many substances that irreversibly react with H2S and mercaptans. Some are solid substances. Others are liquids that could be oxidizing chemicals, aldehydes, formaldehyde reaction products, high valence metal chelates, and other chemicals. The physical chemistry of the H2S scavengers is important in determining the scavenger for the given application. Many different methods are used to introduce chemical scavengers in a process stream such as in-line injection, gas lift injection, injection using a capillary string, contact tower applications, crude oil storage tanks, and rail cars. Some challenges associated with the use of H2S scavenging are hard to remove reaction byproducts, scaling induced by the basicity of the scavenger, corrosivity of the H2S scavenger, formation damage, emulsion problems with oil and water separation, and the inability of the H2S scavenger to remove H2S to a low enough concentration at economical levels. Automation is increasing in different scavenging applications. H2S scavengers are applied to mixed products in the well tubing. New nontriazine chemistries have been developed for different applications. Some new H2S scavengers created from fermentation processes have been developed for oilfield applications.

Nomenclature BSR cDNA CO2 GC H2S IRR L/kg lb M

Bacterial sulfate reduction Complementary DNA Carbon Dioxide Gas Chromatography Hydrogen sulfide Internal Rate of Return Liters per kilogram Pounds Molar concentration (or molarity)

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CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

MMscfd Mscfd mpy Na2CO3 NaOH ppb ppm SCD SOx SS SSC TLV TSR ZnO

Million standard cubic feet per day Thousand standard cubic feet per day Mils (thousandths of an inch) per year Sodium Carbonate Sodium Hydroxide Parts per billion Parts per million Sulfur chemiluminescence detector Sulfur oxide Stainless steel Sulfide cracking corrosion Threshold Limit Value Thermochemical sulfate reduction Zinc oxide

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[75] S. Lehrer, V, Jovancicevic, S.C. Braman, et al., New hydrogen sulfide scavenger development for downhole mixed production applications  Lab and field data. Paper presented at: SPE-173788, SPE International Symposium on Oilfield Chemistry, 2015 April 1215; The Woodlands, TX, USA. [76] S. Ramachandran, S. Lehrer, V. Jovancicevic, et al., Sour corrosion reduction using new hydrogen sulfide scavengers  Lab and field data. Paper presented at: NACE4320, Corrosion Conference & Expo; 2014 March 913; San Antonio, TX, USA. [77] R.B. Cassinis, W.A. Farone, Improved H2S caustic scrubber. Paper presented at: SPE-38273-MS, SPE Western Regional Meeting; 1997 June 2527; Long Beach, CA, USA. [78] A.P. Lien, D.A. McCaulay, B.L. Evering, Hydrogen fluoride extraction of sulfur compounds and high sulfur petroleum stock. Paper presented at: WPC-4211, Proceedings of 3rd World Petroleum Congress, Section III, pp. 14555. [79] P. Dhulipala, J.K. Jani, M.R. Wyatt, et al., Bio-molecular non-corrosive hydrogen sulfide scavenger. Paper presented at: SPE-190908-MS, SPE International Oilfield Corrosion Conference & Exhibition, 2018 June 1819; Aberdeen, Scotland, UK. [80] P. Dhulipala, J. Jani, M. Wyatt, et al., Sour fluids management using non-chemical H2S scavenger. Paper presented at: SPE-194990-MS, SPE Middle East Oil and Gas Conference; 2019 March 1821; Manama, Bahrain. [81] P. Schuber, J. Su An, C. Arceo, Novel oxidized sugar compounds: a cheap, low dose, environmental friendly replacement for triazine. Paper presented at: URTeC-2727, Unconventional Resource Technology Conference; 2020 July 2022; Austin, TX, USA. [82] Y. Wang, Z. Wang, J. Pan, et al., Removal of gaseous hydrogen sulfide using Fenton reagent in a spraying reactor, Fuel 239 (2019) 7075. [83] S. Ramachandran, S. Lehrer, S. Chakraborty et al., Novel scavenger technology for effective removal of H2S from produced gas in oilfield applications. Paper presented at: SPE-192885, Abu Dhabi Petroleum International Petroleum Exhibition and Conference; 2018 November 1215; Abu Dhabi, UAE. [84] G.N. Taylor, J.J. Wylde, T. Muller, et al., Fresh insight into the H2S scavenging mechanism of MEA-Triazine vs MMA-Triazine. Paper presented at: SPE-184529MS, SPE International Conference on Oilfield Chemistry; 2017 April 35; Montgomery, TX, USA. [85] E.B. Whipple, The structure of glyoxal in water, Journal of the American Chemical Society 92 (1970) 71837186. [86] R. Peschla, B.C. Garcia, M. Albert, et al., Chemical equilibrium and liquid - Liquid equilibrium in aqueous solutions of formaldehyde and 1-butanol, Industrial & Engineering Chemistry Research 42 (2003) 15081516. [87] ASTM D5705, Standard Test Method for Measurement of Hydrogen Sulfide in the Vapor Phases Above Residual Fuel, ASTM International, West Conshohocken, 2008. [88] G.N. Taylor, R.N. Matherly, The laboratory evaluation of hydrogen sulfide scavengers using sulphur specific flame photometric gas chromatography. Paper presented at: SPE-140401-MS, SPE International Symposium on Oilfield Chemistry; 2011 April 1113; The Woodlands, TX, USA. [89] C. Mendez, A.A.O. Magalhaes, P.A. Ferreira, et al., H2S scavengers injection: A novel evaluation protocol to enhance the integrity of production lines. Paper presented at: OTC-20168, Offshore Technology Conference, 49 May, 2009, Houston, Texas, USA.

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34

CHAPTER 1 Chemical scavenging of hydrogen sulfide and mercaptans

[90] C. Mendez, A.A.O. Magalhaes, P.A. Ferreira, et al., Development of a novel testing protocol and equipment for the evaluation of H2S scavengers. Paper presented at: NACE-2012-1466, Corrosion Conference & Expo, 1115 March, 2012, Salt Lake City, UT, USA. [91] D.M. Lillard, J. M. Macias, Improved treatment methodology for removing low concentrations of H2S in South Texas natural gas systems. Paper presented at: SPE94303-MS, SPE Production and Operations Symposium; 2005 April 1719; Oklahoma, Oklahoma, USA. [92] C. Buchan, Modeling H2S scavenger efficiency cuts costs, Journal of Petroleum Technology 59 (2007) 2629. [93] M.E.M. Tawik, B. Al-Matar, F.A. Al-Faresi, et al., Unlocking reservoir potential using an automated multistage H2S technology  first in onshore Kuwait. Paper presented at: SPE-182941-MS, Abu Dhabi Petroleum International Petroleum Exhibition and Conference; 2016 November 710; Abu Dhabi, UAE. ´ vsthus, et al., Design principles for H2S scavenger [94] M.G. Lioliou, C.B. Jenssen. K. Ø injection. Paper presented at: Oilfield Chemistry Symposium, 2018 March 47; Geilo, Norway. [95] J.G.R. Eylander, H.A. Holtman, T. Salma, et al., The development of low sour gas reserves utilizing direct injection liquid hydrogen sulphide scavengers. Paper presented at: SPE-71514-MS, SPE Annual Technical Conference Centre, 2001 30 September3 October; New Orleans, Louisiana, USA. [96] F. Chaudhry, G.L. Mobley, Y.H. Tsang, et al., Laboratory development of a novel non-triazine based hydrogen sulfide scavenger and field implementation in the Haynesville shale. Paper presented at: SPE-164077-MS, SPE International Symposium on Oilfield Chemistry; 2013 April 810; The Woodlands, TX, USA. [97] T.M. Elshiek, H.A. Elmawgoud, S.A. Khalil, et al., Simulation for estimation of hydrogen sulfide scavenger injection dose rate for treatment of crude oil, Egyptian Journal of Petroleum 24 (2015) 469474. [98] J.J. Howell, M.B. Ward, The use of acrolein as a hydrogen sulfide scavenger in multiphase production. Paper presented at: SPE-21712-MS, Production Operations Symposium; 1991 April 79; Oklahoma City, Oklahoma, USA. [99] T. Salma, Cost effective removal of iron sulfide and hydrogen sulfide from water using acrolein. Paper presented at: SPE59708, 2000 SPE Permian Basin Oil and Gas Recovery Conference; 2000 March 2123; Midland, TX, USA. [100] M. Lehmann, T. Pojtanabuntoeng, Y. Long, et al., Evaluation of hydrogen sulfide scavengers for use with static mixers. Paper presented at: OTC-30057-MS, Offshore Technology Conference Asia, Offshore Technology Conference (Virtual); 2020 October 27; 2020. [101] A.A. Dugar, M.J. Villalobos, Case history: using triazine efficiently to scavenge hydrogen sulfide in coiled tubing drilling operations. Paper presented at: SPE/ IADC-166761-MS, SPE/IADC Middle East Drilling Technology Conference; 2013 October 79; Dubai, UAE. [102] J. Draper, J. Stark, Flare enough, Hydrocarbon Engineering (2011). September. [103] S. Ramachandran, S. Lehrer, S. Chakraborty, et al., New, non-corrosive, nonnitrogen containing H2S scavenger use in predominately oil and gas systems. Paper presented at: NACE-202014618, Proceedings of Corrosion Conference & Expo; 2020.

References

[104] S. Lehrer, V. Jovancicevic, S. Braman, et al., Downhole sweetening of mixed production by application of a new hydrogen sulfide scavenger. Paper presented at: SPE-173613-MS, SPE Production and Operation Symposium; 2015 March 15; Oklahoma City, Oklahoma, USA. [105] S. Chakraborty, S. Lehrer, J. Jani, et al., Fit for purpose hydrogen sulfide management for mixed production. Paper presented at: NACE-202014842, Proceedings of Corrosion Conference & Expo; 2020. [106] A. Konczvald, The eight best H2S removal methods in federal waters, offshore California. Paper presented at: OTC-8505-MS, Offshore Technology Conference; 1997 May58; Houston, TX, USA. [107] J.E. Kattner, A. Samuels, R.P. Wendt, Resolutions of problems with iron oxide slurry removal of H2S gas from natural gas  case history, Eva O’Byrne gas unit, Upshur County, Texas, Journal of Petroleum Technology 40 (1988) 12371243. [108] D. Leppin, Natural gas production: performance of commercial technology for removing small amounts of hydrogen sulfide. Paper presented at: SPE-29743-MS, SPE/EPA Exploration & Production Environmental Conference; 1995 March 2729; Houston, TX, USA. [109] M. Schulz, Development in H2S scavenging, Petroleum Technology Quarterly (2018) 1524. June Issue. [110] T. Salma, M.L. Briggs, D.T. Hermann, et al., Hydrogen sulfide removal from sour condensate using non-regenerable liquid hydrogen sulfide scavengers: a case study. Paper presented at: SPE-71078-MS, SPE Rocky Mountain Petroleum Technology Conference; 2001 May 2123; Keystone, Colorado, USA. [111] C. Passucci, G. Grazoli, C. Carniani, Innovative mercaptan scavengers: first field application. Paper presented at: IPTC-13362-MS, International Petroleum Conference; 2009 December 79; Doha, Qatar. [112] Z. Xiao, L. N. Morgenthaler, M.K. Cowan, et al., Evaluation of H2S scavenger for matrix acidization treatment. Paper presented at: SPE-121698-MS, SPE International Symposia on Oilfield Chemistry; 2009 April 2022; The Woodlands, TX, USA. [113] Y.K. Al-Duailej, S.H. Al-Mutairi, A.Y. Al-Humaidan, Evaluation of triazine-based H2S scavengers for stimulation treatments. Paper presented at: SPE-136915-MS, SPE/DGS Annual Technical Symposium and Exhibition; 2010 April 47; AlKhobar, Saudi Arabia. [114] A. Patel, J. Green, J. Conteras, et al., Physical fouling inside a crude unit overhead from a reaction byproduct of hexahydro 1,3,5-(2-hydroxyethyl)-s triazine. Paper presented at: NACE-20156054, Corrosion Conference & Expo; 2015 March 1519; Dallas, TX, USA. [115] G.N. Taylor, P. Prince, R. Matherly, et al., The formation and chemical nature of amorphous diathiazine from use of hexahydrothiazine based hydrogen sulfide scavenger and the use of an alternative scavenger to minimize solid formation in sour gas applications. Paper presented at: SPE-164134-MS, SPE International Symposium on Oilfield Chemistry, 2013 April 810; The Woodlands, TX, USA. [116] L. Kamin, Leveraging the latest advancements in automation and digital technology to improve efficiency and safety in a production field: a journey towards unmanned operations. Paper presented at: SPE-203198-MS, Abu Dhabi Petroleum International Petroleum Exhibition and Conference; 2020 November 912; Abu Dhabi, UAE.

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CHAPTER

Natural gas sweetening

2 Sebastien Duval

Research & Development Center, Saudi Aramco, Dhahran, Saudi Arabia

Chapter Outline 2.1 2.2 2.3 2.4

2.5

2.6 2.7

2.8

2.9

Introduction ................................................................................................. 38 Gas conditioning to satisfy sales gas quality .................................................. 39 Natural gas sweetening methods ................................................................... 42 Chemical absorption ..................................................................................... 43 2.4.1 Chemical reactions between H2S and CO2 and amine .....................44 2.4.2 Amine process overview ...............................................................47 2.4.3 Design best practices ..................................................................49 Physical absorption ...................................................................................... 51 2.5.1 Propylene carbonate process ........................................................52 2.5.2 Dimethyl ether of polyethylene glycol (DEPG or DMEPEG) solvents ..53 2.5.3 N-Methyl-2-pyrrolidone ................................................................53 2.5.4 Refrigerated methyl alcohol (methanol) .........................................53 2.5.5 Combined physical and chemical absorption ..................................54 Adsorption ................................................................................................... 54 Permeation or membrane based technologies ................................................. 55 2.7.1 Principle ....................................................................................57 2.7.2 Polymeric membrane type ............................................................58 2.7.3 Membrane module types ..............................................................59 2.7.4 Gas pretreatment ........................................................................59 Sulfur recovery ............................................................................................. 60 2.8.1 Thermal section ..........................................................................61 2.8.2 Catalytic section .........................................................................62 2.8.3 Major equipment .........................................................................63 2.8.4 Quality of the acid gas .................................................................64 2.8.5 Reduction absorption tail gas treatment ........................................64 Emerging approaches for treating highly sour gas .......................................... 65 2.9.1 Cryogenic distillation ...................................................................66 2.9.2 Membranes for high H2S .............................................................66

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00007-7 © 2023 Elsevier Inc. All rights reserved.

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CHAPTER 2 Natural gas sweetening

2.10 CO2 capture technology at gas plant .............................................................. 67 2.10.1 CO2 capture from flue gas ..........................................................69 2.10.2 CO2 captured from the acid gas stream .......................................70 2.11 Final remarks ............................................................................................... 76 Nomenclature ........................................................................................................ 76 References ............................................................................................................ 76

2.1 Introduction Natural gas currently accounts for 23% of the world’s energy supply, as reported by the 2021 Annual Energy Outlook [1]. In addition, natural gas receives also a lot of attention as a feedstock for chemistry [2]. Eventually, to be commercially attractive and therefore produced, a natural gas reservoir must contain a large volume of extractable paraffinic hydrocarbons [3,4]. The most abundant of these hydrocarbons is methane. The current understanding is that methane is formed by the degradation of organic matter (plants or/and animals) under pressure and temperature when entrapped in sediment layers [5]. Depending on the pressure and temperature history, depth of burial, and plate tectonics, the original macromolecules issued from plants and animals are cracked and form smaller molecules like methane (CH4). Another significant source of methane is the result of microbial activity close to the surface. This is the biogenic gas that has also a lot of attention in the desire to reduce greenhouse gas emissions [5]. The heavier hydrocarbons ethane (C2H6), propane (C3H8), butane (C4H10), pentane (C5H12), hexane (C6H14), and the heavier hydrocarbons are generated only by thermal maturation of organic matter. The different stages of the evolution of kerogen, e.g., diagenesis, catagenesis, and metagenesis, as well as the analysis of the isotopic ratio to identify the source of hydrocarbons, are available in dedicated books for crude oil and gas formation. Natural gas also contains helium. Helium is produced by the alpha decay of 238 U or 230Th and migrates to the surface unless blocked by an impervious layer. Helium is a minor component of natural gas and its concentration ranges from a few ppm to several percent [6]. It is a crucial element for modern societies and is used in liquid form to maintain magnetic resonance imaging magnets in a superconductivity state or gas form as a shielding gas in welding works. Another element found in natural gas is nitrogen, which originates both from bedrock as geologic nitrogen and from organically bound organic matter maturation or atmospheric nitrogen released from sedimentary and metasedimentary deposits. The presence of carbon dioxide in the natural gas results from the degradation of organic matter and the elimination of the carbonyl and carboxylic groups. Another source of CO2 is the release of rocks such as limestone or dolomite under acidic conditions, as well as volcanic activity. The origin of H2S is attributed to the thermal cracking of sulfur compounds present in the organic matter in a reductive environment. H2S is also produced by

2.2 Gas conditioning to satisfy sales gas quality

sulfate-reducing bacteria in anaerobic conditions. Above a certain concentration, H2S becomes lethal to the bacteria so this process alone cannot explain the presence of H2S in some reservoirs in tens of percent. The souring of sweet reservoirs due to improper water treatment is worth mentioning. If produced water or seawater with insufficient biocides is injected into a reservoir formation to maintain pressure and productivity, sulfate-reducing bacteria can produce H2S in hundreds of ppm [7]. For a facility designed to operate in non-sour conditions, expensive retrofitting might be required to comply with sour gas production and associated metallurgy and sulfur disposal. The formation of H2S is also attributed to thermochemical sulfate reduction (TSR), which involves the reaction of CH4 with CaSO4 (rock) to form CaCO3, water, and H2S. For the TSR model, H2S concentration as high as 95% can be found [8]. If formed in a permeable layer, the hydrocarbons and other components will migrate upward to the surface unless they are trapped by an impervious layer and accumulate over millions of years. The extraction of these hydrocarbons requires drilling a well from the surface to below the impervious layer where the gas collects. Due to the high pressure, gas flows upwards to the surface. Water can be injected to maintain the gas cap pressure and gas volumetric flow rate. If the gas and oil phases are formed in the impermeable rock, molecules are entrapped in small cavities until released by fracking. For example, oil shale layer fracking is performed by injecting water to enable the interconnectivity of the cavities and create a path to the well and then the surface for further processing [9]. Some examples of raw natural gas compositions are provided in Table 2.1. Gas with a high C2 1 content is categorized as rich gas. The H2S concentration range provides the sourness ranking. In absence of H2S, gas is defined as sweet gas. Although the compositions are on a dry basis, most raw gases are watersaturated.

2.2 Gas conditioning to satisfy sales gas quality Fig. 2.1 presents an overview of the different gas sources in gas processing facilities where raw natural gas will be purified and treated before reaching saleable quality and maximizing the plant revenue [3]. The gas entering a gas plant is routed to a vessel that collects liquids that might have been formed or entrained along with the natural gas from the reservoir. These liquids can be hydrocarbon condensates or water. The vessel can be a three-phase separator that allows water and condensate to separate and be routed to different units. Liquid hydrocarbons are sent to a stabilization plant to strip the light components from the condensate stream and produce a stabilized condensate stream. The water is sent to a sour water stripper if H2S is present in the stream. Downstream of the three-phase separator, natural gas is routed to a cooler to reduce its temperature before entering the sweetening step. The different

39

Table 2.1 Examples of raw natural gas composition (dry basis), heating value, and molecular weight for major components found in natural gas. Component (dry basis) C2 1 content sourness

Methane Ethane Propane Isobutane n-Butane Isopentane n-Pentane Hexane and heavier Hydrogen sulfide Carbon dioxide Nitrogen

C1 C2 C3 iC4 nC4 iC5 nC5 C6 1 H2S CO2 N2

AG, Associated gas; NAG, non-associated gas. a Estimated.

AG

AG

AG

NAG

NAG

NAG

Heating value

Molecular weight

Very rich

Rich

Very rich

Lean

Lean

Very lean

(BTU/ SCF)

(g/mol)

1010 1769 2517 3253 3262 4000 4010 5000a 637 0 0

16.04 30.07 44.10 58.12 58.12 72.15 72.15 est. 90 34.08 44.01 28.01

High

Medium

Medium

None

Medium

High

44.49 20.76 12.53 0.97 3.01 0.54 0.72 0.57 9.15 6.86 0.40

71.48 10.43 6.07 0.73 1.61 0.40 0.40 0.14 4.15 3.97 0.63

48.86 18.41 11.67 1.09 3.64 0.76 1.20 1.07 2.90 10.06 0.34

78.23 9.39 2.92 0.55 0.90 0.40 0.32 3.43 0.00 1.80 2.06

68.43 5.91 2.47 0.43 0.79 0.23 0.23 2.46 4.55 3.12 11.38

51.25 3.32 1.03 0.3 0.32 0.15 0.16 0.45 21.5 7.35 14.17

2.2 Gas conditioning to satisfy sales gas quality

FIGURE 2.1 Overview of natural gas source and processing.

processes of the sweetening step are described later in this chapter. The function of the sweetening unit is to remove H2S or other sulfur contaminants like carbonyl sulfide (COS), carbon disulfide (CS2), or mercaptans (RSH) from the natural stream. CO2 concentration is also reduced in the sweetening unit; the advancement of formulated amines enables control of the CO2 slippage while maintaining complete H2S removal. H2S and CO2 are referred to as acid gases because they act as weak acids when dissolved in water. The pH of condensed water is 4 when exposed to 15 psi of H2S or CO2 [10]. The disposal of a large quantity of H2S, e.g., more than 20,000 lb/day, is carried out in a sulfur recovery plant or a dedicated sulfur recovery unit (SRU) sometimes referred to as a “Claus unit.” that converts H2S into elemental sulfur [11]. For very small quantity # 1000 lb/day, H2S can be disposed of by non-regenerated chemical sorbents such as iron sponge. For an intermediate quantity, elemental sulfur can be produced by a redox process in which H2S is oxidized and then reduced in solution. Upon exiting the sweetening unit, natural gas may require additional steps before reaching sales gas quality. If an amine process is used for the sweetening, then the gas will be saturated with water (  40008000 ppm) and well above the specification of sales gas which requires a water content # 4 or 7 lb/MMSCFD (  80150 ppm) depending on the location and the expected external temperature. This requirement to reduce water concentration aims at minimizing the risk of water condensation in the pipeline since the water dew point will be # 40
F for the water content of 7 lb/MMSCFD. The predominant process for natural gas dehydration is glycol-based absorption. Alternatively, if the gas is routed to a cryogenic natural gas liquid recovery, an adsorbent such as a molecular sieve and silica gel can be used to reduce the water content to below a few ppm.

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CHAPTER 2 Natural gas sweetening

The removal and recovery of heavy hydrocarbons from natural gas are often necessary to meet a sales gas hydrocarbon dewpoint specification or to increase plant income due to the price differential between natural gas and heavy hydrocarbons. Since the price of natural gas is $/MMBTU. The removal of heavy hydrocarbon reduces the heating value of the gas and hence the revenue from the gas sale. Revenue from C2 1 should offset the investment and operating costs associated with liquid recovery. Natural gas liquid recovery is carried out at low temperatures. Therefore, and prior to entering the cryogenic plant, any trace of mercury must be removed to avoid catastrophic liquid metal embrittlement due to the reaction between liquid mercury and the brazed aluminum used in the heat exchangers. This mercury is removed by circulating the natural gas through sulfur impregnated activated carbon beds. The recovery of ethane 1 also requires a very low temperature, which is achieved with turbo expanders in which the natural gas undergoes an isentropic expansion. Part of the energy is recovered to recompress the gas after the liquid is separated from the gas phase. The liquids are then fractionated to separate ethane, propane, butane, and pentane 1 in a series of distillation towers [12]. It may also be necessary to reduce the nitrogen content to lower the overall inert content below the sales gas specification and meet the heating value. Nitrogen content can be reduced with a membrane or adsorbent, but the most prevalent nitrogen rejection technology is cryogenic separation. Due to their difference in boiling point, methane and nitrogen can be separated: a distillation process ensures high purity methane and minimizes the methane loss with the rejected nitrogen. Advantageously, helium is concentrated in the overhead stream of the nitrogen rejection unit and can be separated from nitrogen with a membrane and adsorbent. Helium is then commercialized after liquefaction at 2452
F [6].

2.3 Natural gas sweetening methods Sweetening consists of separating and removing H2S from the gas stream. By extension, the sweetening includes the complete or partial removal of CO2 and minor components like COS, CS2, or mercaptans RSH (R: alkane chain). The process selection depends on the gas flow rate, and the amount, and type of acid gases to be removed. The description of the sweetening process of natural gas is provided in numerous books or book chapters [9,1319]. The sweetening cases can be listed as follows, where: • • • •

Only CO2 is present and needs to be partially or totally removed; Both CO2 and H2S are present and both need to be removed; Both CO2 and H2S are present and only H2S needs to be removed (selective removal); Only H2S is present and needs to be removed, (mostly met in refineries because of the hydrotreatment activities).

2.4 Chemical absorption

The main processes for sweetening can be categorized as absorption (chemical /reactive), physical, and combined (chemical and physical); physical adsorption (physisorption); and membrane gas separation (membrane permeation or membrane separation) Each type is introduced in the following Sections 2.42.8.

2.4 Chemical absorption The general flow of the chemical absorption process is presented in Fig. 2.2. The ideal absorbent is a chemical that can weakly bond with acid gases so that breaking the bond does not result in a high energetic penalty. Among the weak bases that can perform this desired absorption, the aqueous alkanolamine solutions are the most frequently used. Alkanolamine molecules contain an amine functional group (nitrogen atom linked to three carbon or hydrogen atoms) and an alcohol functional group that will favor solubility in water of the amine and formed salts. The strength of the amine and acid gas bond varies with their chemical interaction, which depends on the number of hydrogen atoms linked to the nitrogen atom. The most common alkanolamines used in the sweetening of natural gas are: •

Primary amines (RNH2 type) Monoethanol amine (MEA): 2-aminoethanol or monoethanolamine (HOCH2CH2NH2) Diglycol amine (DGA): 2-(2-aminoethoxy) ethanol or diethylenegycolamine or diglycolamine (HOCH2CH2OCH2CH2NH2)

FIGURE 2.2 Amine-based acid gas sweetening unit.

43

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CHAPTER 2 Natural gas sweetening

•

•

Secondary amines (RR0 NH type) Diethanol amine (DEA): 2,20 -aminodiethanol or dietanolamine (HOCH2CH2)2NH Diisopropanol amine (DIPA): N-propan-2ylpropan-2-amine or diisopropanolamine ((H3C)2CH)2-NH Tertiary amine (RR0 RvN) Methyl diethanol amine (MDEA): 2,20 -(methylazanediyl) di(ethan-1-ol), N-methyl diethanolamine (HOCH2CH2)2NCH3

2.4.1 Chemical reactions between H2S and CO2 and amine From their atomistic composition, H2S and CO2 do not follow the same reaction path towards a base such as an amine. •

Case of H2S H2S can directly exchange proton H1 with a base such as an amine whether it is a primary, secondary, or tertiary amine. 2 H2 S 1 RNH2 5 RNH1 3 1 HS

(2.1a)

2 H2 S 1 RR0 NH 5 RR0 NH1 2 1 HS 0

0

1

(2.1b) 2

H2 S 1 RR RvN 5 RR RvNH 1 HS

(2.1c)

The reactions (Eqs. 2.1a2.1c) are acid-base reactions and are exothermic; water is acting not only as a carrier for amine molecules, a source of the proton but also as a thermal buffer. H2S, converted to HS2 is trapped in the amine solution until the solution is regenerated by heat. •

Case of CO2 CO2 also has an acidic character but it requires first to bind with one water molecule H2O, the formed carbonic acid is a source of a proton, hence its acidic character (Eq. 2.2). CO2 1 H2 O 5 H2 CO3

(2.2)

Then H2CO3 will dissociate in H1 1 HCO32. The large difference in the mechanism between H2S and CO2 is that the reaction path for CO2 requires a hydration step before engaging the reaction with the base, and this step is slow [14]. In addition to the alkanolamine solution, potassium carbonate (K2CO3) is used for bulk removal of CO2 in the presence of low H2S content. Table 2.2 lists the application for each of the most used chemical solvents for gas sweetening, recommended chemical reactions, solvent concentration, acid gas loading, regeneration, degradation as well as reclaiming. To remove acid gases, intimate contact between the natural gas to be treated and the alkanolamine or hot potassium solution is necessary. The contact between the two phases is achieved in a column, called an absorber or

Table 2.2 Main properties of chemical solvents used in the oil and gas business for the natural gas sweetening. Primary MEA Application

DGA

Secondary DEA

Deep H2S Deep CO2 Not H2S-selective even if H2S reacts faster than CO2

Tertiary

Carbonate

DIPA

MDEA

“hot pot”

Deep H2S and COS H2S-selective since CO2 reacts very slowly

Deep H2S and H2S is selective since CO2 reacts very slowly, CO2 when activated (addition of DEA or piperazine)

Bulk CO2 preferably with very low H2S

R2R0 N 1 H2S R2R0 NHSH Bicarbonate salt R2R0 N 1 H2O 1 CO2 - R2R0 NH2HCO3

K2CO3 1 H2S KHS 1 KHCO3 K2CO3 1 H2O 1 CO2 - 2 KHCO3

3550 wt.% Typical at 50 wt.% or 15 mol.%

Typical at 30 wt.% or 5 mol.%

Rich up to 0.45 Lean

Not applicable, loads are in the range of 5 SCF/gallon at 30 wt%

Chemical reactions H 2S

RNH2 1 H2S - RNH3SH

R2NH 1 H2S - R2NH2SH

CO2

Carbamate formation 2 RNH2 1 CO2 - RNH2COORNH3 Bicarbonate salt RNH2 1 H2O 1 CO2 - RNH3HCO3

Carbamate formation 2 R2NH 1 CO2 R2NHCOOR2NH2 Bicarbonate salt R2NH 1 H2O 1 CO2 R2NH2HCO3

Concentration

1020 wt.% Typical at 15 wt.% or 5 mol.%

Typical 65 wt.%

Acid gas loading (mol acid gas/mol amine)

Rich up to 0.35 Lean

Rich up to 0.40 Lean

2535 wt.% Typical at 25 wt.% or 5 mol.% Rich up to 0.40 Lean

3040 wt.% Typical at 30 wt.% or 5 mol.% Rich up to 0.50 Lean

(Continued)

Table 2.2 Main properties of chemical solvents used in the oil and gas business for the natural gas sweetening. Continued Primary

Secondary

Tertiary

Carbonate

MEA

DGA

DEA

DIPA

MDEA

“hot pot”

0.05 for H2S 0.15 for CO2 225
F250
F at 310 psig

0.05 for H2S 0.10 for CO2 260
F280
F at 310 psig

0.05 for H2S 0.05 for CO2 270
F at 15 psig

0.05 for H2S 0.05 for CO2 270
F at 15 psig

0.02 for H2S 0.01 for CO2 270
F at 15 psig

250
F at 15 psig

Degradation

O2 irreversibly degrades amine COS and CS2 tend to form heat stable salts

O2 irreversibly degrades amine COS and CS2 tend to form heat stable salts

O2 irreversibly degrades amine

Reclaiming

Can be partially regenerated

Regeneration

No reclaiming since no degradation by COS and CS2

DEA, Diethanol amine; DGA, diglycol amine; DIPA, diisopropanol amine; MDEA, methyl diethanol amine; MEA, monoethanol amine.

Not susceptible to O2 degradation

2.4 Chemical absorption

contactor, equipped with packing or trays to ensure that the acid-base reaction between the weak acid (H2S and CO2) with the amine or potassium carbonate is taking place. In a regeneration column, the acid gases are stripped from the amine or carbonate potassium solution, resulting in an acid gas stream and a regenerated absorbing solution, which is ready for another cycle of absorption. In the following sections, the amine process is detailed.

2.4.2 Amine process overview Fig. 2.2 provides an overview of the major equipment of amine-based gas sweetening. Although minor variations can occur due to differences in amine concentration, reactivity, and acid gas load, the process follows the following steps: Inlet separator/scrubber: First, the untreated gas enters the inlet separator. The function of the inlet separator is to remove any entrained liquid carried over from upstream vessels or picked up as condensate in the piping due to cooling. The inlet separator is also equipped with a filter to prevent fine particles or entrained solids like detached corrosion products from penetrating the absorber. Fine particles are prone to induce foaming in the unit and impair its scrubbing capacity and operation. The guidance in designing gas-liquid separators can be found in [20]. Due to the expected low liquid flow rate, as compared to the gas flow rate, a double-barrel filter horizontal separator can be considered, equipped with filter tubes to remove fine particles and a mist extractor. Amine process absorber (contactor): The gas, purged of liquids and particles, enters the absorber vessel at the lower part below any tray or packing. The gas flows upward in the column and interacts with the aqueous amine solution. The column is equipped with trays or packing to increase the contact surface between the vapor and liquid phases while minimizing froth and amine entrainment with the gas. The absorption process consists of several steps, of which the transfer of acid gas molecules across the interface between the liquid and vapor phases is critical. The amine-depleted (“lean amine”) acid gas enters the absorber at the top section and circulates downward by gravity, creating a counter-current flow with the gas stream. At the bottom of the absorber, the amine solution exits loaded with acid gas (“rich amine”), while the sweetened gas exits at the top of the absorber. In addition to creating interfaces between phases, the absorber internals ensures that both phases are well stirred to minimize concentration polarization for the reactive components. In order to monitor any foaming or fouling or trays damages, pressure differential gauges and transmitters must be installed and the value closely monitored. Foaming manifestation can be deduced from instable reading. There is also a requirement for controlling the temperature of the lean amine. Contacting the gas with a low-temperature amine favors acid pick-up, but the heavy components of the natural gas may condense in the amine. The condensation of these heavy hydrocarbons, especially aromatics, can have detrimental effects,

47

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CHAPTER 2 Natural gas sweetening

such as inducing excessive foaming of the amine, requiring a larger separator upstream regenerator to avoid sending hydrocarbons to a regenerator and then to the sulfur recovery plant. Conversely, operating with high-temperature lean amine results in water transfer from the amine solution to the sweetened gas, since the water solubility of the natural gas increases with temperature. This transfer necessitates frequent fresh water make-up to the amine system and generates an additional load on the dehydration unit downstream of the sweetening. To reduce the loss of amine, the top of the absorber can be equipped with an additional section, where fresh water is circulated to capture the amine from the gas stream. The gas stream then circulates to a KO drum equipped with a demister pad to collect any liquid droplets from the gas phase. The treated gas is routed to the downstream process, which involves at least a dehydration step to meet the sales gas specification or a liquid recovery or nitrogen removal unless the gas is directly utilized for power generation at the site. The basis for absorber design, selection of the type of internals, and calculation of absorption process is found in numerous books and book chapters, in particular [9,1316]. Amine flash drum: The amine solution is collected at the bottom of the absorber and pumped to the amine flash drum, where the rich solution containing the acid gas and some co-absorbed hydrocarbons is subjected to pressure reduction. This basic unit operation, referred to as flash or partial evaporation, causes the methane, light, volatile hydrocarbons, and a small fraction of the acid gas to be quickly released (“flashed”). The flash gas enters a small absorber to cleanse (“scrub”) any released acid gases with a slip stream of lean amine. The treated gas is either recompressed and sent to the treated gas, or, more usually, used as a plant fuel gas at low pressure. The amine flash drum is also designed to allow physical separation between the light hydrocarbon phase and the aqueous amine solution. Similarly, in a crude oil-water gravity separator vessel, the hydrocarbon phase carries over a baffle and is sent to a dedicated sour condensate handling system (e.g., a condensate stabilization plant). Lean amine/rich amine exchanger (economizer): To reduce the heating and cooling demands a lean amine/rich amine exchanger (“economizer”) is employed on both the regeneration and chilling sides. The lean amine/rich amine exchanger or economizer aims at reducing the heating demand on the regeneration side as well as the cooling demand on the lean amine cooler/chiller. The heat exchanger is, in general, of shell and tube type and the practice is to route a rich solution on the tube side and a lean solution on the shell side to minimize the corrosion erosion phenomenon. Stripper or regenerator: Downstream of/from the economizer, the scrubbing functionality of the amine solution is regenerated by the application of heat to destabilize the acid-base bond and create a stream of steam that strips acid gases out from the amine solution. The process takes place in a column equipped with trays or packing beds. The stripped amine solution is usually heated with

2.4 Chemical absorption

de-supersaturated low-pressure steam in a kettle-type reboiler. Control of the tube temperature is important to avoid amine degradation due to excessive temperature. The regenerated amine solution is then routed to the economizer and either a tank or a cooler (air or water or both) before entering the absorber for another cycle of absorption. Stripper condenser and reflux drum: The overhead stream of the regenerator is a mixture of steam, stripped acid gases, and traces of amine. To recover the water, this stream is cooled with either an air or water cooler, and the condensed water is collected by the reflux drum. The water is then pumped and routed back to the top of the stripper. Reflux ratios of acid gas water per mole of acid gas range from 1 to 3. Filtering system on lean amine: As mentioned in the inlet separator section, the presence of solid particles shall be minimal since it is a source of foaming and erosion in the amine system. A common practice is to send 10%20% of the solution flow rate to the filter system. The filter system shall be sized to remove any particles larger than 10 μm. A good and safe practice is to add an activated carbon bed to remove any other contaminants, such as entrained hydrocarbons and/or an excess of defoaming or surface-active agents. Reclaimer: The amine solution is ’rejuvenated’ by the thermal amine reclaimer which collects impurities that exacerbate corrosion, foaming, and fouling. These high boiling point impurities are accumulated in the reclaimer during a hightemperature semi-continuous batch distillation and then discharged. Surge tank: Safe storage of the amine solution during any maintenance activity or emergency is provided by a surge tank. In addition to mitigating sudden pressure fluctuations, the surge tank is also an integral part of the process, and hydrocarbon blanketing should be employed in the tank headspace. Blanketing with a covering layer of an oxygen-free gas such as nitrogen or sweet gas is required because the amine solution degrades and forms corrosive salts when exposed to oxygen. Water make-up system: Fresh demineralized water injection is required to remediate the water loss in the absorption and regeneration process. The purity of water is critical to avoid the accumulation of corrosive salts in the amine solution.

2.4.3 Design best practices 2.4.3.1 Process unit design The most critical units in the process design of chemical and physical absorption plants are the absorber and stripping vessels that must be sized with sufficient capacity to ensure that gas purification is achieved. Many reference textbooks [1620] guide establishing the design of absorbers and strippers with respect to size (diameter and height), number of trays, or packing types. The loading capacity of the selected solvent (see Table 2.2), along with the required regenerated

49

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purity, is used to set the solvent flow rate. Acid gas loadings are traditionally limited by recommended practices, to prevent excessive corrosion in carbon steel units. Heating and cooling process heat duties are calculated based on system modeling of solvent composition and properties and process parameters such as flow rate and loading. Preliminary design stages may refer to existing plants with comparable flow rates, acid gas compositions, and targeted purity and derive from these data the equipment sizing, flow rate, and process heat duties for reboilers and heat exchangers and their water cooling requirements. With a high-level plot plan and the main equipment listed, a cost can be estimated with an accuracy of 6 30% to support process pre-selection and establish project value [21,22]. With the development of simulations replacing the tedious and necessarily simplified hand calculations, accurate process models can be established. Some software packages provide main equipment sizing and cost estimates for the equipment, interconnecting piping, process instrumentation, and controls. For a more accurate design or an existing plant optimization or troubleshooting, a rate-based modeling approach that includes the resolution of the actual mass and energy balance in each tray or packing height section is preferable. Such retrofit or optimization designs require numerous coupled equations, including thermal (exothermic absorption), change of state due to the temperature variation, chemical, and mass transfer terms. Hence an advanced modeling program is required to solve the complex set of simultaneous equations (equation system) [23,24].

2.4.3.2 Main operation issues The main operational issues in meeting amine process gas specifications are the difficulties that arise as a result of the presence of fine particles in the amine, excessive foaming, and poor hydrocarbon skimming [25], which creates a problem for the SRU which process acid gas streams containing H2S and convert H2S into elemental sulfur. Most of the fine particles in the amine solution are corrosion products detached by high shear stress from carbon steel in the unit. Recommended practice is to design the unit, (heat exchanger and pipe sizing) to keep the flow speed # 6 ft/s to minimize the erosion-corrosion. The monitoring of the corrosion is recommended at adequate locations [13]. The particles are normally removed from the amine solution with a filtering system. Monitoring the plugging rate of filter by monitoring the pressure drop across the filter helps in identifying dirty amine. Recommended practice is that the filtering unit treats 10% of the amine flow rate with one filter and one in standby as with any rotating equipment like pumps. Amine process foaming can be attributed to several factors, including particulates and reagents like surfactants, corrosion inhibitors, cleaning agents after major maintenance, and sometimes an excess of anti-foaming agents. Circulating the amine solution through an activated carbon bed removes such agents and

2.5 Physical absorption

reduces the carry-over of adsorber foaming. Such carry-over can result in failure to meet targeted purification standards, as well as in the regenerator and overhead condenser issues as foaming alters the stripping process. The installation of an amine flash drum is strongly recommended to reduce hydrocarbon slippage in the SRU. A three-phase separation on the rich amine loop is also highly recommended, and a proper residence time of 1520 minutes will allow good separation between the amine solution, the denser phase, and the hydrocarbon phase. Turbulence in a vessel can be effectively managed with baffles to regulate and direct flow and weirs to constrain and restrict flow, reducing carry-over to the regenerator and the SRU. Hydrocarbon condensation inside the absorber should also be prevented. Recommended practice is to set the lean amine temperature at 10
F above the feed gas temperature because the concentration of hydrocarbons in the gas phase increases during the acid gas removal (AGR) phase, lowering its hydrocarbon dew point temperature. At the lower part of the absorber, the heat released by the absorption prevents condensation but higher in the absorber, heavy hydrocarbons may start condensing and eventually become a source of foaming. In addition to proper corrosion monitoring, the amine process absorber should be instrumented with pressure differential transmitters to detect foaming or flow circulation issues within the column. A very important parameter to monitor is the temperature at the different trays. Typically, an absorber has 20 trays, plus 24 for the top water washing. Monitoring the temperature profile of the column will help ensure proper operation. For instance, a small reduction of amine flow rate or an increase in acid gas load may have a dramatic effect as explained by Wieland [23]. The typical temperature profile within the absorption column shows a bulge, which should be maintained, in the lower third to half of the column. It is important to maintain the high temperature relatively low in the column so the lean amine will be able to capture the acid gas and the treated gas will meet the specification. Reducing the amine flow rate even by 5% can move the hightemperature bulge upward, which compromises the complete absorption unit performance, e.g., it fails to achieve 4 ppm of H2S in the treated gas. Reducing the amine circulation rate should be considered with great care to avoid sudden acid gas breakthrough.

2.5 Physical absorption The absorption relies on the ability of certain liquids to solubilize large amounts of acid gas and minimal amounts of the hydrocarbons contained in the natural gas. Solvents considered for the physical absorption of acid gases (CO2, H2S) should exhibit [13,16]: •

Low vapor pressure to minimize solvent loss during absorption and regeneration steps;

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CHAPTER 2 Natural gas sweetening

• • •

Low hydrocarbon solubility to minimize loss of high-value products and impair acid gas disposal option; High chemical and physical stability when exposed to any gas component and regeneration step; Nil reactivity with common metallurgy in petrochemical industries.

Most physical absorption-based processes operate at low temperatures because the solubility of acid gases increases as the temperature is reduced. The untreated gas enters the absorber at the bottom and flows upwards, like the amine process, and the solvent flows counter current and absorbs the acid gases. The solvent is regenerated by successive flash operations at gradually lower pressures, and in most cases in the final step, the solvent is heated to strip most of the acid gases. The initial high-pressure flash gas is rich in hydrocarbons and therefore is either recompressed and returned to the absorber or used as fuel gas if it meets the fuel gas specification. If it is off-specification, the rich high-pressure flash gas can be treated in a dedicated absorber where acid gases are removed with a fully regenerated solvent. Lower pressure flash gases are either vented or flared, depending on their composition. When H2S is present, the generated acid gases are routed to the SRU or flared if the SO2 emissions comply with applicable regulations. Physical absorption processes require less heat than chemical processes, so the energy required to regenerate the solvent is lower compared to that of the amine. Most conventional physical absorption processes are proprietary and are based on a variety of solvents such as: •

•

• •

Propylene carbonate (PC), a polar aprotic solvent and downstream product of propylene glycol (Fluor Solvent process, Fluor SolventSM, Arconate 5000t, and Texacar PC) Dimethyl ether of polyethylene glycol (DEPG or DMEPEG), another propylene glycol downstream product that is noncorrosive and nonfoaming (Fluor EconoSolvSM, Selexol Process, Selexol), N-Methyl-2-pyrrolidone (NMP), a non-volatile polar solvent (Purisol process, Purisol) Methyl alcohol (methanol) in a refrigerated solution (Rectisol).

2.5.1 Propylene carbonate process The use of polypropylene carbonate as an absorbent targeting CO2 for removal from high-pressure gas streams is commonly known as the Fluor Solvent process and has been used for half a century. Since physical absorption follows Henry’s law, which holds that the quantity of dissolved gas in a solution is proportionate to the partial pressure of that gas on the solution, a higher partial pressure for CO2 increases the process capacity to solubilize a large amount of gas at a given absorbent flow rate, making the process economically attractive when the partial pressure of CO2 is greater than 75 psi [15].

2.5 Physical absorption

2.5.2 Dimethyl ether of polyethylene glycol (DEPG or DMEPEG) solvents DPEG is a solvent containing asymmetrical (mixed) ethers that have a strong chemical affinity with H2S, CO2, COS, mercaptans, water, and heavy hydrocarbons. DEPG applications, such as the Selexol process, operate at low temperatures and require that the gas be dehydrated before being cooled to 40
F upstream of the absorber. Without dehydration, hydrate may form and water will accumulate in the solvent reducing its capacity. DEPG applications may include a split flow where only a portion of the solvent is subjected to a simple flash and returned to the absorber at an intermediate point in the absorber and the remaining portion of the solvent is subjected to a hot flash to remove most of the acid gases. This lean solvent is then injected into the top section of the absorber to provide the polishing stage. This split flow arrangement, which can also be utilized with the chemical absorption process, reduces the heating and cooling duties but increases operational complexity and capital expenditure. The DEPG process has limitations in rich gas sweetening because of the large solubility of hexane 1 in the solvent, which will be desorbed along with acid gases and may produce low-quality sulfur in the SRU.

2.5.3 N-Methyl-2-pyrrolidone NMP is a non-volatile polar solvent with a strong affinity for H2S and CO2. The solvent is subjected to an intermediate flash (partial evaporation) within an absorber or contactor (reabsorber), where released hydrocarbons from the rich solvent along with some acid gases are contacted with a lean solvent stream. Acid gases are then stripped from the rich solvent in a regenerator where steam provides the energy required to desorb the acid gases. NMP, like DEPG, has a strong affinity with heavy hydrocarbons and as a result, heavy hydrocarbons may be lost and sent to an SRU.

2.5.4 Refrigerated methyl alcohol (methanol) Low temperature (220
F to 215
F), methanol physical absorption applications for AGR, like the Rectisol process are more complex because deep refrigeration must be employed. Oxygen is highly soluble in methanol, and hence methanol storage tanks require high-quality inerting since any oxygen may react with H2S and result in the elemental sulfur formation and cause localized corrosion. Like the NMP process, the refrigerated methanol process regenerates the solvent in two steps: first, a single flash out of the dissolved hydrocarbons with the semi lean solvent recycled back to the absorber at an intermediate height, and second, a hot regeneration of the solvent to deliver a lean solvent. If the water collected by the methanol is not removed upstream of the unit, a slip stream of solvent is

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subjected to a methanol/water separation to dispose of water and maintain solvent absorption capacity.

2.5.5 Combined physical and chemical absorption Some applications combine the advantages of both chemical and physical absorption, namely the high absorption rate via the chemical bond and the low regeneration duty of the physical solvent. Such dual-function (combined) applications like the Shell plc Sulfinol processes (Sulfinol-D and Sulfinol-M) use a mixture of sulfolane (tetrahydrothiophene dioxide) and, respectively, an aqueous solution of DIPA and MDEA with a typical composition of 15 wt.% water, 45 wt.% amine and 40 wt.% sulfolane. The combined process operation is like physical solvent amine processes in which heavy hydrocarbons absorbed by the solvent may be found in the acid gas stream sent to the SRU and hence impair sulfur recovery operations. A reclaimer is recommended for processes using DIPA, which is subject to degradation with CO2; however, no reclaimer is necessary with processes that use MDEA, as it is not degraded by CO2.

2.6 Adsorption Separation by adsorption refers to the use of a solid material that exhibits a large surface area and affinity with the components to be removed [26]. The adsorbents are used for natural gas dehydration due to their unique capacity to reduce the water content to trace amounts before the cryogenic process. Adsorption can effectively remove H2S and CO2, but the process has a large physical footprint so adsorption is mostly used for the removal of trace contaminants. For water removal or AGR, the process involves two or three columns, one for the adsorption process itself, the second for the regeneration, and a third in standby. The process cycles between adsorption/and regeneration (generally 8 or 12 hours for both sequences). The adsorbent bed is regenerated by circulating either a slip stream of the raw gas or the treated gas itself. The slip stream is heated, (e.g., 550
F) with high-pressure steam and circulates to the bed for regeneration. Because the hot gas can hold a larger amount of water and acid gases than the gas to be treated, the adsorbent is regenerated. Acid gases removed from the adsorbent are sent for disposal, and after a cooling period, the bed is ready for a new adsorption cycle. The adsorption takes place in a fixed bed of microporous, semi-solid material referred to as a molecular sieve (or matrix), a synthetic metal aluminosilicate ˚ ). crystalline structure with well-controlled pore diameters of less than 2 nm (20 A Depending on the cation type of the material (Na, K, or Ca), the pores of a simple ˚. cubic Type-A crystals structure would have respective sizes of 3, 4, or 5 A

2.7 Permeation or membrane based technologies

Alternative arrangements such as a tetrahedral structure, of Type-X crystals result in larger pores: for instance, 10X and 13X have, respectively, pore sizes of 8 and ˚ . The adsorbents are traditionally manufactured as beads with sizes ranging 10 A from 1/16 to 1/2 in. in diameter. The water capacity of molecular sieves depends on the type and manufacturer, but the order of magnitude is 20 wt.%. The appropriate molecular sieve type must be selected for the intended service: ˚ type molecular sieves are suitable since only water for selective water removal, 3 A molecules, due to their size, can penetrate the pore and be adsorbed, unlike other gas components. H2S, CO2, and hydrocarbons molecules do not enter the pores and their adsorption is limited to the binder or the small fraction of large pores within the beads. For acid gas or mercaptan removal, molecular sieves with larger pores ˚, 5 A ˚ , or 13 X type should be selected. The adsorption is based on the like 4 A polarity of compounds therefore the adsorption occurs preferentially in the following order of adsorption: H2O, then H2S, and finally CO2. CO2 is not polar but has a quadrupole that induces some affinity for the molecular sieve surface. The gas enters the vessel at the top and moves downward, and following the order of adsorption, the water is removed first, and, since the flow is continuous, the incoming water molecules displace any adsorbed H2S molecules, which readsorb lower in the column (Fig. 2.3A). Due to its higher affinity with water, H2S breakthrough may occur quickly as breakthrough capacity is met, in which case, a ˚ adsorbent to remove water and a 4 A ˚ adsorbent for the capcombination of a 3 A ture of H2S may prevent a premature H2S breakthrough. During regeneration, the adsorbed H2S is released and concentrated. Typically, the regeneration gas, which is a 5%10% slip flow from treated gas has a concentration pulse in the early stage of regeneration that can be 10 to 50 times that of the untreated gas. In presence of wet CO2 and H2S gas, the co-adsorption of both gases can lead to the formation of COS during the hot regeneration step. As mentioned earlier, if the regeneration gas is sent to an amine plant. the facility needs to be designed to recover this COS. If it is an issue, certain grades of molecular sieves are designed to minimize the formation of COS during the regeneration phase. Alternative handling is presented in Fig. 2.3B, where the acid gas stream is sent to the SRU for disposal. The regeneration gas is the tail gas stream downstream from the tail gas treatment unit (TGTU).

2.7 Permeation or membrane based technologies Solution diffusion permeation (permeation separation) relies on a membrane or a dense nonporous layer of polymeric material across which natural gas components solubilize and then diffuse at different rates. The general flow of a permeationbased separation process is presented in Fig. 2.4 with the high-pressure gas entering the membrane module (membrane feed), the low-pressure gas stream

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FIGURE 2.3 Example of adsorption process (A). Dry gas is used for regeneration of adsorbent (B). Example for H2S removal from raw gas and regeneration (see after for sulfur recovery).

2.7 Permeation or membrane based technologies

FIGURE 2.4 Spiral wound membrane module.

(permeate), and the high-pressure gas stream leaving the membrane (retentate) [27,28]. Membrane separation technology has been implemented since the late 80s for bulk CO2 removal from natural gas with more than several hundred applications worldwide.

2.7.1 Principle The permeation separation process can be summarized as the succession of the following different steps in which gas molecules: • • • •

Adsorb on the membrane surface (high-pressure side) Solubilize within the external layer of the polymer Diffuse through the dense polymer layer through the fractional free volume Desorb from the polymer surface (low pressure).

To be attractive for the removal of acid gas, the polymeric material should have high selectivity for H2S and CO2 and be non-selective for other hydrocarbon gas components to minimize the loss of valuable products. The selectivity between two components A and B is defined as the ratio of the molar flows of the two components over a given membrane surface area, the membrane thickness, and the pressure difference across the membrane. The molar flow is a direct application of Fick’s first law of diffusion, where diffusion is the rate-determining step. It is then clear that the highest is the pressure difference between each side of the membrane and the thinnest is the membrane, the highest is the molar flow thorough the membrane for a given membrane surface area [27].

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2.7.2 Polymeric membrane type With respect to permeation and mechanical properties, membrane materials are classified into two families depending on the operating temperature and the glass transition temperature (Tg) of the selective polymer. The Tg separates two domains: at temperatures lower than Tg, the polymer exhibits rigidity and is referred to as “glassy;” at temperatures higher than Tg, polymer chains are more mobile, and the elastic and flexible mechanical behavior of the polymer is referred to as “rubbery.” This is an oversimplification, since a glass transition is kinetically controlled and not a second order phase transition, so Tg is better defined over a range of temperatures than an exact value. For a polymer operated above its Tg, the major parameter governing the permeation is the condensability of the gas components: the most condensable is the most soluble and has the most permeability. For rubbery membrane, the fastest components are ranked from left to right as follows: H2 O . C5 1 . C5 . C4 . C3 . H2 S . CO2 . C2 . C1 . N2

The removal of H2S with a rubbery membrane will lead to the removal of heavy hydrocarbons. If the membrane is used for fuel gas conditioning, a rubbery membrane may be suitable to simultaneously remove H2S and heavy hydrocarbons. For gas sweetening applications, a glassy membrane is a better option to avoid the loss of valuable heavy hydrocarbons. In fact, for polymers operated below their Tg, the major parameter governing the permeation is the size of the gas component or its molar volume estimated with the van der Waals equation. The smallest gas components have the most permeability. For glass membranes, the fastest components are ranked from left to right as follows: H2 O . CO2 . H2 S . N2 . C1 . C2 . C3 . C4 . C5 . C5 1

Note that the Tg value for a given polymer can vary due to the interaction between the polymer and the gas components. For example, water is known to significantly reduce the glass temperature of a polymer: a rule of thumb is that one percent of the water uptake by the polymer reduces the Tg by 30
F, so a water uptake of 2% will result in a Tg depression of 60
F [29]. Similarly, CO2 and H2S are also known to be plasticizers and therefore reduce the sieving properties of glassy materials [30], in other words, the selectivity in service. Heavy hydrocarbons may also induce plasticization and hence modify the separation performance. A polymer also undergoes physical aging, where the polymer chain packed with a certain pattern will tend to reduce its energy by relaxation, which is facilitated by plasticizers such as CO2 and H2S. Therefore, the actual permeation performance is very difficult to deduce from laboratory data because of the technical challenge in replicating the feed gas composition and maintaining the

2.7 Permeation or membrane based technologies

exposure long enough to make it suitable for assessing long-term performance. In addition, during the development of a gas project, exact gas compositions are rarely known at the time of selection of the sweetening process.

2.7.3 Membrane module types There are two major types of membrane arrangement for natural gas separation: spiral wound membrane and hollow fiber module. A spiral wound arrangement (Fig. 2.4) is a combination of flat sheets coiled around a central pipe that collects the permeated gas components. Spacers inserted between each layer maintain homogeneous gas flow distribution between each sheet and provide vorticity in the flow. This vorticity reduces the effect of the concentration polarization, which is inherent to all membrane processes, and results in lower separation rates at the interface of the gas and the selective layer. This is particularly important when membranes, are used to remove very dilute gas components (e.g., trace level helium) and components with high permeability, must operate with high removal efficiency. A hollow fiber arrangement has the membranes arrayed with hollow cylinders (external diameter u0.5 mm). This type of membrane arrangement offers a higher density of membrane surface area by volume and can be utilized with countercurrent flow, concurrent flow, and radial cross-flow.

2.7.4 Gas pretreatment The enrichment in components and cooling due to gas expansion can result in liquid formation in the membrane module, which can damage the membrane surface and cause loss of selectivity. A gas pretreatment is always required to avoid loss of separation performance or catastrophic failure of the membrane. Minimal pretreatment requires avoidance of carry-over of liquid as well as prevention of liquid formation in upstream piping and within the membrane module on the retentate side, so a coalescing filter is required to merge and collect entrained droplets and particulates. A gas heater is installed upstream of the membrane. Heating the gas is critical for a membrane in a glassy state to avoid condensation of heavy hydrocarbons which permeate at a much slower rate than other components and therefore accumulate on the high-pressure side of the membrane. Pretreatment can include a refrigeration step to condense heavy hydrocarbons and to maintain an optimal membrane operation (100
F140
F), in which case a liquid desiccant system (glycol dehydration) is necessary to avoid hydrate formation. Gas with high levels of acid gases can form hydrate at 70
F, and a safe operating margin calls for a minimal 40
F difference between the gas temperature and the hydrate dew point. Therefore, the pre-removal of water is mandatory to perform the refrigeration and prevent hydrates. The gas stream enters a glycol dehydration unit and is refrigerated to condense the heavy hydrocarbons. After the hydrocarbon collection, the stream is

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CHAPTER 2 Natural gas sweetening

sent to an economizer to cool the feed gas and save cooling duty, and then the gas is filtered and reheated to the desired operating temperature. However, such extensive pretreatment, if required, may reduce the cost advantage of the membrane process. Advantages of membrane-based processes include ease of installation and operation with proper pretreatment, low capital investment, modularity that facilitates process capacity adjustments, the avoidance of a large chemical inventory, and ease of disposal. There is also the likelihood of benefit from continuous product improvement. For example, cellulose acetate-based membranes can be replaced with new chemistry membranes using similar membrane housing. The limitations of membrane-based processes are related to the cost associated with a recycle compressor if required and the necessary pretreatment of the gas before entering the membrane. The cost of the installation is also linearly proportional to the flow rate due to its modularity; hence the little economy of scale is achieved.

2.8 Sulfur recovery Hydrogen sulfide is toxic and must be disposed of after being extracted from natural gas; it cannot be released “as is” and only very small quantities of H2S can be sent directly to a thermo-oxidizer and a stack for dispersion because produced sulfur oxides (SOx) are a major concern, and their emissions are strictly controlled to limit air pollution and acid rain. Preferential disposal relies on sulfur recovery and the modified Claus process to achieve a partial oxidation of H2S to elemental sulfur. This is the most common process to convert toxic H2S into a utilizable product: elemental sulfur [11]. A typical SRU includes two sections: first, a thermal section (temperature $ 1800
F) where H2S is oxidized to SO2 with some of SO2 and H2S reacting to form elemental sulfur, and second a catalytic section where the remaining H2S and SO2 are converted into sulfur at a lower temperature (  400
F650
F). The multi-step reaction identified by Claus, commonly and collectively referred to as the “Claus reaction,” is an equilibrium reaction and according to the Le Chatelier principle, the sulfur recovery is maximized by performing a series of catalytic stages with intermediate sulfur withdrawal. The acid gas content dictates the type of arrangement for an SRU. Fig. 2.5 shows a typical simplified flow diagram in which the thermal stage precedes three catalytic conversion stages. For rich acid gases or high H2S content, the design is straight through where the whole acid gas stream is sent to the reaction furnace. For leaner acid gases, the design is split flow, where a part, up to 60%, of the acid gas bypasses the reaction furnace. Bypassing the reaction furnace enables the reduction of inert content in the reaction furnace while achieving the complete

2.8 Sulfur recovery

FIGURE 2.5 Typical SRU—one thermal stage followed by three catalytic conversion stages. SRU, Sulfur recovery unit.

oxidation of a third of the H2S into SO2 and hence reduces parasitic reactions such as COS formation.

2.8.1 Thermal section Upstream of the reaction furnace, a knockout drum removes any entrained liquids from the AGR unit. These liquids may include condensed hydrocarbons or amines that would alter the operation or reduce the capacity of the SRU. Minimizing the introduction of hydrocarbons into the SRU reaction furnace is important to prevent increased air demand, avoid carryover of non-combusted hydrocarbons to the catalytic section, and minimize the formation of CS2. A high level of hydrocarbons in the acid gas could also reduce sulfur quality. In the reaction furnace, H2S from the acid gas is combusted with O2 from the air or oxygen-enriched air in a proportion that ensures that one-third (1/3) of the H2S in the acid gas is converted into SO2 and all present hydrocarbons and ammonia (NH3) are destroyed. The combustion of hydrocarbons produces CO2 and H2O and consumes O2; the combustion of ammonia produces N2 and H2O also consumes O2. The oxidation of H2S is shown below equation: 3H2 S 1 3=2O2 -SO2 1 2H2 S 1 H2 O

(2.3)

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The combustion of H2S is very exothermic [the higher heating value of H2S is 637 Btu/SCF (see Table 2.1)]. The produced SO2 reacts with the remaining H2S (2/3 of the inlet) according to the Claus reaction [11]. The reaction is an equilibrium reaction, and it is shown in Eq. (2.4): SO2 1 2H2 S 1 H2 O23S 1 3H2 O

(2.4)

The conversion is stoichiometrically sensitive; hence, the 1:2 ratio of oxygen to H2S must be maintained as close as possible. This control is usually achieved with a tail gas analyzer that monitors stream H2S and SO2 and sends a continuous demand signal to back-control the air inflow and maintain the 1:2 ratio. The chemistry inside the reaction furnace is complex and leads to the formation of different species. For instance, H2 is produced and is still present in the tail gas and helps in the reduction of sulfur compounds into H2S in the reduction absorption tail gas treatment. H2S and CO2 can react and produce COS. Hydrocarbons and H2S can react and form CS2. The practice is to operate the first catalyst bed at an outlet temperature of 650
F to hydrolyze COS and CS2 into H2S and CO2. If COS and CS2 are not hydrolyzed at this stage, they will be carried over to the tail gas and negatively impact the overall sulfur recovery. In the refinery, the elimination of ammonia from the sour water stripper is achieved in the SRU reaction furnace. The destruction of ammonia requires a high temperature ($2200
F), which can be achieved either with a good quality acid gas (H2S $ 80%) or a specific two-chamber reaction furnace. The two-chamber design enables the NH3-bearing gas outlet stream from the sour water stripper to reach a sufficient temperature before the remaining acid gas is injected into the second chamber. Downstream of the reaction furnace, the hot gas enters a waste heat boiler (WHB). The gas temperature is reduced to 550
F650
F and high-pressure steam is generated. The hot gas is further cooled to around 350
F favoring the sulfur condensation. Liquid sulfur is withdrawn and routed via the seal pot to liquid sulfur storage.

2.8.2 Catalytic section In the catalytic section, the Claus reaction is conducted at a lower temperature with the help of an alumina-based catalyst. High sulfur recovery is achieved by combining in-series catalytic conversion with interstage cooling, liquid sulfur formation, withdrawal, and then reheating. The general practice is to have two to three catalytic stages. There is a limit on the achievable recovery with this approach, which is due to the sulfur solubility as a vapor in the gas-exiting condenser. Operating the SRU with O2 enriched air reduces the overall tail gas flow rate, and hence, improves the recovery by reducing the sulfur slippage in the vapor phase. Three catalytic stages can achieve a 98% sulfur recovery. If higher recovery is needed to comply with emission regulation, tail gas treatment must be implemented. Because the Claus reaction equilibrium is favored by lower temperatures, the stream temperature should be as low as possible but high enough to avoid sulfur

2.8 Sulfur recovery

deposition in the catalytic bed, since this deposition reduces the catalyst activity. Typically, the first bed operates at an exit temperature of 650
F to favor COS and CS2 hydrolysis and the second and third beds operate at 440
F and 380
F, respectively. High-pressure steam reheats the stream to the desired temperature. Hydrolysis of COS and CS2 is instrumental in achieving high sulfur recovery, especially in the absence of tail gas treatment. Specific grades of catalysts, such as activated alumina and titanium dioxide, are required to improve hydrolysis.

2.8.3 Major equipment Burner: The operation of the reaction furnace and burner are critical for sulfur recovery. The design of the burner should allow sustainable and complete oxidation of a third of the H2S, and destruction of NH3 and hydrocarbons, if present, under the sub-stoichiometric air conditions required by the Claus reaction, which is partial oxidation of H2S into sulfur. In addition, no passage of oxygen to the catalytic stage is permitted, as it will deactivate the catalyst. Similarly, incomplete destruction of hydrocarbons will lead to deactivation of the catalyst and incomplete destruction of NH3 could lead to ammonium sulfate [(NH4)2SO4] deposition on the catalyst and/or the gas outlet demister pads and eventually create flow restriction or blockage. Reaction furnace and waste heat recovery: The reaction furnace is a refractory-lined vessel designed to ensure proper combustion of the contaminant and the desired level of sulfur oxidation. A waste heat recovery system is installed at the end of the reaction furnace. The waste heat recovery tube sheet is located at the end of the reaction furnace, sometimes preceded by a checkered wall to minimize radiation from the burner. The WHB simultaneously reduces the hot gas temperature and produces high-pressure steam up to 600 psi. Sulfur condensers: Sulfur condensers are heat exchangers that reduce gas temperature and enable sulfur condensation. The design should accommodate the proper drainage of condensed sulfur into the collecting pit. Like the reaction furnace and burner, proper design is essential to reduce operational issues such as proper line sizing and jacketing, steam tracing around process pipes, depth of the seal pot against the reaction furnace, operating pressure, and pocket avoidance. Catalytic bed: The catalyst beds are supported inside horizontal vessels. A deflection plate is installed to ensure proper flow distribution on the surface of the bed. Gas enters the vessel by the top and exits from the bottom. The bed should be equipped with temperature probes located at different depths in the bed, traditionally four. The monitoring of the temperature profile through the bed is instrumental in determining the in-situ catalyst performance and eventual deactivation. With a new catalyst, temperature across the bed increases quickly just below the top of the bed. For less active catalysts, the temperature increases at the lower part of the bed. In this case, the analysis of the reason for loss of performance should be performed to assess if it is due to sulfur deposition, which can be remediated by increasing the bed operating temperature.

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Reheater: There are two main methods to reheat the gas downstream of the condensers: direct heating or indirect heating. Direct heating includes using bypassed hot gas, but this reduces the overall recovery as compared to the indirect method. Alternatively, an inline burner can be used, which makes it easy to reach a high temperature, but there is a potential issue of soot formation and deposition on the catalyst bed during the start-up and shut-down phases, with also a potential for oxygen breakthrough. Preferably, the reheating is achieved indirectly with the help of a heat exchanger and steam as heating fluid.

2.8.4 Quality of the acid gas A reaction furnace fed with rich H2S ($55%) exceeds the temperature ($1900
F) required for the complete hydrocarbon destruction including benzene, toluene, and xylene (BTX) compounds. Consequently, the entire acid gas stream is sent directly to the reaction furnace. For a leaner acid gas, (H2S content 40% 55%), a straight-through SRU operation is still possible, but the acid gas stream must be preheated to ensure the reaction furnace temperature reaches the needed value for hydrocarbon destruction (insert temperatures). For leaner acid gas, the amount of CO2 and N2 prohibits reaching such a temperature. Alternative processes, such as the split flow approach, have been designed to operate at low H2S content. This arrangement reduces the dilution of gas in the reaction furnace and enables higher temperature and complete oxidation of up to a third of the H2S. To avoid catastrophic O2 breakthrough into the catalyst, the flow rate by pass is limited to 60% of the acid gas inlet flow. BTX is a strong catalyst deactivator, so the removal of BTX from the acid gas stream entering the catalytic section is crucial, and selective adsorption beds are commonly employed. An alternative to the bypass arrangement is the use of O2 enriched air to reduce the amount of inert gas in the SRU reaction furnace, which also increases SRU capacity and hence avoids building additional trains. However, the utilization of O2 enriched air may require costly replacement (and associated downtime) of a refractory burner with a reaction furnace for higher temperature operation. For lean gas, an option is to use acid gas enrichment (AGE) with an H2S selective amine as in the chemical absorption process. MDEA or hindered amine has the potential to enrich the feed to SRU. The enrichment can be three to fourfold leading to a rich acid gas for straight-through SRU operation. One advantage is the partial CO2 capture upstream of the SRU, where it is the most profitable since the SRU capacity increases. In line with increasing regulation related to CO2 capture, it is an attractive approach. Further optimization using novel and hybrid approaches is presented in Section 2.10.

2.8.5 Reduction absorption tail gas treatment An SRU equipped with three catalyst stages achieves a sulfur recovery of only 98% and will still emit a significant amount of SO2. Therefore, most of the new

2.9 Emerging approaches for treating highly sour gas

SRUs are combined with a TGTU to achieve 99.9% 1 recovery. The predominant process is a reduction absorption tail gas treatment that includes a catalytic hydrogenation step, where all sulfur compounds are converted into H2S at high temperatures (440
F480
F). The gas is then cooled to condense most of the water and meet the design temperature to enter the amine absorber. The TGTU uses a selective amine-absorbing solution that reduces H2S concentration to trace amounts (,75 mg/m3, 30 ppm). The captured H2S is recycled back to the SRU reaction furnace, and the clean tail gas is sent to the thermal oxidizer and stacked to disperse any trace of SO2.

2.9 Emerging approaches for treating highly sour gas More than 15% of proven gas reserves worldwide contain H2S in concentrations exceeding 10%. This large fraction of gas cannot be neglected, but its production is economically and technically challenging. The economic viability of sour gas production is impaired by the presence of high amounts of H2S, which means large investment for the gas sweetening, sulfur recovery, and tail gas treatment. Furthermore, due to the high H2S content, the hydrocarbon content of the raw gas is itself limited, resulting in a smaller sales gas flow rate. Furthermore, the sulfur market is saturated and additional products may be difficult to monetize at attractive prices. The technical and safety-related challenges include drilling and well completion for heavily sour gas reservoirs and the need for expensive downhole material to provide higher resistance to acid gas corrosion. Requirements to ensure large unpopulated buffer areas around the wells may prevent optimized well location or require expensive population relocation to minimize exposure to an accidental H2S release. Pipelines conveying the highly sour raw gas are also a potential risk of leakage and must be designed to resist corrosion (external and internal) and be routed to avoid populated areas. Wet highly sour gas is also subject to hydrate formation, which necessitates the addition of hydrate inhibitors. In turn, the corrosion and hydrate inhibitors must be oxygen-free to avoid the formation of elemental sulfur and related enhanced corrosion and plugging in the pipeline and process facility. At the gas plant, the removal of a large content of acid gases with a gas sweetening unit (absorber, regenerator, and ancillary equipment) requires adequate heat management within the absorber and the installation of absorber side coolers. DGA has been reported successful for highly sour gas purification [31], with H2S, CO2, and most of COS, CS2, and mercaptans removed by DGA operating at high pressure. However, solvent regeneration and circulation require a large amount of energy, which drives the development of alternative processes to the amine process aimed at reducing the cost of gas purification.

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2.9.1 Cryogenic distillation Acid gases can be separated and recovered in a liquid state under pressure using cryogenic distillation applications such as the Sprex (special pre-extraction) process [32] Cryogenic processes are very selective and the removal of acid gases (30%60%) can reduce the cost of H2S removal and reservoir reinjection. In the distillation column, the acid gases are concentrated in the bottom as liquids (110
F150
F). The overhead is maintained at 225
F, and the partially treated off-gas is gas sent on for amine sweetening to reach sales gas specifications. The major advantage of cryogenic distillation is the reduced energy required for the acid gas reinjection; in particular, the compression is limited to the acid gases that are removed at the amine section. The pressure of the acid gases from the bottoms of the column is elevated to injection pressure with pumps, which demand less energy than the gas compressor used to recompress the acid gas collected from the sweetening process. The major disadvantage is that along with acid gases, most C3 1 are concentrated in the bottom stream and hence are reinjected and lost in the formation. Another disadvantage is the accumulation of a large inventory of high-pressure H2S-rich liquid at the bottom of the distillation column. The technology has been piloted but has not been installed at a commercial scale [32].

2.9.2 Membranes for high H2S The bulk removal of CO2 from a natural gas stream with membrane-based processes can be extended to H2S removal. The high concentration in H2S provides a sufficiently large driving force for H2S permeation. Similarly, the polishing step is performed with amine sweetening [33]. A polymer eligible to serve as a selective layer for AGR should be easily synthesized, have high solubility in environmentally friendly solvent (for casting thin layers) with high acid gas permeability and natural gas hydrocarbon impermeability, and great performance stability over time (minimum 34 years). Typical glassy polymeric membranes (modified lower viscosity solution grades of cellulose acetate, polyimide, or polysulfone) used in gas sweetening are permeated by H2O, H2S, and CO2 much faster than methane, therefore the acid gases will be concentrated at the low-pressure side of the membrane. This is not an issue if the permeate can be routed to an SRU operating at low pressure (10 psi). However, routing the permeate gas directly to sulfur recovery may not be practicable, even for highly sour gas, without a second stage to minimize the hydrocarbons slippage to the SRU. Such slippage should be minimized to avoid degradation of the produced sulfur quality, as well as to reduce air demand and consequently reduce the tail gas flow rate and the associated capital and operating cost for both the SRU and the TGTU. If a second stage is required to minimize hydrocarbon slippage, the compressor adds to the overall cost of the separation project.

2.10 CO2 capture technology at gas plant

FIGURE 2.6 Permeation schema 13.

The results of process simulations of varying membrane configurations (Fig. 2.6) and membrane selectivity (heat and mass balance, relative membrane surface area, and compressor duty), see Table 2.3 show that with a given membrane performance, a comparable AGR can result in large variations (. 40%) for requirements such as compressor power and membrane surface area. The costs of the compressor and membranes dictate the best arrangement. A superior membrane separation performance can have a drastic effect on the project economics (as much as doubling the selectivity results while reducing compression by 75% and membrane surface area by 50%). As recently discussed by Hamad [34], in heavy sour gas treatment the membrane and amine cases differ from the CO2 only case (gas free of H2S or at very low concentration), in which the slippage of methane along with CO2 in the permeate stream may be acceptable if the gas is used as low BTU fuel gas, or if the gas is flared when the flow rate is small. However, in a high H2S content case, the handling of the permeate stream is different. In absence of reinjection, the H2S must be recovered in an SRU. In addition, air demand increases dramatically with the amount and type of hydrocarbons that slip in the acid gas. An increase of methane content from 0.5%, (the higher limit for a poorly operated amine process), to 1.5% when recombining the permeate of membrane and amine acid gas streams, increases the air demand by more than 5% for rich acid gas and almost 7% for intermediate acid gas quality, as shown in Table 2.4.

2.10 CO2 capture technology at gas plant Even in an optimized plant, CO2 emissions are unpreventable unless a dedicated CO2 capture technology is implemented to extract and concentrate CO2 instead of venting or discharging it into the atmosphere. The CO2 content, flow rate, pressure, stream temperature, and other component content, will dictate the choice of capture technology utilized for the two main sources of CO2 emissions. The main source of CO2 is natural gas combustion with air in flue gases from boilers and/or natural gas combined cycles (NGCC) that provide overall electrical

67

Table 2.3 Results of membrane-based separation according to the three permeation schema of Fig. 2.6. Medium performance membranea

High-performance membraneb

Schema 1 Composition

Feed

Pressure (psi) Flow rate (lb mol/h) C1 C2 C3 C4 C5 C6 1 H2S CO2 N2 H2O

1015 100 52.00 4.00 2.00 0.50 0.50 0.50 20.00 10.00 10.00 0.50

Acid gas removed (%) C1 slippage (%) Relative membrane surface area Relative compressor power

To amine 1000 81.02 63.51 4.92 2.46 0.62 0.62 0.62 10.10 5.05 12.06 0.06

2 To SRU 25 18.62 2.93 0.08 0.04 0.00 0.00 0.00 63.44 31.72 1.23 0.55

To amine 1000 80.99 63.62 4.92 2.46 0.62 0.62 0.62 9.99 4.99 12.10 0.06

3 To SRU 25 18.65 2.54 0.07 0.04 0.00 0.00 0.00 63.80 31.90 1.08 0.57

To amine 1000 80.95 63.48 4.92 2.46 0.62 0.62 0.62 10.12 5.06 12.04 0.07

1

1 To SRU 25 18.72 3.26 0.09 0.05 0.01 0.00 0.00 63.08 31.54 1.36 0.61

To amine 1000 81.33 63.48 4.91 2.45 0.61 0.61 0.61 9.99 5.00 12.11 0.22

To SRU 25 18.52 1.99 0.06 0.03 0.00 0.00 0.00 64.11 32.06 0.82 0.93

To amine 1000 75.26 68.25 5.29 2.65 0.66 0.66 0.66 5.83 2.91 12.94 0.14

To SRU 25 24.58 2.58 0.07 0.04 0.00 0.00 0.00 63.52 31.76 1.07 0.95

59.1 1.05 100

59.5 0.91 62

59.1 1.17 45

59.4 0.71 50

78.1 1.22 85

100

140

138

27

58

SRU, Sulfur recovery unit. Assuming membrane selectivity as follows: a Medium performance C2/C1 5 0.6, C3/C1 5 0.6, C4/C1 5 0.4, C5/C1 5 0.3, C6/C1 5 0.3, H2S/C1 5 15, CO2/C1 5 15, N2/C1 5 1.5, H2O/C1 5 100. b High performance C2/C1 5 0.6, C3/C1 5 0.6, C4/C1 5 0.4, C5/C1 5 0.3, C6/C1 5 0.3, H2S/C1 5 30, CO2/C1 5 30, N2/C1 5 1.5, H2O/C1 5 100.

2.10 CO2 capture technology at gas plant

Table 2.4 Effect of levels of methane slippage on the sulfur recovery unit (SRU) air demand. Rich acid gases with CH4 slippage

Intermediate acid gases with CH4 slippage

Flow rates (lb mol/h) H 2S CO2 CH4 C2H6 C3H8 C7H8 C8H10 Air

70.00 30.00 0.50 0.10 0.10 0.03 0.03 178.69

70.00 30.00 1.50 0.10 0.10 0.03 0.03 188.23

70.00 30.00 2.00 0.10 0.10 0.03 0.03 193.01

70.00 30.00 2.50 0.10 0.00 0.03 0.03 195.39

55.00 45.00 0.50 0.10 0.00 0.03 0.03 140.50

55.00 45.00 1.50 0.10 0.00 0.03 0.03 150.05

55.00 45.00 2.50 0.10 0.00 0.03 0.03 159.59

9.3

0.0

6.8

13.6

Air increase from 0.5% CH4 slippage (%) Air increase (%)

0.0

5.3

8.0

Bold indicates methane content.

power and the steam needed for the multiple plant operations (amine adsorption, bed regeneration, condensate stabilization, water stripping, and compressor/ blower driving fluid). These flue gases contain a low concentration of CO2 (NGCC 4% and boilers 8%) along with nitrogen, oxygen, and water [35]. The high level of dilution makes the CO2 capture challenging but achievable with existing technologies. The other main source is the natural gas itself. Entrained CO2 is often removed from the natural gas and vented directly to the atmosphere from the amine regeneration process if the gas contains no H2S, from the AGE process if any, or finally from tail gas in the SRU. The stream exiting the sweet gas amine regeneration or the AGE slipped gas is almost pure CO2 on a dry basis. On the contrary, the CO2 concentration in the SRU feed or tail gas is determined by the raw gas CO2 and H2S content and can range from a few percent to more than 50%.

2.10.1 CO2 capture from flue gas In terms of CO2 concentration, capture from the boiler flue gas or NGCC is more challenging than capture from the coal plant flue gas due to the higher dilution of CO2. That is a consequence of the operation which requires a very large excess of air to meet for instance the turbine material operating conditions. Amine absorption processes dominate current technologies, but novel partial exhaust gas recycling, and hybrid membrane-absorption technologies hold great potential. Existing technologies: Currently, the most mature and efficient technology for CO2 capture from combustion gas is the amine absorption process [35]. Similar to

69

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CHAPTER 2 Natural gas sweetening

the AGR process with natural gas, primary amines like MEA can selectively absorb CO2 from flue gas and deliver a high purity CO2 stream at the regeneration step. Captures $ 90% are achieved at a commercial scale, e.g., over 3000 tons/day [36] from coal plant flue gas (CO2 content  12%). An efficiency penalty of at least 10% should be considered when accounting for the overall power plant efficiency [35]. For NGCC, the low concentration of CO2 results in very low partial pressure of CO2, which limits the efficiency of the amine absorption process. The presence of oxygen also tends to accelerate amine degradation and hence special formulations are required along with the reclaimer to minimize amine loss [37]. Novel approaches: The impact of CO2 on global warming drives unprecedented research work and the development of technologies for CO2 capture. A promising approach aims at increasing the CO2 content in the flue gas and, therefore, easing its capture. The simplest option to increase CO2 content is based on a partial exhaust gas recycling as depicted in Fig. 2.7B. This approach can be significantly enhanced by installing a membrane on the flue gas stream, such that the air required for the gas turbine operation sweeps the membrane permeate side and hence extracts CO2 from the flue gas and recycles it back to the gas turbine. With an enriched flue gas containing 17% CO2 and a reduction in flow rate by a factor of 5 (Fig. 2.7C), recovery is facilitated [38]. In a recent study, Song et al. [39] listed and reviewed various hybrid processes for CO2 capture, which may offer cost-effective alternatives to less complex standalone technologies that are simpler to maintain and operate but require significant capital expenditures. Such hybrid approaches can reduce capital expenditure and energy requirements and hence improve the overall system efficiency and the eventual cost of carbon capture (CC). The hybrid and multiple approaches will also help in distributing the load on the manufacture and supply of materials and technologies and hence reduce the risk of bottlenecks.

2.10.2 CO2 captured from the acid gas stream The capture of CO2 from the acid gas stream is of particular significance to gas plants since the acid gas stream itself results from natural gas purification. Regarding CO2 capture, the best scenario is the natural gas stream containing only CO2 as an acid gas component, such as in the case of non-sour non-associated gas (NAG) in the raw natural gas composition (Table 2.1). After selective extraction from the natural gas with an amine-based process, a high purity ($99% on a dry basis) wet CO2 stream is produced, which can be easily dried [40] and compressed for usage and/or sequestration. The strategy is different for sour gases, due to the presence of H2S along with CO2. The classical approach to handling acid gas containing H2S involves a SRU where the H2S is converted into elemental sulfur by partial oxidation with oxygen from air or enriched air (Fig. 2.8), and the CO2 is captured upstream (Fig. 2.8B and C) or downstream of the SRU (Fig. 2.8D and E) as explained below.

FIGURE 2.7 Comparison of the schema for stream CO2 concentrations and volume flows in gas turbines with (A) no modification of normal operation, (B) partial exhaust gas recycling and CO2 capture, and (C) CO2 capture and CO2 selective membrane recycling of exhaust gas. Reproduced with permission from R.W. Baker, B. Freeman, J. Kniep, et al., CO2 capture from natural gas power plants using selective exhaust gas recycle membrane designs, International Journal of Greenhouse Gas Control 66 (2017) 3547.

FIGURE 2.8 Schema for sour acid gas stream handling upstream of the SRU: (A) without CO2 capture, (B) with amine AGE, and (C) with amine AGE and CO2 selective membrane recycling; schema for handling and CC with amine or hybrid cryogenic/membrane (D) upstream of the TGTU, tail gas treatment unit; or (E) downstream of the TGTU. AGE, Acid gas enrichment; CC, carbon capture; SRU, sulfur recovery unit.

2.10 CO2 capture technology at gas plant

2.10.2.1 CO2 capture upstream of SRU An AGE unit captures CO2 in a process that increases H2S concentration in the acid gas stream by removing CO2 [41]. Removal of CO2 upstream of SRU has multiple advantages (Figs. 4.8B), including the production of a higher quality acid gas for SRU. The high H2S content helps stabilize the SRU operation by ensuring complete combustion of entrained hydrocarbons and destruction of ammonia. It increases the capacity of existing SRUs by allowing higher acid gas flow for a similar residence time in the different vessels and reduces the footprint of the SRU and associated TGTU for new plants. The most widespread technology for the AGE is based on absorption, particularly using an H2S selective tertiary type amine like MDEA, formulated or hindered [41]. The lean acid gas with low H2S content (lower than 35%) is fed to the selective amine process in which the H2S is preferentially absorbed by the amine over CO2, resulting in the production of an H2S-rich acid gas stream at the amine regeneration step. The second benefit is the production of a high purity CO2 stream as an absorber vent gas (typically expected enrichment performance shown in Fig. 2.9). Amine-based enrichment is more effective for feeds with low H2S content (#30%35%), than for feed gas with a higher H2S composition. The capture of CO2 with selective amine follows the same trend: for a feed at 10% H2S, the enrichment to 40% H2S leads to a CO2 capture of 85%, but a feed at 25% H2S results in only 75% CO2 capture.

H2S in the enriched gas (mol-%) & CO2 captured (%)

120

100

80

60

40 H2S content after AGE 20

"Capture CO2" with only AGE "Capture CO2" with hybrid AGE/Membrane

0 0

10

20

30

40

50

H2S in feed (mol-%)

FIGURE 2.9 Typical expected enrichment performance: evolution of CO2 capture with H2S selective amine and combination of H2S selective amine and CO2 selective membrane as a function of H2S content in the acid gas feed.

60

73

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CHAPTER 2 Natural gas sweetening

The trade-off for amine-based AGE operation is the acceptable level of H2S slippage in the CO2 for a given enrichment. Models based on mass and heat transfer can accurately predict this H2S slippage, which can be minimized by increasing the circulation rate of the selective amine, but the modeling shows that enrichment is then reduced unless a tertiary amine promoter (“stripping promoter”) like phosphoric acid is added to reduce the H2S content in the lean amine and hence avoid a drastic reduction in the driving absorption force (“lean-end pinch”) in the lean end of the absorber. The reduction of lean amine temperature is also a solution to reduce H2S slippage, but temperatures lower than 80
F might be required to meet a 50 ppm H2S concentration, which may not be practical in some locations. In the context of global emission reduction, Vaidya et al. [42] proposed a new approach that synergizes an H2S selective amine and a CO2 selective membrane to capture CO2 upstream of the SRU (Fig. 2.8C). It is important to remember that with membrane processes it is possible to produce a high purity residue stream for the less permeable components, but the higher the purity, the lower the recovery [27]. For a CO2 over H2S selective membrane [43], H2S can be enriched to virtually any concentration, but economics limits the practicality of the process. Target enrichment of 85%90% H2S is achievable with a highly CO2 selective membrane (Table 2.5). With a CO2-rich and H2S lean permeate gas, the H2S can be recovered with an H2S selective amine while the CO2 is slipped with a very low level of H2S. Without AGE, it is also possible to take advantage of the amine used in the tail gas treatment, which is by nature H2S selective. Simulations show that it is possible to enrich a 60% H2S stream up to 90% while minimizing H2S recycling (Fig. 2.8C). Results for the fraction of H2S rejected by the membrane for a CO2-over-H2S selective membrane and five feed compositions ranging from 10% to 60% H2S are provided in Table 2.5. For simplicity, the acid gas is treated as only a mixture of H2S and CO2. The stream is water-saturated and contains traces of CH4, mercaptans, and BTX, and the selectivity of CO2 over H2S is selected at 10 as measured by Merkel et al. [43]. Table 2.5 H2S rejection with CO2-over-H2S selective membrane.a % H2S in membrane feed

% H2S in the membrane rejected

Fraction of H2S rejected by the membrane

10 10 10 10 30 30 50 50 60

60 70 80 90 80 90 80 90 90

0.796 0.760 0.713 0.625 0.797 0.710 0.863 0.767 0.797

a

Permeation of CO2 is 10 times faster than H2S.

2.10 CO2 capture technology at gas plant

2.10.2.2 CO2 capture downstream sulfur recovery plant Downstream of the SRU, the CO2 capture process can be located at two locations, either immediately after tail gas treatment if any, or further downstream the thermal oxidizer unit (sometimes referred to as a TOx) that converts any residual H2S into SOx. The first location offers a stream with a workable temperature for an amine process, and the CO2 is not diluted by any combustion. It has a lot of similarities to CO2 captured from natural gas. A typical primary or secondary reactive amine for CO2 can be considered. Furthermore, the amine used for the AGR can be utilized for the CO2 capture, which minimizes the types of amines required on site. Capture at the second location, downstream of the thermal oxidizer, faces several challenges since the stream is hot (1200
F) and needs cooling and removal of SOx generated by the combustion of slipped H2S prior to entering the CO2 capture section. In addition, combustion with air generates a larger stream, introduces oxygen, and dilutes the CO2, all of which makes the capture of CO2 downstream thermal oxidizer like CO2 captured from flue gas. Alternative technologies can compete with amine processes because CO2 content can be quite large ( . 50% for lean acid gas) in the tail gas stream. For example, Kulkarni [44] proposed a combination of cryogenic and membrane processes to extract CO2 from the CO2-rich stream. The process involves the removal of contaminants (SOx and NOx) prior to compression and water removal with an adsorbent. Then a CO2 selective membrane concentrates CO2 further if required. The CO2-rich permeate of the membrane is then cooled and liquid CO2 is collected in a vessel. The overhead of this vessel is routed back to the membrane inlet to ensure a high CO2 capture rate. The membrane rejects stream is expanded to recover power from the high-pressure stream and deliver the required cooling to condense CO2 from the permeate stream. Condensed CO2 can be pumped to the desired pressure. This hybrid approach has the potential to outweigh the wellestablished amine process for a CO2-rich stream since produced CO2 is already dry and compressed (Fig. 2.10).

FIGURE 2.10 Hybrid membrane cryogenic for CO2 capture from CO2-rich stream.

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2.11 Final remarks Natural gas will undoubtedly be a pillar of the energy transition by providing cheap energy for power generation, mining, and refining of metals that massive electrification will require, without forgetting the energy needed for the concrete manufacture for onshore wind turbine foundations. It is also anticipated that natural gas will deliver low-cost hydrogen helping in advancing hydrogen infrastructures by expanding steam methane reforming units beyond the current use for refining and petrochemical. Finally, all the knowledge acquired for natural gas processing will also support the deployment of technologies for the capture of CO2 from the combustion of fossil fuels.

Nomenclature AG AGE AGR BTU BTX CC COS DEA DIPA DGA lb MDEA MEA MMSCFD NAG NGCC SCF SRU Tg TGTU TOx TSR WHB

associated gas acid gas enrichment acid gas removal British thermal unit benzene, toluene and xylene carbon capture carbonyl sulfide diethanol amine diisopropanol amine diglycol amine pounds methyl diethanol amine monoethanol amine million standard cubic feet per day of gas non-associated gas natural gas combined cycles standard cubic feet sulfur recovery unit glass transition temperature tail gas treatment unit thermal oxidizer thermochemical sulfate reduction waste heat boiler

References [1] IEA, Annual Energy Outlook 2021, International Energy Agency, Paris, 2021. [2] J.C. Ve´drine, Natural gas as a feedstock, Proceedings of the NATO Advanced Study Institute, July 618, 2003, Vilamoura, Portugal.

References

[3] J.M. Campbell, Gas Conditioning and Processing, in: The Basic Principles, 9th Edition, 1, Campbell Petroleum Series, 2014. [4] R.C. Selley, S.A. Sonnenberg, Elements of Petroleum Geology, 3rd ed., Elsevier, 2015. [5] B. Tissot, D. Welte, Petroleum Formation and Occurrence, 2nd ed., Elsevier, New York, 1984. [6] C.A. Scholes, U. Ghosh, Helium separation through polymeric membranes: selectivity targets, Journal of Membrane Science 520 (2016) 221230. [7] R.J. Johnson, B.D. Folwell, A. Wirekoh, et al., Reservoir souring  latest developments for application and mitigation, Journal of Biotechnology 256 (2017) 5767. [8] L. Jiang, R.H. Worden, C. Cai, Generation of isotopically and compositionally distinct water during thermochemical sulfate reduction (TSR) in carbonate reservoirs: Triassic Feixianguan Formation, Sichuan Basin, China, Geochimica et Cosmochimica Acta 165 (2015) 249262. [9] A. Bahadori, Natural Gas Processing Technology and Engineering Design, Elsevier, 2014. [10] ISO 15156 (Annex D), Petroleum and natural gas industries - Materials for use in H2S-containing environments in oil and gas production - Part 2: Cracking-resistant carbon and low-alloy steels, and the use of cast irons, International Organization for Standardization, Geneva, 2020. [11] J.A. Sames, H.G. Paskall, Sulphur Recovery, Sulphur Experts, Calgary, 2003. [12] S. Mokhatab, J.Y. Mak, J.V. Valappil, et al., Handbook of Liquefied Natural Gas, Elsevier, 2014. [13] A. Kohl, R. Nielsen, Gas Purification, 5th ed., Gulf Publishing Company, Houston, 1997. [14] D.A. Delmer, Gas Treating, John Wiley and Sons, 2014. [15] R.N. Maddox, Gas Conditioning and Processing, Campbell Petroleum Series (1982). [16] A. Kidnay, W.R. Parrish, Fundamentals of Natural Gas Processing, CRC Press, 2006. [17] M. Steward, Surface production operations, Design of Gas-Handling Systems and Facilities, 3rd ed., Elsevier, 2014. [18] W.C. Lyons, Working Guide to Petroleum and Natural Gas Production Engineering, Elsevier, 2010. [19] J.G. Speight, Natural Gas Basic Handbook, 2nd ed., Elsevier, 2019. [20] M. Steward, K. Arnold, Gas-Liquid and Liquid - Liquid Separators, Elsevier, 2008. [21] A.B. Badiru, S.O. Osisanya, Project Management for the Oil and Gas Industry, CRC Press, 2013. [22] H. Baron, The Oil and Gas Engineering Guide, 2nd ed., Technip, Paris, 2015. [23] R.H. Weiland, M.S. Sivasubramanian, J.C. Dingman, Effective amine technology: Controlling selectivity, increasing slip, and reducing sulfur, 53rd Annual Laurance Reid Gas Conditioning Conference, February 24, 2003 Norman, OK. [24] J. Haydary, Chemical Process Design and Simulation, John Wiley and Sons, 2019. [25] N.P. Lieberman, Troubleshooting Natural Gas Processing, Penwell Publishing Company, 1987. [26] R.T. Yang, Gas Separation by Adsorption Processes, Butterworth Publishers, 1987. [27] R.W. Baker, Membrane Technology and Applications, third ed., John Wiley and Sons, 2012. [28] I. Pinnau, B. Freeman, Y. Yampolskii, Materials Science of Membranes for Gas and Vapor Separation, Wiley, 2007.

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[29] P.R. Couchman, F.E. Karasz, A classical thermodynamic discussion of the effect of composition on glass-transition temperatures, Macromolecules 11 (1978) 117119. [30] Y. Matsumiya, T. Inoue, T. Iwashige, et al., Dielectric relaxation of polymer/carbon dioxide systems, Macromolecules 42 (2009) 47124718. [31] D. Schulte, S. Van Wagensveld, C. Graham, et al., DGA treatment of ultra-sour gas, in: Proceedings of the SOGAT conference, April 29May 3, 2017, Abu Dhabi, UAE. [32] F. Lallemand, F. Lecomte, C. Streicher, Highly sour gas processing: H2S bulk removal with the SPREX® process, in: Proceedings of the SOGAT conference, November 2629, 2005, Abu Dhabi, UAE. [33] A. Baudot, Gas/Vapor permeation applications in the hydrocarbon processing industry, in: E. Drioli, G. Barbieri (Eds.), Membrane Engineering for the Treatment of Gases, Vol. 1: Gas Separation Problems with Membranes, Royal Society of Chemistry, London, 2011. [34] F. Hamad, M. Qahtani, A. Ameen, et al., Treatment of highly sour natural gas stream by hybrid membrane-amine process: techno-economic study, Separation and Purification Technology 237 (2020) 116348. [35] B. Metz, O. Davidson, H.C. de Coninck, et al., IPCC Special Report on Carbon Dioxide Capture and Storage, Cambridge University Press, Cambridge, NY, 2005. [36] Kearney Energy Transition Institute, Carbon Capture Utilization and Storage Towards Net-Zero, Gravenhage, 2021. [37] L. Dume´e, C. Scholes, G. Stevens, et al., Purification of aqueous amine solvents used in post combustion CO2 capture: a review, International Journal of Greenhouse Gas Control 10 (2012) 443455. [38] R.W. Baker, B. Freeman, J. Kniep, et al., CO2 capture from natural gas power plants using selective exhaust gas recycle membrane designs, International Journal of Greenhouse Gas Control 66 (2017) 3547. [39] C. Song, Q. Liu, N. Ji, et al., Alternative pathways for efficient CO2 capture by hybrid processes  a review, Renewable and Sustainable Energy Reviews 82 (2018) 215231. [40] J.J. Carroll, Acid Gas Injection and Carbon Dioxide Sequestration, Scrivener Publishing, Salem, MA, 2010. [41] R.H. Weiland, Acid gas enrichment-maximizing selectivity, in: Proceedings of the Laurance Reid Gas Conditioning Conference, February 2427, 2008, Norman, OK. [42] M.M. Vaidya, S. Duval, F. Hamad, et al., Sulfur recovery operation with improved carbon dioxide recovery, US Patent No. 11,325,065, 2022. [43] T.C. Merkel, L.G. Toy, Comparison of hydrogen sulfide transport properties in fluorinated and nonfluorinated polymers, Macromolecules 39 (2006) 75917600. [44] S.S. Kulkarni. CO2 capture by sub-ambient membrane operation, in: Proceedings of the NETL CO2 Capture Technology Meeting, July 912, 2012, Pittsburgh, PA.

CHAPTER

3

Emulsion separation

Thomas Krebs and Mohamed Reda Akdim TechnipFMC, Arnhem, The Netherlands

Chapter Outline 3.1 3.2 3.3 3.4

3.5 3.6

3.7

3.8

3.9

Introduction ................................................................................................. 80 Emulsion formation ....................................................................................... 82 Emulsion stabilization ................................................................................... 88 Theory of emulsion separation ....................................................................... 93 3.4.1 Settling velocity of droplets ..........................................................95 3.4.2 Coalescence rates .....................................................................100 3.4.3 Semi-empirical approaches ........................................................105 Emulsion separation techniques ..................................................................107 Thermal demulsification ..............................................................................110 3.6.1 Effect of heating on emulsion properties .....................................110 3.6.2 Heater technology .....................................................................112 3.6.3 Case studies .............................................................................113 Mechanical internals ..................................................................................114 3.7.1 Separator vessels ......................................................................114 3.7.2 Perforated baffles ......................................................................115 3.7.3 Plate packs ...............................................................................116 3.7.4 Pipe separators .........................................................................119 3.7.5 Case studies .............................................................................120 Chemical demulsification ............................................................................121 3.8.1 Effect of demulsifier on separation rates .....................................122 3.8.2 Mechanisms of demulsifier action ..............................................123 3.8.3 Demulsifier formulation .............................................................125 3.8.4 Case studies .............................................................................128 Electrostatic demulsification .......................................................................129 3.9.1 Droplet migration in electric fields ..............................................130 3.9.2 Droplet collisions in electric fields ..............................................131 3.9.3 Effect of electric field properties on droplet coalescence ..............134 3.9.4 Electrocoalescer technology .......................................................136 3.9.5 Case studies .............................................................................138

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00003-X © 2023 Elsevier Inc. All rights reserved.

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3.10 Concluding remarks ....................................................................................139 Nomenclature ......................................................................................................139 References ..........................................................................................................140

3.1 Introduction The major task of crude oil processing is to provide oil of a quality that is suitable for transport and further downstream processing. The oil that is produced needs to be separated from natural gas, produced water, and solids. This chapter will deal with the specific task of separating crude oil and produced water. In a general manner, liquidliquid separation is the most challenging bulk separation task compared to gasliquid and liquidsolid separation, as the density difference between the oil and the water phase is the smallest of all four phases. Gravity is the major driving force for bulk separation, so the separation rate of water from oil will be the smallest compared to that of the other phases. Another challenge is that the produced fluids may contain interfacially active molecules or nano-scale particles that can stabilize the droplets against coalescence, and thus impede separation. Liquidliquid separation usually occurs simultaneously with the separation of gas and solids, so the design of a separator needs to consider the outlet specs for all four phases. The separation of gas and solids from crude oil will not be covered in this chapter, and especially in the discussion of the fundamentals of liquidliquid separation, we will pretend that gas and solids do not exist in the physical reality of upstream petroleum processing, as no physical model exists yet that can satisfactorily describe the interaction of all four phases inside a separator under continuous flow. Fig. 3.1 displays an example of a high-level scheme for crude oil processing. The fluid enters the processing facility from the wellhead or a manifold at elevated pressure. The system pressure typically is reduced in several steps to release gas dissolved in the liquid phases, and for each pressure stage, the produced fluids pass through a bulk separator where gas, oil, water, and solids are separated. The purpose of the pressure stages is to limit the transfer of heavier ends in the crude to the gas phase and to minimize the power consumption of compressors in the gas processing train [1]. The figure shows a first- and second-stage bulk separator, which is usually sufficient for intermediate feed pressures (2070 bar). Three or even four stages may be used if the wellhead pressure is high ( . 70 bar) [2]. After the gas and the majority of the produced water have been removed, some water remains as a dilute water-in-oil (WiO) emulsion in the crude oil. Correspondingly, the water phase will contain a residual concentration of oil as a dilute oil-in-water (OiW) emulsion (typically up to 1% v/v combined from all separator outlets) and is further processed in the downstream produced water treatment system. The target value for the residual WiO concentration is of the order 0.5% v/v which is required for the oil to be considered suitable for transport to the downstream facilities.

3.1 Introduction

Gas Processing

Gas

From Well

1st Stage Separation Water

Wash Water

Gas Oil

2nd Stage Separation Water

Oil

Dehydration

Oil

Water

Produced Water Treatment

Mixer Desalting

Oil

Water

Waste Water Treatment

FIGURE 3.1 Schematic depiction of a possible arrangement for the processing steps of crude oil. Emulsions can be present in all four vessels shown in the figure.

In many instances, the effluent crude WiO emulsion even from the second or third-stage bulk separator contains more than 0.5% v/v of water, and further separation can be achieved only through additional measures such as heating and/or utilization of electrostatic coalescers [3]. This type of treatment is conducted in the dehydration vessel, after which the WiO concentration should be of the required order of 0.5% v/v. While this concentration is sufficiently low for further processing, the water phase often contains a high concentration of brine, and salt crystals may also be present in the oil phase [4]. The quantities of salt that are typically present in either phase can cause severe problems in downstream processing due to equipment corrosion and deposition of the solids. For that reason, the oil is washed with water of a much lower salinity to extract the salt from the water and oil phase. After mixing the wash water with the crude, the wash water is removed from the desalting vessel. When the desalting step is done onshore, fresh/brackish water or low-salinity waste water from refinery operations is typically used [5]. Offshore desalting is also possible if low-salinity water is available, either from a desalination unit or by using sea water as wash water if the salinity difference between the formation water and the seawater is sufficiently large [6]. In all the processing steps shown in Fig. 3.1, liquidliquid separation is required. From an engineering perspective, the ultimate question is how large a separator needs to be to achieve efficient oilwater separation, and what additional driving forces need to be employed to develop a design that delivers the required performance while also minimizing cost. In the most general sense, separation of the produced fluids is free of charge, as a liquid mixture with different densities of its constituents will eventually separate under the influence of normal gravity. And while in principle no external energy input is required to do so, just collecting all fluids in a giant empty vessel with a large residence time is not an economically feasible solution. Exceptions exist for onshore applications where sometimes this approach is taken, but even then, additional measures may be required to enhance separation performance to not let the required volume of

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the separator escalate dramatically. In an offshore environment, the situation is different, as the weight and space of any processing equipment require expensive structural support. For that reason, striking the balance between performance and cost is a delicate task in the design of offshore processing plants. For liquidliquid separation, this usually means that to optimize the size of the separator, additional driving forces such as chemicals and mechanical internals are exploited to achieve a compact design. For the sizing and design of the separator, it is important to have an idea about the expected separation rate of all fluids in the process flow, from which a required residence time inside the separator can be estimated. The scope of this chapter is to provide an overview of both the fundamental aspects of crude oil emulsion separation and the technologies employed to facilitate their separation in the field. We start with an overview of the mechanisms and forces that govern the formation, stability, and separation of emulsions, as these are the foundations required to understand the driving forces that can be used to break them. We will describe these driving forces for emulsion breakup in more detail and how they are employed in the reality of oilfield processing using selected case studies. The basic concepts of separator units, their internals, and other equipment for emulsion separation will also be covered in this chapter. By necessity, we need to stay brief in the discussion of all these aspects, but we hope that the reader will find a satisfactory overview of the theory and practice of emulsion separation in crude oil processing.

3.2 Emulsion formation The focus of this chapter is on emulsion separation, but we believe it is instructive to briefly discuss the basic aspects of emulsion formation first. After all, it is only beneficial if the creation of emulsions can be prevented, or less stable emulsions can be created during production. In properly designed separators, the properties of the flow field favor coalescence almost in the entire vessel, but upstream of the separator conditions are frequently encountered that favor the formation of emulsions. In an undisturbed reservoir, the gas and liquid phases are residing separated before production starts. The flow that is required for production to be economically attractive is so high that it causes the reservoir fluids to mix to some degree at the wellbore once production has started, and even more in the production flow line. This effect is only exacerbated with increasing field life, as more and more water will be produced together with oil. In addition, the reservoir flow and pressure are usually regulated with a choke valve. This sudden pressure drop is the source of dissipation of turbulent energy into the fluid, with even more intense mixing as a result. Any other geometrical constrictions can also induce the formation of emulsions, just as a poorly designed inlet section of the separators. When pressure needs to be boosted in the production process,

3.2 Emulsion formation

Possible Locations for Emulsion Formation Wellhead

Separator Inlet Constriction

Choke Valve Flowline

Wellbore Reservoir

FIGURE 3.2 Schematic overview of the flow path for a produced fluid from the well to the processing facility.

emulsions can also form in pumps [7]. Fig. 3.2 provides a schematic overview of the different locations in the production process where emulsions can be formed from the wellbore to the gravity separator. In this section, we discuss the basic parameters that govern the creation of emulsions during crude oil production. An emulsion is formed when a homogeneous phase of liquid is broken up into individual droplets (the dispersed phase), which are distributed into a second liquid phase (the continuous phase). Thermodynamics dictates that an increase in interfacial area ΔA of a system, such as through the formation of droplets, requires the addition of Gibbs free energy (ΔG) to the system. The two properties can be related to each other via the interfacial tension (σ) between the two liquid phases: ΔA 5 ΔG/σ [8]. Interfacial tension is one of the key parameters determining the properties of the emulsion, as it governs the range of emulsion droplet diameters. The Gibbs free energy is not a very practical parameter to describe the properties of an emulsion, which is why typically parameters are used that can be derived from the properties of the multiphase fluids and flow. For a crude oil/water flow in a production environment, the energy for droplet formation is supplied by the turbulent flow. The pressure fluctuations in turbulent flow and collisions of fluid parcels with turbulent eddies lead to droplet deformation and breakup. Fig. 3.3 shows a series of photographs taken in situ visualizing the deformation of a droplet and breakup into several smaller droplets in a turbulent flow near an impeller at 300 rpm. The dissipation rate of turbulent kinetic energy was estimated to be of the order B of 350 m2/s3 [9]. The properties that are relevant to assessing the magnitude of droplet breakup are summarized in the dimensionless Weber number We, which can be interpreted as a ratio between inertial forces acting on the droplet surface and interfacial forces resisting deformation, We 5 ρcυ0 2d/σ [10], where ρc is the density of the

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FIGURE 3.3 Photographic snapshots of the breakage of a droplet into smaller daughter droplets in a turbulent flow. Reproduced with permission from M. Ashar, D. Arlov, F. Carlsson, et al., Single droplet breakup in a rotorstator mixer, Chemical Engineering Science 181 (2018) 186198.

continuous phase, d is the droplet diameter, and υ0 is a characteristic velocity, often taken as the fluid flow velocity, but a definition based on a characteristic velocity difference near the droplet as induced by turbulent eddies would be equally valid. If the energy of the flow field surrounding the droplet is large enough, the droplet will break up into two or smaller daughter droplets. The concept of a critical Weber number (Wec) has been introduced to provide an estimate for the maximum droplet size (dmax) in a given flow environment. The value of Wec depends on the nature of the flow and the fluid properties. As an example, for air bubbles in water in a turbulent flow, Wec  2.3 was obtained [11]. Different models can be used for providing a value for Wec, so that dmax can be calculated. A relatively simple equation for emulsions that came forth from such an analysis is shown in Eq. (3.1) as an example [12].  0 μd v 20:6 20:4 dmax 5 c σ 1 ρc E 4

(3.1)

In the equation, c is a constant (specified as of the order of unity [12]), μd is the viscosity of the dispersed phase, and A is the turbulent energy dissipation rate. More sophisticated equations have been derived to estimate the maximum droplet size from basic flow parameters, but Eq. 3.1 illustrates the important dependencies, and usually gives a correct order-of-magnitude estimate. For A, we can use the expression A 5 Δpv/(lρe) [13,14]. In the equation, Δp is the pressure drop over a length scale l, v is the mean fluid velocity, and ρe is the emulsion density.

3.2 Emulsion formation

For regular pipe flow, the term Δp/l can be estimated using the properties of the flow [15]. In the ideal case of fully developed pipe flow upstream of a separator with pipe flow velocities of the order 12 m/s, the expected maximum droplet diameter is of the order B1 mm for a light crude oil/water mixture [16]. Much smaller droplet diameters can be obtained downstream choke valves or inside pumps [7] where relatively large amounts of energy are dissipated into a small volume of fluid. For choke valves, some empirical relations have been derived for specific experimental conditions and orifice geometries [14,17], but we use Eq. (3.1) to illustrate the general trend of droplet sizes. The same expression for A as for regular pipe flow can be used, although this assumption is not strictly valid. In this case, a fixed-length is chosen for the length scale l, which is the length of the dissipation zone downstream of the constriction. A relation of l  2.5dp was determined experimentally by Morrison et al. [18] (dp is the pipe diameter). For the flow through an orifice at 10 m/s, the change of dmax as a function of the pressure drop Δp over the orifice can be approximated by using typical physical properties for a WiO emulsion: σ 5 20 mN/m, μd 5 1 mPa s, ρc 5 870 kg/m, ρe 5 900 kg/m3. For v0 , the relation v0  (Admax)1/3 can be used [19]. With these parameters, Eq. (3.1) is solved, and a curve for dmax versus Δp over the orifice is obtained, as shown in Fig. 3.4. The curve shows that the maximum droplet diameter downstream of a choke valve can be one to two orders of magnitude smaller than that in regular pipe flow. If the oil phase is present in bulk quantities, however, re-coalescence of droplets takes place as the fluid passes through the pipe system downstream of the valve. Hence, if the distance between the choke valve and separator is large enough, droplets will grow significantly in size, and alleviate the effect of the choke valve or other constrictions on the flow. It should be emphasized that we used a very simple model to calculate dmax, with the intention to show the general order of magnitude for the trend of droplet sizes. Roughly speaking, the numbers obtained with Eq. (3.1) are a factor 300

dm ax [μm]

250

200 150 100

50 0

0,1

1

10

100

Pressure Drop [bar] FIGURE 3.4 Maximum droplet diameter downstream of an orifice as a function of pressure drop over the orifice calculated using Eq. (3.1).

85

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23 times lower than those obtained experimentally [14,17,20]. One reason for this is that re-coalescence will always take place in concentrated emulsions simultaneously with droplet breakup, which is not considered in Eq. (3.1). A visualization of the effect of an orifice on the breakup of emulsion droplets is shown in Fig. 3.5 [21]. The pressure drop over the orifice was relatively small with 10 mbar, but for the relatively large droplets, the extent of droplet breakup is significant. The image was rotated 90 degrees clockwise for layout reasons. The emulsion flows towards the orifice from the bottom (left side). The pipe diameter is 3 cm, and the orifice diameter is 1 cm. The emulsion consists of heptane droplets dispersed in water at a flow velocity of B 0.1 m/s. The maximum droplet diameter of an emulsion is a relevant parameter, but those droplets do not pose a challenge for separation, as they tend to separate the fastest under the influence of gravity. The small droplets pose the biggest challenge, and thus knowledge of the actual droplet size distribution is of high importance for being able to estimate the separation rate of the emulsion. When a production facility is designed, an appraisal for the expected droplet size distributions at different locations of the production process should be made. Prediction of the entire droplet size distribution from first principle calculations is impractical and still requires empirical adjustment factors to make the models work [11,22]. Often purely empirical relations based on published data or previous experience are used instead. Another option is to predict the entire droplet size distribution based on an estimate for dmax. This is the most convenient approach when actual data is not available, but arguably the coarsest method. The rate of droplet formation of a given diameter depends on the spectrum of the turbulent kinetic energy of the flow, but experimental droplet size distributions obtained for many different flow fields can be described by the same type of

FIGURE 3.5 Photograph of an emulsion passing through an orifice. Reproduced with permission from S. Galinat, O. Masbernat, P. Guiraud, et al., Drop break-up in turbulent pipe flow downstream of a restriction, Chemical Engineering Science 60 (2005) 65116528.

3.2 Emulsion formation

theoretical distribution, such as the log-normal distribution [23]. While the exact nature of which distribution describes the data best has been the subject of debate, we shall only concern ourselves with the typical order of magnitude that is obtained for the width of droplet size distributions in two-phase turbulent liquid flow. We choose the RosinRammler distribution [24] that calculates the cumulative volume fraction xv of droplets with a diameter smaller than d: "

 δ # d xv 5 1 2 exp 2  d

(3.2)

The distribution has two parameters: the width parameter δ and a characteristic droplet diameter d . The d can be related to dmax by using an approximation. dmax is the diameter for which the volume fraction of droplets with a diameter less than dmax is 100%. As the RosinRammler distribution and other distributions do not have a cut-off value, we need to make an assumption for it, and dmax can be equated with a diameter for which the volume fraction of droplets with a diameter less than dmax is e.g., 95% or 99%. With a chosen value for this volume fraction xv,max, we obtain: d 5 dmax[ 2 ln(1 2 xv,max)]21/δ. The value of the parameter δ is obtained experimentally. We conducted a brief survey of droplet size distributions obtained in regular pipe flow [25,26], in stirred tanks [27], downstream of choke valves [20], and in static mixers [2830]. These experimental settings are a cross-section of typical conditions of turbulence in liquidliquid flows. All data yielded values in the range 2 , δ , 3, which suggests that there is a measurable, but not extreme variation in the width of the droplet size distribution with the properties of the turbulent flow and the fluids. We plot the cumulative volume fraction vs. the droplet diameter in Fig. 3.6 for

1,0 0,8

xv

0,6 0,4

d =2

0,2

d = 2.5 d =3

0,0 10

100

1000

10000

Drop Diameter [µm] FIGURE 3.6 Cumulative volume fraction of dispersed phase xv as a function of droplet diameter d. The curves are RosinRammler distributions calculated with Eq. (3.5) for different values of the maximum droplet diameter dmax and the width parameter δ.

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different values of δ and dmax. The two different values for dmax of 150 and 1500 μm can be considered typical for flow downstream of a choke valve and regular pipe flow, respectively. The different values of δ likewise can be considered to represent the typical bandwidth for this parameter expected in oil processing facilities. These curves give a general impression of what the droplet size distribution entering a liquid-separator could look like. If the incoming flow is a more or less fully developed pipe flow, droplets tend to be fairly large (dmax 5 1500 μm case in Fig. 3.6), and separation will be comparatively easy. If there is a source of significant pressure change just upstream of the separator, like a choke valve or a pump, a significant fraction of droplets will be small enough to pose a challenge for separation. For the case with dmax 5 150 μm shown in Fig. 3.6, this could mean that up to 20% v/v of droplets will have a diameter of 50 μm or less. More detailed and fundamental descriptions of the process of emulsion formation in turbulent flow fields are available [31,32]. Our focus in this section was on providing some order-of-magnitude estimates for maximum droplet diameters and droplet size distributions, and their scaling with the most important fluid and flow properties. We also argue from our own experience that a more elaborate treatment of the droplet breakup process usually does not lead to a more accurate prediction of the feed droplet size distribution. Two major complicating factors are the complexity of the flow path of the crude oil/water mixture in the production flow line and the production facility, and the complex chemical composition of the produced fluids, which can have a strong influence on the rate of droplet re-coalescence upstream of the separator. Except for the most basic cases, it is not possible to include these aspects in a practical way for the modeling of droplet size distributions.

3.3 Emulsion stabilization An emulsion is an unstable system from the thermodynamic point of view. The high internal surface area increases the free energy of the system; and in principle, separation into two bulk phases is favored. Emulsions can display enormous kinetic stability, however, which is induced by surface-active components preventing the coalescence of droplets. In crude oil processing, any separation time inside a separator that is longer than a few minutes is typically considered unfavorable from a process engineering perspective. Droplet separation is essentially governed by two processes. The first process is the settling of droplets under the influence of gravity. This induces a certain degree of separation, as droplets will accumulate at the top or bottom of the emulsion, depending on their density. For complete separation, the droplets need to coalesce with each other to form bulk phases of oil and water. A high viscosity of the continuous phase and small droplet diameters will have a stabilizing effect on

3.3 Emulsion stabilization

the emulsion, as both factors impede settling under the influence of gravity, and increased viscosity of the continuous phase also slows the drainage of the continuous phase film between droplets required for coalescence. Fig. 3.7A visualizes this trend and shows the stability index of WiO emulsions prepared from North Sea crude oils with different viscosities [33]. The second step required for emulsion breakup is coalescence between droplets that are in close contact. An emulsion can be stabilized against coalescence by colloidal forces between the droplet interfaces. The major terms that are relevant for the colloidal stability of crude oil emulsions are summarized in Eq. (3.3) [36]. Other contributions to inter-droplet forces, such as hydration, depletion, and hydrophobic forces also exist [37], but have not been the subject of much investigation in crude oil emulsions, and presumably are often small compared to the terms in Eq. (3.3): Fs 5 Fvdw 1 Fedl 1 Fsteric 1 Fmech

(3.3)

The four terms that combine to the total surface force Fs in the above equation are: •

Fvdw is the van-der-Waals force between the droplet interfaces, a weak electrostatic force of quantum-mechanical origin. For interactions between

FIGURE 3.7 Impact of crude oil viscosity and asphaltene content on emulsion stability. Panel (A) shows a stability index for emulsions prepared from different types of crude oil as a function of the emulsion viscosity which was governed by the crude oil viscosity in these experiments [33]. Panel (B) shows the stability of a model water-in-oil emulsion as a function of asphaltene concentration for two different concentrations of resins also present in the emulsion [34]. The insert in panel (B) shows a possible structure of an asphaltene molecule as an example [35]. Reproduced with permissions from [3335].

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•

•

•

droplets and particles, van-der-Waals forces are always attractive. They are short-ranged with interaction distances of the order 110 nm [37]. Fedl is the electric-double layer (EDL) repulsion between the two interfaces. Particles and surfaces have a surface charge, and the interaction between the two charged layers results in a repulsive force that can prevent droplet coalescence. A bare oil/water interface will acquire a surface charge through the presence of ionized or dipole-aligned water molecules. This contribution to the EDL force often is not significant enough to prevent coalescence. The presence of surfactants, however, either native or introduced as production chemicals such as corrosion inhibitors, can strongly increase interfacial charge and thus create an electrostatic barrier against coalescence [36]. The magnitude and range of the EDL force are much lower if oil is in the continuous phase as compared to water due to the low dielectric permittivity of crude oil [37], and EDL repulsion is thus not considered to play a large role in stabilization of water-in-crude-oil emulsions. In water-continuous crude oil emulsions, EDL forces are of relevance as a mechanism for stabilization, in principle. Formation water has a high concentration of electrolytes of weight percent, though, and at these high electrolyte concentrations, the repulsion between the charged surfaces is strongly reduced [37], which in turn reduces the barrier to coalescence. Fsteric are repulsive steric forces between the two interfaces. These forces arise from the adsorption of relatively large molecules, such as asphaltenes, resins, and fatty acids, which all are native constituents of crude oil [38]. The adsorbed layers can form rigid films [39] or gel-like structures [40] and can present a strong barrier to the further approach of the droplets. Fig. 3.7B shows the stability of a model water-in-oil emulsion as a function of asphaltene concentration for two different concentrations of resins also present in the emulsion [34]. The insert in panel (B) shows a possible structure of an asphaltene molecule as an example [35]. Fig. 3.8 shows two photos of a water droplet attached to a micropipette in an emulsion of 0.1 wt.% bitumen in an organic solvent [41]. The left picture shows the droplet as prepared after sufficient time to reach adsorption equilibrium at the liquidliquid interface. Part of the water phase was then removed with a pipette, and a rigid shell becomes visible. While in this experiment no actual crude oil emulsion was used, it nicely visualizes the type of stabilizing adsorption layers that can be created in crude oil emulsions. Fmech are mechanical forces created by the adsorption of small inorganic particles (such as silica and clay) [42] or organic particles composed of wax that can form a solid-like shell around the droplets.

A data set that can provide information relevant to assessing emulsion stability is the so-called SARA analysis where the concentrations of the crude oil constituents saturates (5 alkanes), aromatics, resins, and asphaltenes are measured [38]. From the SARA components, resins and asphaltenes are thought to play a large

3.3 Emulsion stabilization

FIGURE 3.8 Water droplet attached to a micropipette immersed in a solution of 0.1 wt.% bitumen in an organic solvent. Reproduced with permission from P. Tchoukov, J. Czarnecki, T. Dabros, Study of water-in-oil thin liquid films: implications for the stability of petroleum emulsions, Colloids Surf. A: Physicochem. Eng. Aspects 372 (2010) 1521.

role in stabilizing crude oil emulsions [43,44]. Additional stabilizing molecules are acidic components such as fatty acids and naphthenic acids (which are part of the resins fraction [45]). Their stabilizing effect is pH-dependent [46]. For model systems, a multitude of experiments has demonstrated an effect between asphaltene content and water-in-oil emulsion stability. For systems that contain additional surface-active constituents, and with actual crudes, more complex behavior is observed [34,39,46], and quantitative predictions for the stability of an emulsion based on the composition of the crude alone generally cannot be made. For that reason, predictions for separation efficiencies and emulsion stability for a produced fluid from a well usually follow the rules of thumbs to consider the fluid properties and composition. A high fraction of asphaltenes and resins often signals that the corresponding water-in-crude-oil emulsions may exhibit elevated stability against coalescence. If the produced fluid also contains a substantial concentration of small solids, a stabilizing effect may also occur. The results from these quantitative analyses cannot directly be converted to an emulsion lifetime, but experiments on batch emulsion stability can be conducted that usually give a good idea about the separation rate. The colloidal chemistry of asphaltenes and other surface-active constituents in crude oil is complex, and the considerations presented here are brief. More comprehensive discussions on the subject are compiled in Refs. 38,47. In addition to surface forces, hydrodynamic forces between droplets in close contact can also slow down the coalescence process. The film of the continuous liquid phase between the droplets needs to be drained down to a certain critical thickness before coalescence can occur. The rate of film drainage is impacted by many factors, but two key parameters are the magnitudes of the so-called

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interfacial mobility and droplet deformation. The interfacial mobility relates to the friction at the liquid interface with the two limiting cases of full interfacial mobility (for which film drainage occurs fastest), and zero interfacial mobility [48]. Interfacial mobility decreases with increasing viscosity of the dispersed phase, and typically also with increasing coverage of the interface with surfactants (the Marangoni effect [49]). Droplet deformation can also impact the rate of film drainage, and in the case of the formation of a flat film between droplets, the drainage rate is substantially decreased compared to spherical droplets [50]. For other shapes resulting from droplet deformation, such as dimples, complex relations for the film drainage as a function of the film geometry are obtained [51]. Fig. 3.9 schematically summarizes the colloidal forces and hydrodynamic aspects that contribute to the kinetic stability of emulsions. A WiO or OiW emulsion is a complex fluid, but even more complex systems such as multiple emulsions [52] and microemulsions [53] can occur during crude oil processing posing additional challenges for separators. Other problematic systems such as the rag layer [54] and “schmoo” [55] also exist. The production process can be optimized to minimize the stability of formed emulsions, but often their formation is unavoidable. Hence, the process of

Colloidal Stabilization Electric Double Layer Force

Hydrodynamic Factors Interfacial Mobility Surface Charge (Native or From Adsorbed Species) Drop Viscosity Impacts Continuous Phase Velocity Near Interface

Steric Forces

Adsorbed Single Molecules (Asphaltenes, Fatty Acids)

Drop Deformation

Flattening of Liquid Interface Upon Approach

Mechanical Forces

Organic Particles and Nanoaggregates

Marangoni Effect Adsorbed Species Induce Surface Flux

Solid-Like Films of Adsorbed Molecules

Inorganic Particles

FIGURE 3.9 Overview of the different forces contributing to crude oil emulsion stability.

3.4 Theory of emulsion separation

emulsion breakup also needs to be understood and requires a different theoretical framework as it is not just the inverse of the breakup process. In the next section, we will discuss the basic ideas and models to understand and predict emulsion separation, and how these learnings can be applied to the design and sizing of liquidliquid separators for crude oil processing.

3.4 Theory of emulsion separation When an emulsion enters a separator, droplets need to settle under the influence of gravity and coalesce into bulk oil and water phases for separation to be considered successful. In a perfectly mixed emulsion, one phase will only exist as droplets dispersed homogeneously in a continuous phase, with no free bulk liquid of either oil or water being present. Often a pre-separation of the oil and water phase has already taken place upstream of the separator, and so parts of either phase can be present as a bulk liquid. Even if that is not the case, the settling of droplets and consecutive coalescence events between droplets will eventually lead to the formation of a bulk liquid phase. In most circumstances, at any given vertical cross-section through a separator, there will thus be some amount of liquid bulk phase (both oil and water), and a layer of emulsion above or below that bulk layer. Both WiO and OiW emulsions can occur in crude oil processing, but at a given location in the processing plant rarely simultaneously. Which type of emulsion is formed preferably, is partially dictated by the volume fraction of either liquid. If the liquid volume fraction of one phase is significantly smaller than that of the other phase, it is more likely to form the dispersed phase [56]. Additionally, the phase where emulsifying agents are more soluble (if present at all), tends to be the continuous phase as a rule of thumb [50]. Fig. 3.10 shows a scheme that visualizes the phase distribution of gas, oil, water, and the emulsion phase inside a simple separator with an overflow weir plate as the only mechanical structure added to the interior of the vessel. The situation in Fig. 3.10 is the ideal outcome of the separation process, with stratified bulk oil and water phases, and a WiO emulsion layer in this example, having formed as the liquid moves through the separator. The insert on the right shows a scheme that visualizes a typical distribution of droplets in the WiO emulsion layer in the vicinity of the weir. As the settling of droplets has already taken place, a layer of droplets has formed in the lower part of the emulsion, the so-called dense-packed layer (DPL). The vertical velocity component of droplets is almost zero, and vertical displacement only takes place when droplets coalesce with each other or in the bulk phase. Above the DPL, the droplet concentration and droplet diameter both decrease with increasing height. The distinction between what is a dilute WiO emulsion and a bulk oil phase becomes arbitrary at some point, but if we use the definition of e.g., 0.5% v/v WiO to define a bulk oil phase, we can

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gas outlet

inlet

gas

bulk oil

settling zone

oil water-in-oil emulsion

dense-packed layer water

bulk water

water outlet oil outlet

FIGURE 3.10 Schematic depiction of the phase distribution inside a simple separator. The insert on the right shows the phase distribution in the liquid column during the separation process.

assign an interface between emulsion and bulk oil as shown in the left panel of Fig. 3.10. Even inside the DPL, the phase distribution is not equal. In a WiO emulsion, the volume fraction of water increases from top to bottom, as droplets become more densely packed and deformed due to the weight of the emulsion phase on top of or below them [57]. The best outcome would be that the emulsion is completely broken before reaching the weir, but as long as no emulsion passes over the weir, or gets carried into the water outlet, a phase distribution like in Fig. 3.10 could be considered acceptable. The effect of dispersed gas bubbles on emulsion separation has not been the subject of much investigation so far, it is usually assumed that gas bubbles are large compared to droplets, and are quickly removed for the liquid. Relatively small gas droplets of the order 10100 μm can be created, however, upon degassing of the liquid when a pressure drop occurs somewhere in the system [58], such as at the different stages of depressurization between separators. It is known that gas bubbles can attach to and lift particles and droplets in flotation processes [59], but this process is usually applied in systems where the settling of the particles is more or less unrestricted. In a dense emulsion layer, the situation is less clear, and the presence of bubbles could, in principle, pose a hindrance to droplet coalescence and settling. The finding that stable dispersed systems containing droplets and bubbles (foamulsions) can be created [60], may also be relevant for the separation of crude oil emulsions containing gas bubbles. Because so little is known about the effect of dispersed gas on crude oil emulsion stability, all considerations that will be made in this chapter will neglect the impact of gas bubbles on emulsion separation

3.4 Theory of emulsion separation

rates. Likewise, we neglect the impact of entrained droplets falling from the gas phase back into the liquid on the overall liquidliquid separation efficiency. Solids that are produced from the well usually settle to the bottom of the vessel and are removed with flushing systems. Occasionally small particles, such as scale and corrosion products can get trapped inside the emulsion layer [47], and together with production chemicals and biological materials [61] can lead to the formation of extremely stable emulsions that require the addition of extensive amounts of demulsifiers to be broken. As this behavior also constitutes a rather special case, we shall not discuss the impact of solids on the separation dynamics in this section either. There are three popular concepts to describe the breakup kinetics of emulsions. All of them can inform the design of liquidliquid separators. The first one is based on calculating the settling velocity of droplets inside the separator under the influence of gravity, from which the required dimensions and internals of a vessel to achieve a given droplet removal efficiency can be derived. The concept of using droplet settling velocities to estimate separator efficiencies is the most popular for practical applications. It is simple to use and requires parameters that are usually readily available. It is an extreme oversimplification of the actual separation process, however, as it completely ignores the process of coalescence between droplets and with the bulk phase, and other effects such as elevated emulsion viscosity. For that reason, any predictions based on a model that only considers the settling of droplets need to be treated with some caution. The second concept is to track the evolution of emulsion breakup in time and spatial coordinates by using rates of droplet coalescence. This approach considers droplet interactions which are the key factor to accurately modeling the separation kinetics of an emulsion. While the general concept is fairly rigorous, the equations to calculate the coalescence rate can only be solved with numerous assumptions that have severely limited their practical use until recently. Computational Fluid Dynamics (CFD) is becoming a more and more powerful tool, though, and allows to couple the calculation of a multiphase flow field with the dynamics of droplet populations. Semi-empirical models have been derived to circumvent some of the challenges posed by the coalescence rate and settling-velocity-based models. These models form a sort of middle ground in terms of complexity and accuracy, but require tuning of their parameters based on specific properties of the emulsion considered, which usually need to be determined experimentally. In the following sections, we will discuss the three approaches to assessing and predicting emulsion separation rates.

3.4.1 Settling velocity of droplets In an oilwater separator, the main driving force for separation is gravity if the two liquid phases are of sufficiently different densities. The larger the density difference between oil and water, the faster separation is expected to occur.

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The simplest estimate for sizing of a separator is based on only considering crude oil densities, and rules of thumb for the required residence time of a liquid mixture in a separator when distinguishing between light, medium, or heavy crudes have been proposed [62]. This approach cannot be recommended for the sizing of a separator, but should only serve as a coarse assessment for the expected order of magnitude of the separator volume. A more detailed approach considers more parameters than only the crude oil density and is based on calculating the settling velocity of droplets using a simple hydrodynamic theory, which is derived from balancing the drag and buoyancy forces acting on a hard sphere in a quiescent liquid [15]. The resulting equation for the terminal settling velocity of a droplet is known as Stokes’ law: vz 5

Δρgd 2 18ηc

(3.4)

In the equation, vz is the settling velocity (in the vertical z-direction), Δρ is the density difference between the oil and water phase, ηc is the viscosity of the continuous phase, and g (59.81 m/s2) is the gravitational acceleration constant. Eq. (3.4) is only valid for a single isolated hard sphere moving in a quiescent liquid with a particle Reynolds number Rep # 1 [15]. It is the simplest expression for the settling velocity, and we note that the conditions required for it to be accurate are never met in a liquidliquid separator. Multiple improvements have been introduced to account for the actual physical reality a droplet finds itself in a separator, which are summarized in the list below. • •

•

•

•

For cases with Rep . 1, improvements to Eq. (3.4) have been derived to calculate the particle drag coefficient as a function of Rep [63]. The drag coefficient is also modified if larger droplets deviate from a spherical shape under the impact of gravity, and relations for different types of shapes have been derived [64]. During bulk separation, droplets are not isolated, but in a concentrated emulsion. The presence of neighboring droplets enhances the drag force on the droplets and decreases the rise velocity. This so-called hindered settling effect scales with the droplet concentration. The adjusted rise velocity vz of a droplet can be calculated by: vz 5 vz (1 2 φ)m [65]; where φ is the dispersed phase volume fraction, and m is a parameter related to Rep [66]. Droplets are not hard spheres as assumed in the Stokes equation, but viscous fluids. The flows induced inside and outside the droplets during the settling process also modulate the settling velocity of the droplet. This effect can be calculated as a first approximation from the HadamardRybczynski equation: vz 5 vz (ηc 1 ηd)/(2/3ηc 1 ηd) [67], with ηd as the dispersed phase viscosity. The enhanced droplet concentration in the emulsion layer also modifies the fluid viscosity. Instead of the continuous phase viscosity ηc, an effective emulsion viscosity ηe can be used. ηe is a function of the individual phase viscosities, the dispersed phase volume fraction, and droplet size distribution.

3.4 Theory of emulsion separation

Most expressions to calculate ηe neglect the effect of droplet sizes, and the popular Dougherty-Krieger approximation also ignores the viscosity ratio of the dispersed and continuous phase [68]:   φ ½ηφm ηe 5 ηc 12 φm

(3.5)

In the equation, φm is the maximum packing fraction, which is equal to 0.64 for randomly packed monodisperse spheres but can be higher for poly-disperse emulsions. ηc is the viscosity of the continuous phase, and [η] is the so-called intrinsic viscosity of the emulsion, which is equal to 2.5 for a suspension of hard spheres. For droplets with a finite viscosity, [η] decreases with decreasing ratio of the dispersed and continuous phase [69]. •

The droplet is not suspended in a quiescent liquid, there will be some degree of low-scale turbulence inside the separator vessel. Under typical separator conditions, it is expected that the settling velocity is reduced as compared to the Stokes settling velocity with a rather complex dependence on turbulence intensity and droplet diameter [70,71].

The fluid and flow properties in a separator are not spatially homogeneous. Thus, vz is not a constant. As the respective fluid and flow field properties at a given location inside the separator usually are not very well known, and a spatially resolved calculation of settling velocity typically is not considered practical, mean values for parameters such as the emulsion viscosity and turbulent energy dissipation rate are often used. While all these modifications to Eq. (3.4) can be expected to yield improvement in the prediction of the settling velocity, they still are coarse approximations of the physical reality. Even implementation of all of them would not obtain an accurate quantitative result. But the listed relations are useful to assess the general trend of expected separation efficiencies with properties such as viscosities and dispersed phase volume fractions. Eq. (3.4) can be utilized to estimate the dimensions for a liquidliquid separator. Fig. 3.11 shows the basic geometrical parameters that influence the separation efficiency inside a separator when using the Stokes-type of description for the settling process. A separator is a three-dimensional object, but we consider only a longitudinal cross-section of a separator for simplicity. As a worst-case assumption, we shall say that at the inlet the emulsion will fill the entire liquid phase with no bulk phases being present. As the fluid mixture passes through the vessel, the emulsion layer reduces in thickness, and bulk oil and water phases are formed. For a droplet to be separated successfully, it needs to join its bulk homo-phase. We consider a water droplet dispersed in an oil phase to be considered successfully separated if it settles from an initial starting position and reaches the water/emulsion interface. In principle, a water droplet can be considered successfully separated as long as it does not pass over the weir. In that scenario, a relatively large emulsion layer

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starting position

lsep oil

vax

hsep vz

water-in-oil emulsion

final position water

FIGURE 3.11 Schematic depiction of the key parameters for the Stokes settling model that governs oilwater separation in a horizontal separator vessel.

may remain in the vessel, which poses a risk to the outlet quality in case of flow upsets. We will thus use the more conservative criterion that the droplets must join their bulk homo-phase to be considered separated successfully. To accomplish this, the droplet needs to move a vertical distance ℎsep which marks the position of the emulsion-water interface. As the fluids also move horizontally with an average axial velocity vax, the droplet must reach the oil/emulsion interface before it has traveled the distance lsep between the inlet and the weir. This condition can be expressed as ℎsep/vz 5 lsep/vax. If vz is expressed via Eq. (3.4), Eq. (3.6) is obtained after rearrangement to calculate the droplet cut-off diameter dcut which signifies that all droplets of this diameter or larger will the separated for the given conditions: sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi hsep 18ηc dcut 5 vax lsep Δρg

(3.6)

The axial velocity of the droplet is usually chosen to be of the order of a few cm/s to keep the turbulence level contained and to ensure that the liquid phases are stratified. For simplicity, this value is often taken to be equal to the average axial velocity of the entire liquid volume or an assumed average axial velocity of the continuous phase. The chosen value for vax is based on finding a compromise to keep the diameter of the separator contained while reducing the turbulence level of the liquid flow to a sufficiently low level to not disturb the stratification of the flow. In addition to the length of the separation section (and the total length of the separator that follows from it), the other key geometric parameter is the diameter of the vessel (dves), which can be calculated for a given feed total volume flow rate Qf (gas and liquid combined) from: sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 4Qf
   dves 5 π φl vl;ax 1 1 2 φl vg;ax

(3.7)

3.4 Theory of emulsion separation

Where φl is the liquid hold-up fraction and vg,ax is the average axial gas velocity. In principle, Eq. (3.7) can be further refined by including the individual velocities and hold-up fractions for the water and oil phase [72,73]. The presence of an emulsion phase and evolution of the volume fraction of the three liquid segments throughout the separator is a complicating factor, however. As an example calculation, we assume a vessel of 2 m diameter, a liquid holdup fraction φl 5 0.7, and a water cut of the emulsion xwc 5 0.2. With these values, the cross-sectional area occupied by the oil phase in the separator is calculated from Aw 5 πdves2φl xwc/4. Generally, it cannot be predicted what the volume fraction of the residual emulsion layer is just upstream of the weir. We assume that when the oilwater mixture reaches the weir, all emulsion has disappeared, and correspondingly the height a droplet needs to move vertically to be considered separated (ℎsep), is equal to the height of the oil column in the separator. This height can be calculated from Aw using trigonometric relations for truncated circles [74]. With the above-mentioned parameters, ℎsep  0.4 m is obtained. We now assume that the separator needs to reduce the WiO concentration to 1% v/v. The droplet cut-off diameter corresponding to this performance can be calculated using Eq. (3.6) or any other type of distribution that is considered representative of a given system. We assume parameters typical for regular pipe flow upstream of the separator: d95 5 1500 μm and δ 5 2.5, to calculate a droplet size distribution using Eq. (3.2). With a water cut of 20%, the distribution for the oil volume fraction (xv) is normalized with a factor 0.8, and we obtain dcut  170 μm for which xv  0.01. Using Eq. (3.6) we calculate the required length of the separation section to be B9.5 m. A similar approach can be pursued for calculation of the outlet water quality. The outlined procedure to predict separator efficiencies based on Eq. (3.6) and an assumed feed droplet size distribution can be considered to form the basis of many frameworks for the sizing of separators. Numerous methodologies have been developed for sizing horizontal and vertical separators that also provide recommendations for liquid levels, nozzle sizes and positions, choice of internals, and other mechanical features [1,2,75]. The predicted trends for separation efficiencies as a function of fluid properties and velocities are usually correct, but actual numbers for separator dimensions may vary. The complexity of the calculations can be increased as desired to include factors such as emulsion viscosity, droplet size evolution, the effect of internals, and additional driving forces such as heating and electric fields. An increasing complexity does not necessarily imply higher accuracy, and care needs to be taken when translating the predictions of settling-velocity-based models into a separator design. An oilwater emulsion passing through a separator vessel constitutes a complex multiphase flow, and analytical models to describe droplet movement are a strong oversimplification of the actual situation. The models can be combined with existing empirical knowledge on the separation rate of different types of crude oil emulsion to become more reliable.

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3.4.2 Coalescence rates If the bulk water and oil phases are supposed to grow in volume inside the separator, coalescence of droplets at the emulsion-bulk interface needs to occur. Coalescence is the process where two droplets that are in close contact merge to form one larger droplet, or when a droplet joins its bulk phase. While gravity is the driving force that brings droplets nearby, efficient separation would not be possible without coalescence. The volume growth rate of a bulk phase dVb/dt can be calculated as the total volume of droplets coalescing in a given time interval with the bulk interface: dVb/dt 5 Vi Ni/ti. In the equation, Vi and Ni are the mean volume and number of droplets at the emulsion-bulk interface, respectively; ti is a characteristic time for droplets to coalesce with the bulk phase. In addition to the coalescence of droplets with the bulk-oil interface, coalescence between droplets in the emulsion layer also occurs, which leads to a growth of the mean droplet volume and diameter in the emulsion and corresponding shift of the entire droplet size distribution to larger droplet diameters. The basic parameter that describes these processes is the coalescence rate, which we can express for two droplets of radii ri and rj using a model similar to the collision theory of chemical reaction kinetics [8,11]: 

tc;ij rc;ij 5 θij pij 5 θij exp 2 τ ij

 (3.8)

In the above equation, θij and pij are the collision frequency and coalescence probability; tc,ij and τ ij are the mean coalescence time and the mean contact time between two droplets for given flow and fluid properties. θij can be calculated in a manner derived from the collision theory for gas molecules θij 5 ni nj σij Δvij [8]. where: ni and nj are the concentrations of the droplets, σij is the collision cross-section, and Δvij is a characteristic velocity difference between the droplets. The collision  2 cross-section for spherical droplets can be expressed as σij 5 π ri 1rj =4 [8]. For the coalescence rate rc,i of a droplet with diameter i in contact with its free bulk phase, Eq. (3.8) is still valid, with the exception of a different expression for the collision rate. If we treat the collision of a droplet with the bulk phase in the same way as that of a gas molecule with a wall, we obtain θbi 5 nivi/4 [76]. In a concentrated emulsion with a DPL, this expression is not as useful, as the approach velocity of a droplet in contact with the bulk phase is almost zero. While the general form of Eq. (3.8) is relatively simple, one challenge is to derive accurate descriptions for its parameters. Another challenge is that to calculate the separation rate of an emulsion, we need to calculate rc,ij and rc,i as a function of the spatial coordinates inside the separator. To calculate the collision cross-section in Eq. (3.8), the concentrations of droplets with a given diameter can be obtained by assuming a feed droplet size distribution. The relative approach velocity between two droplets Δvij or the approach velocity of a droplet towards a bulk interface vi, can be estimated with some

3.4 Theory of emulsion separation

assumptions. For turbulent flow, the average droplet approach velocity can be expressed as a mean value derived from the intensity of the flow field: Δv B (Ad)1/3 [48]. For two droplets colliding under the influence of gravity, the relative approach velocity can be obtained from their vertical velocity components vz: Δvij 5 vi,z 2 vj,z. The individual droplet velocity can be calculated from Eq. (3.8), and both the turbulence- and sedimentation-induced velocities can be combined to calculate a net droplet approach velocity. The term pij is the coalescence probability of two colliding droplets with radii ri and rj. It can be calculated from Eq. (3.8) as pij 5 exp(2tc,ij/τ ij). tc,ij is the average time it takes for the two droplets to coalesce while being in close contact (film drainage or coalescence time), and τ ij is the average time the droplets spend in close contact (contact time). Characteristic contact times between droplets can be calculated for a given hydrodynamic flow environment. As an example, in isotropic turbulence, the average contact time will be of the magnitude τ ij B 2=3 rij /A1/3 [77]. rij is an effective hydraulic radius of the two droplets and is calcu

21 [48]. For two droplets of different diameter collidlated as rij 5 0.5 ri21 1rj21 ing with each other in a separator under the influence of gravity and turbulence while in the neighborhood of other emulsion droplets, more complex relations are required. For droplets that are embedded in the DPL of the emulsion, the characteristic contact time can be assumed as equal to the residence time of the droplet in the DPL, as droplets do not leave the DPL once they have entered it before coalescing with another droplet or the bulk phase. An estimation for the coalescence (or film drainage) time tc,ij is also needed. For two droplets to coalesce, the liquid film of continuous phase between them needs to be drained up to a certain critical distance ℎc between the droplet interfaces. At distances smaller than the critical film thickness, the film of continuous phase can spontaneously rupture and coalescence can take place [50]. Different mechanisms for film rupture have been proposed [44], but a key finding of theoretical and experimental investigations is that the critical film thickness typically is of the order of several to tens of nanometers [50]. Calculation of the time it takes two droplets to approach each other from a given initial distance to the critical film thickness can be calculated using hydrodynamic theory. The simplest model to calculate the film thinning rate assumes the head-on approach of two equal-sized droplets that are considered hard spheres. For this case, the rate of film thinning (dℎ/dt) can be calculated from [44]: 2 dh 2 2h ½F 2 Fc ðhÞ 5 dt 3πηc rij2

(3.9)

In the above equation, h is the film thickness, rij is the effective hydrodynamic radius of the colliding droplets, and F is the driving force that brings the droplets together. As an example, for gravity settling, we would use a constant value of F 5 4πrij3 Δρg/3. Another force contribution to Eq. (3.9) is the term Fc that

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signifies the colloidal forces between the droplet interfaces (see Eq. 3.3). These forces are a function of the distance between the interfaces and are composed of several contributions as discussed in the previous section of this chapter. For the simplest emulsion systems, Eq. (3.9) can be solved to calculate the time it takes between droplet approach and film rupture, i.e. the coalescence time. For almost all practical cases involving crude oils, the required parameters such as critical film thickness and accurate terms for the surface forces are not available, however. Other complicating factors are the impact of interfacial rheology and the deformation of droplets during a collision on the film drainage rate [50]. These effects cannot be expressed in a simple analytical expression as Eq. (3.9), but are of crucial importance for an accurate calculation of the film drainage rate, in principle. Fig. 3.12 visualizes the concepts of contact and coalescence times, and shows examples of droplets that coalesce or separate upon collision. The droplets were colliding in a microfluidic channel (panel A) that is described in more detail in [78]. Droplets are created upstream of the chamber in a microfluidic circuit and pass through the collision chamber where they undergo random collisions. The image sequence in panel (B) shows different stages of the coalescence process for the droplet collision that corresponds to the film drainage curve in panel (C) which results in coalescence. For both the case with and without coalescence, the two droplets are initially separated, and then spend some time in close contact before either coalescing or separating. If a sample of the fluids is available, direct measurement of the coalescence time is the best option to obtain reliable quantitative information that can be used

(C)

(A)

h 3 mm

Separation

(B) 100 µm

Coalescence 0 ms

15 ms 35.4 ms 35.6 ms 36 ms

FIGURE 3.12 Visualization of film drainage and coalescence between emulsion droplets. Panel (A) shows a microfluidic channel of a height of 50 μm. Monodisperse droplets are created upstream and enter the channel where they collide and coalesce with each other. The experiment allows measuring the coalescence time between colliding droplets, such as shown in panel (B). Panel (C) shows the approach distance for two pairs of colliding droplets as a function of time. One collision ends with coalescence, and the other one with separation of the droplets induced the external flow.

3.4 Theory of emulsion separation

(A)

(B)

λ= 16.5

λ= 59.3

λ= 6.47

FIGURE 3.13 Experimental results from microfluidic experiments on droplet coalescence. Panel (A) shows a typical distribution of coalescence times obtained from a microfluidic collision channel (see Fig. 3.12). Panel (B) shows the average coalescence time for different mineral oil/water systems obtained from the microfluidic experiments as a function of the Capillary number. Reproduced with permission from T. Krebs, C. Schroe¨n, R. Boom, Coalescence kinetics of oil-in-water emulsions studied with microfluidics, Fuel 106 (2013) 327334.

to predict the separation rate of the emulsion with Eq. (3.8). Multiple experimental approaches exist to measure coalescence times, such as thin-film cells [51], micropipettes [79], settling chambers for single droplets experiments [80], and microfluidic circuits [8183]. Fig. 3.13 shows a distribution of coalescence times that was obtained from a microfluidic experiment conducted with the same experimental setup as shown in Fig. 3.12. Due to the somewhat random collision conditions, a range of impact parameters and a relatively broad distribution of coalescence times is obtained. The average coalescence time for different mineral oil/water systems obtained from the microfluidic experiments is shown in Fig. 3.13 as a function of the Capillary number Ca 5 γdηc/σ. In the equation, γ is the mean local shear rate near the droplets, and the term γd can be considered to be a characteristic velocity difference between the colliding droplets. The different data points in Fig. 3.13 are for three oilwater systems with different viscosity ratios between the dispersed and continuous phase λ 5 ηd/ηc, the lines are power-law fits to the data. Despite the aforementioned complications, the application of Eq. (3.8) to calculate the coalescence dynamics in a dense emulsion based on theoretical grounds alone has had some success when combined with CFD simulations of large-scale separators. The idea is to combine the physical processes that lead to droplet coalescence [84] and breakup [85] with calculations of the velocity and phase distribution of a multiphase flow in a separator vessel. Oshinowo et al. [86] conducted continuous-flow experiments with a separator of 0.7 m diameter and 3 m length,

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(A)

(B) Experiment

CFD

FIGURE 3.14 Schematic depiction of the liquid-separator used to study liquidliquid separation dynamics (panel A). Panel (B) shows a comparison of the experimental and theoretical separation efficiencies for different experimental conditions. Reproduced with permission from L. Oshinowo, R. Vilagines, Modeling of oil-water separation efficiency in three-phase separators: effect of emulsion rheology and droplet size distribution, Chemical Engineering Research and Design 159 (2020) 278290.

and the same geometry was probed in CFD calculations. Fig. 3.14 shows oil water separation efficiencies calculated with CFD and the experimental data for different water cuts. The comparison with pilot-scale experimental results yielded very good agreement between the calculations and experiments. This and some other encouraging results [8789] demonstrate the potential of state-of-the-art CFD to predict the liquidliquid separation performance inside large-scale separators. Aspects related to interfacial chemistry were not considered in the CFD model, as the experiments were done with a light crude where the

3.4 Theory of emulsion separation

concentration of surface-active and emulsifying constituents of the crude oil phase is usually lower than for medium and heavy crudes. It thus remains to be seen if CFD modeling can accurately predict the breakup dynamics of chemically more complex emulsions.

3.4.3 Semi-empirical approaches A third popular approach attempts to circumvent the problems associated with the two preceding takes on the subject and revolves around establishing models to predict separation rates in large separation vessels using experimental data from batch settling tests. The simplest and most straightforward approach to assess the stability of a crude oil emulsion is to conduct a batch settling experiment, also called a bottle test. In this experiment, relatively small volumes of crude oil and a water phase (produced water, or make-up water that mimics the produced water composition) are heated to the desired temperature and then mixed with an intensity that is assumed to be representative of the turbulent and shear forces upstream of the separator. The mixture is allowed to settle, and the positions of the interfaces between crude oil, emulsion, and water phases are tracked as a function of time, from which the separation rate can be calculated. Quantitative and automated techniques such as conductivity [90], turbidity [91], and nuclear magnetic resonance probes [92] are available to provide a convenient and systematic means to track the separation process, and even to determine the vertical distribution of oil and water in the entire liquid column during the settling process. Many different protocols are available which all work equally well, as long as the same procedure is applied every time a test is performed, to ensure results obtained with different fluids are comparable [2,93]. The bottle test has a prominent role in assessing the separation rate of an emulsion. It is also used to investigate the effect of other driving forces for separation, such as heat, chemicals, and even electricity. It is always a good idea to execute a bottle test to assess the order of magnitude for the timescale of emulsion decay. The main challenge in the interpretation of the results is that no generally accepted methodology exists for translating the measured separation rate to the conditions inside a full-scale separator. This problem can be addressed by extracting parameters from a batch experiment describing the internal emulsion breakup dynamics that can be inserted into breakup models developed for a continuous flow environment. The crucial parameters for liquidliquid separation are the formation rates of the bulk oil and water phase. Fig. 3.15A shows a schematic evolution of the settling process in a batch WiO emulsion as an example, with the bulk water phase height (labeled as ℎb), the DPL of droplets and the dilute free settling zone. The height of these zones can be measured as a function of time in experiments. The simplest approach to model the batch separation of emulsions is using

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FIGURE 3.15 Schematic depiction of the phase distribution in a batch emulsion settling under the influence of gravity with the corresponding height of the different sections of the liquid column as a function of time (panel A). Reproduced with permission from M. Henschke, L. Schlieper, A. Pfennig, Determination of a coalescence parameter from batch-settling experiments, Chemical Engineering Journal 85 (2002) 369378 [94].

empirical constants that are specific to a given emulsion system, and a relation to calculate the rate of bulk phase formation (dℎb/dt) is given by [95,96]: 21 1 1 5 1 dhb =dt k1 h k2

(3.10)

From measuring the growth curve of the bulk phase, the parameters k1 and k2 can be extracted. Eq. (3.10) describes batch settling, but an equivalent relation can be written for continuous flow with a volume flow rate Q through a separator with cross-sectional area A [96]: 21 1 1 5 1 Q=A k1 hb k2

(3.11)

Various models have been developed to give a physical meaning to the empirical constants k1 and k2. A model developed by Henschke et al. relates characteristic coalescence times to properties of the DPL [94]: τ5

ð6πÞ7=6 ηc ra7=3 4σ5=6 H 1=6 rf rv

(3.12)

In the equation, H is the Hamaker constant, which determines the magnitude of the attractive van-der-Waals force between droplets [37]. The mean droplet diameter in the DPL is included implicitly in the parameters ra and rf, which describe the geometry of deformed droplets in the DPL under buoyancy pressure, and which can be calculated theoretically if the mean droplet diameter d and the height of the dense-packed zone ℎp are known. rv is a dimensionless parameter also related to droplet deformation, but which is treated as an empirical constant for a given emulsion system. The equation can be used to calculate the coalescence time between drops in the DPL and between drops and the bulk interface with different

3.5 Emulsion separation techniques

assumptions for the parameters describing interfacial deformation. Together with a mass balance for the fluxes of the liquid phases in the liquid column during settling, a set of equations is obtained that provides a means to calculate the formation rate of the free bulk phase (dℎb/dt). The empirical parameter rv is obtained by fitting the model to the decay curve of the emulsion, if information about the droplet diameter is also available from the experiments, ideally with spatial and temporal resolution. If that is not the case, then a second empirical parameter is required to be fitted to the experimental data, which implicitly relates to the mean droplet diameter in the DPL. Fig. 3.15B shows a comparison of the model calculations by Henschke et al. and experimental data for a cyclohexanewater system. The evolution of the sedimentation and DPL was measured together with the mean droplet diameter at different heights of the liquid column. The model calculations showed reasonable agreement with the experimental data when using a value of rv 5 0:073. Comparison of the results from batch settling and continuous flow experiments to validate the approach described in this section has not been done frequently, but there are some studies. Polderman et al. combined data from batch experiments and field operations for the same type of crude oils, and derived envelopes for fluid velocity in the separator versus crude oil viscosity [95]. They did not directly compare the separation rate in the continuous flow system with that of batch experiments, however. Palermo et al. used a model for the batch decay of crude oil emulsions [97] and extracted kinetic parameters that were used to predict the performance of a pilotscale continuous separator. The measured separation efficiencies were in good agreement. There are indications that the measurement of batch emulsion decay can be used to predict the performance of commercial-scale separators, and if more data is emerging in the future, the suitability of this approach can further be probed. In this section, the basic aspects that are relevant to understanding and describing emulsion stability, as well as practical approaches to assessing and predicting separation rates and separator performances, were summarized. Our outline is one possible path from how a fundamental perspective can be used to aid in the design and sizing of commercial-scale oilwater separators. We considered only gravity as the driving force for separation, which in many cases is not sufficient to arrive at economically feasible separator designs. Additional driving forces that can be employed to increase the oilwater separation rate will be discussed in the following sections.

3.5 Emulsion separation techniques Gravity is always the main driving force to induce emulsion separation in crude oil processing. If the required residence time to separate the crude oilwater mixture is larger than a few minutes, however, the required separator volume will be too large to be economical in offshore or subsea installations. For onshore operations, larger residence

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times can be tolerated. But if the emulsion is particularly stable, additional driving forces to boost the separation rate may also be required there. Four major driving forces that are commonly employed to enhance oilwater separation are listed below: 1. Thermal demulsification: An increase in temperature of the liquid decreases the viscosity of the bulk liquid phases and the emulsion, and thereby increases the settling velocity of droplets. 3. Chemical demulsification: The addition of suitable chemicals to the liquid lowers the stability of emulsion droplets against coalescence. 3. Mechanical demulsification: Installation of mechanical internals inside the vessel conditions the flow and provides an additional surface area to increase the rate of droplet coalescence. 4. Electrostatic demulsification: Application of an electric field across the emulsion polarizes and/or charges droplets, which leads to an increase in the collision rate between them, and aids in overcoming forces that stabilize droplets against coalescence. In many processes, several or even all of these driving forces are employed to maximize separation efficiency. For each of them, there are one or more parameters that need to be optimized to balance the incurred expenditures with the enhancement of separation efficiency. What all methods have in common is that their utilization requires no or only a marginal additional pressure drop over the vessel which is usually considered highly desirable in crude oil processing. In this section, we will present the four driving forces to enhance emulsion separation, discuss the mechanisms by which they act on the emulsion in more detail, and show examples of equipment that is utilized for their application. For each of the different methods to enhance emulsion breakup, we will also present a brief survey of interesting field cases reported by operating companies. In a laboratory, experiments can be conducted with relative ease to verify the impact of different parameters on emulsion separation rates. The next step is the execution of a field trial, but the ultimate test for each technology is an application in a production environment at full scale. In this setting, it is usually much more difficult to conduct in-depth systematic studies, as the production process should not be upset or even disrupted. Valuable insights on the effect of a particular driving force on emulsion separation often can be obtained when one or several aspects of the production are changed, as this allows comparison between before/after or on/off scenarios. Operators occasionally publish their experiences with optimizing production processes, and for each driving force, we have selected case studies where their impact is readily visible. For reference, we also do a brief survey below of other driving forces that have shown potential for liquidliquid bulk separation in laboratory tests, but which are rarely used in commercial projects or not at all: •

Filtration: Utilization of filters and membranes to separate the two phases based on size exclusion and/or chemical affinity towards the filter element is

3.5 Emulsion separation techniques

•

•

•

•

not recommended for bulk liquidliquid separation. The two main reasons are the relatively low filtration flux [98] and a high propensity for clogging and fouling of the membrane and filter modules [99]. The former corresponds to a high footprint of the installation, and the latter requires frequent cleaning and/ or filter replacement. Filtration methods are more commonly used in produced water treatment, as dilute OiW emulsions generally are less challenging to process than bulk emulsions [100]. Ultrasound: Acoustic forces have been proposed as a driving force to enhance emulsion separation [101]. Laboratory studies have demonstrated the potential of ultrasound irradiation to enhance the separation rate of emulsions when a particular acoustic pattern is utilized [102,103]. The technology is not used on a commercial level at this time to our knowledge, and it has been argued that ultrasound treatment is more effective for dilute systems rather than concentration emulsions [104]. In addition, a detailed cost-benefit or lifecycle analysis of the technology is not available. Microwave heating: Radiative heating of emulsions using microwaves has been proposed as an alternative to traditional heating methods [105]. While laboratory studies indicate that the separation rate of a small quantity of emulsion can be enhanced using microwave radiation with a lower energy consumption than traditional heating [106,107], large-scale tests have not been conducted. The economic feasibility of the method in commercial systems also has not yet been demonstrated. Enhanced gravity: Centrifuges and inline cyclonic devices are useful technologies to increase the separation rate of emulsions, due to the exertion of forces on droplets that are equivalent to several hundreds of the normal gravity force. For bulk liquidliquid separation in oilfield processing plants, it is rather unusual to utilize either method, however. Centrifuges have a relatively low throughput and thus require a relatively high footprint [108]. In addition, the moving parts require regular maintenance and/or replacement. Inline cyclonic devices, on the other hand, are typically not very efficient for bulk liquidliquid separation due to the low residence time of the liquid mixture in the compact unit, unless the emulsion viscosity and droplet concentrations are sufficiently low [109,110]. Although these limitations hold for bulk emulsions, both centrifuges and inline cyclones are commonly used in the processing of dilute OiW emulsions that are obtained as effluents from the bulk separation process [111,112]. Magnetic forces: Crude oil is weakly diamagnetic and thus does not strongly interact with magnetic fields [113]. The interaction of emulsions with magnetic fields can be enhanced through the addition of magnetic nanoparticles that adsorb at the oilwater interface of the droplets [114,115]. When a magnetic field is applied, droplets are pulled out of the emulsion and form a bulk phase. The magnetic particles can be recovered and reused. This process has been researched in laboratory experiments but has not seen commercial application yet.

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•

Biological demulsification: Bacteria and other micro-organisms produce chemicals that can act as demulsifiers on crude oil emulsions [116]. The addition of nutrients to sustain sufficient bacterial populations in the water phase can be more cost-efficient than the addition of synthetic demulsifiers, in principle. The separation rate induced by biological demulsification usually is of the order of hours or even longer [117], which requires residence times that are unfeasible for most crude oil processing facilities. Generally, this subject has not been researched as much as synthetic demulsifiers, and there are still a lot of unexplored territories both from a scientific and engineering point of view.

3.6 Thermal demulsification 3.6.1 Effect of heating on emulsion properties Heating a liquidliquid mixture inside a separator will impact the viscosity of both the bulk phases and the emulsion layer. The temperature dependence of the viscosity η of a pure liquid can be expressed by Ref. [8]: η 5 aeb/T. This relation is empirical, and should only be used in a sufficiently narrow temperature interval (order of  100K) for given values of the parameters a and b [118]. The exponential relationship enables a strong decrease of viscosity with a relatively moderate temperature increase. For free droplet settling, a decrease in continuous phase viscosity corresponds to an increase in the droplet settling velocity via Eq. (3.4). The temperature dependence of emulsion viscosity can be based on the viscosity of the continuous phase via Eq. (3.8). In the DPL of the emulsion, the droplet settling rate will not be affected much, due to the high droplet volume fraction. But the film drainage rate for two drops in close contact increases with decreasing viscosity of the continuous phase (see Eq. 3.9), and hence an increase in coalescence rate based on this effect can be expected. A detrimental effect of heating is that constituents of the liquid will be boiled off from the crude oil, starting with the light hydrocarbons [119]. This will lead to a volume shrinkage of the crude, and an increase in its viscosity when cooling down. The evaporating components leave the separator with the gas phase and are either lost when the gas is flared or need to be removed in a downstream processing step if the gas is valorized [120]. A temperature change also impacts the properties of the interfacial films that stabilize droplets against coalescence. Temperature affects liquidliquid interfacial tension and quantities of adsorbed molecules, including demulsifiers [121,122], as well as the elasticity and viscosity of the adsorption layer [123]. The observed effects of temperature especially on the interfacial rheology are complex, but one interpretation is that a temperature increase reduces the rigidity of viscoelastic films formed by e.g., asphaltenes, which in turn would be expected to result in a higher rate of film drainage [124]. The interfacial tension between two pure liquids typically decreases with increasing temperature, but for crude

3.6 Thermal demulsification

oilwater systems also more complex relationships can be obtained (see e.g., Ref. [125]). In turbulent flow, a lower interfacial tension facilitates droplet breakup, and thus emulsification. In a separator vessel, flow conditions should favor coalescence instead of a break-up, so an increase in temperature is not expected to lead to a noticeable degree of re-emulsification. Droplets can be stabilized against coalescence through the adsorption of films and aggregates of compounds such as asphaltenes and wax particles [126]. A temperature increase can lead to the dissolution of these structures into the bulk oil phase, thereby reducing the barrier for coalescence. As an example, we show data from work by Thompson et al. [127]. Fig. 3.16A shows the fraction of separated water from a batch crude oil emulsion with 20% water cut after 24 h as a function of temperature. Fig. 3.16B shows the relative concentration of crystalline wax particles in the emulsion. At around 50
C all wax particles are dissolved which coincides with a strong increase in the separation rate of the emulsion. In their work, Thompson et al. provide convincing arguments that there is indeed a causal relationship between the concentration of wax particles in the emulsion (and at the droplet surfaces) and the stabilization of the droplets against coalescence for the crude oil emulsion under investigation [127].

FIGURE 3.16 Separated fraction of water from a batch water-in-crude oil emulsion after 24 h for different temperatures (panel A). Panel (B) shows the relative concentration of crystalline wax particles in the emulsion as a function of temperature. Reproduced with permission from D. Thompson, A. Taylor, D. Graham, Emulsification and demulsification related to crude oil production, Colloids and Surfaces 15 (1985) 175189.

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To assess the combined effect of temperature on all these complex processes and physical properties of the crude oil emulsion, a bottle test conducted at different temperatures may still be the best tool to quantify the relationship between separation rate and temperature. If crude oil samples are not available, estimates for separation rate enhancement through heating have to be made based on modulation of liquid viscosities only.

3.6.2 Heater technology In crude oil processing, heat typically is introduced through a fired tube using fuel gas [1]. A dedicated vessel to pre-heat an emulsion can be used before it is processed in another separator downstream, or a so-called heater-treater unit can be utilized, where heating is combined with liquidliquid separation [3]. Fig. 3.17 shows an example of such a unit in the horizontal configuration. The emulsion is more or less poured onto the fire tube, and rapidly formed free water is separated in the first compartment of the vessel, as well as any gas that is released upon heating. The remainder of the emulsion is separated in the second

FIGURE 3.17 Schematic illustration of a horizontal heater-treater vessel. The emulsion is heated with a fire tube, and free water that is rapidly formed is separated in the first compartment, as well as released gas. The remaining emulsion and oil are separated in the second compartment of the vessel. Reproduced with permission from M. Stewart, K. Arnold, Emulsions and Oil Treating Equipment Selection, Sizing and Troubleshooting, Gulf Professional Publishing, Houston, 2008.

3.6 Thermal demulsification

compartment where more residence time is available. The flow path for the fluid mixture can be complex (such as in Fig. 3.17), but the actual heating element is a rather simple construction with often only a single large fire tube being inserted into the vessel. The advantage is that it can be integrated into the separator with relative ease. On the other hand, such a design does not provide efficient heat transfer, as it causes substantial heat losses and a less uniform temperature profile in the emulsion compared to e.g., tubular heat exchangers and boilers [128], or electric heaters [129]. Even though better alternatives are available, in principle, the heater-treater is a unit that up to the present day is still used widely in the industry.

3.6.3 Case studies A processing facility typically is designed for a relatively narrow range of temperatures in the different process steps. Variation of temperature not only impacts the separation rate of the emulsion, but also the composition of the gas and liquid phases, solidliquid equilibria relevant for flow assurance and material degradation, as well as operating costs. Any deliberate change in temperature thus needs to be carefully considered. Data on systematic variations of temperature in a processing facility are rarely reported, but the effect of temperature on the production process is strongly illustrated by e.g., seasonal variations of the ambient temperature on the fluids while being processed at a topside facility. Al-Ghamdi et al. report the optimization of demulsifier formulations for emulsions treated in several processing plants at the Ghawar oilfield [130]. The temperature of the emulsions arriving at the processing facility was in the range of B 60
C70
C in the period June-October, whereas in January and February the temperature was B 10
C15
C lower. To maintain complete emulsion breakage, demulsifier concentrations could be lowered by a factor B24 during the warmer months, illustrating the strong effect of a higher temperature on facilitating emulsion resolution in this example. For another oil processing facility in Saudi Arabia, Alhajri et al. [131] report the impact of temperature variations during the day and night cycle on the performance of the high-pressure three-phase separator. In one data set, the maximum fluid temperature during daytime was of the order B65
C, whereas during the night the minimum temperature is of the order B55
C. This difference resulted in a change in the water removal efficiency from B80% at the higher temperature to B60% at the lower temperatures. An example of a more direct effect of temperature on emulsion separation rates was reported by Razalli [132] for the Guntong oil field in Malaysia. “Cold” emulsions (B20
C25
C) produced and transported from satellite platforms towards a central processing facility could not be resolved with the existing separation equipment to less than B10%20% water cut remaining in the crude. The processing system was modified to mix the emulsions with the “hot” produced water (B70
C80
C) from the mother platform instead of using a heat exchanger. The flow rates of both process streams were not reported by the

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authors, but as a result, the effluent WiO concentration decreased to B 3%6%, and overall crude production could be boosted by B2%4%.

3.7 Mechanical internals The performance of a separator vessel can be boosted through the addition of mechanical parts. These so-called internals can accelerate the separation of all phases entering the separator, such as mist extractors for the removal of liquid droplets from the gas phase [133], and jetting systems at the bottom of the vessel to flush out accumulated sand particles [134]. In this section, only the internals that directly impact liquidliquid separation are considered, but we would like to stress that there is a strong interplay between all the mechanical internals, and a faulty design of one component can negatively impact the quality of every treated phase.

3.7.1 Separator vessels Fig. 3.18 shows a schematic drawing of a typical horizontal three-phase bulk separator. A separator can be coarsely divided into three sections (inlet, separation, and outlet) whose boundaries are determined by the location of certain internals. The inlet section is the region in the vicinity of the inlet nozzle, where the fluid enters the vessel with a high velocity. The inlet nozzle can be connected to mechanical internals which ensure that the turbulent intensity is dissipated mostly in the inlet region of the separator. Various inlet devices can be employed, such as deflection plates, vanes, and inlet cyclones [135]. The example in Fig. 3.18 shows an inlet cyclone as an entry point for the multiphase flow. Inlet devices also help to create a more uniform flow distribution [136] and provide a boost for gasliquid separation [133]. A device such as an inlet cyclone could be imagined to produce some

FIGURE 3.18 Schematic drawing of a typical horizontal three-phase bulk separator with multiple mechanical internals to enhance separation. The levels of the different phases inside the vessel are also indicated. Courtesy TechnipFMC.

3.7 Mechanical internals

degree of liquidliquid separation, but no conclusive data is published on this effect for gasliquid mixtures. Our laboratory testing on inline cyclonic devices indicated that at best a marginal degree of liquidliquid separation could be obtained in the presence of substantial gas volume fractions ($40% v/v). For liquidliquid mixtures in the absence of gas, inline cyclonic devices have been demonstrated to induce a measurable degree of liquidliquid separation [109]. A conservative estimate would be to assume that inside the inlet device no enhanced droplet coalescence takes place, but that their flow conditioning effect creates more favorable conditions for coalescence further downstream in the vessel.

3.7.2 Perforated baffles After entering the vessel through the inlet device, the fluid passes through one or several (often two) perforated plates. The function of these baffles is to break up large-scale flow circulations [137], mitigate the effect of sloshing in floating production systems [138], and create a uniform axial velocity distribution of the liquid (plug flow profile), which in turn creates a more uniform distribution of residence times of the liquid [139,140]. The emulsion passing through the baffle does not change its droplet size distribution, but the enhanced uniformity of the droplet residence time and elimination of large-scale vortices is expected to enhance the separation performance indirectly. The net free area (NFA) of the baffle and the number of holes are tuning parameters. The NFA and hole diameter should be optimized in a way as to have a sufficiently favorable effect on the downstream velocity distribution while avoiding excessive pressure drop and breakup of droplets if the holes are too small. Fig. 3.19 shows experimental data on the average axial velocity distribution just downstream of a perforated baffle

FIGURE 3.19 Panel (A): Standard deviation of the axial velocity just downstream of a perforated baffle inside a separator with single-phase water flow. For the results shown in panel (B), two baffles were employed. Data points show the results obtained for different distances between the baffles. Reproduced with permission from D. Wilkinson, B. Waldie, M. M. Nor, et al., Baffle plate configuration to enhance separation in horizontal primary separators, Chemical Engineering Journal 77 (2000) 221226.

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in a separator with single-phase liquid flow [141]. The data shows that reduction of the NFA to c. 20% produces a significantly more leveled distribution of the axial velocity as compared to the absence of a baffle. Moreover, there seems to be no clear trend for the axial velocity distribution as a function of hole diameter for the investigated range. When a second baffle is added, the trend for the velocity distribution is also not very clear, but the results suggest that when the two baffles are placed B510 cm apart, the velocity profile becomes slightly more uniform. In Fig. 3.19, the dimensions of the rectangular vessel were 0.88 m 3 0.23 m 3 0.28 m (length 3 width 3 height). Panel (A) shows experimental and theoretical results for different hole diameters and NFA of the baffle. For the results shown in panel (B), two baffles were employed. The data points show the results obtained for different distances between the baffles.

3.7.3 Plate packs After passing through the baffles, the liquid mixture enters the so-called plate pack. The plate pack is a set of parallel plates that intersect the flow (see Fig. 3.20). When the liquid mixture enters the plates, the vertical settling velocity of a droplet remains unchanged, and the axial velocity also remains roughly the same [142]. What changes, though, is that the distance for a droplet to rise to the top or settle to the bottom of the channel is drastically smaller than the vertical distance it would need to pass in a separator without internals. When the droplet reaches the top or bottom of the channel, it is not immediately separated, but it coalesces with other droplets that have settled, and form a thin continuous film. When this film leaves the channel, it may break up into large droplets, which quickly settle to join their bulk phase. If the plates are inclined by an angle θ as shown in Fig. 3.21, the outflow of the film upwards or downwards is facilitated and the accumulation of solids is prevented. The channel height is chosen small enough so that it measurably enhances the separation efficiency, but also enables laminar flow through the channel as opposed to the turbulent flow in other parts of the vessel. Laminar

FIGURE 3.20 Schematic drawing of a plate pack with inclined plates, and its position inside a gravity separator. Reproduced with permission from K. Arnold, M. Stewart, Surface Production Operations, Volume 1: Design of Oil Handling Systems and Facilities, Gulf Professional Publishing, Houston, 2008.

3.7 Mechanical internals

FIGURE 3.21 Illustration of the mechanism of separation enhancement in plate packs with inclined plates [143]. Reproduced with permission from Y. Han, L. He, X. Luo, et al., A review of the recent advances in design of corrugated plate packs applied for oil-water separation, Journal of Industrial and Engineering Chemistry 53 (2017) 3750.

flow under typical flow conditions is expected to be beneficial to not disturb the settling process, and also prevent re-entrainment of droplets from the formed film [144]. The effect of insertion of a plate pack on separation performance as compared to an empty vessel can be estimated by replacing the full separation distance (ℎsep) in an empty separator in Eq. (3.6) for the cut-off diameter of the vessel by the channel height of the plate pack (ℎc), and by addition of the effect of the inclination angle (θ) [144]: sffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi vax hc 18ηc dcut 5 cosθ lsep Δρg

(3.13)

To illustrate the expected impact of a plate pack on liquidliquid separation, we consider the same case as we did when calculating the droplet cut-off diameter in Section 3.4.1. The separation height of the open vessel was ℎsep  0.4 m. For the empty vessel, we obtained a cut-off diameter of dcut  170 μm. A feed water cut of xwc 5 0.2 and outlet WiO volume fraction of xv  0.01 resulted in a length of the separation section of  9.5 m. Assuming a plate pack of channel height ℎc 5 0.02 m and an inclination angle of 45
as typical values, the required separation length (i.e. length of the plate pack section) to achieve the same cut-off diameter is calculated. An lsep value of  0.67 m is obtained, which is more than a factor of 10 lower than the required length of an empty vessel. The actual length of the

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separation section should be significantly higher than the length of plate packs, as some distance to the upstream baffles and downstream weir is required for optimal performance. The estimate made for the plate pack length is very simple, and more complex relations have been derived to describe the separation efficiency of plate packs that consider factors such as spatially resolved velocity profiles and the impact of the formed film on the flow dynamics in the channel [144146]. But the calculation illustrates the enormous effect that utilization of parallel plates can have on outlet qualities. Many studies published on plate pack separators have a commercial background, and independent systematic studies on the effect of plate pack geometry on separation performance are not that abundant. Experimental studies with model systems that can be considered to mimic crude oilwater emulsions indicate that OiW and/or WiO outlet concentrations can be reduced by factors B310 as compared to an empty vessel, when keeping the separator geometry otherwise unchanged [147149]. These data were collected for specific fluid properties and plate pack geometries, and for each given fluid mixture the expected effect of a plate on separation enhancement needs to be assessed again theoretically and/or experimentally. Plate packs with an angle of inclination relative to the vertical axis of the separator (typically in the range of 3060 degrees) are frequently used. Fig. 3.22

FIGURE 3.22 Effect of plate pack inclination and dispersed phase throughput on the separation efficiency of a toluene/water mixture. Reproduced with permission from T. Sekine, Solvent Extraction, Elsevier, Amsterdam, 1992.

3.7 Mechanical internals

FIGURE 3.23 Schematic drawing of the structure of corrugated plate packs with drainage holes for oil and sand. Reproduced with permission from Y. Han, L. He, X. Luo, et al., A review of the recent advances in design of corrugated plate packs applied for oil-water separation, Journal of Industrial and Engineering Chemistry 53 (2017) 3750.

shows a visualization of the effect of both the inclination angle and the flow rate on the separation efficiency of plate packs [150]. The smallest inclination and lowest flow rates provide the best separation performance, which is in line with the predictions made by Eq. 3.12. An alternative configuration is a corrugated plate pack, as shown schematically in Fig. 3.23 [143]. It has been argued that the curved flow path promotes coalescence between droplets, while also preventing the accumulation of solids [143]. For more concentrated emulsions this may be a reasonable assumption, but for more dilute systems where the settling process of individual droplets is more dominant than bulk coalescence, this assumption has been questioned both on theoretical and experimental grounds [144,151]. While there is some debate as to which plate shape is overall the best from a performance and reliability perspective, the very fact that a plate pack is utilized at all brings a strong enhancement to separation efficiency.

3.7.4 Pipe separators Pipe-based alternatives to vessel separators have been proposed (see Fig. 3.24), which exploit a similar effect as a plate pack, by reducing the distance a droplet has to traverse to reach its bulk phase as compared to a vessel-based gravity separator. In addition, due to the velocity slip between the oil and the water phase, a shear force is exerted on the emulsion in between, which is predicted to speed up the emulsion separation rate as long as a critical axial velocity in the pipe is not exceeded (B1 m/s) [153]. The residence time of a fluid mixture in the pipe will be much lower than in a gravity vessel (unless the pipe section is made extremely long), which means that stable emulsions or mixtures containing viscous oils and/ or low water cut cannot be separated with high efficiency [152,154]. For that

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FIGURE 3.24 Prototype of a pipe separator for removal of water from oilwater mixtures. The inlet is on the right side and the flow is distributed into two parallel pipes. Tapping points for removal of the water phase are on an inclined part of the pipe downstream of the separation section. The photograph shows the oilwater mixture inside the pipe near the tapping section [152]. Reproduced with permission from H. Skjefstad, M. Stanko, Experimental performance evaluation and design optimization of a horizontal multi-pipe separator for subsea oil-water bulk separation, Journal of Petroleum Science and Engineering 176 (2019) 203219.

reason, pipe separators are considered to remove only free water that has formed as large droplets drop out from the liquid mixture as it passes through the pipe. A single pipe also is not able to produce the required throughout, and arrangements of multiple parallel pipes have been proposed [155]. The pipe separator concept has been considered attractive especially for subsea separation as the separated water can be re-injected, pipes are more robust and safe structures than large vessels for deployment in deep water [156], and more compact processing equipment can be deployed topside if a substantial volume of water is pre-separated [157].

3.7.5 Case studies The effect of adding mechanical internals to a separator on the emulsion breakup rate is directly illustrated by case studies that involve an upgrade of the separator internals, and where separator performance is reported before and after the modifications. Bell et al. documented the results of a separator upgrade for the processing system of the Hutton Field [158]. The processing facility is installed on a tension leg platform that experienced substantial motion with a corresponding degradation in separator performance. During heavy weather, the outlet liquid qualities of the primary separators were B 5% WiO and B 1000 ppm OiW, as opposed to the target values of 2% WiO and 200 ppm OiW. Attempts to improve liquid level control and increase the concentration of demulsifiers did not alleviate the problem. The original separator was equipped with longitudinal and transversal baffles, which occupied a significant fraction of the separator volume. An experimental campaign conducted with a test separator in a laboratory suggested that the baffles did not have the desired effects to reduce large-scale turbulence and create a uniform residence time for the fluids in the separator. Replacement of the

3.8 Chemical demulsification

baffles with several plate packs gave positive results, and subsequently, the production separator was retrofitted with plate packs and the baffles were removed. The upgrade enabled the operator to achieve the target effluent concentrations, and in addition, demulsifier concentration could be decreased from 50 to 20 ppmv as compared to operation before the replacement. Increasing water production rates and/or addition of produced fluids from new wells often inevitably deteriorates the performance of separators, as the fluid residence time decreases with increasing throughput. In those cases, the addition or upgrade of separator internals may help to debottleneck the vessel. We mention two examples reported by Broussard et al. [159] and Frankiewicz et al. [160]. These studies do not contain much quantitative information on the degree of performance enhancement other than that the specifications for effluent oil and water streams from the separator were met after production was increased. They do describe in detail the procedures for modification, commissioning, and operation of the improved separators, however, and therefore can be recommended for further reading. In the reported cases, the addition of more complex internals such as the replacement of a deflection plate with an inlet cyclone, or replacement of flow straighteners with plate packs yielded an improvement in the separation efficiency.

3.8 Chemical demulsification In Section 3.3, the static and dynamic contributions to repulsive forces between droplets that can prevent coalescence were discussed (see Eq. (3.3) for the relevant force components and Fig. 3.9 for an overview). If these forces become too large, emulsions may be encountered during crude oil processing that cannot be resolved by heating and/or mechanical internals within the residence time available in the separator. This is especially the case for WiO emulsions containing heavier crudes. In addition to a higher viscosity, the concentration of naturally occurring stabilizing agents such as asphaltenes typically increases with increasing crude density [2,38]. OiW emulsions, on the other hand, can be stabilized against coalescence through production chemicals such as corrosion inhibitors [161]. In many cases, significant enhancement of the separation rate is only possible through the addition of chemicals that modulate the interfacial film of the droplets in a way that coalescence is promoted. The topic of crude oil chemical demulsification is complex and a large research field on its own, and in this section, we aim to only provide a coarse overview of the relevant aspects governing chemically induced demulsification. Chemical aides to enhance bulk separation typically are required to resolve WiO emulsions. OiW emulsions tend to separate comparatively fast due to the low viscosity of the continuous phase, and the often lower stability against coalescence for crude oil droplets dispersed in water as compared to water droplets

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dispersed in crude oil. The aqueous effluent phase from the bulk separator should contain less than 1% v/v oil in water, but this concentration needs to be reduced even further to recover residual oil and to obtain a water quality that is safe for discharge or re-injection. The separation of these dilute emulsions takes place in the produced water treatment systems where chemicals such as flocculants [162] and coagulants [163] are added to accelerate droplet aggregation. Produced water treatment is a large technology and research area on its own, and several excellent textbooks have been published on the subject [164,165] to which we would like to refer the interested reader.

3.8.1 Effect of demulsifier on separation rates Chemicals that are intended to destabilize emulsions are selected by their capacity to adsorb at droplet interfaces (such as long-chain amphiphilic molecules or polymers), because they impact the properties of the adsorption layer indirectly, such as through a change in pH [46,166] or by increasing the solubility of adsorbed components in the bulk liquid phase [44]. If droplets are stabilized by an existing adsorption layer, the added demulsifier molecules need to co-adsorb and/or replace part of that layer. The newly formed adsorption film should then have a lower tendency to stabilize emulsions. Fig. 3.25 shows the positive impact of the demulsifier sodium bis(2-ethylhexyl) sulfosuccinate (AOT) on the separation rate of a crude oil emulsion as a typical example [167]. Panel (A) shows the separation rate of the emulsion as a function of time for different demulsifier concentrations. Panel (B) shows the initial separation rate as a function of demulsifier concentration. Panel (C) shows the liquidliquid

FIGURE 3.25 Experimental data on demulsification of a water-in-crude-oil emulsion using the surfactant AOT as demulsifier. AOT, Sodium bis(2-ethylhexyl) sulfosuccinate. Reproduced with permission from R. Aveyard, B. Binks, P. Fletcher, et al., The resolution of water-in-crude oil emulsions by the addition of low molar mass demulsifiers, Journal of Colloid and Interface Science 139 (1990) 128138.

3.8 Chemical demulsification

interfacial tension as a function of demulsifier concentration. The crossed symbols indicate the values for the nominal AOT concentration in the water phase. The circles show the same values for the interfacial tensions, but the x-axis now is the residual concentration of AOT in the water phase after the partitioning of the demulsifier over both the oil and water phase. The water phase contained 0.5 mol/L of NaCl in all experiments. In the vast majority of cases, the liquidliquid interfacial tension between crude oil and water decreases upon adsorption of a demulsifier (see Fig. 3.24C for an example). The concentration of demulsifier (or any species) in crude oil at the interface cannot be accurately inferred from interfacial tension measurements, even though the Gibbs adsorption equation [8] often is erroneously applied as if this was possible [168]. But in any case, a measurable decrease in interfacial tension is generally considered a favorable indication that the demulsifier adsorbs at the interface. A decrease in interfacial tension also decreases the maximum droplet diameter (dmax) in turbulent flow according to Eq. (3.2), but a modest decrease of interfacial tension can be tolerated. The turbulence level inside the separator usually is low compared to the upstream piping, and a slight increase in the droplet breakup rate can be overcompensated by a stronger increase in the coalescence rate.

3.8.2 Mechanisms of demulsifier action A key mechanism of demulsification is that a demulsifier adsorbs at the interface and displaces a large fraction of the material that was initially adsorbed and created a barrier against droplet coalescence in the first place. As the demulsifier is designed in a way as to not stabilize emulsions themselves, droplet coalescence is enabled, and the separation of the emulsion is accelerated. Measurements of dynamic interfacial tension [169] and evaluation of the rheological properties of the interface [170] can give more quantitative information about the degree of disruption a demulsifier has on existing interfacial films. Adsorbed components such as asphaltenes in heavy oils can form rigid, solid-like layers that create a mechanical barrier to coalescence [41]. If a dense adsorption film is broken up by a demulsifier (see Fig. 3.26A for a visualization of this effect), the probability of formation of holes and subsequent coalescence is increased [171]. In doing so, the demulsifier may also decrease the interfacial viscosity, which increases the mobility of the interface and accelerates the rate of film drainage [124]. If particles are contributing to the stabilization of the interface, a demulsifier can also change the wetting behavior of the particle concerning the oil and water. The wetting tendency is usually expressed in terms of the contact angle formed by a liquid meniscus with the particle [37]. The insert in panel (B) of Fig. 3.26B shows the definition of the contact angle for a particle at an oilwater interface. Solids with a contact angle less than 90 degrees stabilize OiW emulsions, and those with a contact angle greater than 90 degrees stabilize WiO emulsions. A demulsifier that causes dewetting of the solid particles can then be used to break

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FIGURE 3.26 (A) Effect of a polymeric demulsifier on the structure of asphaltene films [171]. The left pictures show an atomic force microscopy image of a smooth and closed asphaltene film deposited on a glass surface. The picture on the right shows an asphaltene film that was deposited in the presence of a polymeric demulsifier. The film has broken up into patches with empty areas in between. (B) Impact of demulsifier on oil separation rate and contact angle on an emulsion stabilized by microscopic mineral particles [171,172]. The insert shows the definition of the contact angle for the solid-water-oil interface [173]. Reproduced with permissions from [171174].

the emulsion. As an example, Fig. 3.26B shows the correlation between the separation rate of an emulsion stabilized by inorganic particles and the contact angle modulated through the addition of a demulsifier [172]. Another driver for demulsification can be an adjustment of the pH. A change in pH can have two major effects, the first being a change in the degree of ionization of acidic or basic groups of molecules such as asphaltenes and naphthenic acids at the interface. For a given molecule, a pH exists (the isoelectric point) at which the molecule is neutral, with the consequence that the repulsive EDL force between droplets will be lowest, which in turn is expected to promote coalescence [37]. Charge neutralization of surfactants that are naturally occurring in crude oil can also increase the solubility of these constituents in the crude oil phase, thereby decreasing their interfacial activity [175], and promoting destabilization of the emulsion. The relationship between crude oil emulsion stability and pH thus depends on the crude oil composition. This fact is illustrated in Fig. 3.27A where the relative volume of separated water from different crude oil emulsions is shown as a function of the pH [33]. All curves display a pronounced maximum of the separated volume at pH values as low as B5 and as high as B11, indicating a substantially different susceptibility of the interfacial layer towards a change in pH for the different crude oils. In many cases, there also is an optimal concentration for the demulsifier that maximizes the separation rate of the emulsion. Fig. 3.27B shows an example of experimental results obtained with the commercial demulsifier BJ10 and a 10% water-in-crude oil emulsion [176]. A maximum volume of separated water is

3.8 Chemical demulsification

FIGURE 3.27 (A) Fraction of separated water for different water-in-crude oil emulsions as a function of the pH [33]. (B) Fraction of separated water of a 10% water-in-crude oil emulsion as a function of the concentration of the demulsifier BJ10 [176].

obtained at a demulsifier concentration of B1500 ppm, and for larger concentrations, the separation efficiency decreases. This behavior is often interpreted in terms of a re-stabilization of the interface against coalescence at high demulsifier concentration [177], and/or a decrease of interfacial tension with a corresponding decrease in droplet size, which in turn reduces the separation rate [178]. Overdosing can also be caused by incomplete mixing of the demulsifier with the crude, in which case high local concentrations of demulsifiers are established in certain volume elements, and low concentrations in others [178], both situations which result in an overall reduced demulsifier efficiency [179].

3.8.3 Demulsifier formulation A guide to the selection of a demulsifier to resolve a crude oil emulsion is provided by considering the solubility of the compound in either liquid phase. Surfactants with a large fraction of hydrophobic groups tend to be more soluble in oil, and those with more hydrophilic groups tend to be more soluble in water. For a given class of surfactants, the hydrophiliclipophilic balance (HLB) defines a scale for hydrophilicity and hydrophobicity of a compound: HLB 5 7 1 nℎpℎil 2 nℎpℎob [180]. In the equation, nℎpℎil and nℎpℎob are the number of hydrophilic and hydrophobic groups in the molecule, respectively. The HLB value is also a measure for the tendency of an oilwater film to bend towards either the water or oil phase, i.e. to form an oil- or water-continuous emulsion [181]. If the compound is more soluble in the aqueous phase, it tends to stabilize OiW emulsions, and vice versa [50]. Within this framework, it is predicted that if the component shows significant solubility in either phase (i.e. low interfacial curvature or a partition coefficient of the order unity [182]), neither OiW nor WiO emulsion stabilization occurs, and this is considered the

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“sweet spot” for demulsifier action. This concept, if used with caution, can contribute to the selection of demulsifiers for a given oilwater system. Other scales which do not require a definition of hydrophobic and hydrophilic moieties in a molecule can be used to complement the HLB concept, such as the water/octanol partition coefficient, or the relative solubility number (RSN) that is based on cloud point measurements for different solvent polarities [155]. Fig. 3.28 shows a map of emulsion type for a kerosene/water emulsion as a function of the water cut and HLB value of added surfactant [183]. The dotted line indicates the region of phase inversion, and this is where the ideal HLB value for the demulsifier would be located within the HLB framework. Fig. 3.29 shows the chemical structures of two (of many) types of demulsifier classes that are frequently used. Pluronic P123 is a member of the poloxamer family of non-ionic polymeric surfactants that can be used to destabilize WiO emulsions [184]. Nonylphenol Formaldehyde Resin (NPFR) belongs to another popular class of WiO demulsifiers, the alkylphenol formaldehyde resins [185].

FIGURE 3.28 Generalized scheme illustrating the impact of oilwater ratio and HLB parameter of surfactants added to a water-kerosene mixture, which is a coarse measure for the tendency to form OiW or WiO emulsions. In the transition region indicated by the dashed line, the emulsion is unstable. HLB, Hydrophilelipophile balance. Reproduced with permission from J.V. Kloet, L. Schramm, B. Shelfantook, Application of the hydrophilelipophile balance concept to the classification of demulsifiers and bituminous froth and its components, Fuel Processing Technology 75 (2002) 926.

3.8 Chemical demulsification

Pluronic P123

Nonylphenol Formaldehyde Resin (NPFR) FIGURE 3.29 Base structures of two molecular classes that are often used as demulsifiers for WiO emulsions [184,185].

The general idea behind choosing this type of molecule is that they are surfaceactive in crude oil emulsions, are non-ionic (and thus do not induce repulsive electric double-layer forces between droplets), and are large molecules that form relatively loose adsorption layers. Commercial demulsifiers are mixtures of several compounds that target the described mechanisms to maximize destabilization of the emulsion, and a myriad of compounds has been patented [161,186]. Co-solvents such as alcohols and aromatic compounds are often added as carriers to the formulation which promotes the dissolution of the demulsifier into the crude [44]. A contemporary demulsifier formulation can work at concentrations as low as 10 ppm, which is highly beneficial to reduce the inventory of chemicals, especially in offshore installations [187]. Principles such as the HLB scale can be applied together with experimental data on interfacial tension and rheology to aid in the selection of suitable compounds, but identification of the optimal demulsifier formulation and dosage for a given crude oilwater system always requires some degree of empirical screening. Bottle tests with a crude oil sample are the most common method to identify the most suitable formulation. Another technique is microfluidic experiments that allow rapid screening of emulsion stability against an array of demulsifier types and concentrations [83,188]. This method, in principle, is more systematic and also bears the potential of identifying better formulations, but the bottle test currently remains the standard test method. The maximum separation rate that can be achieved for a given demulsifier type and concentration is a key parameter to assess the efficiency of chemical destabilization, but the kinetics of the process are equally important. When a demulsifier is added to a continuous flow of oil and water, time is required both for homogeneous distribution of the chemicals into the bulk and for adsorption at the interface [189]. A homogeneous distribution can be achieved by injection of the demulsifier into high-turbulence zones e.g., those created by pumps and shear valves, or by utilization of dedicated mixing devices [190]. Concerning adsorption to the interface, time scales of minutes to hours can be required to maximize

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demulsifier action. This time scale depends on the rates of ad- and desorption of the different compounds creating the interfacial film. Enhanced mixing helps to facilitate the transport of molecules between the bulk and near-interfacial region, but to achieve a sufficient adsorption time of demulsifier in the flowing liquid mixture, the injection should take place sufficiently far upstream of the targeted separator vessel [191] or flowline [192].

3.8.4 Case studies Optimization of a separation process to achieve the specifications for effluent gas, oil and water usually involves adjustment of dosage protocols for production chemicals. For liquidliquid separation, demulsifier type and concentrations are the key factors. A pre-selection of these parameters can be achieved in laboratory tests with samples of crude oil and produced water, but ultimately verification of an improved demulsifier formulation needs to be done in the production environment. As the field matures, an increase in the properties of the produced fluids (most notably the water cut) may also require adjustment of the demulsifier formulation. A particularly challenging case was reported by Duke et al. [193] on the demulsification of produced fluids obtained from micellar and polymer flooding in tertiary oil recovery. These systems are characterized by very low interfacial tension and the possible occurrence of microemulsions that do not separate under gravity. Three liquid phases were produced from the M-1 project, two microemulsions, and one conventional “macroemulsion.” The mechanical separation train was incapable of resolving these emulsions, and extreme quantities of an undisclosed commercial demulsifier had to be added (up to 5000 ppm per barrel) to achieve an acceptable degree of separation (B90% emulsion breakage in a single pass through the separator), which resulted in the production process being rendered uneconomical. A series of laboratory tests with novel types of demulsifiers were conducted and polyoxypropylene amines (POPA) were identified as a molecular class to resolve the microemulsions. Implementation in the processing plant enabled complete resolution of the emulsions, with residual WiO B1% at a dosage of 1000 ppm. Utilization of POPAs reduced the treatment cost to B0.55 USD/barrel compared to B1.50 USD/barrel for the previously employed demulsifier. This example highlights the importance of conducting a thorough screening of the demulsifier before deployment, especially when chemically complex emulsions are expected to be treated in the separators. Demulsifier addition should take place sufficiently upstream of the separator to provide enough time for mixing and adsorption of the chemicals to the droplet interfaces. This can be done in the production facility, but also upstream of a flowline with the added benefit that the emulsion viscosity (and thus pressure drop) is reduced during transport by creating large droplets. Demulsifiers can even be injected downhole to achieve pre-separation of oil and water and to decrease the viscosity of the produced fluid. An especially problematic situation

3.9 Electrostatic demulsification

can be encountered when the water cut rises to a value near the inversion point, but the emulsion remains oil-continuous. In that case, a dramatic increase in the viscosity can occur, and based on results from a field study, Gilbert et al. argued that in this water cut regime demulsifier addition would provide the highest economic benefits [192]. Downhole injection of demulsifier for several wells in the Pyrenees development mitigated slugging, reduced emulsion viscosity, pressure drop in the flowline, and enabled a 17% increase in the oil production rate for the whole production facility. Dutta et al. [194] report a study for the Bahrain oil field where continuous downhole injection of 200 ppm demulsifier resulted in a significant increase of the oil production rate by B 50% for one well, along with a significant decrease in emulsion viscosity and casing pressure, thereby greatly enhancing the economics of operation. In the BC-10 subsea project as another example, the injection of 90 ppm of demulsifier into a flowline from a well that produced very viscous and stable emulsions, resulted in a reduction of emulsion viscosity from 900 to 500 cP, thereby enabling to reduce the inlet pressure by B 50 bar, and increasing the production from a single well by 500 barrels/day [195].

3.9 Electrostatic demulsification The three driving forces for demulsification we have discussed so far all help to enhance separation by favorably modulating the conditions required for settling an emulsion under the influence of gravity. An additional mechanism to enhance the rate of emulsion breakup can be introduced through the application of an external electric field. Traditionally, electrostatic treatments of emulsions have been applied in the dehydration and desalting stages of crude oil processing. One reason is that the feed for these two stages is always oil-continuous, which is a requirement for electrocoalescers to avoid short-circuiting between the electrodes if the continuous phase would be saline water. In these applications, the concentration of water droplets in the feed is also relatively low (maximum of the order of 20% v/v, but often lower) and droplets are smaller compared to the upstream bulk separators. These factors result in a reduced coalescence rate and settling velocity, and become even lower if the continuous oil phase is highly viscous. The application of an electrostatic field to the emulsion is an excellent means to treat these more challenging emulsions. In this section, we will outline the principle mechanisms of electrostatic demulsification, and how the interaction of emulsion droplets with an electric field can be translated into designs for full-scale electrocoalescer units to separate WiO emulsions. Electrostatic treatment is a more energy-efficient alternative to heating [196] and could also aid in reducing demulsifier consumption, in principle, to achieve the same separation performance. In many cases, though, the crude oil processing train is highly optimized so that all driving forces including

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electrocoalescence are employed and often pushed to their operating limits to maximize the performance of the system.

3.9.1 Droplet migration in electric fields To better understand the effect of an external field on a water-in-crude oil emulsion we consider the cases of a single water droplet and a pair of water droplets dispersed in an oil phase. A crude oil phase that consists only of hydrocarbons is a weak electric conductor [197], and also only weakly attenuates the electric field due to its low dielectric constant (B23) [198]. We assume that the droplets and the oil are exposed to a homogeneous electric field (i.e. parallel field lines) created by two parallel electrodes as shown in Fig. 3.30. An external electric field

FIGURE 3.30 Overview of the two major mechanisms that promote emulsion breakup in a uniform external electric field. A droplet can acquire a charge at one electrode and then migrate to the electrode of the opposite charge (electrophoresis). Droplets can accumulate at the electrode of opposite charge, in principle, and can also coalesce with droplets in the bulk emulsion while passing through the liquid. Water droplets also get polarized in an external electric field due to the alignment of the permanent dipoles in water molecules with the electric field. This charge redistribution causes an attraction between neighboring droplets which can induce coalescence. The insert in the lower right corner schematically shows the orientation of water molecules in the electric field together with the partial charges on the atoms.

3.9 Electrostatic demulsification

can impact the charge, shape, and movement of a water droplet dispersed in an oil phase through several mechanisms. The first mechanism is electrophoresis, which refers to the movement of a charged droplet or particle towards an electrode with an opposite charge. A water droplet immersed in an oil phase possesses a native charge, but for the droplet diameters and electric field strengths encountered in dehydration vessels, this charge is not sufficient to induce an appreciable electrophoretic velocity [36,37]. When a droplet resides in the vicinity of a charged electrode, it can acquire an electrical charge [199] that is substantially larger than its native charge. This charge will repel the droplet from the electrode with a force Fcℎ 5 qE, where q is the charge of droplet, and E the electric field strength [199]. The droplet will then migrate towards the opposite electrode. The electrophoretic velocity υep of the droplet can be obtained from balancing the electrophoretic force with the drag force on the particle Fd 5 3πdηcυ under the assumption of Stokes flow: υep 5 qE/3πdηc. We calculate the electrophoretic velocity of a droplet with the following parameters as an example: E 5 1 kV/cm, d 5 50 μm, ηc 5 20 mPa s. To estimate the charge acquired by the droplet, we use a scaling relation that was obtained experimentally [200]: q 5 q0(r/r0)1.59(E/ E0)1.33, with q0 5 1.03 3 10211 C, r0 5 363 μm and E0 5 2 kV/cm. With these values we calculate an electrophoretic velocity of  0.62 mm/s. The Stokes settling velocity for the same parameters when assuming Δρ 5 100 kg/m3 is equal to  6.8 μm/s, which is roughly two orders of magnitude lower. This coarse comparison indicates the benefits of applying an electric field to induce electrophoresis of droplets. As the droplet moves towards the opposite electrode, it collides and coalesces with other droplets. If the droplet would retain all of its charges during its trajectory into the oil phase, it could reach the opposite electrode and segregate from the emulsion there, in principle (see Fig. 3.31A for a series of photos showing migration of a droplet between two electrodes). The electrical conductivity of typical crude oils often is of such a magnitude [203], however, that the charge relaxation time, which is a parameter that describes the timescale of charge dissipation from the droplet into the surrounding fluid, is less than one second [44]. This is much lower than the time needed for a droplet to traverse between electrodes in commercial electrocoalescer units.

3.9.2 Droplet collisions in electric fields The second mechanism that induces separation is caused by the polarization of water molecules in the electric field. A water molecule has a permanent dipole moment, i.e. a section with a net negative charge and another section with a net positive charge (see insert in Fig. 3.30). When subjected to an electric field, these dipoles align with the direction of the electric field. In a droplet, all water molecules are subjected to this force, which results in a buildup of net charges at opposing sides of the droplet as shown schematically in Fig. 3.30.

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FIGURE 3.31 Visualization of the impact of external electric fields on water droplets dispersed in oil. Panel (A) shows the migration of a droplet that acquired a charge at the electrode on the left, to the electrode of opposite charge in a DC electric field [200]. The droplet diameter is B0.9 mm and the electric field strength B2.5 kV/cm. Panel (B) shows the periodic deformation of a droplet exposed to a pulsed DC electric field with a very low frequency of 1 Hz [201]. The droplet diameter is B1 mm and the electric field strength B2.5 kV/cm. Panel (C) shows the effect of the electric field strength on the coalescence time between two droplets [202]. A water droplet with a diameter of 175 μm dispersed in oil was settling under the influence of gravity onto a larger droplet of 2 mm diameter, and the time it took from close contact to coalescence was recorded experimentally. The three different data sets were recorded for different viscosities of the continuous oil phase. Reproduced with permissions from Refs. [200202].

This polarization effect creates a force between two droplets, that is attractive if the line connecting the droplet centers and the electric field direction is parallel or a sufficiently small angle is formed between them [204]. If the line connecting the centers of the mass of the droplets is parallel to the electric field line, the attractive force between two droplets due to polarization can be calculated from Fpol 5 24πAr,cA0E2r6/ℎ4 [205]. In the equation, Ar,c and A0 are the relative dielectric permittivity of the continuous phase and dielectric permittivity of vacuum, respectively. r and ℎ are the droplet radii and distance between droplets, respectively, assuming two droplets of equal radius. The equation is valid only for ℎ . . r. The dipole force has a strong inverse dependence on the distance between the droplets. In a dilute emulsion (assuming 5% water cut, xw 5 0.05) with drops of a diameter of 50 μm um the average distance (ℎ) between droplets is roughly equal to four times their diameter, when using the relation ℎ 5 r(4π/3xw)1/3 [206]. Hence, in these systems, the dipole force will be relatively weak and the migration of drops towards each other will be very slow initially if we consider a stagnant emulsion. The approach velocity of droplets is significantly enhanced if an external flow is present, though, especially if it is turbulent [48]. The velocity fluctuations will induce collisions between droplets and bring them

3.9 Electrostatic demulsification

close enough for the effect of the dipole-induced electrostatic attraction to become the dominant force component. It can thus be argued that for efficient electrostatic demulsification to occur, it should happen as a two-step process, the first step is flow-induced contact between the droplets, and the second step is electric-field-enhanced film drainage and rupture. The same mechanism can also be assumed to be beneficial for enhancing the impact of electrophoretic forces on the separation rate. The dipole force can bring the droplets together, but for coalescence to occur, the thin film of continuous phase that is formed between the droplets upon close contact must rupture. We compare the effect of the electric field as a driving force for coalescence with the case of gravity-induced film drainage. For a quantitative estimate, we use Eq. (3.8) to calculate the film drainage rate for two hard spheres. We assume for this calculation that the colloidal force term Fc is zero. The attractive force between two droplets with centers of mass aligned parallel with the electric field direction can be approximated for two hard spheres in close contact [44]: Fpol 5 πAr,cA0E2r3/2ℎ. We calculate the time it takes the droplets to approach from an initial distance (ℎ0) down to a critical thickness (ℎc) where film rupture takes place. Insertion of the term of Fpol into Eq. (3.8) yields after rearrangement for the coalescence time: tc 5 3ηc(ℎ0 2 ℎc)/ Ac,rA0E2r. This simple relation illustrates the strong effect of the external electric field on the coalescence time. For the initial film thickness, ℎ0 5 d/8 was used, which has been argued as the distance from which Eq. (3.8) becomes valid [48]. For the critical film thickness, 0.1 μm was assumed [49]. With the same parameters as for the calculation of the electrophoretic velocity in the previous paragraph, we obtain tc  68 ms. For comparison, we calculate the film drainage time for two droplets colliding under the influence of gravity. In that case, r becomes the equivalent hydrodynamic radius of the droplet pair, as two drops of the equal radius would not collide under the influence of gravity only. The gravity force can be expressed as Fg 5 4πΔρr3/3. Insertion of the force term into Eq. (3.8) yields: tc 5 tTln(ℎ0/ℎ), with tT 5 9ηc/8rΔρg. For the same conditions as in the case of the calculation for the electrostatic dipole attraction, a coalescence time of  5.9 s is obtained, which is two orders of magnitude higher than the value obtained for the case of electrostatic attraction. While we have considered two simplified cases, they still illustrate the significant impact of electrostatic attraction between the two drops on the rate of film drainage. The trend of the coalescence time with electric field strength has been confirmed by numerous experiments. Fig. 3.31C shows some experimental data as one example [202]. A water droplet with a diameter of 175 μm dispersed in oil was settling under the influence of gravity onto a larger droplet at rest of 2 mm diameter, and the time it took from close contact to coalescence was recorded experimentally. A third mechanism affecting drops in an electric field is called dielectrophoresis. In a non-uniform electric field, a single dielectric particle can be polarized and is attracted to the location with the highest electric field gradient [207]. A

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non-uniform electric field will occur near the edges of plate electrodes (or if the electrodes are rods), and also within the emulsion, as the polarized droplets themselves are the source of a non-uniform electric field. From that point of view, dielectrophoresis can be interpreted as an effect that typically will aid in promoting collisions and coalescence between droplets. The considerations and calculations in this section indicate that the separation efficiency should only increase with increasing electric field strength. Another effect of polarization, however, is the deformation of drops, which manifests itself in stretching along the direction of the electric field [208]. If the electric field strength exceeds a critical value Ec, droplet breakup may occur [209], or partial coalescence may take place [210] with smaller daughter droplets emerging from the event. The concept of a critical electric Weber number can be introduced just as for the case of turbulent flow, which results in a relation between Ec and the droplet radius [211]: Ec B (σ/2Ac,rA0r)0.5. For a typical value of the electric field strength applied in electrocoalescers (E 5 1 kV/cm) and a liquidliquid interfacial tension of σ 5 25 mN/m, droplet diameters for the breakup are of the order of millimeters or larger, which is a droplet size that is not problematic to separate in dehydration or desalting vessels.

3.9.3 Effect of electric field properties on droplet coalescence In a DC electric field [207,212] or AC electric field of low frequency [213], the polarized droplets can form chains that align with the direction of the electric field as shown in Fig. 3.32A as an example [213]. Transient chain formation is unproblematic if it is followed by droplet coalescence. If the attractive dipole force cannot overcome the barrier to coalescence, however, the chains can keep growing in length, and in the worst-case form a conductive channel between the electrodes, thereby short-circuiting them. This behavior is considered undesirable, as it significantly increases power consumption and/or decreases the electric field strength. For that reason and also to avoid electrolytic corrosion [201,215], electrodes often are covered with an insulating material [216,217]. Besides the impact of the electric field strength on the rate of emulsion breakup, an equally important choice is whether stationary (direct current, DC) or nonstationary (alternating current, AC) electric fields are employed. the features of both types of fields on the dynamics of emulsion separation are summarized below: •

DC field: The advantage of a DC field as compared to an AC field is that the electrophoretic force as a mechanism will be strong. The electrophoretic force causes a larger translational displacement of droplets as compared to an AC field, and droplet collisions are facilitated also at relatively low WiO concentrations. The downsides of a DC electric field are potential shortcircuiting of the electrodes due to droplet chain formation [214], and power

3.9 Electrostatic demulsification

500 Hz

(A)

(B)

50 Hz

5 Hz

0.5 Hz

FIGURE 3.32 (A) Alignment of a droplet ensemble in an emulsion of 20% v/v of water in crude oil under the influence of an AC electric field of 50 Hz [213]. The droplet diameter is B0.9 mm and the electric field strength B5 kV/cm. The scale of the image was not reported by the authors. (B) Volume of water dropped out after 5 min from 30 mL of a model water-in-oil emulsion in a pulsed DC electric field as a function of applied voltage [214]. The four data sets were acquired at different frequencies. The distance between the electrodes was 1.5 cm. Reproduced with permission from T. Chen, R. Mohammed, A. Bailey, et al., Dewatering of crude oil emulsions 4. Emulsion resolution by the application of an electric field, Colloids and Surfaces A: Physicochemical and Engineering Aspects 83 (1994) 273284; Y. Zhong, L. Siya, Y. Yaochuan, et al., An investigation into the breaking-down of water in-oil type emulsions by means of pulsed voltage, Desalination 62 (1987) 323328.

•

losses if the crude oil has a high electrical conductivity [203]. Even if insulated electrodes are used, a large voltage drop can occur across the insulation due to the accumulation of charges near the electrodes [177]. Pulsed DC fields have been proposed as an improvement, and it has been argued that alternating intervals of high and low electric field strengths without changing the direction of the electric field are highly beneficial to induce droplet coalescence [218]. AC field: A sinusoidal waveform field with a frequency of 50 or 60 Hz is the global standard pattern for an AC electric field. Conversion of the line voltage to the high voltages used in electrocoalescers can be accomplished with a transformer. Due to the rapidly switching polarity of the electrodes, droplets will not move comparatively large distances but perform more of a “wiggling” motion as their polarization changes in tune with the applied frequency [208]. In addition, the droplet will undergo periodic deformation [201,214]. Fig. 3.31B is a visualization of this effect. These effects can induce coalescence between droplets if the distance between them is small enough, but this requires the emulsion to be fairly concentrated or an additional mechanism that brings the droplets in close contact needs to be present, such as turbulent flow. If the WiO concentration is low and the frequency of the AC field is sufficiently high, chain formation of droplets is reduced, and

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short-circuiting is unlikely. If a high-frequency transformer is utilized, power losses due to the high conductivity of the crude oil can also be minimized, if the AC frequency f fulfills the criterion f . κ/2Ac,rA0, where κ is the electrical conductivity of the crude oil [177]. The impact of a nonstationary electric field on the separation rate of an emulsion is shown in Fig. 3.32B [214]. A pulsed DC field was applied to a model WiO emulsion for 5 minutes, and the volume of formed bulk water during that time was recorded. In this experiment, the separation rate increases both with increasing voltage and frequency, and other experiments have also indicated a beneficial effect of increasing the frequency and/or voltage of a nonstationary electric field [219221]. Several studies report an optimum value for both the frequency and the electric field strength at values even below the critical field strength for droplet breakup [201,222224]. These results have been interpreted as a resonance of the applied frequency with dielectric or mechanical properties of the emulsion [218,225,226]. Currently, no theory is available to satisfactorily predict the separation rate of an emulsion in an electric field, even though basic frameworks have been proposed [177,206]. The main reason for this is the enormous physical complexity that is encountered when considering the interaction of a bulk emulsion with electric fields and turbulent flow at the same time.

3.9.4 Electrocoalescer technology Sinusoidal AC electric fields are the most commonly used type of external electric fields for oil dehydration and desalting. Fig. 3.33A shows an illustration of a typical electrocoalescence unit integrated into a dehydration vessel [1]. The emulsion enters the vessel from the bottom and is distributed via a spreader. One or several grids of plate or rod electrodes are placed in the upper part of the vessel. The electric field permeates the entire vessel but will be strongest near the electrodes. The line voltage is converted to a high voltage with a standard transformer at 50 or 60 Hz frequency. Electrocoalescers have seen multiple improvements from the traditional designs to intensify the interaction of the electric field with the emulsion. A design with insulated electrodes and a strong local electric field that treats only the emulsion layer can be used in three-phase separators to facilitate bulk separation of oil and water [227]. Other designs have implemented more complex electric field patterns such as a combination of AC and DC fields [5], or highfrequency AC fields [217], all of which have been reported to enhance separation efficiency. Another feature of modern electrocoalescers is the implementation of advanced electronics that enable optimization of the electric field properties based on the electric and dielectric properties of the emulsion [90,227,228]. Another class of electrocoalescers is compact stand-alone units. Multiple designs have been described in the literature [216,222,228,229] and one example

3.9 Electrostatic demulsification

FIGURE 3.33 Illustration of an electrocoalescence unit integrated into a dehydration vessel (panel A, [1]). Panel (B) shows a scheme of a prototype for a compact electrostatic coalescer for intense contacting of the emulsion with the electric field [223]. Separated water is drained at the bottom of the cell with a valve in this experimental unit. Reproduced with permission from K. Arnold, M. Stewart, Surface Production Operations, Volume 1: Design of Oil Handling Systems and Facilities, Gulf Professional Publishing, Houston, 2008; J. Eow, M. Ghadiri, A. Sharif, Electrostatic and hydrodynamic separation of aqueous drops in a flowing viscous oil, Chemical Engineering and Processing 41 (2002) 649657.

is shown in Fig. 3.33B [223]. The electrode assembly is located in a flow cell with a geometry that ensures that the entire fluid is continuously exposed to a strong electric field while passing through the unit. In addition, a higher level of turbulence as compared to a large vessel can be established, which is expected to promote electrocoalescence by increasing the collision rate between droplets. The turbulence level needs to be tuned to not cause an increased breakup of droplets, however, which puts an upper limit on the maximum flow capacity. The module also has a drain valve and buffer volume at the bottom to remove water that has dropped out from the emulsion. The design shown in Fig. 3.33B is well suited for experiments under continuous flow in a laboratory environment due to its small size, but the idea of a flowthrough electrocoalescence cell has also been adapted in commercial designs [216,229]. These units are not intended to replace vessel-based electrocoalescers, but rather provide a pre-treatment of the emulsion compactly, with the main purpose to induce growth of the droplets in the feed so that separation in the downstream vessel is facilitated. Features such as insulated electrodes and tunable electric fields can equally be applied for compact electrocoalescers just as for vessel-based units.

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A design goal for any electrocoalescer is to maximize separation efficiency while keeping power consumption and hardware cost within reasonable boundaries. The additional investment that is required for advanced electrocoalescence technologies often enables the processing plant to operate more economically as a whole, but the suitability of a given technology to optimize the production process always needs to be evaluated based on each case. Electrocoalescence has become a standard and crucial method to enhance the separation of emulsion in all stages of the processing plant. The general principles of water droplets interacting with electric fields have been established and verified experimentally, and have helped greatly in informing the design of contemporary electrocoalescer units. At the same time, the understanding of the interaction of dense emulsions with an electric field and a turbulent flow, as well as the advancement of electrocoalescer technology, remain challenging and interesting subjects for engineers and scientists.

3.9.5 Case studies The case studies we report in this section are all about upgrades of the production system where an electrocoalescer unit was added to debottleneck or optimize the production. These studies possibly illustrate best the effect of electrostatically induced coalescence as they allow comparison of the system performance before and after the unit was taken into operation. Heatherly et al. reported on an upgrade of the Ewing Bank 873-A platform with a three-phase separator containing electrostatic grids [230]. The reason for the addition was an anticipated capacity increase from two tiebacks, and a less than satisfactory performance of a coalescer-separator (upflow coalescer) vessel already installed on the platform. Both the electrostatic treater and the coalescer-separator were approximately of the same dimensions and installed for side-by-side operation, which provided an excellent situation for direct performance comparison between the units. The upflow coalescer performance provided 1% residual WiO concentration at 15,000 bpd throughout with a feed water cut of 5%. In comparison, the electrocoalescer vessel was operated with a feed flow of 40,000 bpd at 35% feed water cut and achieved a residual WiO concentration of ,0.4%. While the operating and installation costs of the electrostatic unit were not disclosed, the authors considered the project an economic success and mentioned that utilization of the electrocoalescer provided savings of 100,000 USD/year in demulsifier consumption. Another interesting case was provided by Wang et al. where the performance of two electrostatic dehydrators was studied in simultaneous parallel operation in a production environment when processing fluids from polymer flooding in the Daqing field [231]. One unit was operating with a superposition of an AC (50 Hz) and DC field inside the vessel, whereas the other unit was operating with a pulsed DC electric field (15 kHz). Pulsed DC treatment reduced demulsifier consumption from 30 to 1520 ppm as compared to AC/DC treatment, as well as reducing residual OiW (pulsed DC: B600 ppm, AC/DC 1200 ppm) and WiO

Nomenclature

concentrations (pulsed DC: B 0.2%, AC/DC: B 1%). No data on power consumption and other costs were given by the authors, but according to their analysis, the pulsed DC mode of operation was more economical for the fluids produced from this particular well.

3.10 Concluding remarks This chapter aimed to present an overview of the key scientific and technological aspects regarding the separation of emulsions that are encountered during crude oil processing. We intended to provide an outline for how an understanding of the fundamental aspects governing emulsion formation, stability, and separation can inform the selection, design, and efficient application of suitable separation technologies, which in turn can be integrated into a full process scheme. For commercial applications, the engineering of the processing system is also of paramount importance. The science and engineering of crude oil emulsion processing is a challenging but fascinating area that requires an interdisciplinary effort between mechanical and chemical engineering, fluid dynamics, as well as physical and organic chemistry. In many other industrial applications, emulsions are deliberately created in a controlled way. In petroleum engineering, operators, technology suppliers, and service companies have to work with whatever nature provides in form of the reservoir constituents. There is a broad spectrum of properties for the produced fluids, and this uncontrollable aspect introduces additional challenges and research questions that are profoundly different as compared to when dealing with food or health care emulsions. But even though there are crucial differences with other industrial segments, there is also a strong overlap with them, as the principles of emulsion formation and stability, their multi-scale nature, and the complex composition of the fluids are aspects that are relevant for many industrial applications. With the profound advancements in the understanding of the science of crude oil emulsions and optimization of separation processes in the last decades, it may be fair to say that the impact of emulsions on the production process can be handled quite well these days, and situations that prevent separation of emulsions are rarely encountered, at least from a technology perspective. We hope this chapter can provide insights into the most relevant aspects of this important segment of crude oil processing. By necessity, we had to keep many explications brief due to the broadness of the subject, but the numerous references mentioned in this chapter may provide a starting point for further reading on the individual topics.

Nomenclature AC AOT

alternating current sodium bis(2-ethylhexyl) sulfosuccinate

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CFD DC DPL EDL HLB m NFA nm NPFR OiW ppm rpm RSN s SARA v/v WiO

computational fluid dynamics direct current dense-packed layer electric double-layer hydrophiliclipophilic balance meters net free area nanometers nonylphenol formaldehyde resin oil-in-water parts per million revolutions per minute relative solubility number seconds saturate, aromatic, resin, and asphaltene volume to volume ratio water-in-oil

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CHAPTER

Foam control

4

Kalman Koczo1, Mark D. Leatherman1 and Jonathan J. Wylde2,3 1

Momentive Performance Materials, Tarrytown, NY, United States 2 Heriot-Watt University, Edinburgh, Scotland, United Kingdom 3 Clariant Oil Services, Clariant Corporation, Houston, TX, United States

Chapter Outline 4.1 Introduction and overview .............................................................................154 4.1.1 Foam basics ...............................................................................155 4.1.2 Oil-based versus water-based foams ..............................................159 4.1.3 Antifoaming versus defoaming ......................................................159 4.1.4 Antifoaming versus deaeration ......................................................160 4.1.5 Solid-stabilized foams .................................................................161 4.1.6 Overview of foam stabilizer and antifoam chemistries .....................163 4.2 Oil-based foams ...........................................................................................165 4.2.1 Defoaming versus demulsification ................................................165 4.2.2 Nonaqueous foaming ...................................................................165 4.2.3 Nonaqueous foams of crude oil ....................................................168 4.2.4 Chemistry of antifoams for oil-based foams ...................................175 4.3 Water-based foams .......................................................................................185 4.3.1 Chemistry of antifoams for water-based foams ...............................185 4.3.2 Water-based applications .............................................................189 4.4 Mechanical defoaming ..................................................................................198 4.5 Defoaming by chemical reaction ....................................................................199 4.6 Mechanisms of antifoaming action ................................................................199 4.6.1 Antifoaming of nonaqueous foams ................................................200 4.6.2 Antifoaming of aqueous foams .....................................................204 4.6.3 Breaking solid stabilized foams ....................................................211 4.7 Concluding remarks ......................................................................................212 Nomenclature ......................................................................................................213 References ..........................................................................................................213

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00002-8 © 2023 Elsevier Inc. All rights reserved.

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4.1 Introduction and overview Foams are thermodynamically unstable dispersions of gas bubbles entrained in a liquid. As we will see throughout this chapter it is the kinetic stability that matters most for oilfield foams. Sometimes even a foam that is stable only for a matter of seconds can cause operational challenges, for example in a gas/oil separator vessel. It is the control and breaking of foams in oilfield applications that form the basis of this chapter. Foams in the oilfield are generated in a multitude of ways. Gas can be incorporated into a continuous liquid phase by mechanical action, such as mixing, shaking, stirring, pumping, jetting, or bubbling. Another way is by the formation of bubbles from degassing of liquids via the evolution of dissolved gas by reducing the pressure; this is commonplace in separator vessels. There are even intentionally made foams, such as those that are useful downhole, which can become a problem on the surface and need to be eliminated (e.g., foamers for gas well deliquification). In oilfield systems there are often at least three (and usually four) phases present at the same time: a nonaqueous or oil phase (i.e., crude oil), an aqueous phase (i.e., produced water or brine), and a gaseous phase (natural produced gas), and these phases are at least partially insoluble in one another. A fourth, and solid phase can be also present in the form of particles, for example, sand, formation fines, clays, etc. These solid phases can accordingly interact to stabilize or destabilize foams. Fig. 4.1 illustrates a typical three-phase foam structure. It shows a highly stable foam containing air, an aqueous phase (brine), and a nonaqueous phase (in this case hexane), after a long drainage time. The bubbles form a cell structure and are separated by foam films and three foam films meet in Plateau borders, which in this case are thickened due to the presence of emulsified oil droplets. In this case, the oil is not detrimental to foam stability so does not impart any antifoaming action. Initially, when there are only small and individual bubbles (termed bubble dispersion) inside the liquid without contact with each other, this cannot be considered a true foam structure. When the bubbles rise due to buoyancy and approach one another, liquid films form between the bubbles, and a foam structure forms. The stability of this foam depends on the stability of these foam films (or lamellae). The most important physical feature of foam is that its specific volume is much larger (that is, its density is much lower) than that of liquids. This can be an advantage, for example, when a large volume of material must be pumped downhole, or the material downhole has to be lighter, such as in underbalanced drilling or gas well deliquification. Foams can, however, also be a challenge, particularly on the surface, in equipment that cannot handle large volumes, and often these foams are required to be broken and controlled before further processing.

4.1 Introduction and overview

FIGURE 4.1 Image of a highly stable foam prepared by bubbling nitrogen into an aqueous-continuous emulsion of 3% NaCl and a silicone polyether surfactant (0.1 g/100 mL), with 30 vol.% hexane.

Foam is most often eliminated by adding a small amount of liquid chemical called antifoam (also known as defoamer or foam control agent). During the use of an antifoam, it is the introduction of this insoluble liquid (for example a silicone oil) that then interacts with the foam to destabilize it. Mechanical methods also exist for foam control and will also be discussed in this chapter, with the major focus being chemical foam control, or antifoaming. There are a few special cases where chemical reaction causes foam rupture vs the more common physical antifoaming action for oil-based foams, water-based foams, and solid-stabilized foams.

4.1.1 Foam basics When gas is dispersed into a liquid, it generally exists in the form of spherical bubbles, which, unless the liquid is intensively mixed, will then rise due to

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buoyancy. As a rising bubble approaches the liquid surface or when two bubbles get compressed, they form first a thick lamella (Fig. 4.2). Initially, a spherical (wet) foam forms containing relatively uniform shaped bubbles separated by relatively thick liquid films (Fig. 4.3). Gravity then drains the liquid and gradually a polyhedral structure is formed, which is composed of air cells separated by thin liquid films [1,2]. At the meeting point of three lamellae, so-called Plateau borders (GibbsPlateau channels) form (Fig. 4.4). These channels meet in wedges and form a drainage system inside the foam, which is where most of the liquid is located in a polyhedral foam. Plateau borders can play an important role in the antifoaming mechanism (see Section 4.6.2). Foam drainage is driven by gravity and capillary pressure. Due to surface tension, the pressure inside a curved liquid surface is higher than outside of it, and the difference is the capillary pressure. The YoungLaplace equation defines the _

FIGURE 4.2 Formation of foam lamella (A) as a bubble approaches the liquid surface and (B) when two bubbles are compressed.

FIGURE 4.3 Picture of a slightly drained, wet foam (left) and a well-drained polyhedral foam (right), made from an aqueous 1% Triton X-100 surfactant solution.

4.1 Introduction and overview

FIGURE 4.4 Cross-section of a Plateau border; RPB is the radius of curvature in the plane of the drawing.

capillary pressure (pc) as:

  pc 5 σ 1=R1 1 1=R2

(4.1)

where R1 and R2 are the principal radii of curvature and σ is the surface tension. For a bubble (sphere) the radii of curvature are equal, and the equation simplifies to: pc 5 2σ=R

(4.2)

where R is the bubble radius. In the Plateau borders the capillary pressure, PcPB is: Pc;PB 5 σ=RPB

(4.3)

since the Plateau border is almost perfectly perpendicular to the plane of Fig. 4.4. Since the capillary pressure is close to zero in the adjoining liquid films (which are more or less flat), the pressure inside the Plateau borders is lower, thereby draining the liquid from the films into the Plateau borders. In real-world foams, bubbles typically have a size distribution, and the pressure inside smaller bubbles is higher than in larger ones (Eq. 4.2) causing gas to diffuse from smaller bubbles to larger ones through the films. Small bubbles eventually disappear and cause larger bubbles to grow. This phenomenon is called Ostwald ripening, and the rate of this inter-bubble gas diffusion is generally slow. If the gas has a high solubility in the liquid (e.g., CO2 in water or methane in crude oil) the rate of inter-bubble gas diffusion is much faster. The rate of drainage of fluid from the foam is also influenced by the bulk viscosity of the liquid. A fluid with higher viscosity will drain slower and therefore a highly viscous fluid can form a foam with considerable kinetic stability, even if other surfactant-driven stabilizing mechanisms are not present. Another effect of fluid viscosity is that it is more difficult to incorporate small bubbles into a more viscous fluid.

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FIGURE 4.5 Gibbs-Marangoni effect in a draining liquid film containing surfactants. The viscous drag created by the outflow of liquid from the films towards the meniscus creates a surface tension gradient that opposes film drainage [3]. Reproduced with permission from D.T. Wasan, AD Nikolov, L.A. Lobo, et al., Foams, thin films and surface rheological properties, Progress in Surface Science 33 (1992) 119154.

The Gibbs-Marangoni effect is another important film stabilizing phenomenon when surfactants are present. As the foam film between two bubbles drains, by either gravity or capillary pressure, the viscous drag by the outflow of liquid shears the film lamella surfaces and thus sweeps adsorbed surfactants towards the adjoining Plateau borders and creates a surface tension gradient. This gradient opposes film drainage, as shown in Fig. 4.5. If the surfactant concentration is either low or very high, then the film surfaces will be mobile, and drainage will be fast. If, however, the surface tension gradients are large, the film surface will be immobile, and drainage will be slow [3,4]. Analogous to this phenomenon, the surface dilatational viscoelasticity, that is change in surface tension when the surface is expanded or compressed, can be expressed as the Gibbs elasticity (Eo) where σ is the dynamic surface tension, and A is the surface area [5]: E0 5

2dσ dlnA

(4.4)

The viscous response to dilation or compression of a surface is driven by the exchange (adsorption or desorption) of surfactant between the surface and the bulk fluid, or by rearrangement of surfactant molecules at the surface [6]. High surface viscosity and elasticity, imparted by surfactants, act to resist deformation of the lamellae and are therefore associated with more stable foams [7]. It is important to note that the relative magnitude of surface tension change due to adsorbed surfactant is much less in oil-based systems compared to waterbased ones. This is due to the lower surface tension of the “clean” (surfactantfree) oil surface (generally 2535 mN/m), compared to the clean water surface (B72 mN/m). The consequence of this is that the phenomena described here

4.1 Introduction and overview

(Gibbs-Marangoni effect and surface viscoelasticity) are generally much weaker in oil-based systems. As the liquid film drains, opposite surfaces of the thinning films approach each other, and when film thickness approaches 100 nm, the two surfaces (which may have layers of adsorbed surfactant) start to interact with one another. The interactions typically are via Van der Waals attraction, ion-ion repulsion (with ionic surfactants), and steric repulsion. The net interaction is defined as the disjoining pressure, Π [3,8,9]. In a stable foam, equilibrium is achieved when the capillary pressure is equal to the disjoining pressure, thus drainage stops from the films to the Plateau borders and the capillary pressure in the Plateau borders is equal to the (negative) hydrostatic head measured from the bottom of the foam [10]. Foams therefore generally start rupturing from the top where the capillary pressure is the highest and the films are the thinnest. As the thin liquid film drains further, it can reach a critical thickness (often around 30 nm in aqueous liquids) where it either ruptures or becomes a stable, common, or Newton black film (about 10 nm thickness) [3,9,11]. The objective of antifoam addition is to accelerate this process so that the film ruptures much earlier before having to reach this critical thickness (see also Section 4.6.1.2).

4.1.2 Oil-based versus water-based foams All aspects of foam stability highly depend on the nature of the continuous liquid phase. Water-based and oil-based liquids behave very differently, partly because foam (and emulsion) phenomena are highly influenced by the surface tension of the continuous phase, and to some extent by the interfacial tension between two insoluble liquids. Crucially, the surface tension of water (about 72 mN/m at 25
C) is much higher than the surface tension of nonpolar liquids. The surface tension of crude and mineral oils is 3035 mN/m. Therefore, surface-active materials can cause a much larger change in aqueous solutions (dropping the surface tension by 3050 mN/m), while only the most surface-active materials can reduce the surface tension of nonpolar oils, and even then, only by a few mN/m. These surface tension phenomena result in major differences between antifoam chemistries used for aqueous and nonaqueous liquids. Such is the difference that this chapter deals with. Each specifically and in turn for these two fundamentally different kinds of systems (Section 4.2 vs. Section 4.3).

4.1.3 Antifoaming versus defoaming Foam control agents can be classified as antifoams and defoamers. The terms are specific and significantly different: antifoams are the foam control agents which are added to prevent foam generation before its formation; defoamers are added to break the foam after it has been generated. Antifoams often also work well as defoamers, and although there are chemicals that cause defoaming that also

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function as antifoams, there are many chemicals that are good defoamers but exhibit poor antifoaming performance (see also Section 4.6.2). In practice, and as the authors often see put into practice in the industry, these two terms are often used incorrectly. For example, foam control agents (often common silicone-based chemistries) are typically injected in cokers onto the foam to break it (which is a defoaming situation) but they are still often called coker antifoams. In this chapter, we will mostly use the term antifoam since they are more common. Other terms, such as foam suppressors, foam breakers, and foam inhibitors can be also found in the literature with similar meanings but will not be dwelt on here.

4.1.4 Antifoaming versus deaeration Antifoams function by breaking foam structure-supporting bubbles separated by foam lamellae and Plateau borders. Typically, foam formation is preceded by the formation of separate bubbles that have no contact with one other, followed by the ascent of bubbles to the top of the liquid and finally attachment to one another, thus creating a foam structure. This is the case, for example, in crude oil foaming where bubbles form after depressurization and subsequent dissolved gas evolution in a separator, or when the bubbles are introduced by bubbling through a gas filter frit, for example in the ASTM D892 test [12]. If bubbles are small, ascension time can be critical, and bubbles (and foam) can cause operational challenges, for example, pump cavitation. It is important to realize that foam control agents generally are not suitable to help with the ascent of individual bubbles since they work by breaking foam films and not by affecting bubble movement (see also Sections 4.6.1 and 4.6.2). The rate of bubble rise (v) can be expressed by Stokes’ Law, shown in Eq. (4.5). This expression assumes that the rising bubble behaves as a rigid sphere, is not smaller than a few microns, is in a large container, and that the Reynolds number is , 1. v5

D2 ðρ1 2 ρ2 Þg 18η

(4.5)

where D is the bubble diameter, ρ1 is the density of the liquid, ρ2 is the density of the bubble, g is the gravitational acceleration and η is the dynamic viscosity of the liquid. Stokes’ Law shows that a bubble rises more slowly if the viscosity of the liquid is high (for example, in a viscous crude oil) and much more slowly if the bubbles are small. For example, in crude oil with 0.9 g/mL density and 100 cP (0.1 Pas) viscosity, a bubble with 1000 μm (1 mm) diameter would rise 10 cm distance in 3 minutes; this would take 5 hours for a 100 μm bubble, or 21 days for a 10 μm bubble! In reality, a gas bubble does not have a rigid and uniform surface, but rather a mobile one, and as the bubble rises, a circulation forms inside the bubble, and

4.1 Introduction and overview

flow is generated at the gas-liquid interface. Due to this circulation, a bubble (or drop) with a completely mobile surface can rise 50% faster than predicted by Stokes’ Law, and if surface active materials are in the liquid, then the surfactants can immobilize the surface due to the Marangoni effect [13,14]. The surfaceactive materials can be surfactants in water and practically only polydimethylsiloxane (PDMS) in oil. Ross et al. measured the rate of bubbles ascending in nonpolar liquids and found that as low as 10 ppm PDMS (1000 cSt) will cause bubble surface immobilization and therefore about one-third, slower bubbles rise [13,14]. It is important to understand that PDMS contains a mixture of polymers with various molecular weights (MWs) and that it is only the smaller molecules that are soluble in the oil and thus act as surfactants (see Section 4.2.4.1.1). It is for this reason that higher viscosity PDMS grades (e.g., 12,000 cSt and higher) are used as antifoams, as these contain only very small amounts of crude oilsoluble polymer species [15]. In lighter condensates, PDMS is generally soluble and thus it behaves more as a foaming agent, and therefore cannot be used as antifoam in the separators (see Section 4.2.4.1 for further details).

4.1.5 Solid-stabilized foams Foams can be stabilized not only by surfactants but also by solid particles, commonly found in oilfield systems. Moreover, solid-stabilized foams can be more stable than surfactant-stabilized ones. Solid particles can act as foam stabilizers via various mechanisms; if the liquid can sufficiently wet the particles, and there are a high-volume fraction of solids, then the particles can collect in the Plateau borders (similar to the emulsion drops shown in Fig. 4.1), thus slowing liquid drainage and stabilizing the foam. Another, more effective way of foam stabilization can occur if the particles are wetted only partially by the liquid. The main parameter to characterize the degree of wetting is the contact angle, θ, as shown in Fig. 4.6. If θ 5 0 degree, then the liquid can be described as completely wetting the particle, and if θ 5 180 degrees, then the liquid is completely nonwetting. In equilibrium, the relationship between contact angle and interfacial tension of the two fluids (one of them is gas for a foam system, and both are (insoluble) liquids for an emulsion system) can be expressed by the Young equation: σSG 2 σSL 5 σGL cosθ

(4.6)

The particles will act best as foam stabilizers if the contact angle is medium, ideally 90 degrees. In this case, the particles adsorb at the liquid surface on the two sides of foam films, separating the bubbles, as shown in Fig. 4.7 for irregularly shaped particles. Binks and others studied in detail the particle-induced foam stabilization mechanism [16,17]. The contact angle can be artificially varied by treating the (silica) particles with various silanes or by other means. In oilfield systems, hydrophobization generally occurs naturally; for example, crude oil can adsorb on

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FIGURE 4.6 Contact angle of a liquid drop on a solid surface. σGL: surface tension of the liquid; σSG: surface tension of the solid; σSL: solid/ liquid interfacial tension.

FIGURE 4.7 Adsorption of particles at the surfaces of bubbles.

hydrophilic inorganic particles (such as silica) thus making them hydrophobic. Additionally, partially hydrophobic particles based on clay and iron oxide can also act as foam stabilizers by this mechanism (as well as being known to stabilize emulsions). Foaming of polystyrene lattices increased steeply with increasing contact angle (by adding salt), above about 85 degrees. Particle-stabilized emulsions (sometimes called Pickering emulsions) occur in many places other than in the oilfield, such as in the food and beverage industry [1618]. It is interesting to compare foam stabilization by surfactants versus solid particles: • • • •

Surfactants stabilize mostly aqueous foams, due to the low surface tension of oils. Solid particles, however, can stabilize both water and oil-based foams if the corresponding contact angle is intermediate. Surfactant molecules can form micelles, while particles cannot (although sometimes they form aggregates). The adsorption of surfactants is dynamic, that is the molecules are constantly moving between the solution and the surface. The adsorption of particles,

4.1 Introduction and overview

FIGURE 4.8 Energy required to detach a single spherical particle (in kT units) with a 90 degrees contact angle from a planar oil-water interface (interfacial tension: 50 mN/m) as a function of the particle radius at 298 K [16]. Reproduced with permission from B.P. Binks, Particles as surfactants  similarities and differences, Current Opinion in Colloid & Interface Science 7 (2002) 2141.

•

however, is irreversible: once they adsorb, they will not leave the interface easily. The adsorption energy of solid particles is much higher than that of surfactants. Binks [16] calculated the energy necessary to remove a particle from the surface, assuming 90
contact angle, as a function of particle size, as shown in Fig. 4.8. It can be seen that very fine particles, or molecules (such as surfactants) have small adsorption energy (typically a few kT), and the energy increases with the particle size; for 100 nm particles, it is very high, above 100,000 kT.

4.1.6 Overview of foam stabilizer and antifoam chemistries In pure liquids with low bulk viscosity (e.g., pure water or low-viscosity solvents) foam films are very unstable, rupturing quickly without the formation of a stable foam structure. Although a foam structure can form from viscous liquids alone (for example, from a low API gravity crude oil), for the foam to become stable they are required to contain foam stabilizers. Table 4.1 summarizes the various types of foam stabilizers, foam stabilizing mechanisms, and typical antifoaming chemistries. Oil-based foams are sometimes stabilized only by their viscosity, which slows drainage of the foam films before they reach a critical thickness and rupture.

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Table 4.1 Summary of typical foam stabilizers. Foam stabilizer

Foam stabilizing mechanism

Typical antifoam chemistry

Foam stability

Oil-based foams (crude oil, lubrication oil, etc.) Viscosity

Film drainage

Solid particles

Intermediate contact angle Various

Asphaltenes, resins etc.

PDMS, F-Sil, OMS Wetting agents PDMS, F-Sil, OMS

Low/medium Depends on contact angle and particle size Medium

Aqueous foams Surfactants Proteins, macromolecules Latex, coatings, slurries Solid particles

Surfactant adsorption layers Protein adsorption

Hydrophobic particle-filled oils Polyethers

High

Various

Polyethers

Medium

Intermediate contact angle

Wetting agents

Depends on contact angle and particle size

Medium



PDMS, Polydimethylsiloxanes (see Section 4.2.4.1.1); F-Sil, Fluorosilicones (see Section 4.2.4.1.3); OMS, Organo-modified silicones (see Section 4.2.4.1.2)  Organic or silicone polyethylene oxide-polypropylene oxide copolymers (see Section 4.2.4.1.2).

Macromolecules (e.g., proteins) can also adsorb on liquid films and various polyethers can destabilize them, although these are less common in the oilfield. Foam stabilizers, especially in water-based systems, are often amphiphilic surfactants, containing hydrophilic (ionic or nonionic) and hydrophobic (lipophilic) parts, typically one of each. Examples of polymeric (macromolecular) foam stabilizers are proteins in water, which also must have both hydrophilic and hydrophobic sections, typically many of each. In aqueous systems, these surface-active molecules adsorb at the water surface due to their hydrophobic nature. The intramolecular attraction between the water molecules alone is stronger than that between the water molecules and the hydrophobic section of the surfactant (or ions), resulting in decreased surface tension. At surfactant concentrations above the critical micelle concentration (CMC), the surface tension becomes constant and the surfactant molecules spontaneously aggregate into micelles. In aqueous systems, the hydrophobic groups orient towards the inner core of the micelle while the hydrophilic groups orient on the outside. In nonpolar liquids reverse micelles form, in which the hydrophilic groups are inside the micelles and the hydrophobic groups are outside, in the oil. Antifoams used for foams stabilized by surfactants typically contain a blend of hydrophobic particles (e.g., hydrophobized silica) and nonpolar oil (mineral oil, silicone oil).

4.2 Oil-based foams

Solid particles can themselves adsorb on the foam films and, depending on the contact angle, form (sometimes extremely) stable foams, as discussed in Section 4.1.5. These solid-stabilized foams can be typically eliminated by wetting agents instead of traditional antifoams (see Section 4.6.3). The most important antifoam chemistries (both silicones and nonsilicones) will be described in detail in Section 4.2.4 for oil-based foams and Section 4.3.1 for water-based foams. Foam stabilization by solid particles was discussed in Section 4.1.5.

4.2 Oil-based foams 4.2.1 Defoaming versus demulsification Demulsification is an important step for processing crude oils and the separation of water. Defoaming is foam elimination by breaking foam films between bubbles, while demulsification is the breaking of emulsion films between liquid drops, by adding small amounts (10500 ppm) of (blended) additives in both cases. It is tempting to assume that the two processes are similar, especially as they occur together in the oilfield often at the same time in the same locations, however, there are fundamental differences between them. Crude oil demulsifiers are mostly surface-active, surfactant-like molecules, containing polymers with both polar (for example polyethyleneoxide) and nonpolar (e.g., hydrocarbon) groups, and generally dosed as a solution in a nonpolar solvent. The surfaceactive demulsifier molecules then adsorb at the oil/water interface, reduce interfacial tension, and destabilize emulsion films. Antifoams, however, are not necessarily surfactant-like. For example, in PDMS (which is perhaps the most commonly used antifoam for crude oils), there are no polar groups, and so it cannot be considered a surfactant in the classical sense. Moreover, antifoams generally work as insoluble droplets and not as molecular solutions. If PDMS dissolves in a light condensate then it becomes a foamant, exacerbating the foaming problem. Antifoaming is often a much faster process than demulsification, acting in seconds versus minutes or hours in the case of demulsifiers. Due to all these differences, there are only very few additives that can act both as an antifoam and as a demulsifier. Such exceptional molecules include organo-modified silicones that are surfactants and can act as crude oil defoamers and demulsifiers at the same time (see Section 4.2.4.1.2).

4.2.2 Nonaqueous foaming Non-aqueous foams, and in particular foams of crude oil, are far more complex than their typical aqueous counterparts. It appears though that nonaqueous foams are not anywhere near as common as aqueous foams, yet they play a critical role across many industries such as the construction, personal care, and particularly

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the oil and gas industry [1,19,21,23,24]. The occurrence of unwanted nonaqueous foams, such as those found in crude oil and gas separator vessels are critical issues. It is therefore important to understand not only the formation mechanisms of nonaqueous foams in general but also the way they are stabilized, and then how they can be controlled. We shall first deal with some general theory around nonaqueous foaming before we apply this to crude oil foaming. The most significant difference between nonaqueous foam to that of aqueous equivalents is the much lower surface tension that exists at the liquid-gas interface as was discussed in Section 4.1.2 [23]. The significance of this is that in aqueous systems, added surfactant readily migrates to the interface thus strongly lowering surface tension; this does not happen in nonaqueous systems as the surface tension is already very low, therefore adsorption of (oil-soluble) surfactants is poor [24]. It is for this reason that surface tension has been deemed a poor tool for analyzing nonaqueous foam stability [23]. Another major departure from classical aqueous foam stability is the insignificance of the electrostatic double-layer repulsion in nonaqueous foams due to the very low dielectric constant of, for example, hydrocarbons; this prevents meaningful electrostatic stabilization [21]. Work on nonaqueous systems, and in particular crude oil, is limited when compared to the ample literature on aqueous foams. There are three generally accepted phenomena for stabilization of nonaqueous foams: • • •

Presence of surfactants Multi-phase condensed media Particle adsorption

There are physical mechanisms associated with these stabilizing phenomena, namely: • • •

Modification of surface rheology Steric stabilization Particulate layer formation.

4.2.2.1 Presence of surfactants This is a huge area to cover and the literature resource is vast on the various surface-active agents (surfactants) that can stabilize nonaqueous foams (e.g., [2628]). Despite nonaqueous liquids having very low surface tensions, many different surfactant molecules can still adsorb to, and therefore modify, the hydrocarbon-gas interfacial rheology. In general, nonaqueous foam-stabilizing surfactants can be categorized into hydrocarbon-type surfactants, siloxanes, and fluorocarbons. The foam stabilizers in previous sections are, however, not part of produced crude oils, which can contain further foam enhancers, such as asphaltenes.

4.2 Oil-based foams

4.2.2.1.1 Modified hydrocarbon-type surfactants The presence of hydrocarbon-based surfactants on the oil-gas interface is well known [19,20,26]. Generically these molecules can be described as long-chain modified hydrocarbons with acid, alcohol, or amine functionality. Early work in this area was on mineral oil and glycol foams stabilized by surfactants based on polyethylene glycol and stearyl alcohol ethoxylate [29]. This work showed that foam stability was directly related to the solubility of the surfactants, with highly oil-soluble surfactants unable to produce stable foam, insinuating that surfactants precipitate at the surface and act as solid stabilizers. Parallel studies have shown similar effects for fatty acid ester surfactants in various nonpolar and vegetable oil systems [30,31]. Foam stability was directly proportional to surfactant concentration and also noted was that the particle size plays an important role in stability, where the smaller the particle the greater the stability [31].

4.2.2.1.2 PDMS and organomodified silicones It is well known that as the MW of PDMS increases then its solubility in oils decreases. Low MW PDMS is soluble in organic solvents and acts as a foam stabilizer in such systems. High MW PDMS (i.e., when the MW is above the marginal solubility limit) acts as an antifoam agent [19,20,32]. PDMS can be modified with organic groups, often with polyethers. The chemistry, properties, and usage of these organomodified silicones will be discussed in Section 4.2.4.1.2. The strong surface affinity of PDMS can be exploited to create specialty surfactants to stabilize nonaqueous foams; this is used, for example, in polyurethane manufacturing when PDMS is combined with polyols [24]. Polypropyleneoxide-modified silicones can also act as foamants of oils.

4.2.2.1.3 Fluorocarbons Fluoroalkyl surfactants are capable of reducing the surface tension of liquids to ,20 mN/m. Independent studies have reported the use of fluoroalkyl surfactants on organic liquids, in particular relation to enhanced oil recovery operations [19,33]. In both literature examples, foam stability was concluded to come from steric forces that arose from overlapping layers of the adsorbed surfactants.

4.2.2.2 Multiphase condensed media This phenomenon describes the stabilization of foam by liquid crystals or even (condensed) solid components at the interface, resulting in long-term foam stability. These types of foams have been investigated by numerous researchers: Friberg [23] elucidated the importance of multiphase condensed media on foam stability alongside Sanders [29] who, as we already learned, expressed the importance of (insoluble) surfactants and liquid crystals on foam stability. It has also been shown that surfactant size is influential on foam stability, with the most stable foams being produced by smaller-sized surfactants [23,31].

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There are further components in crude oil that can greatly affect its foaming and those will be discussed in subsequent sections.

4.2.3 Nonaqueous foams of crude oil Foams that incorporate crude oil are arguably one of the most important and certainly intriguing nonaqueous industrial foam types encountered and have been reported in the industry literature as far back as 1943 [34]. It is the complexity of the continuous phase, that is the crude oil itself, that in turn drives the intricate nature of the foam. Crude oil, by its very name, is itself a mixture of many things (not just hydrocarbons) and can vary drastically from field to field, from well to well, and sometimes even over the lifetime of an individual well. Crude oil is a mixture of hydrocarbon fluid and other organic components and almost always has an associated aqueous phase and trace solids. The hydrocarbons are varying amounts of alkanes, cycloalkanes, and aromatic compounds; other organics include species like resins, asphaltenes, naphthenic acids (and soaps), and other hydrocarbons that contain sulfur, oxygen, nitrogen, and heavy metals. A light crude oil (e.g., a typical Brent Blend) may be . 90% hydrocarbon, whereas a heavy crude oil (e.g., Orinoco Blend) may be , 50% hydrocarbon. All of these components can become incorporated into crude oil foams, which may be complicated further still by the presence of other additives used to treat the crude oil (e.g., corrosion inhibitors, demulsifiers, biocides, scale inhibitors, hydrate inhibitors, asphaltene dispersants, etc.). Many of these production chemicals are themselves surfactants and are therefore relevant to foam formation and stability in a given system. It could therefore be fair to say that nearly all crude oil foams are, by definition, uniquely different from one another. It is possible to encounter foam at just about every stage of the oil and gas production and processing train. As we will learn, foams can be useful, such as those used for drilling [e.g., 35] or gas well deliquification [36], however, they are often undesirable [3740], and it is these types of foams that will form the basis of our discussion in this section. The major driving force for crude oil foam formation is the evolution of gas from the crude during the production process caused by its inevitable depressurization as crude is transported downstream. As soon as crude oil leaves the reservoir, where it has been at equilibrium for many millions of years, depressurization begins; the once-dissolved gas is expelled by the crude oil, and nucleation and subsequent growth of these evolved bubbles create foam. Practically speaking, however, it is not really until the processing of crude oil (i.e., downstream of the wellhead) that the problems begin. Flowing up the tubing (whilst it may result in degassing) typically is not impaired due to foam and may in some cases even be increased due to more favorable well hydrostatics. Foam breakout usually occurs in the first stage separator of an oil process, where liquid and gas separation occurs either in a two- or three-phase vessel. This can create a multitude of challenges, the first of which is a major economic problem as crude oil can carry over into the gas plant creating a shutdown of the

4.2 Oil-based foams

system due to specification requirements and/or possible damage to equipment [38,41]. The presence of foam in a separator reduces its effective capacity and residence time and likely therefore also the efficacy of separation, as foam requires time to break appropriately. Associated with this is an inevitable loss of level control in the separator which further affects the ability to operate the vessel following process design criteria. The previously mentioned carry-over of liquids into the gas plant causes smoky flares and potential equipment (compressor, pump) damage [40]. Carryover can also result in contamination of sensitive gas dehydration equipment such as glycol contactors and amine units for gas contaminant removal (mercaptans or H2S for instance; see also Sections 4.3.2.5 and 4.3.2.6) [4244]. We often forget about the influence of gas carry-under, or rather the presence of entrained gas bubbles in the crude oil that enters the pipeline, which can cause pump cavitation issues and create further trouble downstream in terminals or refineries caused by processing these crude oils above their vapor pressure specification [45,46]. There are a variety of treatment types for crude oil foams and the latter sections deal with these in far more detail; however, just to mention the major two types of treatment here as mechanical and chemical. Mechanical devices, e.g., specific separator internals such as inclined plates or cyclonic inlet devices can be installed along with heating capability. Several excellent reviews and case histories exist of separator internal design and specific aspects to counter foam presence and will not be dwelt on further in this current review [4749]. See further details also in Section 4.4. Chemical control methods on the other hand are described in much more detail in the latter sections of this chapter and needless to say, chemistry provides arguably the most economically robust method for dealing with foam and injection of antifoams [20,39,50,51].

4.2.3.1 Factors determining crude oil foaming 4.2.3.1.1 Volume and properties of the dissolved gas The amount of foam a given system can generate depends upon the amount of gas that is associated with the crude oil, expressed as Gas Oil Ratio (GOR), typically in cubic feet of gas per barrel of oil. A high GOR crude ( . 100,000 cf/bbl) can liberate a greater volume of gas per unit of crude oil liquids produced when compared to a low GOR crude oil [52]. It is not just the volume of gas that is important, as there are plenty of examples of highly foamy, low GOR (heavy) crude oils having foam challenges [5355]. In addition to the volume of gas that can be liberated by a given oil, the composition of the gas has a further bearing upon the stability of the foam and even the ability to create gas-liquid interfacial films. Gases that are not able to dissolve into a crude oil tend not to form stable foams whereas if the gas is soluble in the oil then stable foams can form upon the liberation of this gas, depending on temperature and pressure changes [56].

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One example in the literature postulated that increased foaminess was attributable to the basic components in crude oil reacting with CO2 (and not CH4) when used as an agitation gas [21]. The observation of CO2 creating more stable foams than CH4 was corroborated independently by the work of Chen et al. who found the solubility of CO2 to be twice that of CH4 in the crudes they tested [57]. As different blends were used the foamability increased predictably as the ratio of CO2 increased. Furthermore, Sun et al. determined that CO2-oil films exhibited higher gas-oil interfacial stability when compared to 100% N2 and mixed gas composition (90 mol% CH4, 10 mol% CO2) [58]. Work by Hutin et al. however noted that there is a more complex behavior when comparing N2 and CO2 as the pressurization gas because not only is the solubility different, but the dissolution of CO2 into the crude tested resulted in viscosity changes to the crude which also contributes significantly to foaming [53].

4.2.3.1.2 Crude oil composition Equally important to the influences of the gaseous phase is the compositional influence that the crude oil itself has on foam formation and stability. Several components have been identified in the literature as key promoters and subsequent stabilizers of foam. A good example of a series of work exists from Callaghan and coworkers [37,45,59] who identified short-chain carboxylic acids along with phenolic compounds (all with a MW , 400 AMU) that were responsible for crude oil foam stabilization and formation of an evanescent (short-lived) foam. It was presumed that a surface tension gradient was created and/or a surface viscosity that altered the nonslip boundary condition during film drainage. The other important phase separation problem of crude oil is the formation, stability, and elimination of water-in-crude oil emulsions. There is a vast literature on crude oil emulsions, and we just quote a few articles [6066]. The main crude oil components responsible for emulsion stability are: • • •

Asphaltenes, resins Naphthenates Inorganic (metal oxides, clay, etc.) and organic (waxes, etc.) solids (particles).

Although we discussed in Section 4.2.1 that there are major differences in the mechanisms of antifoaming versus demulsification, nevertheless these components can affect (mostly enhance) foam stability as well. The effect of solid particles depends on their hydrophobicity and subsequently contact angle and some solids may enhance foam stability while others suppress it. Several authors focused on asphaltenes and resins as the primary cause of foam [58,67]. Asphaltenes and resins are natural components of crude oil that comprise multiple polyaromatic hydrocarbon rings with heteroatoms of sulfur, oxygen, and nitrogen (often associated with other interstitial metals, e.g., vanadium). These components are recognized as being responsible for foam formation in crude oil, and it has been suggested that asphaltene’s presence enables bubble nucleation and inhibits coalescence [28,68]. A systematic study of synthetic crude

4.2 Oil-based foams

oil systems (toluene-based) involved adding different concentrations of asphaltene and resin to study foam formation and stability mechanisms [69]. When the asphaltene concentration exceeded 10% in toluene it was noted that a significant change (for the worse) occurred in the foam formation and stability. This was attributed to the formation of asphaltene clusters due to the fact when resin concentration was increased the foamability decreased (as the resins destabilized the asphaltene clusters). The work of Poindexter et al. focused specifically on the aspects controlling crude oil foam formation and stability and included evaluation of viscosity, density, and oil-gas interfacial tension as well as asphaltene and resin composition (MW) and content [28,70]. This work found that for asphaltenic crudes, foam collapse was correlated directly to asphaltene content as well as crude density, viscosity, and surface tension. Foam volume, however, did not relate to any asphaltenic crude oil property. For low asphaltenic crude oils, foam volume (but not collapse rate) was strongly related to the crude oil surface tension, and further investigation yielded the resin composition influence. The presence of naphthenic acids and naphthenates has also been suggested to influence foaming as observed in flow-loop testing using various Malaysian EOR field crudes [71]. In this work, the carboxylic and naphthenic acid components were isolated via an extraction technique, and their influence on the separability and foaming nature was investigated using a combination of static bottle tests and dynamic flow loop tests. The presence of naphthenic (and other long-chain fatty) acids was determined to increase foaminess significantly and it has been suggested that crude oil antifoams combined with reagents that will react with some acidic crude components may have better performance [57].

4.2.3.1.3 Liquid-gas interfacial properties As classic foam theory (for both aqueous and nonaqueous systems) teaches us, surface and interfacial properties of crude oil foam liquid-gas interfaces are critical to understanding foam stability. It is the surface rheology in particular that has the major influence on the stabilization of foam films in crude oil foams, as the surface tension of a crude oil foam (being a nonaqueous foam) is a poor measure of stability [19,59,72].

4.2.3.1.4 Presence of other (solid) phases and water As we discussed in Section 4.1.6, solid particles can create very stable foams, and nonaqueous foams are certainly no exception to this. In the so-called three-phase foams, the coalescence of the bubbles is inhibited by the presence of solid particles partially embedded in the liquid phase [1]. The composition (and therefore wettability) and size of the particles have a profound effect on the stability of the foam. The smaller the particles, the more stabilizing they are to a given foam [19]. It has been observed that particles must possess the correct degree of wettability with the liquid phase so that it remains at the gas-liquid interface rather than sinking into the liquid itself [73]. Stable foams are caused by particles that are partially oleophobic and possess contact angles between 40 and 90 degrees

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[16,74]. We have already seen the example of asphaltene molecules stabilizing foam and here too, insoluble asphaltene nanoaggregates have been seen to influence the stability of crude oil foam [69]. As we already learned, the presence of solids can also induce foam stability if they can concentrate at the interface [52,69]. Sand, silt, clays, corrosion deposits, scales, organic solids, etc. can collect and stabilize foam [16]. It is not just the presence of solids but also aqueous phases that can influence foam stability, and crude oil separators typically contain not only gas and oil but (emulsified) water as well. One particular study was able to show that when . 2% water was added to the synthetic crude oil system a stable foam would form; the absence of water meant no foam would form [75]. It is proposed that water, when present in the system, surrounds the bubbles, and disperses in the oil, producing an air/water/oil system stabilized by the added surfactants. Other studies have found that water had no measurable effect on foam stability in crude oil  again displaying the complex behaviors and interactions that exist in any given system, making generalizing and “rules of thumb” challenging [76].

4.2.3.1.5 Viscosity The influence of viscosity on the stability of any foam is intuitive as it is directly relatable to the drainage rate of the interstitial fluid comprising the foam [52,77,78]. This is often the case with lubrication oils, especially base oils, which do not contain foam stabilizing components [15]. Aside from this (more obvious) drainage effect, viscosity can also influence the ability of the gas to diffuse between bubbles and reduce the Ostwald ripening effect leading to greater foam stability. Such effects have been observed in experimental studies of real field crude oil, where crude oil with a viscosity lower than 150 cP at 37.8
C generated little to no foam [70]. Conversely, there are also “real field” studies that do not support this theory of viscosity impact on foam stability, where high viscosity oils (even with added foam stabilization components) did not generate foam, displaying the complexity and interaction of multiple factors influencing foam stability [79]. Crude oils, however, often contain asphaltenes, resins, naphthenates, and solid particles, as was discussed in Section 4.2.3.1, which can have foam stabilizing effects and thus can override the viscosity effects.

4.2.3.2 Testing methods Any test method aims to distinguish and quantify the two (related but different) terms of foamability (or foaminess) and foam stability [80]. Foamability is the volume of foam formed and foam stability is how the foam remains after formation. Both these properties are easily directly (visually) observable and therefore relatively straightforward methods exist. It is the study of foam generation and foam collapse that is fundamental to understanding which physicochemical properties are responsible for the phenomenon of foamy oils. Two very broad techniques can be distinguished: firstly sparging, where gas is typically introduced directly to the crude oil [70,25,81]. This is sometimes

4.2 Oil-based foams

referred to as the “Bikerman’s gas sparging tube” after the seminal work in 1973 [1]. Crude oil is placed into a cylindrical (glass) vessel outfitted with a fritted glass plug at the base to disperse a constant stream of gas into the bottom of the crude oil at a rate sufficient to sustain a stable column of foam. ASTM methods are in existence that have been adapted by the oil industry that were originally developed for lubricating oils [12,82,83]. The second major method is the depressurization method, and it is generally regarded as the superior of the two techniques as it can incorporate field conditions of pressure and temperature [21,26,34,52,79,84]. In this method the principle is different from sparging as gas is introduced to crude oil in a pressurized vessel, that is the gas is used to charge the vessel to high pressure. Typically, the pressurized cylinder is then introduced to a roller oven for some time to homogenize at temperature. The homogenized fluid is then rapidly depressurized and released into a vessel and the foam height and persistency are noted. How foam is generated is critical to studying its properties: several studies focused on the foaming properties at ambient and low pressures [45,28,70], which is not always representative of what happens in real conditions. Depressurization is generally considered more representative of the phenomena which occur during oilfield production and separation processes. To that end, methods should be focused on approximating all the conditions experienced pressure, temperature, phase compositions, and then changes to these conditions. Pressure effects of foam stability have been reported and multiple studies have varied pressure and found increasing foam stability with increased pressure [85,86]. Bla´zquez et al. studied the effect of the saturation pressure and the nature of the gas on crude oil foamability and stability [21]. They found, generally, that the higher the saturation pressure, the higher the foam stability for the multiple oils tested, which makes perfect sense as it is following Henry’s law and the increased amount of dissolved gas at elevated pressure. The foamability of the crude, however, was not directly relatable to the saturation pressure as this is more strongly influenced by the crude oil composition. In addition, the authors noted that the type of saturation gas has a profound influence on foaming not only due to obvious solubility differences, but also because gases can react with certain chemicals in the crude. Gas composition was recognized much earlier by pioneers of this branch of oilfield process technology [34]. Most foam generation methods rely on optical techniques to determine foam volume and stability. When it comes to crude oils this gets somewhat complicated due to their opaque nature. This has led to the development of alternatives to optical sensors such as the use of electrical capacitance sensors to determine foam levels [87]. The work of Poindexter et al. used sparging methods in combination with a radio frequency/microwave-based probe [27,88]. There are several foam level monitors reported in the literature that have been used in the oilfield; some of the more common devices include electromagnetic radiation, differential pressure, neutron backscattering, sonic echo (ultrasonic) devices, and capacitance methods [89,90].

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Some studies have used more advanced techniques such as interfacial rheology to complement more classical sparging methods [67]. The rheology measurements of the crude oil/air interface utilized a BiCone accessory and were correlated to classical sparge tests. Other studies have utilized controlled stress rheometry under pressure [76]. Recently, Sun et al. provided a comprehensive study of gas-oil interfacial properties focused on three different heavy crude oils from around the world [58]. In this work, a new experimental set-up for a single film stability determination was used to investigate the gas-oil stability under reservoir conditions complemented by a series of micromodel experiments conducted in high-pressure etched glass windows coupled with microscope viewing capability. Bespoke flow loop tests have also been developed to generate foam and generally comprise a recirculating loop controlled by a variable rate pump and a foam cell/chamber for directly visualizing the foam. Some techniques also have incorporated density and flow meters for more accurate observations and data tracking [71,91,92].

4.2.3.3 Crude oil foaming field case histories There are several cases of field foaming available in the literature from a host of different geographies covering a variety of crude oil types, reservoirs, and extraction methods. Some early literature examples have already been highlighted [34,38,45]. An additional early study evaluates foaming in the El Furrial Field in Eastern Venezuela where serious carry-over and carry-under were experienced caused by foam which was being stabilized by abundant asphaltenes [55]. The paper goes on to describe the evaluation of various antifoam products, and an interesting linkage to the asphaltene control program was made insinuating the interplay between these two challenges and chemical programs. A case history from Oman highlights the complexity of determining the root causes of foam formation and describes how a series of upsets on the PDOoperated field related to fluid foaming caused a series of major shutdowns [93]. An integrated, multidisciplinary team was required to solve these issues and identified that excessive CO2 content from the reservoir was the major influencer with further contributions from the presence of sodium metasilicate, carboxylic acids, and shearing from ESP driven artificial lift practices. The key to success was another combination approach where the correct selection of antifoam and emulsion breaker products was determined to be the best solution. Another case history from the Middle East (Majnoon oilfield) elaborates on some simple field-based tests that can be used to determine foaming efficacy; this has been married to more classical laboratory evaluations where, once again, asphaltene content and type are linked to excessive foam formation in the field [40]. One very detailed laboratory and field study elucidate modeling, tempered by both laboratory and field data, to optimize the production of horizontal wells using crude oils from Canada and Venezuela [54].

4.2 Oil-based foams

It is not just land-based operations that suffer from foaming of course, and the next field case history from the literature draws on experiences in the deepwater production of Brazil from the Parque das Conchas Field in block BC-10 [94]. In this work, the complexity of flow assurance and production chemistry challenges associated with deepwater production is explained in the context of foaming challenges in the subsea separation system. A subsea deployable fluorosilicone-based antifoam was developed (utilizing an atmospheric pressure test based on ASTM D892) that was regarded by the authors to be the first of its kind to be deployed as such. On the use of fluorosilicone antifoams see also Section 4.2.4.1.3.

4.2.3.4 Crude oil lift Similar to gas wells, oil wells can also become liquid-loaded to the point of being hydrostatically overbalanced. The difference is that with gas wells it is most often produced, or condensed, water that causes the liquid loading, whereas, in oil wells, it is more typically the oil itself that causes loading to occur. This is particularly a challenge with low pressure, heavy oil reservoirs, but these wells can be also unloaded by adding foamants into the crude oil. In gas wells the foam is water-based and in most cases, water foamant surfactants are used (e.g., betaines, sultaines, alkyl ether sulfonates, etc.) and are generally not usable as Black Oil Foamers (BOF) [95100]. The foamant instead tends to be a silicone resin-type chemistry (such as MQ-resin) or short-chained fluorine-chemistry surfactants. These can be used alone or in combination with other foamants, such as sultaine salts, betaines, etc., and are required to be tailormade to a given situation or application [99,100]. Since these foamant-stabilized crude oil foams can be different from traditional crude oil foams (and can be much more stable), they can cause specific challenges upon reaching the surface and in the production process, such as in separators [95]. It is one of the important criteria of oil foamant selection to also select a backup traditional crude oil antifoam. Lab and large-scale trials are required to determine the antifoam dosage and confirm that the crude oil foam can be readily eliminated [97,98].

4.2.4 Chemistry of antifoams for oil-based foams As already discussed, there are significant differences between aqueous (waterbased) and nonaqueous (oil-based) foaming systems. It should not be surprising therefore that different materials are used for foam control in the two types of situations. The criteria for successful antifoaming action are: 1. The antifoam should have lower surface tension than the foaming liquid (a more accurate definition is that the bridging coefficient has to be positive, this will be discussed in Section 4.6.1.1) 2. The antifoam should be insoluble in the continuous phase

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3. The antifoam liquid should disperse into small droplets. For oil-based foams, the commonly used antifoams are either various types of silicones (PDMSs, fluorosilicones, and organomodified silicones) or organic (nonsilicone) additives, most often polyacrylates. Silicones are most often used because they, in general, have much higher efficiency than nonsilicones. In the next sections, the chemistry, properties, and use of these materials will be reviewed in detail.

4.2.4.1 Silicones (siloxanes) For both water and oil-based systems, silicone-based antifoam chemistries are very important and widely used. Silicones are siloxane polymers where the silicon atoms are connected via oxygen atoms, such as SiOSiO, and most of the remaining valences of the Si atom are bound to methyl groups [101,102]. (Noteworthy here is the different spelling of the silicon element and the silicone (siloxane) polymer.) To simply describe all the many possible branched siloxane structures with methyl groups, a shorthand notation is generally used, as shown in Fig. 4.9, with group symbols for monofunctional (M), difunctional (D), trifunctional (T) and tetrafunctional (Q) groups. Using these notations, all the possible siloxane structures (with no other groups than methyl on the siloxane chain) can be written as M aD bTcQ d. The unique physical properties of silicones can be attributed to the high flexibility of the siloxane chain. This is the result of the very low barrier of rotation of the SiO bonds (much lower than for CC bonds), which allows almost completely free rotation along the siloxane polymer backbone [103]. Many other materials can be made by replacing one or more of the methyl groups in PDMSs with other organic groups. The preparation and properties of these organomodified silicones as antifoams will be discussed in Section 4.2.4.1.2. These materials also have applications beyond that of antifoams and include use as lubricating oils and greases, insulation materials, heat transfer media, release agents, and a multitude of personal care applications; the list continues for a wide range of industries.

4.2.4.1.1 Polydimethylsiloxane PDMS is the most common silicone containing a linear SiOSi backbone and pendant methyl groups in all the remaining valences of the Si atoms (Fig. 4.10). Using the shorthand notation, PDMS is MDnM and the siloxane chain can also be branched by using T and Q groups. PDMS synthesis. PDMS is manufactured with a multistep process [101,103] typically from silicon metal via the so-called Direct Process [104] with methylchloride gas, in the presence of a copper catalyst or other metal promoters.

4.2 Oil-based foams

FIGURE 4.9 Typical monomeric units of siloxane polymers and the typical shorthand notations.

FIGURE 4.10 Structure of the siloxane chain and polydimethylsiloxane.

The reaction yields a mixture of mono-, di- and trichloro methyl silanes: Si 1 2CH3 Cl-ðCH3 Þ2 SiCl2 1 CH3 SiCl3 1 ðCH3 Þ3 SiCl 1 SiCl4

(4.7)

From the reaction mixture (CH3)2SiCl2 (dimethyldichlorosilane) is obtained by fractional distillation and is then converted into a blend of PDMS and cyclic siloxanes via hydrolysis:  
 ða 1 bÞðCH3 Þ2 SiCl2 1 ða 1 b 1 1ÞH2 O-HO ðCH3 Þ2 SiO a H 1 ðCH3 Þ2 SiO b 1 2ða 1 bÞHCl linear cyclic

(4.8)

The linear, noncapped siloxanes with SiOH end-groups are called silanols, which are also used in many applications but are not common in antifoams.

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To obtain the polymer end-capping groups trimethylchlorosilane is converted into hexamethyldisiloxane: 2ðCH3 Þ3 SiCl 1 H2 O-ðCH3 Þ3 SiOSiðCH3 Þ3 1 2HCl

(4.9)

Then the hydrolysis products from both reactions above are reacted, in the presence of acid or base catalyst, to obtain PDMS via an equilibration polymerization mechanism to yield a mixture of linear and cyclic species:  

 HO ðCH3 Þ2 SiO a H 1 ðCH3 Þ2 SiO b 1 ðCH3 Þ3 SiOSiðCH3 Þ3 -ðCH3 Þ3 SiO ðCH3 Þ2 SiO n SiðCH3 Þ3 1 H2 O

(4.10)

The final polymer contains a distribution of MWs. For industrial applications, the degree of polymerization of PDMS is normally characterized by kinematic viscosity, and a wide range of grades, from 0.65 cSt up to millions of cSt are manufactured commercially. For antifoaming of crude oil or other hydrocarbon oils, generally, PDMS viscosities in the 1000100,000 cSt (from 28,000 to 139,000 Da MW) range are used. PDMS is highly stable, resistant to oxidation, and has low reactivity under normal conditions. The primary degradation reaction of PDMS is hydrolysis since the siloxane bond can be split by acids or bases due to the wide bond angle. For the hydrolysis reaction to occur in the presence of a strong base or strong acid, high temperature and/or extended reaction times are necessary. Physical properties of PDMS. PDMS is a clear liquid with a very low pour point (about -70
C) and glass transition temperature (-120
C). At .10,000 Da MW, the PDMS chains start to entangle and their viscosity increases steeply with the MW above this critical value. Low MW PDMS grades are Newtonian fluids up to high shear rates. With higher MW grades ( . 10,000 Da), the viscosity remains constant with the shear rate up to a critical point above which the polymers show shear thinning behavior. PDMS has negligible vapor pressure (except for the smallest molecules and the cyclic siloxanes), and thus cannot be distilled, but decomposes slowly at high temperatures (200
C), and faster above 260
C. The density of PDMS at ambient temperature is about 0.97 gcm23, which is higher than most hydrocarbon oils, and as a result, PDMS droplets tend to phase separate to the bottom of oils. Solubility and surface tension are key properties of PDMS that affect their foaming/antifoaming ability. Despite the relatively polar nature of the SiO bond, PDMS is a nonpolar liquid with a low dielectric constant, due to the presence of the high number of methyl groups that shield the siloxanes. It is soluble in small molecule, nonpolar solvents, such as C5 to about C10 aliphatic, and in most aromatic hydrocarbons (hexane, toluene, xylene, etc.) as well as kerosene and ketones, and can be soluble in high API gravity crude oil, especially light condensates. PDMS is generally insoluble or sparsely soluble in both polar solvents (water, most alcohols, glycols, acetone, etc.) and long-chain hydrocarbons (mineral oils, etc.). Most surfactants (including even the silicone surfactants) are also insoluble in PDMS.

4.2 Oil-based foams

The surface tension of PDMS at ambient temperature is about 21 mN/m (smaller for the lower MW polymers and increases slightly with the MW for the larger molecules), which is lower than that of most oils, surfactants, and other liquids. The interfacial tension between PDMS and water is high (42.7 mN/m) [105], and although PDMS itself is not a surfactant in the traditional sense as it does not contain hydrophilic groups, it is surface active and readily spreads on the surface of water or mineral oil. PDMS is highly efficient at controlling foam in oil-based liquids, such as crude oils, and thus is a commonly used antifoam. However, it has poor efficiency to break foams in water-based liquids. Such liquids it is generally blended with hydrophobic solid particles (See Section 4.3.1.1). The largest industrial use of antifoams for nonpolar liquids is in crude oil processing, where foam control is often crucial (see Section 4.2.3). In particular, in an early step of processing, the produced crude oil flows into a gas/oil separator where an associated pressure reduction occurs. As a result, large amounts of gas are liberated, which can quickly form a high volume of foam and decrease the efficiency of separation, or in the worst case completely block the process, unless antifoam is added. Silicones are commonly used in this application and PDMS is the most typical, with 10,000100,000 MW, at 0.520 ppm concentrations [106,107]. Another large-scale usage of PDMS antifoam in nonpolar liquids is downstream, in delayed cokers, where the temperature is extremely high (450
C500
C) and high viscosity (105106 cSt) PDMS is often the only option. Generally, high-temperature performance improves with increasing PDMS MW/viscosity, however, handling, incorporation, and creation of small PDMS droplets become more challenging with increasing PDMS viscosity. The easiest way to disperse even the most viscous PDMS in crude oil (or in other oils) is to dose it from a solution in a suitable solvent, for example, in aromatic solvents (xylenes, aromatic naphtha, etc.). Small PDMS molecules and cyclic siloxanes. As mentioned above, PDMS contains a distribution of molecules with various chain lengths; commercial grades of PDMS typically contain smaller siloxane molecules of differing surface properties than the larger polymers. It is important to note that the very small polymers (below about 100 cSt viscosity) have a higher solubility in hydrocarbon oils (such as crude oil and mineral oil) and act more akin to surfactants when in solution. A negative effect of this surfactancy is decreased ascension rate of small bubbles and thus a reduced deaeration rate, as explained in Section 4.1.4. Another family of small molecule silicones is cyclic siloxanes, especially the ones with 4, 5, and 6 silicon atoms, respectively octamethylcyclotetrasiloxane (D4), decamethylcyclopentasiloxane (D5), and dodecamethylcyclohexasiloxane (D6). The structure of D4 is shown in Fig. 4.11. Commercial PDMS grades typically contain these cyclic structures as they readily form at elevated temperatures and during manufacturing via equilibration polymerization (see Eq. 4.8). This is a concern in refineries because cyclic siloxanes can gradually cover the surface of hydrotreating catalysts and then block

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FIGURE 4.11 Structure of octamethylcyclotetrasiloxane (D4).

them from functioning. Despite this concern, high viscosity PDMS (which can decompose at high operating temperatures, forming such cyclic species) is the main antifoam used in cokers. The rate of cyclic formation can be reduced by using very high MW and viscosity PDMS (100,000 cSt or higher) [108]. However, this challenge can be exacerbated if significantly more PDMS is used in crude oil separators (typically 500060,000 cSt) further upstream of the refinery and gets carried with the crude oil. The presence of PDMS in crude oil can also reduce the quality of bitumen [109].

4.2.4.1.2 Organomodified silicones Organomodified siloxanes (OMS; organopolysiloxanes) represent a class of antifoams less known than PDMS. In OMS, one or more methyl groups of the PDMS are substituted by organic groups. Fig. 4.12 shows a general structure of this polymer type which can be derived from PDMS with mono-functional organic groups (R1 and R2) [110]. The organic groups can be in various parts of the siloxane chain: on the side (R1-groups; pendant, comb, or rake type) or at the ends of the siloxane chain (R2-groups; ABA-type, linear) or in both positions. The R1- and R2-groups in OMS antifoams most often are nonreactive groups: • • • • •

Polyethers: polyethylene glycol (PEG), polypropylene glycol (PPG), or combinations thereof Alcohols, diols, polyols Alkyl, aryl groups Fluorinated alkyl groups or combinations.

Fluorinated siloxanes (fluorosilicones) have special antifoaming properties and therefore will be discussed as a separate class of OMS antifoams (see Section 4.2.4.1.3).

4.2 Oil-based foams

FIGURE 4.12 General formula of organomodified siloxanes with monofunctional organic groups (R).

Reactive groups (e.g., amines, epoxies, and vinyls) can be also added, but these are not typically used for antifoam applications. While in PDMS the main variable is the number of D-units, with OMS several more variables are introduced to the polymer structure (see Fig. 4.12), such as the number of the various substitution groups, their ratio to the dimethylsiloxane (D) units, etc., and thus a large number of different polymers are possible. The organic substitution can fundamentally change the properties of the silicone molecule, especially its solubility, compatibility, surface activity, and reactivity. The most common process to make these copolymers is by hydrosilylation (hydrosilation), which is an addition reaction between hydride functional silicone polymers (SiH fluids), and olefins, which are typically vinyl- or allyl-started organic group derivatives, in the presence of a catalyst (typically platinum-based). Organomodified siloxanes may also be made by co-equilibration polymerization of dimethylsiloxy groups with organofunctional silanes or siloxanes. Since PDMS products (like most other polymers) contain a distribution of molecule sizes, the OMS also contains a distribution of “x” and also “y” values (except for the smallest molecules). These polymers are often not blockcopolymers (as Fig. 4.12 would indicate); rather, generally, the two (or more) kinds of siloxane groups are randomly distributed in the molecule. Silicone polyether copolymers. Silicone polyether (polyglycol, co-polyol) copolymers (SPE) are important OMS antifoams, and most of the nonfluorinated, OMS-type antifoams for oils are an SPE type. The polyether is generally polyethyleneoxide (EO, polyethyleneglycol), polypropyleneoxide (PO, polypropyleneglycol), or sometimes polybutyleneoxide (BO, polybutyleneglycol), connected to the ends of the siloxane chain or grafted as side chains, in many possible combinations. Fig. 4.13 shows the structure of typical pendant (graft) polyethers, with EO and PO, illustrating that there are five variables (x,y,M,N,Z) in these structures giving multiple degrees of freedom to the design of such polymers. The Z-end group is typically H, CH3, or C(O) CH3 [102]. The main effect of the polyether substitutions is that they change the polarity, and thus solubility of the siloxane polymer. By functionalizing with ethyleneoxide

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FIGURE 4.13 Structure of typical pendant silicone polyethers with EO (polyethyleneoxide) and PO (polypropyleneoxide) substitutions.

the completely water-insoluble PDMS can become water-dispersible or even soluble in water, depending on the ratio of polyethyleneoxide to dimethylsiloxane. Unlike PDMS, these silicone polyether copolymers can be considered surfactants, since they contain both hydrophobic and hydrophilic sections and provide physical properties that are typical of standard nonionic surfactants, including: • • • • • •

reduction of surface tension of water or other liquids formation of micelles and liquid crystals exhibiting a cloud point in aqueous solutions adsorption on surfaces acting as foamants or emulsifiers acting as antifoams or demulsifiers

depending on the type and ratio of the substitutions. Due to the low surface tension of PDMS, SPE copolymers tend to have lower surface tension than traditional surfactants, often not much higher than that of PDMS: 2226 mN/m versus 21 mN/m for regular PDMS at 0.1 wt.%. The reason for this is the flexibility of the siloxane chain and the low energy of the methyl groups in the siloxane [111]. For antifoaming only some of the possible SPE copolymers and other OMS structures are usable, namely those with limited solubility in the foaming liquid since that is generally a requirement for antifoaming action. Therefore, in waterbased systems, SPE copolymers should not be too hydrophilic (i.e., high EO content) and in oil-based systems, such as crude oils, the polyether should not contain too much oleophilic functionality, (i.e., PO or BO).

4.2 Oil-based foams

Use of organomodified silicones as antifoams. OMS polymers have been used as antifoams for nonaqueous liquids since the late 1980s. Callaghan et al. [112] patented water-insoluble polyether copolymers for crude oil separators, suggesting that they cause less air entrainment than PDMS, though experimental evidence was not presented. Many OMS structures have been patented for antifoaming of hydrocarbons, especially for fuels [113]. In diesel fuels and similar fuel oils, OMS antifoams are used almost exclusively [112,114116]. Such antifoams can be cross-linked [116] and contain not only polyether groups, but also polyhydric, aliphatic, aromatic, unsaturated, and alkylphenolic groups [118127]. These patents specify diesel fuel or nonaqueous liquids in general for the applications. OMS polymers are also used as antifoams in lubrication oils. Specific OMS compositions that work as antifoams have little solubility in the oils (lube oils, diesel fuel, crude oil) and similar hydrocarbons, but they are soluble in many polar and aromatic solvents (alcohols, etc.). Since OMS polymers are surfactant structures, some of them can act as demulsifiers and emulsion breakers. It is also possible that the same OMS structures function as a crude oil antifoam and a crude oil demulsifier at the same time.

4.2.4.1.3 Fluorosilicones Siloxanes with fluorine-containing groups (fluorosilicones, FS) are also organomodified PDMSs and generally can be described by the formula in Fig. 4.12, where the R1 and R2 groups contain fluorocarbons, most often only side-chain modified (R2 5 CH3 or OH). The fluorocarbons contain CF3 and CF2 groups (perfluorinated sections) only. The simplest such groups on organomodified PDMSs would be CF3; however, such polymers are unstable, and only polymers with fluoro substitution only in the γ-position, that is furthest from the Si atoms (CH2CH2CF3 fluoropropyl, etc.) are usable. Callaghan and Taylor described FS with CnF2n11(CH2)2O groups for crude oil [128]. From a large number of possible polymers, only a few structures are commonly used. The most important ones are homopolymers with no D siloxane units, only with a trifluoropropyl and a methyl group on all the siloxanes [129]. Combinations of fluorocarbons and other types of organic (polyethers etc.) are uncommon, although can be found in the patent literature [130]. The most common method to manufacture FS is from the cyclic fluoro trimer [129] followed by ring-opening with acid or base catalysis, analogously to PDMS production. The process is, however, significantly more complicated and costly than PDMS manufacturing due to the raw material costs associated with the fluorinated monomer, as well as the thermodynamic favorability of the cyclic trimer species relative to the linear polymer. Thus, commercial production of FS polymers typically requires specialized catalysts and reaction conditions. Fluorinated siloxanes can be also copolymerized with Dx units in the presence of acidic or basic catalysts [131134].

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Properties of fluorosilicones. The fluorinated groups strongly increase the viscosity of siloxanes, although for antifoaming purposes only viscosity grades no more than 50,000 cSt are generally available. The fluorocarbon groups strongly change the solubility, compatibility, and surface energy of siloxanes. FS are generally not soluble in any hydrocarbons (even for small molecules), only in somewhat more polar solvents (especially in ketones and esters), and they are also insoluble in PDMS or SPE. A possibility to avoid solvents is to emulsify the FS, for example into mineral oil [135] or even into an aqueous solution [136], but this is not commonly practiced. Although the CF3 groups can potentially provide much lower surface energy to FS [137], surprisingly the surface tension of liquid poly[methyl(3,3,3-trifluoropropyl)siloxane] is somewhat higher (2324 mN/m) than that of PDMS (2122 mN/n) [129] Similar to PDMS, FS are also not surfactants in the traditional sense, since they also lack hydrophilic groups. Use of fluorosilicones as antifoams. FS tends to be much more effective than PDMS in antifoaming applications, especially in crude oil separators, where the typical use level of PDMS is in the 0.5100 ppm range, compared to 0.15 ppm for FS. Such efficient performance and reduced dosage rates can provide enhanced total cost-performance relative to PDMS, especially in difficult foaming cases. Evans [138] patented first the use of FS (containing trifluoropropyl substitutions) for foam control of crude oils. The advantage of FS antifoams is particularly strong with light condensates, especially in deepwater applications. These oils contain smaller hydrocarbon molecules, in which PDMS is soluble and therefore tends to promote foam rather than controlling it [139,140]. In such oils, FS is generally the only material that can effectively control foam generation due to their incompatibility with hydrocarbons. Similar to PDMS, FS antifoams are widely used in lubricant fluids as well [141]. They are not commonly used as antifoams in aqueous systems, and FS are not usable even in mixtures with hydrophobic solid particles.

4.2.4.2 Nonsilicone antifoams for oil-based foams For crude oil antifoaming, silicones are used primarily, but there are attempts to also use nonsilicone antifoaming agents. These antifoams comprise a wide range of chemistry types and historically have been based on polyacrylates, but there are different additional types in both the scientific and patent literature. Polyacrylates are well-known polymers for many applications, including foam control. The chemistry of polyacrylates and polymethacrylates is shown in Fig. 4.14. They have several versions which differ in the R-group, which are typically linear and branched short-chain alkyl groups ranging from C2 to C12 [142]. Newer versions are also fluorinated, such as 2,2,2-trifluoroethyl acrylate [143] which can boost efficiency. Polyacrylate antifoams are often used with lubrication oils [15], especially when there is a challenge with the use of classical silicone chemistry (incompatibility or regulatory issues), and sometimes the two types are used together. An advantage of

4.3 Water-based foams

RO

RO

(A)

(B)

FIGURE 4.14 Structure of (A) polyacrylates, and (B) polymethacrylates.

polyacrylates is that they do not cause deaeration problems and provide good air release. Another study reports on synthesized alkylacrylate homopolymers (polybutyl acrylates) via emulsion polymerization, providing good antifoaming efficiency with heavy crude oil using high (relative) dose rates of between 250 and 750 ppm [144,145]. The efficiency was a function of molecular mass and there was an intermediate optimum range above and below which antifoam efficacy dropped off. A major problem with the use of polyacrylate antifoams is that they have to be carefully tailored (e.g., by varying the R-group) for a given system so that the basic requirements of antifoaming are met: insolubility in the oil, lower surface tension than the oil, and the ability to form small droplets. This can be achieved in a given (lubricant) oil system, but it is much harder for crude oil due to subtle differences in crude composition even from well to well, the age of a given well, possible well treatments flowing back, or even by the season of the year. Sodium sulfosuccinates, such as dioctyl sodium sulfosuccinate (DOSS) (wetting agents and surfactants in water-based systems) were also found to give a good performance, although the efficiency strongly depended on the water and salt content of the crude oil [39,146]. Further nonsilicone antifoams for oil-based systems, such as oil-based drilling muds, include polypropylene glycols (typically 10005000 AMU) and alcohol esters with 2-ethylhexanol (iso-octanol), separately or in combination, also with silicones.

4.3 Water-based foams 4.3.1 Chemistry of antifoams for water-based foams 4.3.1.1 Silicones, silica-filled polydimethylsiloxane As is the case for oil-based foam control, antifoams used in aqueous systems in almost all applications generally must meet the following criteria: • •

Lower surface tension than the foaming phase Insoluble in the foaming phase

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• •

Chemically inert concerning other components in the formulation Dispersible into small droplets.

There are, however, significant differences between the surface chemistries of water-based and nonaqueous or oil-based systems (see Section 4.1.2). Therefore, different materials are used as antifoams in the two types of liquids. In the case of oil-based foams, commonly used materials that meet these criteria include silicones (PDMSs, fluorosilicones, and organomodified silicones) and organic (nonsilicone) polyacrylates. Aqueous foams on the other hand can be controlled best by mixtures of nonpolar oils (including silicones or organic oils) and hydrophobic solid particles. Although many other types of chemicals have foambreaking effects and may meet these criteria, in most cases are not utilized because they have much lower efficiency. In the next sections, the chemical structure, manufacturing, properties, and use of these materials will be reviewed.

4.3.1.1.1 PDMS as antifoam for aqueous liquids PDMS is highly efficient at controlling foam in oil-based liquids (including crude oils) and thus is a commonly used antifoam in these systems (see Section 4.2.4.1.1). However, PDMS (as well as fluorosilicone) has poor efficiency for breaking foams in water-based liquids, especially if the foam is stabilized by surfactants. Therefore, PDMS is normally mixed with hydrophobic solid particles when used in water-based systems; the particles add functionality to the antifoam by destabilizing the pseudo-emulsion film, which tends to be stable in (surfactantstabilized) water-based systems, as will be discussed in Section 4.6.2.

4.3.1.1.2 Mixed (oil 1 solid) antifoams To efficiently control the foaming of aqueous liquids, it is most typical for combinations of nonpolar oils and hydrophobic solid particles to be used [147]. Mixedtype antifoams have been known for several decades and have a wide spectrum of use, that is, can break a wide range of foam types and conditions and have the highest efficiency with surfactant stabilized foams [148]. In the mixed-type antifoams the nonpolar oil component is typically: • • •

Silicone-based oil, often PDMS Organic oil, such as mineral oil, vegetable oil, fatty acid esters Nonpolar polyethers. The hydrophobic solid particles are typically:

• • •

Hydrophobized silica Waxes Other hydrophobic solids, such ethyl-bis-stearamide, etc. The mixed-type antifoam actives are often called antifoam compounds.

4.3 Water-based foams

The silica used in antifoams is amorphous and of high purity, typically either fumed or precipitated, and must be fine, typically less than a micrometer in average particle size. Fumed silica is typically made by heating silicon tetrachloride (SiCl4) with oxygen and hydrogen, with HCl as a by-product. Very fine, roundshaped, nonporous particles form with primary particle sizes as low as a few nm, but they melt together to form aggregates, which cannot be separated by mixing [149]. These aggregates tend to stick together to form larger agglomerates, which can be broken up by mixing if high enough shear is applied. Precipitated silica is made from alkaline silicate solution, by reacting it with mineral acids, such as sulfuric acid, often followed by grinding. The precipitated silica particles are porous, and their primary size is very small but can aggregate into larger units. Hydrophilic silica can be hydrophobized by heating it with PDMS or with silanes, such as hexamethyldisilazane ((CH3)3SiNHSi(CH3)3) or dimethyldichlorosilane. A method of reacting silica with PDMS at high temperatures is called dry roasting (at about 250
C). The hydrophobization is either performed by the silica manufacturer [150], or by the antifoam manufacturer, who can treat the silica in a similar way or in-situ, during antifoam manufacturing. In the latter case, the oil and the particles can be a simple mixture, or heated, in the presence of acid or base catalyst. This reaction mixture can also contain silicone resins, which yield high-efficiency antifoam compounds and are sometimes called three-dimensional antifoams and contain branched siloxanes. The concentration of the solids in the mixture is typically between 3 and 10 wt.% and the viscosity of the antifoam compound is typically 5003000 cP but can occasionally be higher (10,000 cP to as high as a million cP). The reason for the high efficiency of the oil and hydrophobic solid combinations is that a strong synergy exists between the two components since each of the components separately does not typically control foam, as will be further explained in Section 4.6.2 on antifoaming mechanisms. The efficiency of this type of antifoams is high, especially against a wide range of surfactant-stabilized foams, and is therefore much more commonly used for aqueous systems than any other kinds of antifoams. An important phenomenon with antifoams, and especially with mixed-type antifoams, is that they have finite durability. That is to say, during prolonged foam generation in the application their efficiency gradually diminishes, and after a critical time, their antifoaming efficiency disappears altogether. The mechanism of this antifoam deactivation phenomenon will be discussed in Section 4.6.2. Silicone-based antifoams very often have much longer durability than mineral oilbased products. The reason for this is the high incompatibility of PDMS with water, such that PDMS does not get solubilized into surfactant micelles or easily emulsified as fine (and ineffective) droplets. Since these compounds are insoluble, they are also very hard to disperse into the water-based foamant. Therefore, they are often emulsified into oil-in-water emulsions, and this will be discussed in Section 4.3.1.3. Fluorosilicones are not commonly used as the oil component

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(in place of PDMS or other oils) in mixed-type antifoams for aqueous systems. Organomodified silicones can also break foams of aqueous systems, although with surfactant solutions they often have worse efficiency than the PDMS-based ones.

4.3.1.2 Nonsilicone antifoams The use of nonsilicone antifoam agents in aqueous foam prevention in the oilfield covers only a limited range of chemistries and is arguably a semi-commoditized area of production chemistry. There are also very limited literature resources available on this topic in the oil and gas literature, as most studies in the oil and gas literature focus on crude oil foaming [39,144,22,151]. Nonsilicone-based antifoams are becoming more widely used in the oil industry, especially in applications that involve slurries where surfactant adsorption is not the main foam stabilizing mechanism, such as in drilling and cementing. These types of antifoams have some advantages over silicone-based products; they first and foremost (generally) are much more biodegradable and therefore environmentally acceptable for certain oil-producing basins, for example, the North Sea. Nonsilicone antifoams are much less likely to poison catalysts at refineries because they are less persistent and thermally stable. Nonsilicone antifoams do tend to have lower efficacy than silicone-based products and are more soluble in the oil product, and therefore can interact with the oil and other production chemicals. Examples of several family types are phosphate esters, metallic soaps of fatty acids, sulfonated compounds, amides, polyglycols, glycol ethers, and alcohols [20,146,152].

4.3.1.3 Antifoam formulations Most antifoam actives (PDMS, mineral oil with hydrophobic silicas, etc.) are insoluble in water and therefore it is difficult, if not impossible, to mix them directly into aqueous liquids. To solve this problem, the actives are formulated into other forms, which are easier to disperse in water, including • • •

Emulsions Water dispersible concentrates Powders.

The most important form is an antifoam-oil-in-water emulsion, with various actives content, typically ranging from 5% to 50%. To emulsify the actives, several other components have to be added, besides water, including • • • •

Emulsifiers Viscosifiers/stabilizers (thickeners), which are typically polymers such as cellulose derivatives, polyacrylates, etc., to ensure stability against separation Antimicrobial agents (preservatives) Other additives (dispersants, co-emulsifiers etc.)

4.3 Water-based foams

There is a range of requirements for a good antifoam emulsion, including • • •

High efficiency in the intended applications Good dispersibility in the process of fluid Storage stability against separation, bacterial growth, and long-lasting antifoaming effect.

For best emulsification, the formulator must be aware that antifoams are often shear sensitive and that the droplet size range of the emulsion is also an important parameter (see also Section 4.6.2.6). Therefore, it is important that proper mixing is applied during emulsification and that the order of component addition is also optimized. Antifoaming efficiency tends to improve with increasing oil viscosity (PDMS), but at high viscosities, the emulsification becomes more challenging, and high shear or possibly the use of mechanical homogenizers may be necessary. The typical emulsifiers used in antifoam emulsions are nonionic surfactants (such as EO/PO polyethers), with anionic surfactants less commonly included. The polyethers can be organic or silicone-based (see also Section 4.2.4.1.2), but these polyethers generally have different polarities than the ones used for antifoaming (see also Sections 4.2.4.1.2 and 4.6.2.7). Often a combination of emulsifiers, with low and high polarities, HLB respectively, provides the best results. Antifoam concentrates have similar compositions (with emulsifiers and other components mentioned above), but no or only small amounts of water. Such concentrates are designed to be dispersible in the aqueous foamant.

4.3.2 Water-based applications In the next sections, we will discuss several water-based oilfield applications, concentrating only on the antifoaming aspects of these processes. Beyond these applications, there are still many others where foam control is needed.

4.3.2.1 Aqueous foams in produced water and seawater injection systems Water injection systems (including injection of seawater, produced water, or mixtures of the two) are commonly used to dispose of water and provide pressure support to the reservoir. It is very typical for very large quantities of water to be processed in these systems and velocities of fluid can be very high. Furthermore, there can often be other production chemicals added to the injected waters to bring them into the specification, for example, scale inhibitors, corrosion inhibitors, oxygen scavengers, biocides, etc., which when used can create additional foam persistency compared to that of the native injection water [20]. One of the most common chemistries in use for antifoaming in water injection systems is polyglycol-based products. Only a few parts per million (, 5) are typically needed as the foam is neither persistent nor high forming. Silicone emulsion type antifoams have also been used (in the experience of the authors) for water

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injection applications and have usually been specified by particular manufacturers of specialized water injection systems (e.g., MINOX system) [153].

4.3.2.2 Field examples of injection water system foaming The formation of foam in oil and gas water injection systems is a phenomenon that is relatively commonplace yet has received very little attention in the literature. Perhaps this is because it is rather “unexciting” and utilizes quite commoditized chemistry that has been established for many years. In the literature, the use of antifoams in water handling systems get mentioned in several places [153156]. These papers only mention the use of antifoam and do not describe in detail much chemistry, usage, or application information.

4.3.2.3 Cementing Cementing is an important step in good construction. As the cement, additives, and water are mixed and then pumped downhole [157] air can get entrapped into the slurry and this can be detrimental, causing pump cavitation, falsifying the slurry density measurement and control, and gas migration [158,159]. If there is a significant amount of air in the slurry its density will be higher downhole (under elevated pressure) due to the compressibility of the air bubbles [157]. The slurry “foam” generally cannot be considered a foam in the traditional sense because its gas content is too low to form foam films, Plateau borders such as in Fig. 4.1. It is rather a gas dispersion [159]. Because of this, the mechanism of defoaming in cement slurries is difficult to explain based on traditional antifoaming mechanisms, such as bridging (see Section 4.6.2) of foam films. The high concentration of solid particles also stabilizes the “foam,” due to their effect on the rheology of the slurry. These particles are hydrophilic, and thus probably do not adsorb on the bubble surfaces, as in particle-stabilized foams (see Section 4.1.4). Many additives are used to optimize cement slurries, and the main ones which are responsible for foam stabilization are: • • •

Fluid loss and gas migration control agents, such as latex dispersions and PVA Retarders, such as sodium gluconate, sucrose, calcium lignosulfonate, Dispersants, such as sodium and calcium lignosulfonates [159], and polycarboxylates [160].

4.3.2.3.1 Testing cement slurry antifoams The first test of evaluating a cement antifoam should be its deaeration efficiency. A laboratory test method is the blender test (Fig. 4.15), such as the method based on ASTM D351988 [161]. The gas content of the foamed cement slurry can be calculated by measuring the volume (Fig. 4.16) or by measuring the slurry density with a density cup (Fig. 4.15) versus time. Another method is using a mud balance (mud scale) which is often the method of choice in field conditions.

4.3 Water-based foams

FIGURE 4.15 Blender and density cups to create and measure foamed cement slurry.

Bava and Mahmoudkhani [159,162] used a more realistic, continuous loop to measure foam and entrained air content. The air entrainment content is also visually evaluated in the slurry as well as in the cured cement. After measuring the efficiency, the effect of antifoam on the rheology of the cement slurry, thickening time, fluid loss, and mechanical properties (compressive strength) of the cured cement has to be measured to make sure that its strength is not compromised by the foam control agent [163].

4.3.2.3.2 Chemistry of cement slurry antifoams A cement antifoam can be a liquid that is dosed into the slurry during mixing, or it can be also in solid, powder form that is blended with the dry cement and then disperses inside the mixture as the slurry is created. The typical antifoam dosage is 0.05%0.1% based on the weight of cement (BOWC). It is important that the antifoam should be stable during storage and should work during the entire cementing process. The most common cement antifoams are polyalkylene glycols and silicones, chosen generally among those which are insoluble in water, and thus polypropylene glycols and nonpolar silicones are favored [158]. The actives in silicone

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FIGURE 4.16 Measuring slurry foam volume.

antifoams can be organomodified silicones (see Section 4.2.4.1.2), PDMS alone, or, most commonly, silica-filled PDMS (see Section 4.3.1.1). Bava and Wilson [164] found that nonsilicone antifoams worked well in cement slurry applications, and they are preferred over silicone-based antifoams for offshore cementing, partly because of stricter environmental regulations. Several other chemistries can be used, including tributyl (or tri-isobutyl) phosphate [160], lecithin [165], fatty alcohols, sulfonated oils, silicone emulsions [162], and combinations of actives (polyols, fatty alcohol-alkoxylates, hydrophobic silica, etc.) [166]. There are also efforts to switch from the traditional nonbiodegradable antifoams to biodegradable ones, based on vegetable oils, and nonbioaccumulative materials [159]. Powdered cement antifoams consist of a solid carrier(s) and the antifoam actives, and they must be stable during storage alone or after blending into the cement. Liquid antifoams are sometimes emulsified as oil-in-water emulsions to be easily dispersible in the slurry [157].

4.3.2.4 Drilling and completion Purposeful foaming of drilling and completion fluids can be also detrimental to operations elsewhere in the oilfield, similar to cement. As the drilling mud is circulated downhole through the drill string, the drill bit, and the annulus, it can

4.3 Water-based foams

easily pick up gases, such as produced hydrocarbon gas and CO2. The chemical additives in the muds help to hold not only the cuttings but the gas bubbles as well [167]. As the mud comes to the surface, the pressure drops, and the expanding gas can cause foam gushing [168]. The entrapped air probably cannot be considered a real (polyhedral) foam, similar to cement slurries, but can cause problems, including: • • • •

Pump cavitation Flow problems and altered rheological properties Reduced mud density Incomplete mixing

that can be detrimental for the drilling fluid to function properly [168,169]. A wide range of chemistries are used for foam control of drilling muds for both oil- and water-based systems: • • • • • • •

Silicone antifoams Polyglycols, especially polypropylene glycol Ester alcohols with 2-ethylhexanol Aluminum stearates (especially aluminum tristearate) Polyethylene glycol mono(2-ethylhexyl) ether Phosphates, especially tri-butyl or tri-isobutyl phosphate Blends thereof.

An important consideration in antifoam selection is cost minimization. Similar to their application in cement slurries, polyalkylene oxides (polalkylene glycols, polyethers), such as PPG [170] or polyoxyethylenepolyoxypropylene (EO-PO) copolymers, and silicones [171] are commonly used antifoam chemistries with drilling fluids. The actives in silicone antifoams for drilling and completion fluids can be only PDMS, but most commonly silica-filled PDMS due to the synergy of the silicone oil and hydrophobized silica (see Section 4.3.1.1). Bava and Wilson [164] found that silicone antifoams worked better than polyalcohol-based nonsilicones for offshore drilling.

4.3.2.5 Gas dehydration Water removal is an important step in the purification of natural gas since water condensation can cause corrosion, line blockage, and other operational problems. The most common method of gas dehydration is based on water absorption with glycols, most typically with diethylene glycol (DEG) or triethylene glycol (TEG). First, the water is removed by absorption in the contactor column at low temperature (typically around 20
C40
C for TEG) and high pressure (6070 bars for TEG), and then the water is removed from the rich glycol stream by heating (up to 180
C204
C with TEG) and pressure reduction. In the contactor column, the gas flows countercurrent against the lean glycol stream, through a packed column of trays, and this is where foaming problems can typically occur [172].

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Pure glycols do not foam, however, contaminations (including hydrocarbons, salts, and pH shifts) can initiate foaming, which is indicated by increased differential pressure in the absorber. Foaming can cause reduced water removal and increased glycol losses [173]. Foaming problems are less serious with packed absorber columns than with tray columns [172]. The foaminess of the glycol can be checked by running a sparging test with a sample, and this test can be also used to qualify antifoams [174]. Antifoams can be used to mitigate foaming problems, although they will not eliminate the causes of foaming, and antifoam has to be kept on hand during operation. Sometimes preinhibited glycol is used which contains an “optimum amount” of antifoam and pH buffers [175]. The most important antifoam chemistries include polyalkylene glycols, such as polypropylene glycol (PPG), polyethylene glycol polypropylene glycol copolymers (EO/PO), and the siloxane versions of these glycols (see also Section 4.2.4.1.2). These are often formulated into proprietary blends and yield a typical antifoam dose in the 10250 ppm range.

4.3.2.6 Gas sweetening An important part of natural gas production, including LNG trains, is the gas sweetening (gas scrubbing) unit where acid gases, CO2, and H2S are removed. This is important, since these gases can cause corrosion and other operational problems, and H2S is highly hazardous and toxic. A typical final sour gas content should be , 50 ppm CO2, and , 4 ppm H2S [176]. The sour gases are most often removed by reacting them with aqueous alkanolamine solutions, a technology known for several decades. Various amines are usable, and the most common ones are: • • • • • •

MEA (monoethanolamine) DEA (diethanolamine) ADEG (aminodiethyleneglycol) DGA (diglycolamine) DIPA (diisopropanolamine) MDEA (methyldiethanolamine).

The amine contents are in the 15%65% range, and mixtures of amines are also used. In the process, the natural gas is passed through the amine solution in an absorber column, which is equipped with trays (generally around 20) or packed beds (generally 2 or 3 beds), where the sour gases are captured at 60
C90
C by forming amine salts. The amine liquid and the gas move in counter current, with the gas entering at the bottom, the lean amine introduced into the absorber column at the top, and rich amine then leaving at the bottom to a regenerator. In the regenerator, the amine salts are decomposed at a higher temperature (110
C130
C) back to the original (lean) amine, and the sour gases leave the regenerator at the top; the lean amine is then moved back to the top of the absorber through a heat exchanger.

4.3 Water-based foams

During the process, continuous gas bubbling through the aqueous amine (both in the absorber and the generator) occurs, which can easily cause foaming if even minor amounts of foam stabilizers are present.

4.3.2.6.1 Foaming problems in amine units The most frequent operational problem of amine units is foaming, which can cause: • • • • • • • • •

increased pressure drop in the absorber, decreased liquid level in the absorber increased solvent loss increased sour gas content in the outlet gas loss of production low liquid level and gas carry under in the flash drums higher pressure, leading to gas pockets in the filters pressure drops when antifoam is added pump cavitation.

A simple indication of foaming problems is that a sample of the amine fluid is easy to foam, while the pure amine solutions do not form stable foams [177] or form foams that are easy to control, but contaminations and added components can make the foaming much more serious. Since aggressive foaming is often caused by contaminations, the foaminess of the system is not easy to predict, and flare-ups can also happen. The foaming contaminations can get into the system from outside, with the input gas or other inputs, or could come from components added to the aqueous amine itself, or can even form during the process, for example, from amine degradation. With the input gas, typically contamination with heavier hydrocarbons and solid particles is common as they can get into the system and cause foaming. Additives typically found in the amines used include corrosion inhibitors, antioxidants, and even antifoams. If the incorrect antifoam is used, as we have already learned, the foaming situation can be exacerbated, especially if any surface-active components of lubricants (e.g., from pumps, etc.) enter the system. During the operation of the amine unit, the amine and amine salts can degrade and decompose to a large number of materials (depending on the amine used and the conditions), most of which enhance foaminess [178]. The amine decomposition not only removes some of the amine and increases foaming problems but can also cause serious equipment corrosion and fouling and reduce production. These contaminations can be removed simply by replacing the amine (which is an expensive solution), or more often with filtration. Other purification techniques include vacuum distillation to remove all contaminations, [179] and foaming of the amine solution in a separation column, in a similar way as in flotation, where the foam enriched with the surface-active contaminations on top flows over to a chamber and gets removed from the amine solution [177,180].

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4.3.2.6.2 Use of antifoams in amine units Beyond typical measures and careful operation of the amine unit, antifoam addition is the general remedy for foaming challenges. The foam control agent can be added occasionally, that is a batch of chemicals added periodically, or injected continuously  it depends on the nature of the specific system and what is even possible to deploy. Even with continuous dosing extra antifoam may be necessary during operational excursions. None of the chemicals typically present in a gas sweetening amine system are surfactants, including the amine solution itself, the amine salts, and even most of the contaminants. Therefore, it is not just silica-filled PDMS-type antifoams that are effective but also polyglycols (Polyalkylene glycols, such as polypropylene glycols) and EO/PO copolymers are efficient and commonly used [181]. Silicone antifoam actives are generally used in emulsified form and polyglycols are generally used neat or with a co-solvent, typically added to ease handling [44]. Other antifoam chemistries reported to be used are polyglycol fatty acid esters and longchain (C16C20) alcohols, such as oleyl alcohol and stearates [177,182]. Antifoam emulsions are often viscous, due to the use of thickeners that prevent the antifoam drops from separating (the drops cannot be very small, otherwise the efficiency will be poor, see Section 4.6.1). Higher viscosity can cause pumping problems, and emulsified products are often too concentrated to dose directly, and therefore require dilution with water. Even a relatively small amount of water can decrease the viscosity substantially, but at the same time creates additional stability challenges: the emulsion will be less stable, allowing the emulsion droplets to separate faster. The problem can be mitigated by agitating (with moderate speed) the diluted emulsion. A further, longer-term challenge is that any added preservative in the antifoam may be rendered less effective with the additional water and therefore additional preservative has to be added if the diluted emulsion is to be stored effectively for a prolonged amount of time (in the order of months). If the emulsion is used within a short time (days or a week maximum) dosing of additional preservatives may not be necessary.

4.3.2.7 Water reinjection Surface water from various sources (e.g., seawater, lake water, and river water) has to be purified by removal of certain materials before it is injected into wells; this is especially true for offshore [183]. The purification process includes filtration to remove particulate material and oil and then an additional processing step to remove any dissolved gases. Hydrogen sulfide, carbon dioxide, and oxygen are the typically dissolved gases and it is particularly important to remove oxygen to a level , 20 ppb because it can cause severe corrosion and growth of microorganisms. A typical method of oxygen removal is via a deaeration tower, either by vacuum or by stripping (with natural gas). Surface waters (especially seawater) can have high foaminess and therefore antifoam must be added before the vacuum

4.3 Water-based foams

tower. Silicone antifoams, polyglycols, or PEG-esters [184] are most typical. Since this is not an aggressive foaming system, only low levels of antifoam (a few ppm or less) are generally sufficient. An alternative oxygen removal method is membrane deaeration [185].

4.3.2.8 Steam regeneration For the production of heavy oils, steam flooding techniques are often employed. One specific methodology used particularly in Alberta, Canada in the development of the oil sands is a process called Steam Assisted Gravity Drainage (SAGD). This is a process where a pair of wells is used to inject hightemperature steam into the reservoir to create a steam chamber. The oil is mobilized as the steam reduces the viscosity of the heavy crude allowing it to flow to the second well via gravity drainage, and then the condensed water and oil are pumped to the surface and separated, followed by oil dewatering and water deoiling. The costliest part of this extraction technology is the generation and regeneration of the steam since several parts of steam are required to produce one part of the oil. The water used for steam generation must be processed to very high purity [186,187]. Steam regeneration requires large plants where several chemicals are dosed, including caustic, scale inhibitor, and antifoam. Foam can trigger carryover of brine to the compressor causing fouling [188]. Antifoam is usually dosed into the feed of the evaporator plant and performance is monitored in the evaporator with additional antifoam occasionally added if necessary; typical antifoam use is up to 10 ppm [188]. Here again, silicone antifoams (emulsions) are most effective, but sometimes silicones are not desirable, and only organic antifoams are used (see Section 4.3.1.2).

4.3.2.9 Foam assisted lift A common problem of mature gas wells is liquid loading where foam is intentionally produced but then may have to be eliminated to avoid handling problems further downstream. This is where fluid builds up in vertical (or deviated) gas wells that slows or possibly even stops gas production. The loading liquid can be aqueous (e.g., formation water or condensed water) or nonaqueous (e.g., hydrocarbon condensates), or a combination of the two fluids in a wide range of ratios. Well-known remediation of the problem of liquid loading is a foam-assisted lift (also known as gas well deliquification). In this method, foamants are added to the liquid so that foam with a low specific gravity is created by the bubbling gas, which enables it to be carried to the surface, thus unloading the liquid from the well and restoring gas production. Hydrocarbon condensates do not foam well and therefore generally a water-based foam is created with various amounts of oil inside [189,190]. A wide range of anionic, cationic, or amphoteric types of surfactants [191] or even silicone surfactants [192] can be used as foam stabilizers, depending on the

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well conditions. When the surfactant stabilized foam comes up to the surface it can be very stable and cause severe operational and handling problems thus defoamer dosing is necessary [193]. Since these foams are generally water-based, silica-filled oil-type antifoams (preferably antifoam emulsions) are used (see Section 4.3.1.1). When a substantial amount of oil phase is also present in the foam (Fig. 4.1) then hydrophobic silica can work, although it is often not practical to dose solid particles. Wang et al. [194] recommended the injection of defoamer (remotely) into another additive (ethylene glycol antifreeze) and then into the produced fluid stream to continuously unload wells without foam problems at the surface.

4.4 Mechanical defoaming While antifoams and defoamers are used primarily for foam control, mechanical methods are also applied in some applications, especially where the presence of defoamers is unwanted, such as in fermentation processes [195]. Aqueous surfactant solutions respond to mechanical action, such as expansion or contraction of the surface area by the Marangoni effect, that is on expansion the surfactant concentration decreases resulting in increased dynamic surface tension, thus resisting expansion, and conversely, upon compression, the surfactant concentration increases and the surface tension decreases. This process stabilizes foam films against mechanical deformation, however, if the expansion is too great or the foam film is too thin (the foam is drained) then the film cannot survive and ruptures. The most important mechanical methods are: • • • • •

Ultrasonics Rotary devices Orifice defoamer Pressure pulse technique Centrifugal tools (hydrocyclone),

and some of these can be found in oilfield applications [196]. Jakkulwar et al. [197] applied a pressure pulse technique to control foaming in gas sweetening plants. The pressure inside the absorber was pulsated by increasing and then decreasing the sour gas feed rate. After the pulsation, the pressure drop in the column was reduced, indicating foam rupture and thus reducing or eliminating the potential need for antifoam dosing. Various explanations were offered as foam rupture mechanisms. In gas/oil separators a cyclone type separator can act as a mechanical foam control device and the use of thermal shock and ultrasonic methods was also suggested [38]. These devices are generally combined with the use of antifoams [48].

4.6 Mechanisms of antifoaming action

Similarly, a combination of a cyclone and defoamer injection was used for offshore foam drilling after the foam passed through the annular back pressure control valve [198] and in underbalanced foam drilling with coiled tubing [199].

4.5 Defoaming by chemical reaction Most antifoaming processes involve only physical phenomena, such as entering of oil droplets to the liquid surface, bridge formation, and dewetting (see Section 4.6). However, there are some cases where a chemical reaction, often initiated by pH change, causes defoaming. These chemical methods are mostly suggested to defoam foamed drilling fluids or foamed fracturing fluids. An advantage of these methods is that they are reversible, that is the foaming and nonfoaming properties can be switched. Perhaps the earliest patent on this is by Thomas [200] who used a mixture of an amphoteric foamant (including betaine, imidazoline derivatives, and amine oxide) and an anionic surfactant (a fatty acid or its salt) to create a stable foam at high pH (. 9.5), and it helped to remove the particulates in the wellbore and move them to the surface, where the pH was reduced to four and the foam collapsed and the particles could be removed. The main reason for this is that the free fatty acid is an antifoam (low pH), but its salt is a surfactant (high pH). Then the pH is raised again above 9.5 and the foaming system is reused. Argillier and Roche [201] disclosed a similar system. Chatterji et al. [202] described a foamed fracturing fluid containing a foaming surfactant  a tertiary alkyl amine ethoxylate with pH-dependent foam. The foam can be generated at high pH and it collapses at low pH. Welton et al. [203] described a more complex chemical foam control: a foamed fracturing fluid that contained not only a tertiary alkyl amine ethoxylate but also an orthoester. After the foam is pumped downhole, the unstable orthoester hydrolyzes, and in this way lowers pH and collapses the foam thereby letting the suspended proppant settle into the fractures [32]. Another option to assure that the fracturing foam does not cause handling and disposing problems after it came to the surface is to use foams with low stability called “self-degrading foams” [204].

4.6 Mechanisms of antifoaming action As was discussed in the introduction, foaming characteristics of oilfield liquids strongly depend on whether the liquid is oil-based or water-based, and on the foam stabilizers present. Water-based foams are often stabilized by surfactants that cause a strong reduction in the surface tension and form adsorption layers on the foam film surfaces. In oil-based liquids, such as crude oil, there is no traditional surfactancy,

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however, there can be several components, such as asphaltenes, resins, and naphthenates, which can act as foam stabilizers. Moreover, both water and oilbased liquids can contain small particles that can cause highly stable foams if the contact angle of the particles at the given liquid is intermediate (see Sections 4.1.5 and 4.2.3.1). Researchers have tried to understand how antifoams work for over seventy years, and major progress was made in the past thirty years [205]. The main criteria for foam control (except for solid stabilized foams) are [196]: • • •

The antifoam must have lower surface tension than the foaming liquid The antifoam must be at least partially insoluble in the liquid and form small drops in it The antifoam drops should enter the liquid surface.

After entering the foam film, the antifoam droplets enter the film surfaces forming a liquid bridge inside the foam lamella. This bridge is unstable and will de-wet and break the adjoining bubbles, as will be described in more detail in Sections 4.6.1 and 4.6.2. An important step of the process is that before the antifoam enters the liquid surface, a special liquid film forms between the drop and the air, which is called a pseudoemulsion film (PF). This film must break before antifoaming action can occur. The PF tends to be stable for antifoam oils (PDMS, etc.) in water-based systems, due to the presence of surfactants. For antifoam action, therefore, fine hydrophobic solid particles must be blended into the antifoam oil, which will then break the PF. In oil-based systems the PF is generally unstable, so no particles are needed with the antifoam oil, and the oil alone will work well without them. In these mechanisms, the antifoams are in a heterogeneous phase and are typical in most antifoaming situations. Solid particle-stabilized foams can be broken by a different mechanism. Since these foams are stabilized by the intermediate contact angle, they can be destabilized by reducing the contact angle to (near) zero by the addition of wetting agents (surfactants) causing the particles to submerge into the liquid, and then the foam will collapse. Here the antifoam is the wetting agent, which is dissolved in the foaming liquid, so it is a homogeneous mechanism; this is a less common foam control situation. In the next sections, all the above mechanisms will be discussed in detail.

4.6.1 Antifoaming of nonaqueous foams 4.6.1.1 Thermodynamic coefficients An antifoam oil drop inside the foaming liquid can enter the surface and have various configurations, as shown in Fig. 4.17. After entering the droplet can: • •

form a lens (partial wetting); or spread on the surface and then form

4.6 Mechanisms of antifoaming action

FIGURE 4.17 Configurations of antifoam oil after entering the liquid surface: (A) Partial wetting; (B) Pseudo-partial wetting, (C) Complete wetting.

• •

a thick film (complete wetting) or lens(es) and a thin film around it (pseudo-partial wetting) [205].

Entering and then spreading the antifoam oil can spontaneously occur only if the process is thermodynamically favorable, that is if the interfacial energy of the system decreases. Let us assume that we have an antifoam drop and oil- or a water-based foaming liquid. Various thermodynamic coefficients have been defined using the surface and interfacial tensions present in the system to describe the change in surface energy. The entry (entering) coefficient is defined as: E 5 σAir=Foaming liquid 1 σAntifoam=Foaming liquid 2 σAir=Antifoam

(4.11)

where σ Air/Foaming liquid and σ Air/Antifoam are the surface tensions of the foaming liquid and the antifoam oil, respectively, and σ Antifoam/Foaming liquid is the interfacial tension between the two liquids. The entry of the drop is energetically favorable if E . 0. The two liquid phases (antifoam and foaming liquid) can be in phase equilibrium (saturated) with each other, or a nonequilibrated state before the entering occurs, and thus different entry coefficients can be calculated. A second parameter, the spreading coefficient, S determines how the drop will spread on the liquid surface after entry: S 5 σAir=Foaming liquid 2 σAntifoam=Foaming liquid 2 σAir=Antifoam

(4.12)

For the entered drop to spread S . 0 is required, and S can be also defined with nonequilibrium (initial) or equilibrium values. Since from the equations above E S 5 2 σ Antifoam/Foaming liquid, the entry coefficient is always larger than the spreading coefficient. If E . 0 and S , 0, then the antifoam oil will form only lenses. In some cases, after equilibration S can become negative, and therefore the spread film retreats to a lens or lenses (for example, benzene on water) [205]. Due to its low surface tension (roughly 21 mN/m at room temperature), PDMS has strongly positive E and initial S values, and zero or near zero equilibrium S values on water, most surfactant solutions, and most oils. After spreading

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it can form a duplex film (complete wetting), lenses, or a combination of both (pseudo-partial wetting) depending on the surfactant solution [205207]. If an antifoam oil drop enters both sides of a foam film, then a liquid bridge forms. Garrett [205] found two possible bridge configurations and calculated that one of them [with a low antifoam/water contact angle (ΘOW , 90 degrees)] cannot be mechanically stable, and the one with high contact angle (ΘOW .90 degrees) can be stable only if the film has a particular thickness. A bridging coefficient was also defined to characterize the instability [205]: B 5 σ2 Air=Foaming liquid 1 σ2 Antifoam=Foaming liquid 2 σ2 Air=Antifoam

(4.13)

and concluded that the bridge will be unstable if B . 0, although there are still discussions about how universal this rule is (note that B has a different unit to E and S) [196]. It is important to emphasize that the value of the positive E, S, and B coefficients often cannot predict the antifoaming efficiency in a given system. An example of this is the work of Rezende et al. [84], who studied the antifoaming effect of four organic EO-PO copolymers and two silicone polyethers (also with EO-PO), with two different (medium heavy) crude oils, and also measured their solubility and interfacial tensions (probably the first studies with nonaqueous systems in the literature). Most of the calculated E and S values were positive and the calculated bridging coefficients were also positive, greater than 1000. One of the silicone polyethers (SP2) performed far better as an antifoam than the other silicone and the four organic polyethers, with both crudes, and this performance showed no correlation with the E and S values. SP2 also had the least effect on the surface tension of the crudes. However, SP2 was more polar than the others and thus less soluble in the crude oil, which offered a possible explanation for its higher efficacy. In general, it can be concluded that antifoam tests cannot be substituted by surface and interfacial tension measurements to predict antifoaming actions, especially with organomodified silicones (see also Section 4.2.4.1.2). Beyond the tests above only little published data are available in the literature for E, S, and B for oil-based systems, especially for oilfield liquids such as crude oil. However, there are a lot of similarities between lubricating oils and crude oils, which offer a reasonable comparison between them, since lubricant oils are often crude oil-based hydrocarbons and formulators generally avoid adding foamants into them. The E, S, and B coefficients were determined by Koczo et al. [15] for group II, III, and IV base oils, with either PDMS (30,000 cSt viscosity) or a fluorosilicone antifoam, and it was found that all the thermodynamic parameters were highly positive. This agrees with the fact these silicones are good antifoams for both crude and lubricating oils, thus supporting the theories of the coefficients above.

4.6.1.2 Mechanism of nonaqueous antifoaming Most of the current antifoaming theories are based on the bridging (or “bridgingdewetting”) mechanism; so-called since the antifoam bridge (see Section 4.6.2.4)

4.6 Mechanisms of antifoaming action

in the foam film is generally very unstable, and the antifoam drop dewets the film, a small hole forms, and then the film (and the surrounding bubbles) ruptures immediately [205]. With oil-based liquids, it often happens that there are no foam stabilizers present and their foam films break after drainage when they reach a critical thickness. Therefore, the lifetime of these foam films and foam is controlled by the drainage time, which strongly depends on the viscosity of the liquid (such as crude oil), assuming that there are no other foam stabilizers present. Then it is logical to ask: why do we need to add antifoams to these liquids? The conceptual diagram in Fig. 4.18 explains. Without antifoam, the drainage time is long, since the drainage rate slows down considerably when the film is thin (B ,1 μm) and typically submicron film thickness must be reached for rupture. However, in the presence of antifoam drops (with diameters typically several micrometers) the drop forms a bridge and breaks the film at a much higher thickness and thus much shorter time, that is the antifoam “short-circuits” the film drainage process. It is often critical in gas/oil separators with short residence times to accelerate foam rupture.

FIGURE 4.18 Conceptual diagram of the effect of antifoam drop on the film rupture time. The antifoam drop bridges and breaks the film at a much larger thickness and much shorter time than the drainage and rupture time without antifoam.

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4.6.2 Antifoaming of aqueous foams As discussed above, a major difference between aqueous and nonaqueous liquids is that the aqueous foamants often contain surfactants that cause major changes in the surface and interfacial tensions versus pure water, and also major differences in the antifoaming mechanisms. In the following sections, we will review all the major aspects of these theories, including: • • • • • • • •

importance of the pseudoemulsion film effect of hydrophobic solids importance of the penetration depth rate and location of the antifoaming action various bridging mechanisms antifoam deactivation role of drop size “cloud point antifoaming.”

There are also several extensive review publications where additional information can be found on the topic [3,50,196,205].

4.6.2.1 The pseudoemulsion film The early theories tried to explain how antifoams work by using the thermodynamic coefficients E, S, and B, as discussed in Section 4.6.1.1, but there were many exceptions to these rules, especially with water-based systems. A major problem was that insoluble oils (such as PDMS and mineral oil) alone did not act as antifoams even if E, S, and B were positive. It was also discovered that blending hydrophobic solid particles into the antifoam oil gave a major boost to the antifoaming effect, without affecting the relevant surface or interfacial tensions. The discrepancy between these theories and the observed antifoaming action could be explained by the role of the so-called pseudoemulsion film. Nikolov and Wasan [208] realized first that when an antifoam oil drop approaches the water surface, an antifoam oil/water/air film forms between the drop and the air and this has a key role in the effect of the oil drops on foam stability. The (antifoam) drop can enter the surface only if this film breaks (Fig. 4.19). They called this asymmetrical film a pseudoemulsion film since it is neither an emulsion (oil/water/oil) nor a foam (air/water/air) film. The PF can be defined analogously for antifoams in oil-based liquids as well. Due to its asymmetry (because of the differing interfacial tensions on its two sides), this film is practically never flat, but curved. For antifoaming action, the pseudoemulsion film must break first. Therefore, the concept of PF became a central element of the current antifoaming mechanism theories [205,209,210]. Since pure oil-based systems do not contain surfactants, their pseudoemulsion film probably breaks readily, and it plays no role in the antifoaming mechanism. With aqueous liquids, surfactants are typically present, which can stabilize not

4.6 Mechanisms of antifoaming action

FIGURE 4.19 Formation of (aqueous) pseudoemulsion film.

only the foam films but in an analogous way they can also stabilize the PF, and then its rupture often becomes the rate-controlling step of antifoaming action. Nonpolar liquids (including PDMS, hydrocarbon oils, etc.) alone (without solid particles in them) are generally ineffective in breaking aqueous foams (especially those that contain surfactants) because the surfactants stabilize the pseudoemulsion film. Nevertheless, using oil (such as long-chain alcohols) as a defoamer (see also Section 4.1.3), for example by spraying it onto a foam, can cause quick foam knockdown, because the oil is added directly to the surface from the air phase, so there is no need for entering. However, when the same oil is preemulsified, it generally becomes ineffective as an antifoam. If the PF is stable then not only the antifoaming action is lacking, but the oil drops will stabilize the foam. It was observed that large amounts of emulsified oil can be incorporated into foams of surfactant solutions (without rupturing the foam) if the system contained stable pseudoemulsion films, even with positive entry or spreading coefficients [3,211]. In this case, the emulsion drops, being unable to enter the bubble surfaces, collect inside the Plateau border network this way, slowing liquid drainage and film thinning and causing very high foam stability. Such foam is illustrated in Fig. 4.1. When some hydrophobic silica was added to this system the foam collapsed immediately and could not be reformed. Surfactants used for foam assisted lift probably stabilize the foam by this mechanism (see also Section 4.3.2.9) [192].

4.6.2.2 Effect of hydrophobic solids and penetration depth As discussed in Section 4.3.1.1, the most efficient antifoams for water-based applications contain a blend of nonpolar oils and hydrophobic particles. One of the main roles of the hydrophobic solid particles in these mixed-type antifoams is that they destabilize the PF. Koczo et al. [212] showed this effect using experiments involving the formation of a single pseudoemulsion film. Bergeron et al. [205] also observed that the pseudoemulsion film breaks at much lower capillary pressure if hydrophobic particles are present in the system. Denkov et al. [50,213] developed an experimental technique that allowed them to quantify the stability of the PF by its entry barrier, that is the minimum

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FIGURE 4.20 Illustration of the penetration depth of (A) hydrophobic particle, (B) antifoam oil lens, and (C) mixed type antifoams into an aqueous film.

capillary pressure for breaking it, which was for PDMS (in a sodium dodecylbenzene sulfonate solution) without particles very high (. 3000 Pa), while with mixed-type antifoams it was very low, less than 10 Pa. This explained well that PDMS is a poor antifoam with aqueous surfactant solutions. The antifoaming process is related to the formation of antifoam oil bridges in the foam. Frye and Berg [214] suggested first that a function of the particles in the mixed-type antifoams is that they increase the penetration depth of the oil lens into the aqueous phase, as illustrated in Fig. 4.20B. With deeper penetration, the lens can reach the other side of the foam film easier, and thus the bridge formation is faster. It can also be seen that hydrophobic particles alone (Fig. 4.20A) have low penetration depth, which is a possible explanation for why solid particles alone are generally poor antifoams [212]. The deepest penetration is with the oil 1 particle blends (see Fig. 4.20C). It is also important that part of the hydrophobic particles in an antifoam oil drop are at the surface of the antifoam oil and partially penetrate the water phase, piercing the pseudoemulsion film with their sharp edges leading to rupture. Using fluorescent labeling, Wang et al. [215] have shown the particles at the surface of the antifoam oil drops. Therefore, it is important that the particles are not smooth, but have a high degree of surface roughness or sharp edges. For the best effect, the particles should not have complete hydrophobicity (180 degrees contact angle), or be near to it, because their penetration depth into the water will be near zero [216].

4.6.2.3 Rate of antifoaming and location of oil drops inside the foam The antifoaming mechanism also depends on the rate of foam rupture. If the antifoaming action is fast, the antifoam drops act directly inside the foam films via bridging, as discussed above. If the action is slow (due to lower antifoam activity or low antifoam concentration) then a different mechanism can occur. In a freshly generated foam with antifoam oil drops in it, the spherical bubbles first approach each other and as the bubbles get compressed (forming foam films) they begin to drain due to gravity (buoyancy). During film drainage, a powerful flow inside the films is moving the liquid into the adjoining Plateau borders also sweeping out

4.6 Mechanisms of antifoaming action

the antifoam oil drops. Koczo et al. [212] suggested that as the oil drops and hydrophobic particles collect in the Plateau borders, and as the foam drains and the Plateau borders are shrinking, they get trapped in them and the bridging happens inside them. Denkov et al. [50] showed systematically that with highly effective antifoams (typically with mixed types) there is no time for the drops to leave the draining films (or for the films to get thinner than micrometers) and therefore they concluded that the antifoams (which act fast) must bridge the foam films directly when their thickness is still in the few μm range. The extremely low entry barrier of the pseudoemulsion film (10 Pa or less) with the antifoam oil and particle antifoams is a reason for this fast action. “Slow” antifoams, which are typically liquids (such as PDMS or long-chained, branched alcohols or esters [217] without particles) form pseudoemulsion films with high entry barriers. It is this barrier that makes them act slowly enough (after the foam drained substantially) that they collect and then act in the Plateau borders, also by bridging. The antifoam bridge breaks in both mechanisms either by the dewetting or the stretching mechanism (see Section 4.6.2.4). Also emphasized was the importance of the prespread antifoam oil (PDMS) layer on the film surfaces. In the presence of the antifoam oil film, the entry barrier of the pseudoemulsion was significantly lower than without it. Some details of the bridge formation and rupture are still not well understood [148]. The mechanism of antifoaming in the Plateau borders could explain that for effective foam control the antifoam drop sizes have to be larger than about 35 μm so that they can get trapped inside the Plateau borders. It is well-known that antifoams with small sizes, less than 12 μm, have poor efficiency in aqueous systems.

4.6.2.4 Bridging A liquid bridge from an antifoam drop can form inside the foam film by entering first one side and forming a lens and then the lens will enter the opposite side, as shown in Fig. 4.21. Wang et al. [215] showed a different way of bridge formation. They recorded how a layer of bubbles (between two glass slides under a microscope) coalesce

FIGURE 4.21 Schematic of bridge formation and film rupture inside a foam film.

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FIGURE 4.22 Bridge formation by the coalescence of two antifoam oil lenses.

inside a water phase containing a commercial antifoam emulsion. Their movie pictures show first that some of the antifoam drops entered the bubbles and formed lenses, and then the lenses of two bubbles meet, and as their lenses touch the bubbles immediately coalesce, that is the lamella between them ruptures. At this point the two lenses coalesced, forming a bridge which then ruptured immediately. Thus, here the last step before bridge formation is the rupture of an emulsion film between two lenses, as illustrated in Fig. 4.22. There are various theories about how an antifoam bridge breaks. A commonly cited theory is that the aqueous foaming liquid film de-wets the bridge and it pinches off [205,209,214,215,218,219]. Denkov suggested the possibility of a different mechanism, where the oil bridge ruptures by stretching [50]. A reason for this is that the antifoam oil surface is not stabilized against stretching by the Gibbs-Marangoni effect. This mechanism, however, requires that the antifoam drop is easily deformable, and therefore it is less likely to work with highly viscous oil and solid mixture antifoams (which generally have the highest efficiency) than with low viscosity antifoams.

4.6.2.5 Antifoam deactivation (durability) It is well-known that after prolonged foam generation (mixing, bubbling, etc.) with aqueous foaming solutions the antifoaming action works only for some time and then it diminishes, and the antifoam becomes inactive [220]. During storage, however, the antifoaming ability does not change, hence the diminishing efficiency should be caused only by processes that occur during the antifoaming action. The phenomenon is also called durability (persistence) of antifoams and it is an important parameter of antifoam performance. There are great differences between the durability of various mixed-type antifoams (see also Section 4.3.1). Another key antifoam parameter is the initial antifoaming action, also called “knockdown.” The knockdown performance of an antifoam is generally not related to its durability; either one can be good while the other could be still poor or good. It is not clear if the deactivation occurs with oil-based liquids (see Section 4.6.1).

4.6 Mechanisms of antifoaming action

It was observed that during deactivation an emulsification process takes place, where the antifoam drops become smaller and the solution becomes hazy [212,221,222]. Another important observation was that the smaller drops that form during deactivation do not contain particles, only antifoam oil, and the particles get concentrated in a few large clusters (particle/oil disproportionation). This dispersion is an ineffective antifoam since the oil drops contain no particles (the pseudoemulsion films will not break easily) and the large drops are too few and solid-like. The phenomenon was explained by the spreading of the antifoam [221,222]: when the antifoam drop or lens enters the water/air surface, the surface-active antifoam oil (but not the particles) spreads on the surface, and after the rupture, this oil film rolls up to a small oil drop. This way the antifoam drop gradually loses oil and gets more concentrated in particles, until it is so concentrated that it is solid-like and can stick to other such particles on the surface of the solution. Denkov et al. [223,224] offered a different explanation for the disproportionation phenomenon. They found that the size of the antifoam drops does not decrease (or not much) and it is not important during the deactivation, and that the deactivation point coincides with the disappearance of the (PDMS) oil film on the surface of the aqueous solution, and also that the critical capillary pressure to break the pseudoemulsion film is significantly higher without the oil film. They suggested that the particle-free oil drops form during the stretching of the bridge, causing oil emulsification.

4.6.2.6 Practical aspects: effects of antifoam viscosity, drop size, and mixing The effect of antifoam viscosity on antifoaming efficiency and deactivation (durability) is of great practical importance. Increasing the antifoam oil viscosity often improves the efficiency and significantly slows the deactivation of antifoams, assuming that the drop size is controlled. Nevertheless, the emulsification of more viscous compounds is a higher challenge to the formulator (see Section 4.3.1.3). Koczo et al. [212] studied the efficiency and deactivation (durability) of PDMS and hydrophobic silica blends with PDMS viscosities ranging from 5 to 60,000 mPas and found that both improve with increasing viscosity. A logical explanation would be that the rate of antifoam oil spreading decreases with the molecular size, however, it was reported by Bergeron et al. [206] that the rate of spreading of PDMS is constant up to about 1000 mPas and then starts to decrease with increasing PDMS viscosity. Hence, the improved durability cannot be explained by the spreading rate for oils under 1000 mPas. Koczo et al. [212] found that the pseudoemulsion film becomes less stable with increasing oil viscosity and therefore antifoam action is faster and there is less time for spreading and losing oil from the antifoam drops. The drop size of the antifoam is also an important factor. In aqueous systems, small drops (i.e., , 13 μm) are generally ineffective, which is explained in various ways, either by their inability of getting trapped in the Plateau borders [212]

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or that it takes a longer time for the films to drain to their size [206]. On the other hand, if the drop size is too high, the number of drops becomes drastically smaller, which can slow the foam breaking action. Based on these considerations and experience, the optimum drop size range for aqueous systems is around 530 μm [50]. The droplet size of commercial antifoam emulsions is typically in the 350 μm range (see also Section 4.3.1.3), which has important practical implications. Filtration in the product line can completely remove the antifoams if fine filters (, 1 μm) are used (potentially also plugging the filter), and even 550 μm filters can take out a substantial amount of the antifoam drops. Another important implication of drop size is storage stability. These droplets are large enough to separate in water due to gravity if the viscosity is low, and the antifoam drops will sink or cream depending on the density of the antifoam drops relative to the aqueous phase (higher with increasing silica content). Therefore, most antifoam emulsions are thickened (typically with polymers) to 5003000 mPas viscosity, so that the yield stress of the emulsion will prevent separation. If, however, the antifoam emulsion is diluted with water then the viscosity will drop, and separation will occur with time (within hours or a few days). The user can avoid this problem by either using the diluted antifoam for a short time, continuously agitating the dilution or by adding supplemental thickener and antimicrobial agent if longer, stationary storage is needed. It is well-known that antifoam emulsions are shear-sensitive: if subjected to high shear, their efficiency can be severely reduced, which can be explained by the above arguments on the optimum drop size. This can be an important consideration when high speed, high shear pumps are used to move the antifoam alone or inside the process liquid, especially if the liquid is circulated for a prolonged time. Gear pumps, for example, can significantly decrease antifoam activity.

4.6.2.7 Cloud point antifoams At increasing temperatures, the micelles of nonionic surfactant solutions grow, and the temperature at which the micelles become large enough to scatter light and thus become visible is known as the cloud point. Surfactant solutions above the cloud point become turbid, and then the “micelles” become so large that they also separate, leaving behind a solution with low surfactant concentration. Bonfillon-Colin and Langevin [225] studied the mechanism of this foam destabilization and found that small drops from the surfactant-rich phase that separates at the cloud point can bridge the foam films. As discussed in Section 4.2.4, silicone and organic polyethers (EO-PO copolymers, etc.) are nonionic surfactants and can act as antifoams. If the polyether surfactant is water-soluble, it generally acts as an antifoam only above its cloud point, and this is the reason that these materials are called cloud point antifoams. There are only a few studies published on the mechanism of cloud point antifoams. Nemeth et al. [226] studied the foaming of nonionic Triton X-100 surfactant solutions, in the presence of water-soluble silicone polyethers, versus

4.6 Mechanisms of antifoaming action

temperature. In these solutions, the micelles of the antifoam and foamant surfactants, respectively, combine and mixed micelles form having only one cloud point between the cloud points of the two nonionic surfactants. They suggested that above the cloud point small drops form and they get trapped in the thinning foam films and break them by bridging. There are also organomodified silicones that are insoluble in water (and therefore cannot have a cloud point) that can be even better antifoams than the cloud point antifoams (which are soluble in the water below the cloud point). These water-insoluble silicones always form drops and their antifoaming behavior resembles that of other insoluble oils such as PDMS (especially in oil-based systems, see Section 4.6.1).

4.6.3 Breaking solid stabilized foams As was shown in Section 4.1.6., foams can be stabilized by solid particles if the contact angle is intermediate, that is, the particles are partially wetted by the liquid (see Fig. 4.7). An effective way to destabilize such foams is to change the contact angle to zero, as this may drive the particles into the liquid (or out of it) so that no particles are safeguarding the foam films, and thus the foam collapses, as shown in Fig. 4.23. The easiest way to change the contact angle to near zero is by adding wetting agents. In water-based liquids, good wetting surfactants include water-soluble ionic surfactants, such as sodium dodecylsulfate (SDS), and dioctyl sulfosuccinate sodium salt (DOSS), alkyl sulfonates, and many similar structures. Moreover, nonionic surfactants such as alcohol ethoxylates, EO/PO copolymers, and some of the silicone polyethers (see Section 4.2.4.1.2), especially the trisiloxanes (the socalled “superspreaders”) [227,228] can serve as effective wetting agents. However, it is important not to overdose on the wetting agent, since these surfactants can also act as foam stabilizers, causing the formation of another type of foam. An interesting feature of this antifoaming method is that in this case, the foam control agent (the wetting agent) is soluble in the foaming solution, unlike in any

FIGURE 4.23 Destabilization of solid stabilized foams with wetting agents.

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of the other mechanisms in the previous sections, making this a homogeneous antifoaming mechanism. The antifoaming action of wetting agents can be also enhanced (or possibly replaced) by traditional antifoam emulsions, and this way the danger of overdosing them can be minimized. In oil-based systems, oilsoluble surfactants and demulsifiers can act as antifoams by a wetting mechanism. This is the most observed practical way in which solids are dealt with in the oilfield, that is demulsifier injection upstream of separation vessels so that the solids are removed before being able to form a stable foam. Often the demulsifier injection is in tandem with an antifoam (often even into the same vessel) and a practical balance in the field can be struck between the injection of the two chemicals.

4.7 Concluding remarks The process of foaming and antifoaming from an academic sense has been very well studied and it is fair to say from a practical sense there is not much remaining to learn. However, from an academic sense, there is much to learn and the continual pursuit of knowledge in that regard will go on. From an industry perspective, the application of antifoams is very often considered to be a “fit for purpose” or “semi-commoditized” area, and end-users are not overly keen to sponsor innovation efforts. The existing solutions are often good enough in terms of performance. A wide range of situations regarding foam generation may be encountered in oil and gas production processes, requiring the development of diverse antifoam chemistries to ensure stable, cost-efficient operation. As noted in this chapter, such chemistries can be customized to address specific challenges, considering key factors in the system including the type of continuous phase (water-based or oil-based), the chemistry of the foamant(s), the presence of stabilizers, pH, temperature, and filtration. The mechanisms of antifoaming in most situations are based on bridging by antifoam oil drops, and they are well-understood now thanks to major research conducted in the field during the past four decades [195]. The types of chemistry employed for foam control may require a deeper review and further innovation in the future given the intense focus on sustainability and legislation changes driving toward more environmentally acceptable materials. The silicone and fluorosilicone chemistries remain the most widely used chemistries for crude oil foam control, despite significant (but as yet unsuccessful) efforts to replace them with reliable, nonsilicone alternatives. Although the degradation mechanisms of silicones in the environment have been widely reported [229233], it is expected that further development will continue, and a range of sustainable antifoam chemistries will eventually be available for use in oilfield applications.

References

Nomenclature ADEG AMU BO BOF BOWC CMC DEA DEG DGA DIPA DOSS EO EOR F-Sil GOR HLB MEA MDEA MW OMS PDMS PEG PF PO PPG SAGD SDS SPE TEG

Aminodiethyleneglycol Atomic mass unit Polybutyleneoxide Black oil foamer By weight of cement Critical micelle concentration Diethanolamine Diethylene glycol Diglycolamine Diisopropanolamine Dioctyl sulfosuccinate sodium salt Polyethyleneoxide Enhanced oil recovery Fluorosilicones Gas to oil ratio Hydrophilic lipophilic balance Monoethanolamine Methyldiethanolamine Molecular weight Organo-modified silicones Polydimethylsiloxane Polyethylene glycol Pseudoemulsion film Polypropyleneoxide Polypropylene glycol Steam assisted gravity drainage Sodium dodecylsulfate Silicone polyether copolymer Triethylene glycol

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CHAPTER

Polymeric drag reduction in pipelines

5

Yung N. Lee1 and Ray L. Johnston2 1

Liquid Power Specialty Products, Inc., Houston, TX, United States Liquid Power Specialty Products, Inc., Ponca City, OK, United States

2

Chapter Outline 5.1 Drag-reducing agent history ..........................................................................228 5.2 Basic pipeline hydraulics tutorial ..................................................................229 5.2.1 Reynolds number ......................................................................230 5.2.2 Laminar flow .............................................................................231 5.2.3 Turbulent flow ...........................................................................231 5.2.4 Pressure drop ...........................................................................231 5.2.5 Static head ...............................................................................232 5.2.6 Friction pressure .......................................................................233 5.2.7 Gradient ...................................................................................235 5.2.8 Profile ......................................................................................236 5.2.9 Pipeline pumps .........................................................................237 5.2.10 Operating point .........................................................................237 5.2.11 Calculating drag reduction performance in a pipeline system ........238 5.3 Drag-reducing agent chemistry ......................................................................240 5.4 Drag reduction mechanism ............................................................................240 5.4.1 Misconceptions ...........................................................................242 5.5 Application to the pipeline—drag-reducing agent theory .................................243 5.5.1 Applications in oil/water or multiphase pipelines ............................247 5.6 Utilization of drag-reducing agent in pipeline operations ................................249 5.6.1 Example cases for utilization in pipelines ......................................251 5.7 Conclusion ...................................................................................................257 Nomenclature ......................................................................................................257 References ..........................................................................................................258

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00005-3 © 2023 Elsevier Inc. All rights reserved.

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5.1 Drag-reducing agent history In 1949 British researcher B.A. Toms published data showing his discovery that a dilute solution of poly(methyl methacrylate) in chlorobenzene would flow faster in pipes and could flow with reduced drag by up to 50% [1]. Toms concluded, partially correct, that this behavior was due to a “wall effect” or region near the wall where polymer molecules would be excluded because of their large size. This phenomenon of reduced drag affected by low concentrations of polymer became known as the “Toms Effect.” At the same time, Mysels and his associates were conducting research using gasoline thickened with aluminum soaps [2,3]. Their results, which weren’t published until a few years later, showed that the treated gasoline appeared to have lower frictional resistance to turbulent pipe flow than the less viscous untreated gasoline. The finding of these two pioneers in the hydraulics world was a distinct surprise. Traditional hydrodynamics would suggest that a solvent treated with low concentrations of polymer would require if anything, a higher pressure drop for flow due to its slightly higher viscosity. But, yet, these treated solvents were showing much less turbulent friction resistance than their untreated pure counterparts. The work of Toms and Mysels set off a wave of academic research for the next several decades into the phenomenon of drag reduction [46]. Much of the early work was conducted in water using water-soluble polymers, due to the availability of the polymers, both from natural sources and industry and due to the safety and availability of the flow medium. Research in the 1960s showed that polymer solutions could also drag reduce external skin friction, allowing for higher velocities of selfpropelled vehicles in water [7,8]. Much of the academic work focused on understanding the mechanism for drag reduction and defining the effects of various solvent, polymer, and flow parameters on the level of drag reduction achievable [913]. This work has shown the importance of polymer molecular weight and hydrodynamic volume of the polymer in the solvent (good vs poor solvent) to the drag reduction performance. The work of Virk [14] and others showed that there is a theoretical limit to the overall drag reduction that can be achieved. This maximum drag reduction (MDR) asymptote is largely defined by the turbulence of the system. In the latter half of the 1960s, the production division of Continental Oil Company conducted research into the effectiveness of various high-molecularweight polymers to drag reduce aliphatic hydrocarbons such as crude oil. This work led to the eventual development and testing of an experimental drug reducer (code-named CDR) in field pipelines in Oklahoma. Their results for performance in 8- and 12-inch pipelines demonstrated that sizable drag reduction could indeed be achieved in these large-diameter, long-length pipelines. Their results also discredited some of the ongoing voiced concerns about the ultimate degradation of the long polymer chains due to long-term exposure to the shear conditions of the pipelines. These test results were published and presented at an SPE meeting in Houston in 1970 [15]. This work also led to the eventual granting of the first US

5.2 Basic pipeline hydraulics tutorial

patent in 1972 for the application of high-molecular-weight polymers in hydrocarbon pipelines to increase the flow rate of a pipeline [16]. However, the technical and cost issues associated with scaling up the production volumes did not warrant continued internal full-scale use or commercial application of the technology. That all changed in late 1977 with the temporary loss of one of the operating pump stations on the Trans-Alaska Pipeline. The urgent need for a quick solution to increased throughput led to cooperative testing and development of the first commercial pipeline drag reducer, CDR101. During testing, Alyeska realized the potential for increasing throughput well beyond the existing level without the need to build further pump stations [17]. Commercial injection began in 1979 at three pump stations and continued uninterrupted for more than one decade. The injection levels of drag-reducing agents (DRAs) at each pump station were in the range of 20200 ppm. With the advent of a successful commercial product for drag-reducing pipelines, the oil industry began to quickly define other opportunities for using DRA [1821]. Applications of CDR rapidly grew to pipeline sites over five continents. Many of the applications were to increase the throughput of offshore pipeline systems where oil production was exceeding the original design capacity of the main to-shore pipeline. In 1983 applications of the DRA were increased to include refined product pipelines [22]. The early commercial DRAs were produced by solution polymerization which yielded a final product that was a very viscous and elastic gel (“liquid bubblegum”). These products were difficult to handle, required pressurized containers to inject, and required specially designed nozzles for optimum injection into the pipeline. With mounting costs for worldwide delivery in the early 1990s, Conoco Inc. developed, patented, and introduced suspension DRAs [23]. These next-generation DRAs were a suspension of bulk-polymerized polymer in a low-viscosity carrier fluid. Not only were the suspension DRAs higher-performing (25 times more efficient), but they also represented much more ease in handling as they did not require pressurized vessels for site delivery or injection. As the use of suspension DRAs replaced the use of conventional gel DRAs around the world, many customized versions of the suspension DRAs were developed for specific applications or areas such as arctic climates, the North Sea, and more viscous crude oils, and refined products. In 2007 ConocoPhillips introduced and patented a new type of DRA to specifically meet the increased demand for movement of high-asphaltene laden heavy crude oil from Canada and South America to US refineries [24]. This DRA was in a latex dispersion form and utilized a non-PAO polymer.

5.2 Basic pipeline hydraulics tutorial One cannot discuss pipeline drag reduction without first understanding pipeline hydraulics. The science of pipeline hydraulics is credited to Osborne Reynolds

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(18421912) for starting this field. He is the engineer that came up with the dimensionless number we now know as the Reynolds number (NRe). He was the first engineer to discover that liquids flow in three distinct manners depending largely upon how relatively fast they are flowing. These three types of flow (flow regimes) are laminar, and turbulent, with a transition flow going from laminar to turbulent. Laminar flow is the most efficient way a liquid will flow in a pipe and usually is at very low rates. Turbulent flow happens when the flow rate exceeds a certain point and the liquid flow pattern becomes unstable and begins to contain many violent eddies. As one can expect, this type of flow is very inefficient. But this also is the most common way oils tend to flow in a pipeline due to the need or desire to move the oil at a reasonably high flow rate. Contrary to intuition, turbulent flow in a pipeline also minimizes the intermixing between sequenced batches of oil; so, this is another reason for preferring turbulent flow over laminar flow.

5.2.1 Reynolds number The NRe is dimensionless and is defined as: NRe 5

VDρ VD 5 μ v

(5.1)

Where: V 5 velocity D 5 inner diameter of the pipe ρ 5 density of the fluid μ 5 absolute viscosity ν 5 kinematic viscosity In metric units, the NRe can be calculated as: NRe 5

1000VD v

(5.2)

Where: V 5 velocity (m/s) D 5 inner diameter of the pipe (mm) ν 5 kinematic viscosity (cSt) In US/English units, the NRe can be calculated as: NRe 5

2214 3 BPH ID 3 v

Where: BPH 5 Flow rate in Barrels per hour ID 5 inner diameter of the pipe, inches ν 5 kinematic viscosity, cSt

(5.3)

5.2 Basic pipeline hydraulics tutorial

The traditional definition for the occurrence of the three flow regimes is: Laminar Flow: NRe , 2100 Transitional Flow: 2100 , NRe , 4000 Turbulent Flow: NRe . 4000 The major difference between the laminar flow and the turbulent flow is how the fluid flows in the pipe.

5.2.2 Laminar flow Laminar flow is a “quiet, smooth” flow with all flow molecules flowing in a unidirectional manner down the pipe. This type of flow is viewed as being much more energy-efficient. The following diagram (Fig. 5.1) depicts the velocity profile of laminar flow viewed through a cross-section of a pipe. The fluid at the center of the pipe is moving at the highest velocity, whereas, the fluid near the wall is moving very slowly.

5.2.3 Turbulent flow Turbulent flow is very chaotic and has turbulent eddies that originate from the pipe walls and grow as the fluid flows down the pipe. Because of the turbulent eddies, the fluid mixes throughout the cross-section of a pipe and tends to have a more uniform (more plug shaped) cross-section profile for average velocity as shown in Fig. 5.2. This type of flow is very inefficient in terms of utilization of energy to move the fluid down the pipe due to all the eddy or vortex motions which are counter to the overall direction of flow.

5.2.4 Pressure drop The driving force for liquid flow in a pipe is the energy imparted at the beginning of the pipe in the form of pressure. As the liquid flows down the pipe, energy

FIGURE 5.1 Laminar flow cross-section velocity profile.

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FIGURE 5.2 Turbulent flow cross-section velocity profile.

loss occurs due to “frictional” effects in the pipeline. This yields a “frictional” pressure drop. At higher flow rates, the frictional pressure drop will be higher. Pressure is defined as a force per unit area. The SI unit of pressure is the Pascal (Pa), which is equivalent to one Newton/meter2. Other commonly used metric units of pressure are the kilopascal (kPa) (where 1000 Pa 5 1 kPa), bar (where 100 kPa 5 1 bar), and psi (where 6.895 kPa 5 1 psi 5 1 pound-force/inch2) Atmospheric pressure (or barometric pressure) is the absolute pressure that exists in the open air at any given place. Absolute pressure uses a perfect vacuum as its reference. The normal barometric pressure at sea level is 101.325 kPa (1.01325 bar) or 14.696 psi. Atmospheric pressure decreases as elevation increases above sea level. At 1000 meters of elevation above sea level, the atmospheric pressure will be about 90.113 kPa, or 11% less than that at sea level. Gage pressure uses atmospheric pressure as its reference; it is the pressure above atmospheric pressure. Pipelines will typically utilize gage pressures in all of the operational monitoring and calculations.

5.2.5 Static head A static head is defined as the height of a column of a given fluid above a given reference point. The static head can be given in units of pressure. For an incompressible fluid, the gage pressure at the bottom of a column of the fluid can be calculated by: p 5 ρgh

Where p 5 pressure ρ 5 density of fluid g 5 acceleration constant due to gravity h 5 height of a column of fluid

(5.4)

5.2 Basic pipeline hydraulics tutorial

FIGURE 5.3 Gage pressures in a liquid column.

By rearrangement of the above equation, the static head or equivalent height of a given fluid can be calculated from the gage pressure: h 5 p=ρg

(5.5)

The pressure at the bottom of the fluid column is a function of the specific gravity of the fluid as well as the fluid column height. The increase in pressure due to a rise in liquid is linear with the liquid depth (Fig. 5.3).

5.2.6 Friction pressure First, what is frictional pressure loss? It is the loss of energy or pressure due to the friction within the flowing fluid and the interaction with the pipe wall. Several factors influence the frictional pressure loss that is experienced in a pipeline system. Those factors are: • • • •

The viscosity of the fluid The velocity of the fluid The internal diameter of the pipe Surface roughness of the pipe

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The basic equation used in estimating or calculating frictional pressure losses in pipeline systems for either laminar or turbulent flow is known as the DarcyWeisbach or the Darcy equation. It is expressed as: hf 5 f

L V2 D 2g

(5.6)

where: hf 5 frictional losses (meters) f 5 friction factor (dimensionless), can be read directly from the Moody chart L 5 pipe length (meters)V 5 average pipeline velocity (m/s) D 5 average inside diameter (m) g 5 gravitational constant (9.8 m/s2) For laminar flow (NRe below 2100) the surface roughness of the pipe does not affect the frictional losses in the pipe and the friction factor can be expressed as: f5

64 NRe

(5.7)

Please note that this friction factor is the “Darcy” friction factor, which is 4 times the value of the “fanning” friction factor. A lot of modeling programs use the fanning friction factor, so the user must be sure of which friction factor is being used in calculations. One has to be consistent in the friction factor definition to ensure that the correct pressure drop or head is being calculated from the equations. For turbulent flow, numerous equations have been derived to calculate the friction factor. All of these equations consider the surface roughness (ε or e) of the pipe which is measured in units of length. Drawn metal tubing would have the lowest absolute roughness (0.000005 ft), whereas galvanized iron pipe would have a roughness approximately 100 times higher. The absolute roughness of commercial steel used in modern-day pipelines is about 0.00015 ft or 0.046 mm. Once one knows the roughness factor and, therefore, the relative roughness (e/d) for the pipeline pipe and they have calculated the NRe for the flow conditions, they can estimate the friction factor in several different ways. The simplest way is to look up the friction factor from a Moody diagram as shown in Fig. 5.4. With the advancement of computer systems, instead of visually estimating the friction factor from this chart, many people have proposed equations to approximate the friction factor. Some of the most prominent equations in use today in modeling programs are the following: The Jain equation [25]: Fanning Friction Factor 5 

1 

0:9 2 2:2824 log Dε 1 29:843 NRe

(5.8)

5.2 Basic pipeline hydraulics tutorial

FIGURE 5.4 Moody chart for (Darcy) friction factors. From https://commons.wikimedia.org/wiki/File:Moody_diagram.jpg.

The Swamee & Jain equation [26]: 1 Fanning Friction Factor 5  

2 0:9 ðDε Þ 6:97 1 3:7 4 log NRe

(5.9)

The Churchill equation [27] is the best equation for the transition flow region into laminar flow:  Fanning Friction Factor 5

8 NRe

12 1

1 ðA1BÞ3=2

!1=12

32

(5.10)

where: !!16 

A 5 2:457 ln

1 0:9 ð7=NRe Þ 10:27 e=D

!16

B5

37;530 NRe

5.2.7 Gradient A gradient is a graphical representation of the pipeline pressure loss due to friction. A gradient will plot the head or pressure of the pipeline relative to the distance down the pipeline, as illustrated in Fig. 5.5. Generally, it is easier to

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FIGURE 5.5 Illustration of a gradient line for a pipeline.

convert the gradient from pressure units to head loss (see static head) so that the profile (discussed in the next topic) can be drawn using the same scale. The slope of the gradient is equivalent to the head (or frictional pressure) loss per length and can be in units of head length/pipe length or units of pressure/pipe length. The slope will change with the flow rate of the pipeline and will increase as the flow increases. Some level of backpressure is required at the end of the pipeline for sufficient pump suction at the next pump station, or to provide pressure to get the petroleum into the terminal. Therefore, the discharge pressure at the pump station will need to increase to meet either an increase in flow rate or an increase in backpressure.

5.2.8 Profile A profile is a graphical representation of the actual elevation of the pipe along the path of the pipeline. Profiles, gradients, and the maximum allowable operating pressure (MAOP) can all be placed on the same graph, as long as all are in the same units. Generally, it is easier to convert the gradient and MAOP from pressure units to length units (see static head) so that the profile can be drawn using the same scale as the elevation. Fig. 5.6 is a depiction of a pipeline hydraulic profile including the gradient and the MAOP line. A pipeline system operating under normal preferred conditions will have the gradient curve passing between the elevation profile and the MAOP profile with no intersections of either profile. If the gradient curve falls below the elevation profile, then this will either represent an unachievable flow rate or can represent slack-line operation over a mountain top. If the gradient curve passes through the MAOP line, then this represents a situation where the local pipeline pressure

5.2 Basic pipeline hydraulics tutorial

FIGURE 5.6 Illustration of a hydraulic profile map for a pipeline.

would exceed the MAOP of the pipe at that point. A pipeline operator must be operationally prudent to make sure that the gradient is passing between the elevation and the MAOP profiles.

5.2.9 Pipeline pumps We use pumps to impart energy and increase the pressure of the liquid at the start of the pipeline to move it down the pipeline and deliver it to the terminus at the required pressure and rate. There are two basic types of pumps: 1. Positive displacement pump: which gives a constant flow rate regardless of pump discharge pressure. This type of pump will deliver whatever pressure is required to move the liquid at that flow rate in the system. 2. Centrifugal pump: which produces a constant static head (blocked in; no flow) and a varying flow rate with varying pressure. This type of pump imparts velocity energy to a liquid using an impeller. This velocity energy is then transformed into pressure energy as the liquid leaves the pump. Centrifugal pumps are used in nearly all mainline pipeline pumping stations. Pump curves for both a positive displacement pump and a centrifugal pump are depicted in Fig. 5.7.

5.2.10 Operating point The operating point is the intersection of the pump curve with the system curve when the system curve is overlayed on top of it. Since the pump will always operate on the pump curve for that impeller size, and the system curve will always

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FIGURE 5.7 Pump curves for different types of pumps.

stay on the curve representing the fluid in the system, the predicted operating point can be determined by the intersection of the two curves. Fig. 5.8 shows an example of a pump curve overlayed with two system curves (Line A and Line B). We can predict the resulting pump operating point for the Line A system curve as follows: Flow rate 5 B55,000 BPH Head 5 B800 psi Let’s say the pipe size has been increased, and the system curve shifts to Line B. We can now predict the operating point in the larger pipe as follows: Flow rate 5 B81,000 BPH Head 5 B630 psi The reasons for a pump curve shift could be a change in impeller size or change in pump speed; while the system curve shift could be caused by changes in a liquid property (viscosity, specific gravity), line length, line size, or destination pressure.

5.2.11 Calculating drag reduction performance in a pipeline system The performance metric for any application of DRA in a pipeline is the percentage drag reduction (%DR) that is achieved in the pipeline segment being treated. The %DR value is based solely on the frictional pressure drop in the pipeline

5.2 Basic pipeline hydraulics tutorial

FIGURE 5.8 Operating curves for a pipeline.

(total gage pressure drop less the elevational head pressure loss) and is determined by Eq. (5.11). %DR 5

ΔPf

ðuntreatedÞ 2 ΔPf ðwith DRAÞ 3 100 ΔPf ðuntreatedÞ

(5.11)

The frictional pressure drop values used in the above equation must be for the same volumetric flow rate. Sometimes the treated flow rate is not the same as the baseline flow rate. In this case, the pressure drop due to friction for the untreated case (baseline) must be normalized up to the treated flow rate to make an even comparison. Use the following formula to normalize a pressure to a given flow rate (Q):  ΔPf ðcorrected baselineÞ 5 ΔPf ðbaselineÞ

QðtestÞ QðbaselineÞ

n where nB1:8

(5.12)

The above equation is derived from the Blasius equation, and incorporates the changing values of the friction factor with flow rate changes:  n ΔPf 2 Q2 5 ΔPf 1 Q1

where nB1:8

(5.13)

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It is possible to calculate “n” for a specific pipeline and fluid, if given Q and Pf for two different flow rates using the following equation:

n5

log

ΔPf 2 ΔPf1





ln

ΔPf 2 ΔPf1



5

Q2 2 log Q ln Q1 Q1

(5.14)

The previous equations are using ΔPf and it is critical when analyzing pipeline to make sure that the ΔPf values being used are truly differential frictional pressure losses. The frictional pressure losses in a pipeline will be the gage pressure difference plus the elevational pressure difference as shown in Eq. (5.15):     ΔPf 5 Pðpump station gageÞ 2 Pðend gageÞ 1 ρg hðpump stationÞ 2 hðendÞ

(5.15)

In the previous equation, the change in gage pressures and the elevation change are calculated as the pump station value minus the pipeline end values. As an example, a pipeline moving diesel (specific gravity 5 0.85) that drops 300 feet over the pipeline distance with a discharge pressure of 1000 psi and an end pressure of 100 psi will have a frictional pressure drop equal to: 900 psi 1 110.6 psi 5 1010.6 psi.

5.3 Drag-reducing agent chemistry To function as a DRA, a single molecule must be extremely large and be able to associate with thousands of the liquid molecules that are being drag reduced. Therefore, all DRAs utilized in the oilfield have high-molecular-weight polymers as the active ingredient. The majority of these polymers are made from alpha-olefin monomers. The resultant PAO polymer has a flexible linear pure-hydrocarbon backbone with alkane side chains off every other carbon atom in the backbone. Fig. 5.9 depicts a typical DRA polymer. The side chains give the required association with the hydrocarbonbased liquid in the crude oil or fuel pipelines (no functional groups are necessary). The backbone must be extremely long to give the molecule its size and its ability to associate with a large number of liquid molecules. The polymer must have a molecular weight exceeding approximately 5 million to function effectively as a DRA. For some oilfield applications such as waterflooding or production of high-water content crude oils, DRAs which are water-soluble are utilized. These DRAs are typically polyacrylamides or acrylamide/acrylate copolymers. These polymers also contain a long flexible linear hydrocarbon backbone. The side chains, in this case, contain the hydrophilic structures that allow association with the water molecules. Again, the molecular weight must exceed several million to function effectively as a DRA.

5.4 Drag reduction mechanism The phenomenon of active drag reduction being discussed in this chapter only occurs in turbulent flow conditions. Therefore the mechanism by which drag

5.4 Drag reduction mechanism

FIGURE 5.9 Typical drag-reducing agent polymer for hydrocarbon liquids.

reduction occurs must be tied to the dynamics of turbulent flow. Turbulent flow is a very complex phenomenon consisting of random fluid motions such as streaks, bursts, and vortices on both small and large scales and with varying intensity. Because turbulent flow, itself, is not fully understood and cannot be defined rigorously, then understanding the exact mechanism by which drag reduction occurs becomes even more difficult. However, there have been many studies of drag reduction in which the data gathered show some effects of the drag-reducing polymer on the turbulent flow regime and suggest some aspects of the mechanism. In a turbulent fluid flowing within a pipe, three distinct regions exist near the wall and are defined by their velocity profile. The (very thin) viscous sublayer occurs at the wall and is characterized by a “near laminar” flow which contains a streak pattern of slow and high-speed fluid. These streaks run parallel to the streamflow. The turbulent core is a core of fully developed turbulent flow which occurs from near the pipe wall and extends to the center of the pipe (99 1 % of the total cross-section is the turbulent core). Between the turbulent core and the laminar sublayer exists a buffer zone. Within this buffer layer, turbulent production and dissipation occur. The turbulent production and dissipation are maximum at the edge next to the sublayer and extend to the other edge of the zone. Essentially all the turbulent production occurs during “bursting.” In the bursting phenomenon, a streak of low-speed fluid near the pipe wall is gradually pulled away from the wall in a kind of swelling process. This streak then becomes unstable due to the higher kinetic energy away from the wall and suddenly disintegrates or “erupts,” throwing the fluid in a variety of directions. This bursting process occurs on a fairly regular interval in the buffer zone and appears to be quasi-cyclic. The bursting process, and the resultant radial flow and vortices,

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account for a significant fraction of the energy losses associated with turbulent flow. In turbulent flow, the pressure/energy requirement relative to the flow rate is an exponential relationship (need nearly four times the energy input to achieve a doubling of the flow rate); whereas, in laminar flow, the relationship is linear. Turbulent flow is much less efficient than laminar flow for translating energy input to mass flow. When a DRA polymer is added to a flowing stream, the very high molecularweight polymer dissolves into the flowing fluid and, because the polymer is designed to be of similar makeup to the flowing fluid, the polymer associates well with the fluid. Because of its relatively huge size, a single DRA molecule can associate with thousands of fluid molecules. Now, fluid molecules which would have normally been randomly colliding and transferring energy, are part of a collective fluid domain associated with the polymer. Studies have shown that under drag reduction conditions, the viscous sublayer and the buffer layer are substantially increased in dimension. Other tests have shown that drag reduction doesn’t occur until an added polymer is present in an annulus whose coordinates correspond to the buffer layer [28]. With drag reduction, the quasi-cyclic or regular intervals of the bursting process are interrupted. The time and spatial distances between the turbulent bursts are increased (i.e., the frequencies are greatly reduced). This reduction in the turbulent bursting cycle yields a reduction in the radial energy transfer. The result is a decrease in the amount of energy lost to turbulence and, therefore, a reduction in the friction factor. Simply put, DRAs take an inherently inefficient system for translating energy input to fluid flow and make it much more efficient.

5.4.1 Misconceptions There are several misconceptions about DRAs and their mechanism or effects on the pipeline system. Partly because of the appearance of the DRA when it is in extremely high concentrations in solvent (which has a slimy appearance) there has long been the misconception that the DRA performs by coating the pipeline wall and creating a slickened surface. This is not true. To perform as a drag reducer, the DRA must be in solution homogeneously within the pipeline fluid. Therefore, the DRA dissolves into the fluid, moves with the treated fluid, and leaves nothing behind. Because they are added at much lower ppm levels to the pipeline, DRAs do not alter the viscosity of the treated pipeline fluid (they are not viscosity modifiers), and they do not change the density of the fluid. The wall coating misconception is also counteracted by the fact that most pipelines are made of materials such that the pipe wall is near to what would be called “hydraulically smooth.” Traditional engineering calculations for the fanning friction factor in most pipelines would utilize a wall roughness value (e/d) at or very close to the hydraulically smooth value. Therefore, calculations for the pressure drop requirements for that pipeline would already yield a value for pressure drop which is nearly equal to the value which would be achieved with a mirror-finish

5.5 Application to the pipeline—drag-reducing agent theory

wall. But, the use of DRAs can yield a pressure requirement that is more than 50% lower than the smooth-wall value. There is also the misconception that the use of DRAs will significantly alter the radial velocity profile of the pipeline taking it from the traditional turbulent profile to a near laminar profile. This is not the case for pipelines flowing in well-developed turbulent flow (beyond transition flow). Even with substantial drag reduction levels ( . 50% DR), the pipeline is still flowing with a turbulent velocity profile that is nearly equal to that of the nontreated flow. If this were not the case, then the use of DRAs would not be feasible in product pipelines, which require fuel batch segregation [29]. In today’s industry, DRAs are used widely in product pipelines. There is a misconception that the active DRA polymer can fall out of solution in the pipeline or the terminal tankage. Once dissolved, the DRA polymer will not come back out of solution. There is no feasible method on a commercial pipeline scale to drive the polymer out of solution and recover the polymer.

5.5 Application to the pipeline—drag-reducing agent theory Because of their high-performance capabilities and applicability to almost all types of liquid hydrocarbon, DRAs are used widely around the world in the oil pipeline industry. Many advances have occurred during the past four decades in terms of making the DRA product easier to handle and inject. Today, most DRA is injected from a self-contained injection skid which includes a supply tank and metered injection pumps. The skids are usually tied remotely into the pipeline control centers such that the DRA can be turned on and off automatically as needed. The injection rates are also controlled at set ppm ratios to the pipeline flow rate and can be adjusted as needed to give the most efficient use of the DRA product. As indicated in a previous section, when applying DRAs to pipelines, the metric for performance is % drag reduction (%DR) that is achieved in the pipeline. The %DR value is an indicator of the percentage reduction in the frictional pressure demand of the pipe flow that has occurred with the addition of the DRA. The equation for calculating %DR is shown again for reference. %DR 5

ΔPf

ðuntreatedÞ 2 ΔPf ðwith DRA Þ 3 100 ΔPf ðuntreatedÞ

(5.16)

Fig. 5.10 indicates the levels of %DR performance that can be achieved in today’s pipelines dependent upon the type of hydrocarbon being transported. Once injected into a pipeline, the DRA additive goes through its performance “life cycle.” This life cycle consists of three phases. The first phase is dissolution where the additive in its neat form (usually highly entangled polymer molecules)

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FIGURE 5.10 Performance potential for drag-reducing agents in oil pipelines [30]. Reproduced with permission from R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018.

must be wetted by the pipeline fluid and allowed to disentangle the polymer molecules. Once disentangled, each separate polymer molecule can then be elongated within the flowing fluid and begin creating the drag reduction effect. The DRA will be designed to achieve dissolution very quickly. Some of the DRA polymer molecules are released almost immediately upon wetting, and dissolution progresses with typical full dissolution times on the order of 1020 minutes. For a pipeline segment with a typical linefill time of 10 hours, the dissolution phase is a very small fraction of the linefill time. As such, the DRA can often appear to be performing almost instantaneously upon injection. The second phase in the life cycle is where the polymer is extended within the flow field. This is the fully active state for the DRA. Users would prefer this phase to exist over the complete pipeline segment. The efficiency performance of DRA during this phase pretty much determines the overall pipeline %DR. The third phase is the degraded polymer/inactive polymer state. Mechanical degradation of the polymer occurs when, on a molecular scale, one end of the polymer is accelerated faster than the other end by its association with accelerating fluid molecules. This stretching of the molecule creates forces within the backbone of the polymer and eventually breaks a carbon-carbon bond, yielding two smaller polymer molecules. Smaller molecules are less effective for %DR.

5.5 Application to the pipeline—drag-reducing agent theory

Once the original DRA molecule is broken several times, this renders the DRA molecules ineffective as a drag reducer. Wall shear forces and fluid acceleration will cause mechanical degradation. Fortunately, within a typical crude oil or products pipeline, wall shear rates are not extremely high. As such the degradation rate of the active DRA in a typical pipeline is not dramatic. It does occur, but it is usually not dramatic or overpowering. For most typical long-distance pipelines, the effective loss of the original fully active DRA molecules is on the order of 20%50%. In some instances, however, there are localized in-line shear events within a pipeline system. These localized events are generally caused by a sudden pipeline diameter decrease, a floating check valve, a partially closed mainline valve, or similar in-line devices. These devices can cause enhanced and sudden degradation of a significant portion of the active DRA polymer due to the acceleration and velocity increase that occurs through the short length of the device. Mainline centrifugal pipeline pumps at a pump station will impart extremely high shear rates to the pipeline fluid and will cause complete degradation of any active DRA polymer. That is why re-injection of new DRA is required for drag reduction to occur downstream of any operating pump station. Many may consider the degradation of the DRA polymer in the pipeline system as an undesirable aspect. However, eventual degradation of the polymer is a necessity considering the end use of the hydrocarbon streams being treated. Active DRA polymer not only imparts pressure reductions in turbulent flow, but it also imparts significant reductions in heat transfer effects in turbulent flow. Heat transfer is a very important integral part of the refining of crude oil. Active DRA also imparts resistance to passage through extensional flow fields such as filters, which are crucial to the hydrocarbon fuel industry. Fortunately, in the eventual passage of the hydrocarbon streams from terminal storage to final use or processing, any present DRA is sheared. In the degraded state, DRA polymers have been shown not to impact refinery operations processing treated crude oil and several tests have shown that there is no impact on engine performance for gasoline and diesel engines with the fuels containing fully sheardegraded polymer. The performance of DRAs in the fully active state (second phase) is most impacted by the concentration of the DRA additive. The typical performance curve for DRA relative to ppm concentration is shown in Fig. 5.11. Performance increases with concentration, with the %DR/ppm slope being the greatest at the lowest concentrations. As the ppm level increases, the slope decreases, and the curve asymptotically approaches a maximum limit for %DR. This limit will be different for any given application depending upon several factors such as the system turbulence (NRe) and the shear intensity of the system. P. S. Virk established a correlation for the MDR that can be achieved in conduit flow: this is based largely upon the NRe for the given flow and the viscosity of the fluid [14]. Virk’s correlation is shown in the following equation. Interestingly,

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FIGURE 5.11 Performance curve aspects for drag-reducing agent.

even though this correlation was developed in small diameter tubes, this correlation has appeared to hold up well for applications in large diameter pipelines.

f -1=2 5 19:0 Log10 NRe f 1=2 2 32:4

(5.17)

The typical concentration performance curve for normal pipeline applications will fit a linear inverse relationship as shown in the following equation and in Fig. 5.12. 1 A 5 1B %DR ppm

(5.18)

The helpfulness of this relationship is that one can conduct a performance test at two DRA concentrations (one high, one low) and generate a reliable performance curve over a large range of concentrations for the system being tested. The inverse of “B” will indicate the theoretical maximum achievable %DR in the system. Another advantage of this relationship is that one can conduct a test at more than two concentrations and determine whether the pipeline system behaves “normally” or whether there are deviations from “normal.” If the fitted curve for multiple points is not linear but shows curvature, this can indicate issues such as localized high-shear events or partial nondissolution of the DRA. The fluid property which most significantly affects DRA performance is the flowing oil viscosity. For higher viscosity fluids, the DRA efficiency will be decreased. The effect is strong. One can consider this to be a result of higher viscosity fluids generating lower NRe values or less turbulence. Lower turbulence means less opportunity for impact. A doubling of the oil viscosity in a pipeline (which could easily be the result of colder weather cooling the oil) will generally cause the DRA injection requirement to increase by more than 50% to achieve the same level of drag reduction. DRA performance is also affected by the pipeline’s flowing velocity. This effect is twofold. Increasing the velocity increases the turbulence (NRe) which should yield more DRA efficiency. However, increasing velocity also means

5.5 Application to the pipeline—drag-reducing agent theory

FIGURE 5.12 Inverse relationship for drag-reducing agent performance in a normal pipeline system.

higher shear rates within the pipeline. These higher shear rates will increase the rate of degradation of the DRA polymer and will, thus, cause a lowering of efficiency. The resultant performance curve is a result of the counterplay between these two factors. At lower velocities (typically less than about 8 ft/second) an increase in velocity will yield higher DR performance. At higher velocities (typically more than about 10 ft/second) an increase in velocity will lower DR performance. At the moderate velocities in-between these values, the effect of velocity is fairly flat and can be referred to as the plateau region for velocity effects. Fig. 5.13 shows the general effect. This effect is general and the actual plateau region for any given pipeline application is a function of several pipeline factors including pipeline diameter and DRA concentration. The effect of pipeline diameter on DRA performance is minimal. In general, the larger diameter pipelines are slightly less efficient in terms of DRA response. But the effect is small and is generally hard to evaluate considering the effects of all the other variables which do not remain constant when applying DRA to different size pipelines.

5.5.1 Applications in oil/water or multiphase pipelines Turbulent drag reduction using polymeric DRAs requires treatment of a continuous liquid phase. In a production pipeline moving both oil and water, to get significant drag reduction, the DRA added must be soluble in whichever component is the continuous phase. Adding a DRA to the dispersed phase will not yield any substantial turbulent drag reduction effect. For most production systems that are producing less than about 40% water fraction, the DRA of choice would still be oil-soluble. But, as the water percentage increases from about 5% to the 40% level, the overall effectiveness or efficiency of the DRA will decrease dramatically. The contribution of the water phase to the overall pressure drop in the

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CHAPTER 5 Polymeric drag reduction in pipelines

FIGURE 5.13 General velocity effects on drag-reducing agent performance in a pipeline.

pipeline is not altered by the oil-soluble DRA. And, as the water phase contribution becomes a larger and larger fraction of the overall pressure drop, the portion of pressure drop available for drag reduction becomes less and less. When the water fraction becomes large enough and approaches the point of phase-inversion (where water becomes the continuous phase), the drag reduction from the oilsoluble DRA will become negligible. For a production pipeline that is producing a very high percentage of water and it is not evident which phase (oil or water) is continuous, it would be best to separately test both an oil-soluble DRA and a water-soluble DRA. These tests will help to determine which phase is the continuous phase before injection and/or which phase will become continuous under the influence of the DRA. Significant levels of drag reduction can also be achieved in multiphase (oil, water, and gas) systems. These systems are much more complex. The total pressure drop in the pipeline not only includes the turbulent wall friction factor but can also include acceleration pressure losses (slugging and waves) and energy losses through interfacial mixing. DRA testing in multiphase systems has shown not only profound effects on the pressure but also dramatic changes in the existing flow regime or flow pattern. As such, the major benefit of using a DRA in a multiphase production line may not be the reduction in pressure requirement or the resultant flow increase but may be the smoothing out of the flow (reduction or elimination of slugging) coming into the production handling systems. The different types of flow regimes that can occur in two-phase flow in a horizontal pipeline are illustrated in Fig. 5.14. For oilfield production flowlines, the flow regimes most encountered are either stratified flow, slug flow, or annular flow. The liquid volume fraction plays an important role in determining the potential magnitude of drag reduction that might be achieved in multiphase flow. Therefore, in general, the potential for drag reduction in the various flow regimes depicted in Fig. 5.13 is highest at the top and decreases as one goes down the

5.6 Utilization of drag-reducing agent in pipeline operations

FIGURE 5.14 General Dr potential in multiphase flow regimes.

figure. Likewise, from a gas-oil ratio (GOR) standpoint, in general, as GOR increases, the potential for drag reduction decreases. However, keep in mind that these are only general trends. Other factors such as absolute liquid phase velocity can have an impact on the magnitude of drag reduction that might be achieved. A higher GOR can mean higher liquid phase velocities. The work of Fernandes [31], and Kang and Jepson [32,33] have shown that the addition of DRA to multiphase flow can significantly alter the flow regime and will likely result in a transition to another regime. Slug flow will change to stratified flow (wavy or smooth) or will result in slug flow with a much lower frequency of slugs. Annular-entrained flows have been shown to transition to stratified-wavy flow with or without entrainment. And, stratified wavy flows can easily change to stratified smooth flows. The presence of the DRA polymer in solution in the oil provides for more “cohesiveness” to the liquid phase and reduces the potential for high shear to alter or remove portions of the liquid phase and create oil droplets, entrain gas, or form waves.

5.6 Utilization of drag-reducing agent in pipeline operations DRAs provide flexibility in the operation of a pipeline [30]. Because DRAs can be applied over a range of dosage levels in the pipeline, they provide a range of options in terms of operating points at higher flow rates or lower pressures. DRAs can be used to substantially increase the flow rate in a capacity-limited pipeline or maintain the flow rate in a de-rated line. They can be used to bypass pump

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stations. They can be used to minimize capital and operating costs in new pipeline designs. DRAs can bring substantial value to pipeline operations through aversion of flow loss, and through minimizing capital and energy costs. DRAs can be applied to almost all areas of the oil industry. DRAs can be used to increase the production flow of crude oil to the wellhead, from the wellhead to gathering sites, and from offshore platforms to shore. DRAs have been used to increase loading rates from land tankage to offshore tankers. The largest volume used for DRAs is for treating long-distance land-based pipelines moving crude oil from production areas to refineries. These pipelines often involve many pump stations and may span a distance of more than one-thousand miles. Because of their high-performance capabilities, especially in fuel products, DRAs are often injected into product pipelines to allow the bypassing of an intermediate pump station or to reduce the number of operating pumps at a given station. This can yield a significant reduction in operating and energy costs. Fig. 5.15 illustrates the effect and utilization of the drag reduction effect within a pipeline system. In this figure, the hydraulic system curve for the pipeline fluid (untreated) is shown. This is a single curve with the head or pressure requirement increasing exponentially with the fluid flow rate. This is the total

FIGURE 5.15 Operating curves showing drag-reducing agent utilization within a pipeline system [30]. Reproduced with permission from R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018.

5.6 Utilization of drag-reducing agent in pipeline operations

pressure required to move the given fluid down the pipeline. For a given pump curve (hydraulic head available from the pump impeller at a given flow), there is only one operating point, where the two curves intersect. This single point represents the maximum flow rate that would be achievable in that given pipeline system. With variable pump speeds or throttling, the pipeline can operate only along the single system curve. The pipeline is limited to that curve. The system curves for the pipeline fluid with DRA treatment are also shown. Because the DRA can be injected at multiple ppm levels, there are multiple system curves available depending upon the treatment level. With DRA, there is a multitude of operating points available along the pump curve. Likewise, there is more than one maximum flowrate that might be achievable in the system. With variable speed pumps or throttling, there is a field of operating points for the pipeline. Depending upon the situation desired, the pipeline can be operated at a new operating point which represents almost total flow increase with very little decrease in pressure. Or, the pipeline can be operated at the original flow rate, with much less pressure head requirement. The availability of DRA to the oil industry provides for tremendous flexibility in the pipeline operating systems. The pipeline operator can tune the DRA injection level to achieve an operating pressure/flowrate operating point that fits his current needs. As these needs change, the operator can adjust the DRA injection levels to optimize his system. He can achieve the desired flow rates with an optimized power or energy requirement on his system. If an unexpected bottleneck occurs within a pipeline (such as a sudden derating of the pipe), the pipeline operator can use DRA to continue to operate at the normal flow rate with much lower pressure. If production barrels at a field site exceed the pipeline designed capacity to move those barrels, then DRA can be injected to allow the increased flow rate while still operating under the other design limits. If the operating costs or issues with an intermediate pump station become too high, then injection of DRA can be considered to shut down the intermediate pump station (allow it to be bypassed). DRAs also provide flexibility in the fact that they can be implemented within a short period. Injection skids designed for the considered DRA along with a supply of the DRA can be placed at a site quickly. This can also be done with virtually none or very little CAPEX requirement. Likewise, if the need for a DRA goes away, then the injection can be stopped. The DRA can be injected only as the need arises.

5.6.1 Example cases for utilization in pipelines Following are a few example cases to illustrate these opportunities in more detail.

5.6.1.1 Multi-station pipeline de-rated in the middle segment Consider a long-distance pipeline with four operating pump stations moving crude oil from production sites to a refinery. During normal operation, the pipeline

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FIGURE 5.16 Hydraulic profile for derated middle segment in a pipeline [30]. Reproduced with permission from R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018.

moves about 100,000 barrels per day of crude oil and operates at a pump station discharge pressure near the MAOP rating of the pipeline. An inspection of the line reveals that the second segment of the pipeline has to be de-rated due to some local corrosion issues. The deration is 60% of the original MAOP. With the deration, the second segment becomes the pipeline bottleneck. The pipeline can now only move about 75,000 bpd while operating at the reduced MAOP. All four segments are forced to operate at a lower rate than the second segment. However, a DRA can be implemented at the second pump station within a few days. With an injection of a few ppm of DRA, the pipeline can increase its flow rate back to the original 100,000 bpd rate and still operate at the reduced MAOP in the second segment. The injection can continue until the pipeline can replace the affected pipe. The hydraulic profile for this application is shown in Fig. 5.16. Pump stations A, C, and D continue to operate as normal with a discharge pressure near the original MAOP, and the normal pressure gradients at the high flow rate. Pump station B has a discharge pressure at 60% of MAOP and the pressure gradient for the second segment is less steep with the effect of the DRA. All segments continue to move crude oil at the 100,000 bpd rate.

5.6.1.2 The production rate exceeds the design capacity An offshore field produces about 350,000 bpd of crude oil. An offshore gathering platform and pipeline system are designed to move the produced oil to shore tankage. The subsea pipeline has an MAOP which allows for the current production flow while running with a discharge pressure at the MAOP. A new production zone is implemented which increases the field production capability to 450,000 bpd. The offshore pipeline is still limited to moving 350,000 bpd with no

5.6 Utilization of drag-reducing agent in pipeline operations

FIGURE 5.17 Hydraulic profile for production rate exceeding design capacity [30]. Reproduced with permission from R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018.

feasible capital option for increasing the throughput. The extra 100,000 bpd of oil cannot be produced without increasing the existing MAOP of the pipe by nearly 1.6x factor or by supplying some alternate means to ship the oil to shore. If a DRA injection is implemented onto the platform, however, the extra oil can be moved down the pipeline. Treating the crude oil with a few ppm of DRA can substantially lower the system pressure curve for the pipeline. The treated system can easily pump the 450,000 bpd rate down the pipeline with the platform discharge pressure staying below the MAOP. The hydraulic profile for this application is shown in Fig. 5.17. The pressure gradient with DRA at 450,000 is nearly the same as that without DRA at the lower rate of 350,000. If the field production capability were to increase even further, even this extra flow could be handled by injecting at a higher ppm level of DRA. A flow rate more than double the original maximum could likely be achieved with DRA, as long as the current MAOP of the pipeline could be achieved (or close to it) with the existing pumping system.

5.6.1.3 Pump station bypass An existing cross-country pipeline moves refined products (gasoline and diesel) to a market terminal. The pipeline currently consists of four operating pump stations equally spaced along its 400-mile course. Pumping energy costs and the operating costs of the pump stations continues to increase substantially. The pipeline operator decides to lower his operating costs by injecting DRA and

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FIGURE 5.18 Hydraulic profile for pump station bypass [30]. Reproduced with permission from R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018.

effectively bypassing half of his pump stations. DRA injection systems are placed at the 1st and 3rd pump stations. The flow through the 2nd and 4th pump stations are placed in the bypass mode and their pumps are shut down. By injecting approximately 10 ppm of DRA at the 1st and 3rd pump stations, the operator can maintain the normal flow rate of the pipeline. His pumping energy costs are cut in half. Plus, his other operating costs at the two idled pump stations are reduced. If the operator maintains injection of DRA for monthly periods, the energy demand charges for the bypassed stations can also be dramatically reduced or eliminated. The hydraulic profile for this application is shown in Fig. 5.18. The pressure gradient for the treated product flow is more than 50% less than that for the untreated case. This allows the gradient to cover the equivalent of two pipeline segments and eliminates the need for intermediate pumps.

5.6.1.4 New pipeline design As indicated earlier, nearly 40 years of application of DRA to pipelines, and the advancements that have occurred in product form and performance have proven its reliability to the industry. Today, injection systems can be installed that continuously inject with virtually no downtime and product supply can be maintained with certainty. With this high reliability, the designers of new pipeline systems can consider the use of DRA in their design. The use of DRA gives them much more design flexibility. If the anticipated flow capacity or need for the new pipeline is not certain; if the need is defined as a range of flow rates with levels of uncertainty, then the pipeline can be designed to meet a lower flow rate. This can be done with the knowledge that DRA can be used to obtain a higher flow rate than the design capacity if the need arises.

5.6 Utilization of drag-reducing agent in pipeline operations

Likewise, because the use of DRA at various injection levels gives the pipeline an unlimited array of operating points for a given system, the designer of a new pipeline has much more flexibility in choosing pipeline parameters such as pipe diameter, wall thickness, and segment length in his design options. Once a base design is established to meet the design basis, the designer may then choose designs with less capital investment/costs (less steel in the ground) and plan for DRA injection to allow the base design flow rates to still be met. The designer may opt for a thinner pipe wall that has a lower MAOP and lower pipe cost. The designer may opt for smaller pipe diameter. This would yield a slightly higher MAOP and lower pipe cost, but the overall capacity, without DRA, would be substantially lowered due to the higher pressure gradient. The pipeline designer may also consider designing for longer station-to-station segments which will result in fewer total operating pump stations. This would yield much lower capital costs for the pumping stations and yield long-term lower operating costs. The use of DRA would allow for the much lower pressure gradient needed in this design. Table 5.1 lists more specific parameters around pipeline design considerations as discussed in the previous paragraph. These values are approximate and are shown as an illustration of the options which DRA use allows in the design. The design calculations are based upon ANSI/ASME Standard B31.4 Code. The basis for this example is a 500-mile cross-country pipeline moving a 5-centistoke, 35
API crude oil. Laying of 24-inch standard wall pipe (seamless, Grade B pipe) with the calculated MAOP would allow for a maximum flow rate of 250,000 bpd with 5 operating pump stations (100-mile segments). A design considering a thinner wall (0.312-inch wall thickness) would put less steel cost into the ground, but have a design MAOP that is slightly lower. The lower MAOP would mean that the untreated capacity of the pipeline would be less at 218,000 bpd. However, a DRA injection of less than 2 ppm at each pump station would allow the capacity to reach the original 250,000 bpd. A similar design could consider an even thinner pipe wall (0.281-inch wall thickness) and would yield an untreated capacity of 206,000 bpd. In this case, a DRA injection of about 3 ppm at each pump station would allow the capacity to be back to the original level. Designs considering lower pipe diameter would also put less steel cost into the ground. These designs would have slightly higher MAOP, but much less untreated capacity due to the higher pressure gradient in the pipe. Putting a 22inch diameter pipe into the ground would lower the untreated capacity to about 206,000 bpd. However, including a DRA injection of approximately 4 ppm at each pump station would allow the flowing capacity to be back at the original level. Putting even smaller 20-inch diameter pipe into the ground would cause a lower untreated capacity of 166,000 bpd. This smaller pipe would require an injection level of about 10 ppm DRA at each pump station to attain the original capacity of 250,000 bpd. Lastly, a design that considers fewer pump stations and longer segment lengths between stations is shown in the table. Three pump stations, instead of five, would mean substantially fewer capital costs for the construction of stations.

255

Table 5.1 Example pipeline design options with drag-reducing agent [30]. 500-mile pipeline, 24-inch O.D., seamless grade B pipe 5 centistoke/35
API crude oil 250,000 BPD capacity Pipe diameter (inches) Base design Thinner wall 1 Thinner wall 2 Less diameter 1 Less diameter 2 Fewer pump stations

Wall thickness (inches)

Calculated maximum allowable operating pressure (psi)

# Pump stations

Segment length (miles)

Cumulative drag-reducing agent injection (ppm total)

24

0.375

788

5

100



24

0.312

656

5

100

9

24

0.281

591

5

100

15

22

0.375

859

5

100

20

20

0.375

945

5

100

50

24

0.375

788

3

167

22

Reproduced with permission from R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018.

Nomenclature

And it would mean substantially less operating costs during the life of the pipeline. The capacity of the pipeline with the longer segments between stations would only be 189,000 bpd for the untreated crude oil. However, with an injection of DRA at the three planned stations at a level between 7 and 8 ppm, the capacity of the pipeline would be 250,000 bpd.

5.7 Conclusion The discovery of a flowing fluid phenomenon by a couple of researchers over 70 years ago led to a wave of research into what is known today as drag reduction technology. This research showed the complexity of the drag reduction phenomenon, which occurs in turbulent flow and which is still not fully understood even today. But this research did reveal some basic understanding of the requirements for drag reduction and the potential for dramatic lowering of frictional pressure drop in conduits through polymeric interaction and reduction of the generation of turbulence. Subsequent entrepreneurial work in large-scale field-size pipelines and a subsequent loss of a pump station on the Trans Alaska Pipeline system led to the first commercial use of DRA in a hydrocarbon pipeline. This application fueled the development of even more effective DRAs and spurred worldwide use of DRA. As the use of DRAs increased, new variants of DRA were developed to meet just about any need in most crude oil or product pipelines. These applications have shown the potential for very high drag reduction in long-distance pipelines. The methods for measuring performance in a pipeline are straightforward based on recorded pressures and flow rate. The numerous applications have also revealed some general performance trends that result from additive concentration, flowing velocity, and oil properties. DRA can be used for a major flow increase in a bottlenecked pipeline. A pipeline operator can more than double the throughput capacity in many cases by using this technology. DRA can also be used for maintaining current flow rates in a derated pipeline, generating energy and operational savings by eliminating the need for intermediate pump stations, and for reducing capital expenditures in new pipeline construction. It is estimated that more than 80% of the pipelines in North America use this technology today. The use of DRA has become a well-known and readily used tool in the pipeline industry to effect increases in throughput capacity or to reduce energy requirements.

Nomenclature ANSI ASME bpd CAPEX

American National Standards Institute American Society of Mechanical Engineers Barrels per day Capital expenditure

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CDR DRA GOR MAOP MDR PAO ppm SPE % DR

Conoco drag reducer Drag reducing agent Gas-to-oil ratio (multiphase parameter) Maximum allowable operating pressure Maximum drag reduction (performance asymptote) Poly-alpha-olefin (polymer type) Parts per million Society of Petroleum Engineers Percentage drag reduction

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References

[17] E.D. Burger, W.R. Munk, H.A. Wahl, Flow increase in the trans Alaska pipeline using a polymeric drag reducing additive, in: SPE-9419-MS, SPE Annual Fall Technical and Exhibition, Dallas, TX, USA, September 2124, 1980. [18] W.R. Beaty, R.L. Johnston, R.L. Kramer, et al., Drag reducers increase flow in offshore pipelines without additional expansion, Oil & Gas Journal 82 (1984) 7174. [19] C.L. Muth, C.J. Stanberry, G.J. Husen, et al., Powerful new tool for pipeliners  flow improver, Pipeline & Gas Journal 83 (1985) 3840. [20] R.J. Schoneberger, K.S. Erickson, J.M. Curry, Drag reducer helps boost crude thruput during hydrotest, Pipe Line Industry (1992). June. [21] H.A. Wahl, W.R. Beaty, J.G. Dopper, et al., Drag reducer increases oil pipeline flow rates, in: Offshore South East Asia 82 Conference, Singapore, February 912, 1982. [22] W.R. Carradine, G.J. Hanna, G.F. Pace, et al., High-performance flow improver for products lines, Oil & Gas Journal 81 (1983) 9298. [23] W.R. Dreher, R.L. Johnston, K.W. Smith, Field use supports move to suspensionbased DRA, Oil & Gas Journal 104 (2006) 5560. [24] R. Johnston, J. Pierce, P. Lauzon, Heavy crude oil production increase through the use of a novel flow improver product  a case study, in: World Heavy Oil Congress, Edmonton, Alberta, Canada, March 1012, 2008. [25] A.K. Jain, Accurate explicit equation for friction factor, Journal of the Hydraulics Division 102 (1976) 674677. [26] P.K. Swamee, A.K. Jain, Explicit equations for pipe-flow problems, Journal of the Hydraulics Division 102 (1976) 657664. [27] S.W. Churchill, Friction-factor equation spans all fluid-flow regimes, Chemical Engineering 84 (1977) 9192. [28] W.D. McComb, L.H. Rabie, Local drag reduction due to injection of polymer solutions into turbulent flow in a pipe, AIChE Journal 28 (1982) 547557. [29] C.F.L. Goudy, C.L. Muth, Line tests show DRA’s don’t cross-contaminate, Oil & Gas Journal 87 (1989) 4142. [30] R.L. Johnston, Y.N. Lee, Maximizing pipeline flexibility with drag reducing agents, in: SPE-192878-MS, Abu Dhabi International Petroleum Exhibition & Conference, Abu Dhabi, UAE, November 1215, 2018. [31] R.L.J. Fernandes, B.M. Jutte, M.G. Rodriguez, DR in horizontal 2-phase flow, International Journal of Multiphase Flow 30 (2004) 10511069. [32] C. Kang, W.P. Jepson, Multiphase flow conditioning using drag-reducing agents, in: SPE-56569-MS, SPE Annual Technical Conference and Exhibition, Houston, TX, USA, October 36, 1999. [33] C. Kang, W.P. Jepson, Effect of drag-reducing agents in multiphase, oil/gas horizontal flow, in: SPE-58976-MS, SPE International Petroleum Conference and Exhibition, Villahermosa, Mexico, February 13, 2000.

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Natural gas storage by adsorption

6

Yuguo Wang and Rashid Othman Research & Development Center, Saudi Aramco, Dhahran, Saudi Arabia

Chapter Outline 6.1 Introduction .................................................................................................262 6.2 Fundamentals of adsorption ..........................................................................265 6.2.1 Definition ...................................................................................265 6.2.2 Adsorption forces ........................................................................265 6.2.3 Adsorption separation and storage mechanism ...............................266 6.2.4 Adsorption processes ...................................................................267 6.3 Industrial adsorbents ....................................................................................268 6.3.1 Adsorbent selection .....................................................................268 6.3.2 Silica gel ....................................................................................268 6.3.3 Activated alumina .......................................................................269 6.3.4 Zeolites ......................................................................................270 6.3.5 Activated carbons ........................................................................272 6.3.6 Potential novel industrial adsorbents .............................................275 6.3.7 Summary of natural gas storage adsorbents ...................................275 6.4 Case study: screening activated carbon for natural gas storage ......................276 6.4.1 Experimental ..............................................................................276 6.4.2 Method of determining the amount of methane adsorbed ................278 6.4.3 Experimental results ....................................................................280 6.4.4 Empirical modeling with adsorption potential theory .......................282 6.4.5 Isosteric heat of adsorption modeling ............................................284 6.5 Heat management modeling ..........................................................................286 6.5.1 Mathematical modeling ...............................................................286 6.5.2 Performance analysis through thermal simulation ..........................288 6.6 Summary ......................................................................................................294 Nomenclature ......................................................................................................294 References ..........................................................................................................295

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00010-7 © 2023 Elsevier Inc. All rights reserved.

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CHAPTER 6 Natural gas storage by adsorption

6.1 Introduction Natural gas is one of the cleanest burning fossil fuels; its combustion generates less CO2 when compared with the combustion of petroleum-derived fuels (e.g., gasoline, diesel) [1,2]. With the broad availability of natural gas, electricity generation using gas turbine technology has been on the increase during the past few decades. However, there is a fundamental issue with natural gas transportation networks. This issue manifests itself in the fact that there is a swing in demand for natural gas (sales gas) between day and night. This is called a diurnal demand swing [3]. The higher rates of consumption form a “peak period of demand” and the lower rate of consumption creates a “non-peak period of demand.” This swinging electrical and natural gas consumption, not only in daily use but also in seasonal differences, results in variability across the natural gas transportation and production system. Electricity generation facilities prefer constant, high-pressure natural gas as a feedstock. Pressure swings in natural gas feed can damage the electricity generation equipment, especially rotational equipment (e.g., gas turbines) due to sudden inappropriate feed-to-fuel ratios that cause equipment slowdowns while under load. The use of in-line compressors within the transportation system has been used to mitigate pressure swings in natural gas transportation systems. Compressed natural gas (CNG) booster compressors can be operated during peak demand hours to maintain gas pressure. Unfortunately, using compression equipment can substantially increase operating and maintenance costs. In-line compressors do not steadily operate: they start when system pressure is at a low threshold value and stop when system pressure is at a high threshold value. Moreover, compressors often experience an inadvertent breakdown, despite the best maintenance practices. Rotational equipment breakdowns are sometimes catastrophic, requiring weeks of downtime. The sudden loss of natural gas feed pressure due to a malfunction can result in immediate downtime for downstream electrical generators and long-term dissatisfaction from gas consumers. An alternative to in-line compressors is to adsorb natural gas on adsorbents inside vessels. Adsorbed natural gas (ANG) on microporous adsorbents, e.g., activated carbon (AC), offers its advantages. Fig. 6.1 shows that pressure being equal, the volume of gas stored inside a vessel filled with adsorbents can be three times the volume stored with an empty vessel of the same dimensions. Fig. 6.2 presents a schematic for an adsorption-based natural gas storage facility that is operated close to a natural gas power plant. In comparison with CNG, ANG offers the advantage of storing gas at lower pressure while still maintaining high storage capacities via the gas adsorption on microporous materials [424]. To maximize the performance of an ANG storage system, it is essential to rapidly remove adsorption heat to fully utilize its design storage capacity. Various thermal modeling studies on small-scale natural gas storage systems for mobile applications were carried out. These used small-size adsorber vessels. Gu¨tlein

6.1 Introduction

140

Storage Amount (V/V)

120

CNG ANG

100 80 60 40 20 0 0

5

10

15

20

25

30

35

40

45

50

Pressure (bar) FIGURE 6.1 Comparison between volume stored in adsorption-based natural gas storage vessels and empty vessels (red line) at the same pressure and temperature.

Higher pressure pipeline

Higher pressure pipeline

Lower pressure pipeline

Power Plant

Pressure Control Valve Filtraon and metering staon Delivery Valve Receiving Valve

Releasing valve

Gas Adsorpve Storage Vessel Compressor

Heang/Cooling Unit

FIGURE 6.2 A diagram for an adsorbed natural gas storage facility.

et al. [12] suggested the use of a gas cooling loop for a lab-scale adsorber and demonstrated the effectiveness of fast adsorbent cooling for the full utilization of adsorptive methane gas storage potential. Basumatary et al. [8] linked gas inlet temperature and rate of charging with the maximum bed temperature and time

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required to fill the cylinder, when they modeled, thermally, a natural gas storage system, employing AC. Sa´ez and Toledo [15] investigated the thermal effect resulting from adsorption heat on both charging and discharging performances. Rahman et al. [16] used a fin-tube-type heat exchanger for thermal management during the adsorption and desorption processes. Ridha et al. [17] studied the thermal behavior of an adsorptive natural gas vessel and found out that extreme thermal fluctuations can occur, resulting in a 26.9% reduction of the dynamic storage capacity. Their predictions agree with the results reported by Gu¨tlein et al., which indicated that limited heat conduction could lead to a temperature rise of as high as 110
C [12]. The aforementioned studies mainly focused on the charging process, we will now shift our attention to the discharging process. Zhou [18] demonstrated a slow depressurization process to prevent the vessel wall temperature from dropping below the ductile-brittle transition limit of the steel. Vasiliev et al. [19] studied different sources of energy (exhaust gases, waste engine cooling liquids, solar and other types of energy) to avoid the temperature drop in the storage tank due to the enthalpy of desorption and thus stimulate the gas desorption process. Ridha et al. [20] conducted experiments to investigate the thermal behavior of an ANG system under dynamic discharge conditions and found that the enthalpy of desorption could cause a 15.2% reduction in methane dynamic delivery capacity when compared with the isothermal delivery capacity. To minimize the dynamic delivery drop, Chang and Talu [21] installed a central perforated tube in a cylindrical vessel. This perforated tube enhanced heat transfer from the wall to the central region, and the performance loss was reduced to 12% from 22%. For both charging/adsorption and discharging/desorption processes for ANG vessels for natural gas vehicle application, Ybyraiymkul et al. [22,23] investigated the effect of thermal management on the storage capacity of the ANG vessel, through experimental and numerical study. In their simulation, a 3D numerical model was used to analyze the influence of thermal control to understand the heat transfer phenomena. Their results showed only a 0.1% of deviation between experimental and numerical results. They also found that by optimizing the heat exchanger’s design, the storage capacity of the ANG vessel increased 10% while heat exchanger volume decreased 3 times. This is a significant improvement in the design of ANG vessels for vehicle application due to its limited onboard space. More recently, Grande and Vistad [24] published their work in which a modeling approach is used to understand the different needs for designing a device for methane storage using porous adsorbents. They used 1D, 2D, and 3D models of the storage tank, and different model results were compared to experimental data of methane storage in a 150 mL tank filled with AC. They found that 2D and 3D models are more appropriate for designing devices for thermal management. It is noteworthy to mention here that the literature provides valuable insight into a heat management strategy for many ANG storage systems; however, this insight is all about small-scale storage vessels suitable for vehicle applications.

6.2 Fundamentals of adsorption

Later in this review, we will present the results of a study that focused on how to minimize the loss in adsorbent storage capacity during the charging process by studying heat removal in industrial-scale stationary ANG systems. The study involved using the adsorption isothermal model to calculate the isosteric heat of adsorption and integral heat of adsorption. Internal heat generation rate was then derived by averaging the integral heat of adsorption over the off-peak charging time (810 hours), and the temperature profile of the adsorber bed was simulated by changing adsorption vessel aspect ratio, bed thermal conductivity, and vessel cooling style (e.g., use of water jacket and central cooling tube). The results of this study have confirmed that ANG can be a viable technology for natural gas supply and demand peak shaving purposes [25]. The remainder of this chapter will focus first on the fundamental of adsorption, then on experimental methods to determine the adsorbent capacity (which is used for adsorbent selection), adsorbent characterization, and modeling of adsorption isotherms, integral heat calculation, and the different methods for large scale heat removal methods. This will be followed by focusing on adsorbent development for the improvement of ANG technology and its potential integration with renewable energy for efficient heat management purposes.

6.2 Fundamentals of adsorption 6.2.1 Definition Adsorption is the adhesion of atoms, ions, or molecules from one phase onto the surface of solid material in another phase by physical or chemical interaction. The atoms, ions, or molecules are called adsorbate and the solid material is called adsorbent. This term is used to differentiate from the term absorption which means that the gas molecules penetrate the mass of the absorbing solid. And nowadays, absorption generally means the penetration of a chemical species in one phase into the volume of another liquid or solid to form a homogeneous mixture at equilibrium.

6.2.2 Adsorption forces Adsorption is a physical or chemical phenomenon depending on whether or not a chemical bond is formed between the adsorbate and the adsorbent species. Typically, physical interaction forces include van der Waals forces which are correlated with adsorbate molecular polarizability, and electrostatic forces such as polarization forces, surface field-molecular dipole interactions, and surface field gradient-molecular quadruple interactions. For example, hydrocarbons adsorption onto ACs is mostly due to van der Waals forces while water adsorption onto molecular sieves is due to surface field-molecular dipole interaction, mainly. Some of the key physio-chemical properties of molecules common in natural gas

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Table 6.1 Physical-chemical properties of some molecular species present in natural gas. Molecule

σ (Å)

α (Å3)

μ (D)

Θ (DÅ)

Tc (K)

CH4 N2 CO2 H2O H2S He H2 CO

3.80 3.64 2.30 2.65 2.60 60 2.89 3.76

2.448 1.710 2.507 1.501 3.630 0.208 0.787 1.953

0.000 0.000 0.000 1.850 0.970 0.000 0.000 0.112

0.02 1.54 4.30 2.30 3.74 0.00 0.43 2.04

190 126 304 647 373 5 39 133

H2 and CO are listed for comparison purposes.

streams are listed in Table 6.1, where σ is their kinetic diameter (in Angstrom, ˚ ), α is their polarizability (in A ˚ 3), μ is their dipole moment (in Debye, D), Θ is A ˚ for quadrupole moment (in DA), and Tc is their critical temperature (in K).

6.2.3 Adsorption separation and storage mechanism Adsorption-based storage and separation involve the action of similar mechanisms. In storage, the mechanism involves the attainment of an equilibrium amount of adsorbate adsorbed on the adsorbent. Thus, storage involves equilibrium effects. In separation, three basic mechanisms are in action. These are based on steric, equilibrium, and kinetic effects [26]. For steric separation, only molecules with a kinetic diameter smaller than the adsorbent pore diameter can penetrate the pores and get adsorbed, while larger molecules are excluded from adsorption. For example, drying natural gas with a 3 A molecular sieve and the separation of normal paraffins from iso-paraffins and cyclic hydrocarbons using a 5 A molecular sieve are adsorption-based separation by steric effect mechanism. In contrast, ˚ molecular sieve with its kinetic diameH2O is adsorbed inside the pores of a 3 A ˚ ˚ ter less than 3 A, and other components whose kinetic diameter is larger than 3 A ˚ are excluded. Similarly, n-paraffins are adsorbed inside the pores of a 5 A molec˚ while other compoular sieve because their kinetic diameters are less than 5 A nents (e.g., branched paraffins) are not adsorbed because their kinetic diameters ˚. are larger than 5 A Kinetic separation is based on the different rates of diffusion of different species into the pores of a solid adsorbent. This mechanism is involved in the separation of nitrogen from air by a carbon molecular sieve. The difference in the kinetic diameters of nitrogen and oxygen molecules makes oxygen molecules diffuse at a relatively high rate into carbon molecular sieve than nitrogen molecules do; therefore, by controlling the residence time, the faster diffusing nitrogen molecules are preferentially adsorbed. Molecular Gate adsorbents are a good example of utilizing

6.2 Fundamentals of adsorption

kinetic separation since the difference in the diffusion rates of nitrogen, methane, and CO2 molecules are great enough to effect separation [27]. Equilibrium separation is based on the solid adsorbents having different abilities to accommodate different species; the stronger adsorbing species become preferentially adsorbed and thereby get easily separated. Most of the adsorption separation processes are equilibrium separation. For natural gas dehydration that involves the use of silica desiccants or molecular sieves, equilibrium-based separation is involved. Likewise, the use of AC for the removal of aromatics from acid gas in sulfur recovery units involves equilibrium-based separation. Another example of an equilibrium-based separation is refining hydrogen purification. This process occurs because hydrogen is adsorbed much less than almost any other component.

6.2.4 Adsorption processes Based on the method of adsorbent regeneration, an adsorptive process can be called either temperature swing adsorption (TSA) or pressure swing adsorption (PSA). In TSA, the adsorbent is regenerated primarily by heating while the adsorption and regeneration pressure does not change much. In natural gas processing, the bulk water content is removed by liquid glycol absorption down to the level of about 10 parts in million by volume (ppmv); however, in a liquefied natural gas (LNG) plant or if there is a downstream cryogenic nitrogen rejection or helium recovery unit, water removal by adsorption on 4 A molecular sieve is used to further dehydrate natural gas to a lower level of water content, namely 0.1 ppmv, to prevent the formation of gas hydrate which may cause blockage in the piping or valves. TSA is almost exclusively used for purification purposes. Regeneration in PSA processes is accomplished by lowering the pressure to release adsorbate molecules from the adsorbent surface. In rapid cycles of minutes or seconds, the throughput is high and in the inert purge cycle, the adsorbent is regenerated by passing a nonadsorbing or very weakly adsorbing gas through the adsorber. In addition, based on the purpose of the TSA or PSA process, it can further be classified as either bulk or purification process. This is based on the concentration of the strongly adsorbed component in the feed gas mixture. According to Keller [28], the process is called bulk separation when 10 wt.% or more of the mixture is adsorbed. When less than 10 wt.% of the strongly adsorbed component is adsorbed, then the process is considered a purification process. Bulk separation processes include those used for the separation of normal paraffins from isoparaffins and aromatics by zeolites, nitrogen and oxygen separation by zeolites or carbon molecular sieves, and separation of water and ethanol by 3 A zeolites. Industrial purification processes include cracked gas purification by silica, alumina, or 3 A zeolite, dehydration of natural gas by 4 A molecular sieves, and indoor air pollutants (volatile organics) removal by AC, silicate, or resins. Bulk separations include xylenes separation by zeolites and chromatographic analytical

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separations which utilize a wide range of inorganic, polymer, and affinity agents as adsorbents. Purification separation processes include inorganics removal from water (As, Cd, Cr, Cu, Se, Pb, F, Cl, radionuclides) by AC and decolorizing petroleum fractions, syrups, and vegetable oils by AC [29].

6.3 Industrial adsorbents 6.3.1 Adsorbent selection The heart of an adsorption process is the porous solid adsorbent, which provides the surface area and micropore volume, and, therefore, the adsorptive sites and space necessary for achieving high adsorptive (separation/storage) capacity [30]. The performance of the adsorbents, no matter what the separation/storage mechanisms in action may be, can be optimized by proper engineering design. This could involve proper selection of adsorption residence time, temperature, and pressure. An adsorbent with good equilibrium capacity but a slow adsorption rate means that the adsorbate will take a long time to reach the interior pores, channels, and internal surface. This in turn will require longer gas residence time in the adsorbing/storage vessel and hence a very low throughput. On the other hand, adsorbents with fast kinetics but low capacity are not good either because a large amount of adsorbent is required for a given throughput. Taking the above considerations into account, an adsorbent may be considered properly designed when it provides high adsorptive capacity and fast kinetics. To meet these two requirements, the adsorbent must have the proper combination of high surface area, high pore volume, and proper pore network for the transport of molecules to the internal surface area, adsorption sites, and pores. To meet these requirements, an adsorbent has to have the right combination of both micropore, mesopore, and macropores. The classification of the pore size which is recommended by the International Union of Pure and Applied Chemistry (IUPAC) are micropores (diameter , 2 nm), mesopores (2 , diameter , 50 nm), and macropores (diameter . 50 nm). The classification is based on nitrogen adsorption at its normal boiling point of a wide range of porous adsorbents [31]. The remainder of this section presents the most common commercially available adsorbents as well as novel adsorbents as they relate to storage applications.

6.3.2 Silica gel Silica gel is the jellylike precipitate that forms during the commercial production of silica gel. The production involves mixing a sodium silicate solution with a mineral acid, which results in the formation of finely dispersed SiO2nH2O particles, widely named silica hydrosol or silicic acid. The precipitate forms when the hydrosol is left standing and it is the result of a polymerization process. Water molecules, which are about 5 wt.%, are present mainly in the form of chemically

6.3 Industrial adsorbents

Table 6.2 Summary of some typical properties of adsorbent-grade silica gels. Surface area (m2/g) Density (kg/m3) Reactivation temperature (
C) Pore volume (% of total) Pore size (nm) Pore volume (cm3/g)

830 720 130280 5055 140 0.42

Adsorption properties H2O capacity at 4.6 mmHg, 25
C H2O capacity at 17.5 mmHg, 25
C O2 capacity at 100 mmHg, 2183
C CO2 capacity at 250 mmHg, 25
C n-C4 capacity at 250 mmHg, 25
C

11 35 22 3 11

bound hydroxyl groups. As a final product, silica gels come in the form of hard, and milky white glassy beads. They were first developed in World War I for use in gas masks. Nowadays, they are used in the medical and electronics industries for dehydration. The presence of polar surface functional groups such as SiOH and SiOSi make this product unique because these functional groups can also adsorb other polar molecules such as alcohols, phenols, amines, etc. through forming hydrogen bonds. They can also be used for the separation of aromatics from paraffins and the chromatographic separation of organic molecules [31]. In natural gas processing, silica gel is used as a desiccant for gas dehydration. At low temperatures, the ultimate capacity of silica gel for water is higher than the capacity of alumina or zeolites. But at low water content dehydration for downstream cryogenic nitrogen rejection or helium extraction units in gas plants or LNG plants, zeolites have to be used due to their higher water removal capacity at low relative humidity levels. Typical properties of silica gel of adsorbent grade are listed in Table 6.2.

6.3.3 Activated alumina Activated alumina is a porous and high surface area form of aluminum oxide, which is commercially prepared by the dehydration of aluminum oxide tri- or mono-hydrate (Al2O33H2O or Al2O3H2O) and recrystallization at elevated temperature. Because of the amphoteric nature of aluminum oxide, the surface of activated alumina is both acidic and basic and their hydroxyl groups are more strongly polar than those of silica gel. Due to this nature, activated alumina displays a higher capacity for water adsorption than silica gel at elevated temperatures and it is used as the preferred adsorbent for warm gases dehydration. Activated alumina demonstrates higher water capacity at high temperatures for gas dehydration and their regeneration requires a higher temperature. For example, the regeneration of silica

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Table 6.3 Summary of some typical properties of adsorbent γ-alumina. True density (kg/m3) Particle density (kg/m3) Total porosity Macropore porosity Micropore porosity Macropore volume (cm3/g) Micropore volume (cm3/g) Specific surface area (m2/g) Mena macropore radius (nm) Mean micropore radius (nm)

29003300 6501000 0.70.77 0.150.35 0.40.5 0.40.55 0.50.6 200300 100300 1.83

gel is by heating to 150
C while that of activated alumina and zeolitic materials requires a heating temperature up to 350
C. Among a variety of alumina commercially available, the most used one is γ-alumina. Some of the physical properties of a typical γ-alumina are listed in Table 6.3.

6.3.4 Zeolites Zeolites are microporous materials that are chemically composed of assemblies of SiO4 and AlO4 tetrahedra joined together through the sharing of oxygen atoms. The empirical formula of a zeolite framework is Mn/2Al2O3xSiO2yH2O where x is greater than or equal to 2, n is the metal cation valency and y is the number of water molecules inside the pores or cages of the zeolite structure. The intersection of pores (channels) forms the cages of the zeolites. The structure is electrically neutral since the negative charge introduced by each aluminum atom on the structure is neutralized by the n/2 metal ions (Mn1). common metal ions used in the synthesis of zeolites are sodium (Na1), calcium (Ca21), and potassium (K1). In crystal form, zeolites are distinct by the nature of the distribution of their pores; the crystals have the uniform channel and cage sizes. Since the porosity of zeolites is high, most of the adsorption takes place internally and the external surface area available for adsorption is insignificant. Zeolites’ pore size is determined by the number of atoms that form the pore mouth (also called apertures, or windows) leading to the internal cages. The pore mouth could be constructed of rings of 6, 8, 10, or 12 oxygen atoms together with the same number of aluminum and/or silicon atoms. Fig. 6.3 shows the 3-d computer model of the structure of zeolite ZSM-5, in which there are 10 oxygen atoms in the ring of the pore mouth. Zeolites have pore rings containing 8, 10, and 12 oxygen atoms having a limiting pore size diameter of 0.42, 0.57, and 0.74 nm, respectively. These pore size diameters give the steric selectivity for adsorption. Only molecules that have a smaller kinetic diameter than these pore diameters can penetrate inside the zeolite crystal and get adsorbed. But molecules

6.3 Industrial adsorbents

FIGURE 6.3 The microporous molecular structure of a zeolite, ZSM-5.

with kinetic diameters slightly larger than the pore diameter of a zeolite can also gain access to the inside adsorption sites of a zeolite because of the vibration of molecules and the crystal lattice. The accessibility of the internal cages and adsorption sites for adsorbate molecules can be changed by changing the metal cations or their location in the zeolite structure. Changing the position and/or type of the cation changes the channel size and properties of the zeolite, including its steric selectivity or molecular sieving effect. The positions of metal cations in the framework of a zeolite structure depend on the number of cations per unit cell; in other words, it depends on the valency number of the metal cations. For example, ion-exchanging Ca21 into Na 1 increases the number of cations per unit cell, and ion-exchanging Ca21 into K1 not only increases the number of cations per unit cell but also increases the cation size and reduces the effective channel diameter. In the case of type-A zeolites, the channel size will be reduced from 0.42 nm down to 0.3 nm when Ca21 is ion-exchanged into K 1 . In addition to the changes to the cationic property of the zeolites, the Si/Al ratio can be varied to change the acidity of the zeolites. The Si/Al ratio can be changed from 1 to more than 1000. In short, zeolites’ adsorptive properties can be tailored by changing their cationic nature, Si/Al ratio, and framework structure, and this change is usually done to achieve the selectivity required for a given separation/storage purpose.

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The ionic nature of most zeolites makes them have high hydrophilic and high affinity for polar molecules such as water, CO2, H2S, and mercaptans. Type A zeolites have been used commercially as adsorbents for water and CO2 removal from natural gas streams; type X zeolites have been used for niche mercaptans removal from specific natural gas streams. With the increase of the Si/Al ratio in zeolites, the material can become hydrophobic. Silicalite, a pentasil zeolite, which has no aluminum in its framework structure, is used to remove hydrocarbons from aqueous systems and humid gases. Table 6.4 shows the different applications of zeolite adsorbents. Since 1756, more than 200 unique zeolite frameworks have been identified, and over 40 naturally occurring zeolite frameworks have been characterized. As synthesized, zeolite crystals are quite small with a typical size of 110 μm. To make such powdery material a useful adsorbent, the crystals are formed into macroporous pellets, usually in the size range of 1.82.5 mm or 2.55.0 mm. This is achieved by using a binder, which usually is a macroporous material that provides transport channels for the adsorbate molecules to access the microporous zeolite crystals; they also provide the mechanical strength for the zeolites. Binder commonly takes 10%20% of commercial zeolite weight. Extrusion processes are used to form cylindrical pellets, while pellet forming, granulation processes, and/or rolling are used to form spherical particles.

6.3.5 Activated carbons AC is one of the most versatile solid adsorbents in the manufacturing industry. It has applications ranging from being used as an adsorbent in liquid and gas separation processes to being a catalyst for catalytic processes [32]. The processes, for example, involve adsorption of iodine, acetic acid, inorganic solutes, and organic solutes from aqueous solutions, gas phase H2S removal, simultaneous dry removal of SO2 and nitrogen oxides (NOx), air separation into nitrogen and oxygen, methanecarbon dioxide mixtures separation, and methane storage. Water treatment by AC to remove organic pollutants such as chlorinated compounds and other volatile organic compounds, metallic ions takes about 55% of the AC market in the United States. The versatility of AC comes from the extremely high surface area and unique pore size distribution of the material. The bimodal or trimodal pore size distribution provides fast transport channels for adsorbate molecules to access the micropores of AC where adsorption occurs. The structure of the AC is composed of an amorphous structure and a graphite-like microcrystalline structure. The arrangement of carbon atoms in the graphitic structure is similar to that of pure graphite. The graphitic structure provides the adsorption space in the form of slit pores compared to cylindrical pores in silica gel, alumina, and zeolites. The size of AC micropores is therefore defined as the half-width rather than radius. The graphitic structure also has the same layer distance of 0.335 nm as that of pure graphite crystals with a deviation in the range of 0.340.35 nm. The graphitic carbon

Table 6.4 Summary of chemical compositions and some applications of zeolite adsorbents. Framework

Cationic form

The formula for typical unit cell

Pore size (number of oxygen atoms)

Effective channel diameter (nm)

A

Na

Na12[(AlO2)12(SiO2)12]

8

0.38

Ca

Ca5Na2[(AlO2)12(SiO2)12]

8

0.44

X

K Na Ca

K12[(AlO2)12(SiO2)12] Na86[(AlO2)86(SiO2)106] Ca40Na8[(AlO2)86(SiO2)106]

8 12 12

0.29 0.84 0.80

Y Mordenite

Sr, Ba Na K Ag

Sr21Ba22[(AlO2)86(SiO2)106] Na56[(AlO2)56(SiO2)136] K56[(AlO2)56(SiO2)136] Ag8[(AlO2)8(SiO2)40]

12 12 12 12

0.80 0.80 0.80 0.70

Silicalite

H 

H8[(AlO2)8(SiO2)40] (SiO2)96

12 10

60

ZSM-5

Na

Na3[(AlO2)3(SiO2)93]

10

0.60

Application Desiccant, CO2 removal from natural gas Linear paraffin and air separation Drying of cracked gas PSA H2 purification Mercaptans removal from natural gas Xylene separation Xylene separation Xylene separation I and Kr removal from nuclear off-gases Removal of organics from water Xylene separation

274

CHAPTER 6 Natural gas storage by adsorption

crystallite usually is composed of about 67 layers, with the average diameter of each unit being about 10 nm. The linkage between graphite units is possible with strong cross-linking and the interspace among the graphitic units forms the pore network of the AC mesopores and macropores. ACs can be made by thermal decomposition of various carbonaceous materials followed by an activation process. The precursors for making ACs include woods, rice hulls, refinery residuals, coal, peat, lignin, pitches, nutshells such as coconut and date seeds, corn cobs, and polymers like polyvinyl chloride, resins, and polyimides. Two types of traditional manufacturing processes can be used to make ACs, gas activation and chemical activation. In the gas activation process, heating in absence of oxygen at 600
C900
C to drive off the volatile materials is applied first, followed by activation with steam at temperatures above 250
C, usually between 600
C and 1200
C. Other gases such as carbon dioxide or flue gases can also be used for activation. Chemical activation process uses zinc chloride, calcium chloride, phosphoric acid, potassium hydroxide or sodium hydroxide to impregnate the raw material, followed by carbonization at 450
C to 900
C. Template directed chemical vapor deposition (TDCVD) is a novel method in synthesis of customized ACs. The ACs synthesized has very large pore volume and surface area, this type of material is customized for natural gas storage. Tables 6.5 and 6.6 list the physical properties of typical ACs and the customized ACs synthesized by TDCVD, respectively. Sample CaX-973P5 shows extremely high micropore volume and this property makes it a suitable material towards maximizing methane storage capacity [33,34]. For methane storage by ACs, methane exists in a pseudo-liquid adsorbed Table 6.5 Summary of some physical properties of typical ACs. True density 3

2.2 g/cm Mean micropore half width 12 nm

Particle density 3

0.73 g/cm BET surface area (m2/g) 1200

Total porosity

Macroporosity

Microporosity

0.71 Micropore volume (cm3/g) 0.44

0.31 Macropore volume (cm3/g) 0.47

0.4 Total pore volume (cm3/g) 0.91

BET, Brunauer, Emmett, and Teller.

Table 6.6 Summary of some physical properties of customized ACs.

Sample name

BET surface area (m2/g)

Micropore volume (cm3/g)

Mesopore volume (cm3/g)

Total pore volume (cm3/g)

CaX-973P5 CaX-973E6 CaX-1073E6

1900 1540 1792

0.83 0.56 0.58

0.42 0.49 0.85

1.25 1.15 1.43

BET, Brunauer, Emmett, and Teller.

6.3 Industrial adsorbents

state in the micropores of ACs and in compressed gas state in the mesopores and macropores of ACs. These mesopores and macropores have more important role in providing transport channels to and from the micropores, which is necessary for fast adsorption and desorption. Numerical simulations have shown that the maximum density of the adsorbed phase is attained within micropores with slit shapes having width of 1.121.14 nm [3537].

6.3.6 Potential novel industrial adsorbents Driven by the stricter environmental regulations for clean fuels, adsorptive methane and hydrogen storage are new areas of exploration for researchers worldwide both in academia and industries. As a result, many novel adsorbents have been synthesized in research labs and tested under pure gas conditions. Among them, are metalorganic frameworks (MOFs). MOFs are polymeric crystalline compounds consisting of metal ions or clusters coordinated with organic molecules to form one-, two-, or three-dimensional network structures. Exceptional methane uptake at room temperature and 35 bars up to 220 cm3/cm3 have been reported for MOF PCN-14 [38], which is the highest methane storage capacity in literature. Because of their exceptional uptake capacities, MOFs have been heralded as the holy grail materials for gaseous molecules storage applications be they green or blue storage applications. However, it is worth pointing out here that these volume-to-volume-based adsorption capacity numbers are based on the ideal densities of MOF crystals. These reported volume-to-volume methane storage capacities only represent the material performance of ideal, single crystal samples [39,40]. In industrial applications, these crystals have to be formed into different shapes using binders, which may cause the blocking of channels and the creation of dead-end pores that reduces the storage capacity. To commercialize MOFs in industrial applications, more research and development work is needed to improve their bulk density, hydrothermal stability, mechanical strength, and susceptibility to chemical attack by impurities in industrial gases such as H2S and water molecules. For methane storage, compared with ACs, MOFs have an ordered crystal structure, which makes them easier to understand the adsorption binding mechanism. In return, this will provide key guidance for further, rational development of new MOF materials.

6.3.7 Summary of natural gas storage adsorbents In natural gas storage applications, the most important parameter for the widespread use of commercial adsorbents is the selection, or development, of commercial microporous materials, that have high storage capacity and are stable under cyclic operation. Potential microporous materials, which can meet these requirements, include ACs, MOFs, and other organic solids [38,4148]. Although MOFs and the other organic solids display attractive sorption properties, with a V/V adsorption capacity as high as 230 of absolute methane adsorption at 290K

275

276

CHAPTER 6 Natural gas storage by adsorption

and 35 bar, which also exceeds the US Department of Energy target of 180 V/V, their deliverable amounts are not well established. The other factors which may affect the applicability of MOFs as widely used adsorbents for natural gas storage include their stability and ability to tolerate common natural gas impurities such as H2S, black powder, mercaptans, etc. ACs currently remain the only commercially viable adsorbent for natural gas storage in terms of adsorption capacity and stability as well as cost.

6.4 Case study: screening activated carbon for natural gas storage 6.4.1 Experimental In screening adsorbents for natural gas storage, the volumetric method was usually used. Our examples are about measuring methane adsorption on granular ACs. The major components of the experimental set-up are shown in Fig. 6.4. A vacuum pump, not shown in Fig. 6.4, is also connected to the system for degassing. The adsorber has a thermal jacket connected to a heater/chiller so that the temperature of the adsorber can be set to the desired temperature for measurement. To measure the adsorption isotherm, the adsorber was first filled up with 50 g of granular AC and was degassed at a pressure of 2.5 3 1024 Torr at 120
C for 4 hours. Then, the adsorber was cooled down and maintained at desired temperatures of 10
C, 21
C, 38
C, or 56
C. The reference vessel has a volume of two

Methane line

Pressure Indicator 1

Pressure Indicator 2

Gate Valve 1 Vent Gate Valve 3

Gate Valve 4

Helium Line Gate Valve 2

Reference Vessel

Adsorber filled with adsorbent

FIGURE 6.4 Schematic diagram of experimental set-up for activated carbon screening.

6.4 Case study: screening activated carbon for natural gas storage

Table 6.7 Physical properties of the granular activated carbon samples. Sample

AC1

AC2

AC3

AC4

AC5

ASTM mesh size Bulk density (g/cm3) Skeletal density (g/cm3)

8 3 16 0.47 2.299

30 3 70 0.39 2.363

2 3 60 0.49 2.402

12 3 40 0.54 2.059

6 3 60 0.50 2.286

0.417 968.6 1235 0.629 18.00

0.412 969.4 1589 0.747 18.70

0.487 1186.1 1426 0.599 17.47

0.292 722.5 999 0.500 20.64

0.452 1082.1 1510 0.682 26.08

Nitrogen porosimetry (77K) t-plot micropore volume (cm3/g) t-plot micropore area (m2/g) BET surface area (m2/g) Total pore volume (cm3/g) BJH average pore width (Å)

Results of DubininAstakhov modeling of methane adsorption data at 10
C, 21
C, 38
C, and 56
C

3

Micropore volume (cm /g) Micropore average pore width (Å)

AC1

AC2

AC3

AC4

AC5

0.380 9.44

0.449 10.62

0.416 9.88

0.2865 10.34

0.445 9.58

BET, Brunauer, Emmett, and Teller; ASTM, American Society for Testing and Materials; BJH, BarrettJoyner-Halenda.

liters and the adsorber has a volume of 120 cm3. The reference vessel was first charged with methane to a certain pressure, then, valve 3 was opened and the pressure between the two vessels was left to equalize. Readings of the two pressure indicators were taken 20 minutes after reaching equilibrium. Five types of commercial granular ACs with different physical characteristics were used in the experiments, they are labeled as AC1, AC2, AC3, AC4, and AC5. The high purity methane and ultra-high purity helium were used without further purification. Nitrogen adsorption at 77K data is presented in Table 6.7. Nitrogen adsorption at 77K was used to determine the surface area and pore volume of AC samples. Fig. 6.5 shows the amount of nitrogen adsorbed versus the relative pressure for the five samples. The five isotherms are of Type-I, which indicates that these ACs are essentially microporous. After the relative pressure reaches 0.8, there is a slight increase in nitrogen adsorption, which is most probably due to the presence of mesopores and the fact that the relative pressure is high for condensation of nitrogen to occur. The correlation between the amount of N2 adsorbed at the relative pressure of 0.8 with Brunauer, Emmett, and Teller (BET) surface area showed a linear relationship, indicating that the amount of N2 adsorbed before condensation occurs is proportional to the BET surface area of the material [49].

277

CHAPTER 6 Natural gas storage by adsorption

22 20

Adsorbed Amount (mmole/g)

278

18 16 14 12 10

AC1

AC2

AC4

AC5

AC3

8 0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

Relative Pressure (P/P0)

FIGURE 6.5 N2 adsorption isotherms at 77K for granular activated carbons.

6.4.2 Method of determining the amount of methane adsorbed After the adsorber is filled with granular AC, it is tapped gently until the level of the granular AC stabilizes by not going down any further. At this point, the granular AC is packed to its bulk density. The total volume of the adsorber cell consists of the following volumes: adsorbent skeletal, micropore, mesopore, and macropore volumes and the void space among the adsorbent particles. In the precise determination of the total volume of micropores, mesopores, macropores of the adsorbent, and the void space among adsorbent particles, helium was used since it is essentially inert as far as adsorption is concerned at our experimental conditions. To determine the adsorption amount, pressure P1 was recorded for the reference cell and thermodynamic equilibrium pressure P2 was also recorded after gate valve 3 was opened. The total volume of the reference cell and the non-skeletal volume of the adsorber is V2. Knowing the individual volumes of the reference cell and absorber, Vr and Va respectively, the total non-skeletal volume of the adsorber is determined from Eq. (6.1): Vvoid 5 V2 2 Vr

(6.1)

Assuming that the methane adsorption only takes place in micropores, methane molecules inside the mesopores and macropores behave like they are in the gas state. The gas phase methane volume, therefore, is (V2 2 Vr 2 Vmic), where Vmic is the micropore volume of the adsorbent. To calculate the amount of methane

6.4 Case study: screening activated carbon for natural gas storage

adsorbed, non-ideal behavior is considered and the below equation is used: nadsorbed 5

P1 Vr P2 3 ðV2 2 Vr 2 Vmic Þ 2 ZRT ZRT

(6.2)

where Z is the methane compressibility factor, T is temperature and R is the ideal gas constant and Vr is the volume of the reference cell. The virial expansion for the first two terms is one way to determine the compressibility factor as in Eq. (6.3): Z 511

BP 511 RT

   BPc Pr RTc Tr

(6.3)

in which BPc 5 B0 1 ωB1 5 B0 1 RTc

0:0104B1

(6.4)

where B0 5 0:083 2

0:422 Tr1:6

(6.5)

B1 5 0:139 2

0:172 Tr4:2

(6.6)

Z can also be determined from the SoaveRedlichKwong (SRK) equation of state. Z and fugacity (f) are also related to each other in the SRK equation of states as follows: ln

  f A B 5 Z 2 1 2 lnðZ 2 B Þ 2  ln 1 1 P B Z

(6.7)

where [ð 5 f =pÞ is the fugacity coefficient, A and B can be determined from three gas parameters: critical pressure Pc, critical temperature Tc, and the acentric factor ω, as in the following equations: A 5 0:4278 3 α 3  B 5 0:0867 3

P=Pc 2

(6.8)

T=Tc

P=Pc T=Tc

(6.9)

  α0:5 5 1 1 m 3 1 2 Tr0:5

(6.10)

m 5 0:480 1 1:574 3 ω 2 0:176 3 ω2

(6.11)

Since the critical parameters of methane are Tc 5 190.6K, Pc 5 4.596 MPa, and ω 5 0.0115, the values of A , B , α, and m can be determined from Eqs. (6.8) to (6.11). Z is obtained by solving Eq. (6.12):   Z 3 2 Z 2 1 Z 3 A 2 B2 2 A 3 B 5 0

(6.12)

279

CHAPTER 6 Natural gas storage by adsorption

1.00 0.95

Compressibility Factor Value

280

0.90 0.85 0.80 0.75

Virial expansion calculation

0.70

SRK EOS calculation

0.65 0.60 0

50

100 150 Pressure (bar)

200

250

FIGURE 6.6 Compressibility values calculated from SRK and Virial EoS. EoS, equation of state; SRK, SoaveRedlichKwong.

Fig. 6.6 compares the Z values calculated from the two methods, the method associated with the SRK equation of state (SRK Z) and the one calculated from the method associated with the virial equation (Virial Z), at a temperature of 56
C. Although Fig. 6.6 displays a comparison of Z values at 56
C, Z values at other temperatures, 10
C, 21
C, and 38
C, are also calculated with both methods. The trend and the difference between the two methods with pressure change are the same for all temperatures. The truncation of higher terms in the virial expansion causes a higher error in the calculation of Z at higher pressure. Z values calculated from the SRK equation of state are, therefore, used in the determination of the amount of methane adsorbed. In Fig. 6.6, the largest difference between the two methods is 0.5503 at 250 bar. Therefore, in this study, Z values determined from the SRK equation of state are used.

6.4.3 Experimental results Fig. 6.7 shows the amounts of adsorbed methane versus pressure at 21
C for the five different ACs. All the plots have the shape of a Type-I isotherm, which also indicates that these samples are microporous materials, typical of ACs. Based on the moles of methane adsorbed per gram of AC (Fig. 6.7A), the capacity of methane adsorption on five ACs increases in the order of AC4 , AC1 , AC3 , AC2 , AC5. Although the surface areas of AC2 and AC5 are close to each other, the micropore volume of AC5 (0.452 cm3/g) is higher than that of AC2 (0.412 cm3/g), which is important for the adsorption of microporous adsorbents

6.4 Case study: screening activated carbon for natural gas storage

9

Adsorbed Amount (mmole/g)

8 7 6 5 4 3

AC1 AC3 AC5

2

AC2 AC4

1 0 0

10

20

30 Pressure (bar)

40

50

60

(A) 120

Adsorbed Amount (V/V)

100 80 60 40

AC1

AC2

AC3

AC4

AC5

20 0 0

10

20

30 Pressure (bar)

40

50

60

(B) FIGURE 6.7 Comparison of adsorption capacity among the five granular activated carbons. (A) Mass to mass (mmol/g) comparison and (B) volume to volume (v/v) comparison.

occurring by the mechanism of pore filling. This may also be since AC5 has more slit pores that are close to the optimal pore size for methane adsorption. Earlier theoretical studies found that for methane adsorption on slit-pore ACs [3537], the

281

282

CHAPTER 6 Natural gas storage by adsorption

˚ to create the maximum density for the optimum pore width is 11.211.4 A adsorbed phase. Based on the volume of methane adsorbed per volume of AC (Fig. 6.7B), the capacity for methane adsorption increases in the order of AC4 , AC2 , AC1 , AC3 , AC5. This change of order in comparison to the mass to mass adsorption capacity in Fig. 6.1 can be attributed to the fact that the bulk density in volume to volume comparison is an important factor for volumetric adsorption capacity. The bulk density for the five ACs are AC1 (0.47 g/cm3), AC2 (0.39 g/cm3), AC3 (0.49 g/cm3), AC4 (0.54 g/cm3), AC5 (0.50 g/cm3). Although AC2 has the highest BET surface area and high micropore volume, it has the lowest bulk density. Therefore, it can be concluded that synergistic effects among BET surface area, micropore volume, micropore size distribution, and bulk density play an important role in determining the volumetric adsorption capacity.

6.4.4 Empirical modeling with adsorption potential theory The adsorption potential theory can be applied successfully to gas adsorption on microporous adsorbents such as zeolites and ACs. According to this theory, the adsorption occurs by a micropore filling mechanism [5052]. In microporous adsorbents, micropore walls are near each other, providing an enhanced dispersive force which consequently causes a higher adsorption potential. The enhanced adsorption potential results in a higher heat of adsorption in the micropore compared to the corresponding one on a surface. For example, the adsorption heat of n-hexane on AC at 20
C at a loading of 0.25 mmol/g is about 15 kcal/mol while that of a nonporous carbon black is only about 10 kcal/mol [53]. The adsorption potential is defined by: A 5 RTln

  Ps P

(6.13)

where P is the adsorption equilibrium pressure at temperature T, Ps is the saturated vapor pressure of the adsorbate at the same temperature T and R is the ideal gas constant. At high pressure, P and Ps should be replaced by the corresponding fugacities, f, and fs, to correct for non-ideal behavior. The adsorption potential is also called the differential molar work and is temperature invariant at a constant degree of filling of the adsorption space. It is usually scaled against characteristic energy for the distribution function related to the micropore volume occupied by the adsorbate. Dubinin and Astakhov chose the functional form of the Weibull distribution and derived the Dubinin and Astakhov (DA) equation [5052] as follows:   n W A θ5 5 exp 2 W0 E0

(6.14)

where W0 is the limiting volume of adsorption or the volume of micropores, n is an empirical parameter that characterizes the heterogeneity of the adsorbateadsorbent

6.4 Case study: screening activated carbon for natural gas storage

system, E0 is the characteristic energy—a measure of the adsorption strength between adsorbate and adsorbent and W is the filled volume of micropores. Eq. (6.15) is also called the adsorption characteristic curve. According to Eq. (6.15), when the experimental data of various temperatures and pressures are plotted against the adsorption potential, the data will collapse on the same curve. For different adsorbentadsorbate systems, at an equal degree of filling of micropores, the scaled adsorption potential must also be equal. Using one adsorbate as the reference, the ratio of the characteristic energy of one adsorbate relative to the reference adsorbate is called the coefficient of similarity or affinity coefficient: E A 5 5β E0 A0

(6.15)

Therefore, Eq. (6.14) can be re-written as Eq. (6.16) in terms of the characteristic energy of the reference vapor: θ5

   W A n 5 exp 2 W0 βE0

(6.16)

Since the characteristic energy is independent of temperature, plots of the fractional loading versus the adsorption potential at different temperatures will collapse into one curve, which is called the characteristic curve. Therefore, it can be said that if the coefficient of similarity is used as a shifting factor, the characteristic curves of different adsorbates on the same adsorbent can be brought into a single curve. In other words, knowing the reference’s characteristic energy and coefficient of similarity, adsorption isotherms of different gases on the same adsorbent can be predicted. However, the calculation of the adsorption potential in Eq. (6.1) is only applicable to sub-critical condition adsorbates. For supercritical gas adsorption experiments, the effective vapor pressure can be defined by Dubinin [54], Amankwah and Schwakz [55], and others as in Eq. (6.17). And the adsorbed phase molar volume for super-critical gases is estimated by Ozawa et al. [51] as defined in Eq. (6.18). Ps ðT Þ 5 Pc

 2 T Tc

(6.17)

vM ðT Þ 5 vM ðTb Þexp ½0:0025ðT 2 Tb Þ

(6.18)

where vM ðTb Þ is the molar volume of the liquid adsorbate at the normal boiling point Tb . The methane and nitrogen’s critical temperature and pressure are listed in Table 6.8, for effective vapor pressure calculation. Table 6.8 Critical temperature and pressure of methane and nitrogen. Critical temperature (K) Critical pressure (bar)

Methane

Nitrogen

190.55 45.95

126.2 33.90

283

284

CHAPTER 6 Natural gas storage by adsorption

The knowledge of the adsorption equilibrium and isosteric heat of adsorption is essential for the proper design and operation of any adsorption-based process [49]. The isosteric heat of adsorption is usually estimated from the temperature dependence of the adsorption isotherms [52,54,56]. Using Eq. (6.17), and the   Van’t Hoff equation ΔH 5 2 RT 2 @lnP=@T , we obtain Eq. (6.7) for isosteric heat of adsorption. 2ΔH 5 A 1 ΔHvap 2

ðβE0 Þn TdCμs =dT Cμs nAn21

(6.19)

where A is the adsorption potential as in Eq. (6.1), ΔHvap is the heat of vapor condensation as defined in ClausiusClapeyron Eq. (6.9) at constant adsorbed amount, in which Ps is expressed as in Eq. (6.17).   @lnPs ΔHvap 5 RT 5 2RT @T c 2

(6.20)

and dCμs =dT is the change of the maximum adsorption capacity with temperature. The relation Cμs 5W0 =vM ðT Þ is used to express the maximum capacity of adsorption in mmol/g. By using Eq. (6.18), dCμs =dT becomes: dCμs 5 2 0:0025 Cμs dT

(6.21)

6.4.5 Isosteric heat of adsorption modeling Since volumetric adsorption capacity is more important for the limited space of AC packing [49], methane adsorption data on AC5 is chosen for further analysis by adsorption potential theory (characterization and adsorption isotherms are in supporting information). Fig. 6.8 shows the modeling of the DA equation with three parameters W0 , E0 , n using non-linear least-square method. The average relative error for this modeling is 1.82% defined as in Eq. (6.22). Average relative error ðAREÞ% 5

P



  Yiexp 2 Yimodeled =Yiexp N

(6.22)

The values of these optimized parameters are presented in Table 6.9 for methane. Eq. (6.20) indicates that the isosteric heat of adsorption can be divided into three components, the adsorption potential or the molar work for filling the micropores, the enthalpy of condensation, and the molar work change due to the maximum adsorption capacity change with the change of temperature. To express the isosteric heat of adsorption in terms of loading, the following equation is

6.4 Case study: screening activated carbon for natural gas storage

Adsorbed Amount (mmole/g)

8 7 6 model data

5

experimental

4 3 2 1 0 0

5

10 Adsorption Potential (kJ/mole)

15

20

FIGURE 6.8 Characteristic curve for methane adsorption on AC5 and DA equation modeling at 10
C, 21
C, 38
C, and 56
C.

Table 6.9 Optimized parameters of DA equation to model methane adsorption on AC5 at 10
C, 21
C, 38
C, and 56
C. Adsorption system

W0 (mL/g)

E0 (kJ/mole)

n

ARE (%)

Methane-AC5

0.444

7.283

1.600

0.05

ARE, Average relative error.

obtained by using Eq. (6.8):  1=θ  2ðn21Þ=n 1 ðβE0 Þn 1 ln 2ΔH 5 βE0 ln 1 ΔHvap 2 θ θ n

(6.23)

Fig. 6.9 shows that the isosteric heat of adsorption changes with loading with different values for the thermal expansion coefficient δ at 298.15K. As expected, the isosteric heat curve decreases with loading, and when saturation is approached the isosteric heat of adsorption increases rapidly, which is due to the change in the saturation capacity Cμs with respect to the temperature. If this saturation capacity does not change with temperature, or the thermal expansion of the liquid is zero, the heat curve decreases monotonically with loading and approaches zero when saturation is reached. This is consistent with the derivation of isosteric heat of adsorption from adsorption potential theory [30].

285

CHAPTER 6 Natural gas storage by adsorption

50

Isosteric heat of adsorption

286

thermal expansion coeff. = 0 thermal expansion coeff. = 0.0005 thermal expansion coeff. = 0.001 thermal expansion coeff. = 0.0015 thermal expansion coeff. = 0.0020 thermal expansion coeff. = 0.0025

40

30

20

10

0 0

0.1

0.2

0.3

0.4 0.5 0.6 Fractional Loading

0.7

0.8

0.9

1

FIGURE 6.9 Isosteric heat of adsorption versus fractional loading with E0 5 7.283 kJ/mol, n 5 1.6, T 5 298.15K for a methane-AC5 adsorption system.

6.5 Heat management modeling 6.5.1 Mathematical modeling Physical description of the storage system: The geometrical model of the ANG storage system under study is depicted in Fig. 6.11 and it consists of a cylindrical vessel of length Z0 and inner radius of R0, which is filled with a homogeneous medium of granular AC AC5. The cylinder is assumed to be made of steel. Since natural gas primarily consists of methane (B95%), it is idealized as pure methane for thermophysical properties. As shown in Fig. 6.10, a two-dimensional axisymmetric model is assumed for the thermal transient thermal behavior study of the natural gas storage system. The thermal properties and dimensions of the idealized adsorber are listed in Table 6.10. The following equations are used in simulating the thermal transient behavior of the adsorber bed by considering the adsorption bed heat conduction as an initial value problem with surface heat convection. 1. The energy equation is the following 2-dimensional equation in the cylindrical coordinate system:    H 1@ @T @2 T @T 1α r 1 2 5 ρCp r @r @r @Z @t

Where

@T @θ

5 0;

α5

λ ρCp

(6.24)

6.5 Heat management modeling

2. Initial conditions: Initial temperature of all points in the storage system is 25
C. 3. Boundary conditions are: At r 5 0;

At z 5 0; @T @z 5

h λ

 @T @T h  5 0; at r 5 R0 ; 5 3 T Ro 2 T N @r @r λ

h 3 ðT 0 2 T N Þ; and at z 5 Z 0 ; @T @z 5 2 λ 3 ðT z0 2 T N Þ

Numerical modeling: All computer programs were written in MATLABs . The non-linear least square method was used to fit the experimental isothermal absorption data for the DA equation parameters E0, n and W0 in Eq. (6.2). Numerical integration is used to calculate the integral heat of adsorption at a certain temperature up to a fixed pressure value. The PDEs, from Eq. (6.11), were discretized on 21 points in the 2-dimensional spatial domain as in the r and z direction in Fig. 6.10 with 4th order finite difference approximation. This resulted

FIGURE 6.10 Schematic diagram of the storage system.

Table 6.10 Storage system dimension and thermal property value. R0 Z0 Mg (molecular weight) Cpg Cp λ hw T0

0.7112 m 15 m 16.04 g/mole 2450 J/(kg K) 650 J/(kg K) 0.5 to 3.3 W/(m K) 500 W/(m2 K) 298.15K

287

288

CHAPTER 6 Natural gas storage by adsorption

in a set of 441 ordinary differential equations (ODEs) which are sufficiently stiff and a stiff MATLAB solver ode15s was used. The accuracy of the 4th order finite difference approximation to the second derivatives was validated by solving a 2 2 partial differential equation solution of  of @u=@t 5 @ u=@x with an analytical 2ðπ2 =4Þt uðx; tÞ 5 e sin πx=2 to meet the error tolerance of 1 3 1024 [57]. To simulate the transient thermal behavior of the cylindrical storage system, non-dimensionalization is performed and the Eq. (6.24) becomes: @2 θ 1 @θ 1 1 2 @γ γ @γ

where: θ 5

T T0

;

γ5

R R0

;

[5

Z Z0

;

R20 @2 θ 1 Z02 @[2

τ5

λ R20 ρCp

R20 H @θ 5 λ @τ

(6.25)

t

6.5.2 Performance analysis through thermal simulation When it comes to studying the thermal behavior of large-scale natural gas storage at a static state, it is important to focus on the thermodynamic state at the end of charging because the temperature of this final state decides the storage capacity. Therefore, the dynamic behavior/path to reach the final thermodynamic state is not studied here. The following is assumed during the simulation to reach the final thermodynamic state: 1. The bed density is the weighted average of the gas mass (adsorbed and compressed) and AC in the storage system. 2. The bed heat capacity is the weighted average of the gas mass (adsorbed and compressed) and AC in the storage system. 3. The effective bed thermal conductivity is within the range of 0.53.3 [6,8,12,33,34]. 4. The contact thermal resistance between the bed and the vessel wall (made of carbon steel with thermal conductivity of 5460 W/(m K) at 25
C) is negligible since the thermal conductivity of the vessel wall is much bigger than that of the adsorbent bed. To calculate the total amount of heat of adsorption (integral heat of adsorption) that is released during the charging/adsorption process, Eq. (6.14) is numerically integrated up to the final charge pressure of 60 bar. The integral heat of adsorption is further divided by the overall charge time to get the heat generation rate, then the thermal temperature profile of the adsorber bed is simulated as an initial value problem using the two-dimensional cylindrical system and governing equations. As far as mass balance is concerned, the total mass of gas stored is the excess amount of gas adsorbed onto the adsorbent surface and the amount compressed in non-adsorbent skeletal space inside the adsorber. The mass balance for the charging process (adsorption), which is not a flow-through process like adsorptionbased pressure or temperature swing separation, is that the amount of gas flowing

6.5 Heat management modeling

through the gas flow meter for charging to the adsorber equals the amount of gas adsorbed and compressed inside the vessel. The mass balance is expressed as: ð tc 0

CRðtÞdt 5

ð tc 0

PRðtÞVvoid dt 1 ZRT

Mac 3

ð tc ARðtÞdt

(6.26)

0

Where CR(t) is the gas charge rate to the adsorber vessel, PRðtÞ is the pressure change rate inside the adsorber, Vvoid is the non-skeletal space volume in the adsorber, ARðtÞ is the adsorption rate, which is the derivative of Eq. (6.5) and its change rate is linked through adsorption potential which is a function of pressure, Mac is adsorbent mass. With a gas distributor and up to 810 hours of charging, the gas mass distribution inside the adsorber is viewed as evenly distributed throughout the whole adsorber for thermal transient behavior simulation. Differently from the other previous reported work in the literature, this simulation is about industrial-scale ANG application for peak shaving in supplying natural gas to power plant literature [625]. Under this circumstance, there is no fast filling requirement like in those applications for mobile vehicular ANG storage tanks. The other advantage is that industrial-scale ANG application is a static storage application, with more space available to design heat removal accessory system for the isothermal adsorptive storage process to meet the design storage capacity. However, the minimization of the energy cost by optimizing the charging process is still the goal. Fig. 6.11 shows how the % storage capacity loss changes relative to the bed temperature rise during storage at the design pressure of 60 bar. The design 50

Storage capacity loss (%)

45 40 35 30 25 20 15 10 5 0 298 308 318 328 338 348 358 368 378 388 398 408 418

Bed temperature rise during storage (K) FIGURE 6.11 Design storage capacity loss due to adsorber bed temperature rise.

289

290

CHAPTER 6 Natural gas storage by adsorption

equilibrium storage capacity at 60 bar and 298K is used as the reference to calculate the storage capacity loss with the increase in bed temperature during storage. If there is a 50
C rise during the charging process, there will be about a 20% storage capacity loss. Higher than 50
C rise in bed temperature causes a higher loss percentage, with about a 45% loss in storage capacity if the rise is about 110
C. Therefore, an optimized way to cool the adsorber within the given off-peak time of 12 hours is very critical to maintaining the ANG facility at close to the design storage capacity. The simulation of the adsorber system thermal transient temperature profile is used to find a technical solution to remove heat promptly for static industrial large-scale non-vehicular natural gas storage systems, which are used for peak shave diurnal demand swing. The goal is to reach no less than 95% of the design capacity of the system. The available space and accessory equipment to harness energy from nonvehicular natural gas storage systems are different from those used for vehicular storage systems. Different options to remove the heat are available such as a water cooling jacket, a central tube (2-in. diameter) in the axial direction, and a cooling coil in the middle of the radial direction (1-in. diameter). In Table 6.11, the results of different cooling methods with or without the use of the central cooling tube and cooling coil are listed; charging time is 8 hours and the final charge pressure is 60 bar for all the different cases, and λ is the bed thermal conductivity and h is the water forced convective heat transfer coefficient. The value of λ changes from 0.5 to 3.3 W/(m K). Typical porous carbons have thermal conductivity values in the range from 0.1 to 1.5 W/(m K), while higher thermal conductivity values are the results of adding thermal conductive graphitic powders in the forming process of monolithic carbon adsorbent [6,8,12,33,34]. Kuwagaki et al. [58] showed that adding graphitic powders could improve the thermal conductivity by 20 times without serious reduction of surface area and n-butane adsorption capacity. Therefore, it is assumed here that in the manufacturing of carbon adsorbents, adding graphitic powders and densifying the formed adsorbent by increasing its bulk density, will not affect the overall volume to-volume storage capacity (Fig. 6.12). Case 1 and 2 show that low temperatures at both ends of the cylindrical vessel do not propagate as deep as the radial direction outer wall low temperature does. This is expected from Eq. (6.25) as by substituting the value of 0.7113 m for R0, the radius of 56-in. diameter adsorber, and 15 m for Z0, length of the      adsorber, we have: R20 =Z02 @2 θ=@[2 5 0:0022 @2 θ=@[2 , which has a shrinking factor of 0.0022 for the second derivative in the axial direction. Therefore, only when the aspect ratio is low, the temperature at both ends of the cylindrical adsorber can propagate into the internal points of the adsorber bed as shown in    Fig. 6.13, with an aspect ratio of L/D 5 2 resulting a R20 =Z02 @2 θ=@[2 5 0:25: Cases 36 show the effect of increasing the thermal conductivity of adsorbent bed within the literature values reported before [33,34,37]. It is obvious that as

Table 6.11 Thermal transient simulation parameters for different case studies. Case study #

Charge time (h)

Charge pressure (bars)

Charge temperature (K)

λ (W/ (m K))

h (W/ (m2 K))

Central tube cooling (Y/N)

Aspect ratio (L/D)

Bed average temperature (K)

Capacity loss percent (%)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16

8 8 8 8 8 8 8 6 6 8 8 9 9 9 10 10

60 60 60 60 60 60 60 60 60 60 60 60 60 60 60 60

298 298 298 298 298 298 278 278 278 278 278 278 278 278 298 298

0.5 0.5 1.0 1.5 2.5 3.3 3.3 3.3 3.3 2.5 1.5 1.0 0.75 0.6 1.5 2.0

500 500 500 500 500 500 500 500 500 500 500 500 500 500 500 500

N N N N N N N Y Y-c Y-c Y-c Y-c Y-c Y-c Y-c Y-c

10.5 2 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5 10.5

407.3 398.2 386.7 371.1 350.5 340 319.8 310 284.6 286.4 291.5 297.8 304 309.7 311.5 308.4

40.7 37.8 34.1 28.7 21.2 17.2 9.3 5.1 0 0 0 0 2.5 5.0 5.8 4.5

Y-c, central tube plus coil cooling.

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FIGURE 6.12 Temperature profile of Case 1 at the end of 6 h of charge time.

FIGURE 6.13 Temperature profile of Case 9 at the end of 12 h of charge time.

6.5 Heat management modeling

the thermal conductivity of the bed increases, heat is removed more effectively and this gives the option of using graphitic layers from exfoliated graphite in the binding process of making carbon pellets or monoliths. Since only increasing the thermal conductivity could not meet 95% of the design capacity, Case 7 simulates the scenario that the gas is precooled differently from the work of Gu¨tlein et al. [12], where a gas cooling loop was used. To reduce the potential of generating fines that may affect the operation of the system, for example by blocking its valves, precooling of gas is preferred. Precooling in Case 7 helped to reduce the performance loss from 17.2% to 9.3%. But that was not enough to reach the target 95% design capacity. Cases 914 show the results of adding a cooling coil in the middle of the radial direction, in all these cases, the 95% design storage capacity is met even with using an adsorbent bed with thermal conductivity of 0.6 W/(m K). Cases 15 and 16 show the results of using room temperature water to remove heat. With a thermal conductivity of 2.0 W/(m K), adsorption heat is removed fast enough and storage capacity at the end of 8 hours of charging time is maintained at over 95% of design capacity. By adding a cooling coil in the middle of the radial direction, there is no need to cool the water temperature down to 278K for removing heat. Using room temperature water at 298K can save extra enough cost for removing heat during the charging process. Fig. 6.14 showed the temperature profile of the bed at the end of the 8 hours charging period for Case 16.

FIGURE 6.14 Temperature profile of Case 16 at the end of 8 h of charge time.

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6.6 Summary Adsorption-based storage of natural gas is potentially a preferred technical option relative to compression-based storage to address diurnal swings in demand for natural gas, which is used as the preferred feed for power generation facilities in many countries. To demonstrate the advantages of such storage systems, this chapter presented some of our recent works related to the adsorbent materials selection, methods of screening adsorbents through the volumetric method, and establishing an isotherm model for adsorbent, modeling the heat management scenarios for process optimization. A simple adsorption-based process that stores natural gas during non-peak hours and delivers it during peak demand hours is described, the heat management components within this process that will maintain at least 95% design storage capacity for the process is discussed, and the performance of different commercial AC materials concerning physical properties such as volume to volume storage capacity and bulk density is compared, and some of the emerging technologies for the production of ultra-high capacity and economic adsorbents for natural gas storage is also briefly discussed. As renewable energy penetrates the utility sector, this will lead to large gas demand swings that the gas network will not be able to cope with. Storage of natural gas by adsorption is a viable option to help build a flexible natural gas infrastructure and integrate power production from intermittent sources of energy with gas-based turbines. This technology will help improve the reliability of the gas supply network and minimize the gas well production rate adjustments and reduce gas wells operation costs and improve relevant equipment lifetime.

Nomenclature AC ANG BET CNG DOE IUPAC LNG MOF ppmv PSA SRK TDCVD TSA ZTC

activated carbon adsorbed natural gas Brunauer, Emmett, and Teller compressed natural gas Department of Energy International Union for Pure and Applied Chemistry liquefied natural gas Metalorganic frameworks parts in million by volume pressure swing adsorption SoaveRedlichKwong template directed chemical vapor deposition temperature swing adsorption zeolite-templated activated carbon

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CHAPTER

7

Crude oil storage

James Alexander McRae, Kennith Van Ness and Chandrashekhar Khandekar Schlumberger, Houston, TX, United States

Chapter Outline 7.1 Introduction .................................................................................................299 7.2 Types of storage ...........................................................................................300 7.2.1 Storage tank ...............................................................................300 7.2.2 Concrete gravity-based structures .................................................301 7.2.3 Floating tanks .............................................................................301 7.2.4 Underground caverns ...................................................................302 7.3 Chemistry-related issues and solutions ..........................................................302 7.3.1 Corrosion ....................................................................................303 7.3.2 Bacteria .....................................................................................308 7.3.3 Emulsion ....................................................................................312 7.3.4 Carboxylate soaps ........................................................................314 7.3.5 Paraffin and asphaltene ...............................................................315 7.3.6 Inorganic solids ...........................................................................319 7.4 Summary ......................................................................................................321 Nomenclature ......................................................................................................321 References ..........................................................................................................322

7.1 Introduction Crude oil storage is an important part of the oil production process. Oilfield developments usually consist of many producing wells connected to fluids handling systems for treating the crude oil to sales specification before export or sale to refineries. The storage tanks act as an initial area for staging crude oil, and in some cases final treatment before transport. In many cases, it is not possible to continuously export crude oil as it is produced, so it is important to be able to store it for some time.

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00008-9 © 2023 Elsevier Inc. All rights reserved.

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This chapter discusses different types and designs of storage tanks as well as the main issues connected to the operation, maintenance, and treatment of large volume crude storage facilities.

7.2 Types of storage 7.2.1 Storage tank Early oil and gas industry storage tanks were often built onsite using various types of wood. As the industry evolved tanks began to be constructed using metal, often being bolted or riveted together, and later with the development of welding techniques, it became common to see welded-steel tanks. Tanks could be built onsite using sheets of steel or as technologies developed, they could be designed and supplied as segmental elements for final assembly on site. Today crude oil storage tanks are available in varying sizes and shapes. Depending on the application, these tanks may take the shape of vertical or horizontal cylinders, spherical or rectangular. Special applications might require tanks to be full pressure storage in which case horizontal cylinders and spherical shapes are generally used. Tanks used can range from pressurized to low-pressure atmospheric which are widely used from the production fields to the refinery. The most common shapes used are the vertical, cylindrical storage tanks and their capacities can range from less than 100 bbl to over 1.5 MMbbl. As tanks’ capacities vary, so may the dimensions of these tanks, ranging from very small vessels commonly found on offshore structures used to essentially serve as a reservoir to feed a LACT (Lease automated custody transfer) unit. They can vary in dimensions from 10 ft diameter welded or bolted steel tanks commonly found on small tank batteries to massive tanks with floating roofs of over 400 ft in diameter. Bolted steel tanks are available in capacities up to 40,000 bbl or more, depending on the storage application. They are especially useful when smaller sizes are needed for onsite field storage. Generally, bolted tanks are fabricated directly either from 12- or 10-ga steel and sometimes contain several nonmetallic materials such as reinforced concrete, polymer plastics, resins, or hard-wearing enamels. If not galvanized or furnished with a protective coating for corrosion protection, bolted steel construction might not have the expected service life provided by welded-steel tanks. Specifications on design, fabrication, erection, and inspection have been developed and updated by American Petroleum Institute (API) for capacities from 100 to 10,000 bbl [1]. Welded-steel tanks are constructed of heavier plate materials that often include a corrosion allowance. API has established three specifications for this type of tank: Spec. 12 F for capacities of 90500 bbl [2], Spec. 12D for nominal capacities from 500 to 10,000 bbl [3], and Standard 650 for larger volumes [4]. The Standard 650 outlines material, design, fabrication, erection, and testing requirements for tanks operating at atmospheric pressures. Design pressures above

7.2 Types of storage

atmospheric and design temperatures exceeding 200
F (93
C) may be permitted when additional requirements are met. Some of the early tanks used in the petroleum industry were open-top tanks. While these tanks provided liquid containment, the exposure of the liquid surface to the atmosphere resulted in high evaporative losses, product odors, and of course the increase of potential fires. Open-top tanks today have only limited use, primarily collection of contaminated run-off or wash water and wastewater. Fixed roof tanks provided improved containment of product vapors and odors and reduced the potential for fires. They can still result in product loss through evaporation and increase the possibility of forming a combustible gas mixture within the vapor space of tanks. As a result, tanks with fixed roofs are normally used for products with vapor pressures of less than 1.5 psia and equipped with pressure vacuum valves to release excessive pressure and prevent collapse due to vacuum caused by movement out of the tank by pumps. Tanks with floating roofs, although not usually found in production operations, are often encountered in pump stations or oil storage terminals where the crude oil has been stabilized to vapor pressure of less than 11.1 psia.

7.2.2 Concrete gravity-based structures Concrete gravity-based structures (GBSs) are often found in some offshore platforms and usually form part of the structure connecting the production platform to the seabed. The weight of the concrete structure keeps it fixed to the ocean floor and usually contains connected production completions to allow for further field development. ExxonMobil Hebron project is located 350 km (217 mi) offshore Newfoundland. The centerpiece is a stand-alone concrete GBS platform in 95 m (311 ft) of water. The GBS is designed to withstand sea ice, icebergs, and the severe meteorological and ocean conditions of the northern Atlantic Ocean. The GBS has a storage capacity of 1.2 MMbbl of crude.

7.2.3 Floating tanks Floating offshore crude oil storage tanks were a common engineering design for new fields, especially where export pipelines are not immediately available. The tanks work on the principle that crude oil is lighter than seawater, and once full, will anchor the tank to the seabed. Head pressure from the seawater and wellhead pressure from the production wells means that the crude oil can enter and exit the tanks without the need for other installed pumps. Dubai’s Kazzans were designed in 1970, the Kazzans’ K2 & K3 are 500,000 bbl oil storage tanks. The tanks were constructed onshore on a beach near the sea’s edge and when complete they were floated out to the location by flooding the construction area and towing them out to the location. As the bottoms of the tanks are essentially open to the sea and filled with seawater, as oil is pumped in

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for storage before shipping the water is displaced out when oil is unloaded onto tankers the oil is displaced by seawater entering the bottom. These types of tanks are not as common today mostly due to concerns about controlling potential leaks where the seawater is directly contacting the produced oil.

7.2.4 Underground caverns An alternative to large onshore or offshore storage tanks above ground is the use of underground caverns built into the natural rock formations. These caverns are usually large spaces, manufactured specifically for this purpose, making use of the natural barrier provided by the rock. They are usually used to store very large volumes of crude oil, for example for a country or state emergency stock [5]. Oil is pumped into the caverns under pressure which allows it to displace the brine inside. The water is then used above the ground in brine storage reservoirs and eventually reused in the storage and distribution process as a displacement for the crude oil when needed. As brine is heavier than crude oil, the oil sits atop the brine and very little mixing occurs. The caverns are always full of brine and/ or crude oil. The enormous capacity of the caverns allows customers to offload large volumes of crude oil at one time and make deliveries to refineries in volumes that meet their specific needs. Louisiana Offshore Oil Port (LOOP)—Crude Oil is shipped from all over the world and as demand has increased the size of tankers plying the world’s oceans has increased to the point that accessing traditional ports has become increasingly difficult. Originally organized in the early 1970s the port utilized several singlepoint mooring buoys and pumping platforms located offshore Louisiana for the offloading of crude tankers. The Clovelly Hub is located twenty-five miles (40 km) inland and is connected to the port complex by a 48-in. (122 cm) diameter pipeline is the largest privately-owned crude oil storage facility in the nation. At Clovelly LOOP operates eight underground caverns with a total storage capacity of 58 million barrels. The cylindrically shaped caverns are man-made having been hollowed out from a vast underground salt formation. While salt dome storage is not an unfamiliar technology, the LOOP Clovelly Salt Dome facility is the only repository in the world continuously receiving and distributing crude oil.

7.3 Chemistry-related issues and solutions Produced crude oil requires a variety of physical and chemical treatments before storage and export for sale. Co-produced water and solids should be removed and the fluids treated to ensure the stability and integrity of the production systems. It is common for the crude oil to contain at least some water and solids at the inlet

7.3 Chemistry-related issues and solutions

of storage tanks, which over time will cause issues that must be addressed when using long-term storage of crude oil. Corrosion and sludge accumulation are two major chemistry-related issues for crude storage. Often the sludges are a multicomponent mixture of organics and inorganics, consisting of paraffin, asphaltene, naphthenate, emulsion, corrosion product, scale precipitates, unconsolidated fines produced from the reservoir, etc. The root cause of the mitigation strategy for each component is discussed in the flowing subsections.

7.3.1 Corrosion One property common to almost all crude storage is metallurgy which in most cases is a variant of carbon steel. This is primarily due to cost considerations, with it being a cheaper material family than more exotic metals. There are examples of tanks being lined with resins, polymers, or simply manufactured from other materials but this is not as common. Corrosion in tanks can broadly be categorized into general corrosion and pitting corrosion. General corrosion tends to be associated with CO2-related attacks and is fairly predictable if the incoming fluids are well understood. Pitting corrosion is usually associated with the formation of solid layers on the metal surface. These solids could be sand, biomass, solid reaction products like iron sulfide, scale, organic deposits such as paraffins/asphaltenes, or even just the disruption of existing corrosion reaction products like iron oxides on the surface. Pitting corrosion is more difficult to monitor and predict as it can happen in very specific, small areas. The best method of protecting against this in a tank or any other production process is to keep the surfaces as clean as possible. The consequences of undetected or untreated pitting corrosion can be serious and occur without warning. Commonly the pit can become a tiny hole allowing oil to leak from the storage tank into other containment or the environment. In the case of multiple pits forming nearby, they can sometimes join together and form weak areas in the metal. This can cause a large reduction in the strength of the structures which is of particular concern when it is at the bottom. The three main physical areas of concern related to corrosion in tanks are the bottom plate, internal tank roof, and tank externals.

7.3.1.1 Bottom plate corrosion The bottom of the tank is likely to have the most contact with solids, water, and other contaminants due to gravity. It will also have contact with the ground, known as the “soil side” which is usually treated separately from the internal surface. Cathodic protection is added to most tanks as standard with different foundation designs also considered limiting direct contact of the soil side with the ground as much as possible. Physical protection like coating is common for tank bottom plates. The use of volatile corrosion inhibitors to treat the soil side has been gaining more exposure in recent years [6,7]. The inhibitors contain

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components that will partition into the gas phase and migrate to the metal surface through the layer between the tank bottom and the ground, which can be concrete, sand, or similar materials. Application is dependent on tank design, but can often be via injection into leak detection ports or similar access points. There are a variety of concerns on the internal side of the bottom plate. Contaminants to be considered for treatment or removal from a corrosion perspective include free water, emulsion, and other solids or sludge as often as practicable. The use of emulsion breaking chemistry can be useful to treat mixed sludges, biocides will reduce the potential for biomass and H2S corrosion, and scale dissolvers can be considered to treat organic deposits. All are discussed in more detail in subsequent sections of this chapter. The use of liquid corrosion inhibitors can also be considered in storage tanks. The most commonly used chemistries for corrosion inhibition in both volatile and standard internal crude tank applications are imidazolines and quaternary ammonium compounds (Fig. 7.1). The accompanying solvent package is usually changed to either target required volatility in the air to allow the product to reach the tank bottom or provide relative solubility in oil or water to treat the tank’s internal fluid streams. In tanks where there is a separate water layer on the bottom, any corrosion inhibitor selected should have components that transition well into the water phase to ensure it is present in high enough concentrations in the dosages needed. In tanks with only small amounts of water, a more oil-soluble formulation may be preferred, with the main mode of corrosion protection being film-forming on the tank walls. The selection of corrosion inhibitors for the crude/water/sludge mix is usually conducted using static immersion tests with metal coupons. Corrosion inhibitors for bottom plate treatment or roof internals would most likely be volatile and the testing required for this would be slightly more complex, as would need to mimic inhibitor transition from the liquid phase, into the air and then contact the metal surfaces [8]. It is often possible to optimize the required dosage of chemical inhibitors when a cathodic protection regime is in place. Cathodic protection should be added as standard to crude storage tanks, either via sacrificial anode in the case

FIGURE 7.1 Generic chemical structures of imidazoline (left) and quaternary amine (right).

7.3 Chemistry-related issues and solutions

FIGURE 7.2 Illustration of cathodic protection in overground storage tanks.

of small structures or more complex cathodic protection via electrical current (Fig. 7.2). Physical removal of solids is also an option that should be considered where solids are not dissolvable and the tank has such a configuration that allows drainage or physical intervention via jetting or similar. Care must be taken not to create new anodes, as removing corrosion deposits often makes the general corrosion worse in the short term until either a new layer of inhibitor or oxidized metal forms to protect the surface.

7.3.1.2 Internal tank roof corrosion Tank roof types can vary, with either internal or external floating, cone roofs, or geodesic domes being the main types found. Some storage tanks employ a double design with an internal floating roof cone or dome. Crude oil tank roofs are often made of aluminum, due to the high strength mass of the metal [9]. One thing almost all tank roofs have in common is the fact they are usually made from smaller pieces of metal joined together, often by metal alloy which can corrode itself or promote galvanic corrosion along the edges of the roof material itself. The tank roofs have contact with any gas still left in the crude. Acidic gas containing either CO2 or H2S can cause a severe attack in concentrated areas.

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Ingress of O2 is also a potential concern related to the potential for explosion and the most effective way of treating both is blanketing using a nonacid gas source providing positive pressure inside the tank, displacing the potential corrosive species and reducing the possibility of air ingress. The source of gas and process to provide it can be different, but usually include volatile hydrocarbon species like methane from the production process, and sometimes a portion of an inert gas like N2 which would be generated onsite. Optimization of this process is important to reduce the risks described above, but also the amount of produced gas vented to the atmosphere and overall environmental footprint [10,11].

7.3.1.3 Tank external corrosion Corrosion on the outside surfaces of crude storage tanks can be expected and the potential can depend on the environmental conditions where the tanks are situated. Tanks situated close to large bodies of water, especially sea coasts, tend to experience more humid or salty air than deserts or other land locations where the air is drier. Locations that have more precipitation in the form of rain or dew also present more risk. In colder locations, it is common to have external insulation on tanks and vessels. It is normal practice to have water-proofing materials added to the insulation to prevent water ingress against the tank walls, providing a localized corrosion risk. Insulated tanks that are situated in locations with low ambient temperatures which lie between the dew and freezing points of water can be especially prone to corrosion on external surfaces [12]. Corrosion on the external surfaces of crude storage tanks is treated by standard maintenance regimes of cleaning and painting, thus creating a physical barrier between the iron-containing metal and the external conditions. There are no specific considerations compared to any other metal structure open to the elements. Continuous maintenance of any protective barrier is extremely important to ensure there are no small, exposed sections that would act as an anode for the corrosion redox reactions.

7.3.1.4 Tank corrosion monitoring Scheduled inspection is the most efficient way of monitoring corrosion. There are relevant API standards that deal directly with this type of inspection protocol, including recommended times, scope, and methods that can be conducted by normal personnel working with these tanks [13,14]. The standards recommend an initial observation step, followed by continuous monitoring with different methods available. Monitoring of tank corrosion protection is usually conducted by a combination of direct measurement of the corrosive environment alongside nonintrusive monitoring of the effect on the tank. The most common direct monitoring methods are coupons and electrical resistance (ER) probes. Coupons are the simplest of all direct measurement devices (Fig. 7.3). They can be inserted and retrieved relatively easily and can be directly examined for evidence of general corrosion from weight loss or localized corrosion indicated

7.3 Chemistry-related issues and solutions

FIGURE 7.3 Example of corrosion coupons.

pits. The coupons should be specimens of the exact tank metal, ideally treated the same way. The main disadvantage is that the data gathered represents current conditions, so any corrosion detected has already occurred. There is also the possibility that localized corrosion could be missed if the coupons are inserted in the wrong area. ER probes can gather continuous real-time analysis of the corrosive environment where they are placed. The probes record the thickness of small metalconducting elements which are exposed to the corrosive environment. The principle behind the technique is that the ER of the conductor element will increase as the cross-section decreases as it corrodes. As with coupons, the metal loops should be made of the same metal as the storage tank. Upgraded versions of ER method are available on the market today. They work on the same principle but can give more instantaneous results when added to the system of interest. There are also nonintrusive techniques that can assess potential corrosion by detecting a metal loss in tanks and vessels. The most common methods to be considered for use in crude storage tanks are ultrasonic (UT), radiography, and thermography testing. UT testing works by propagating short pulses of sound waves at the UT frequencies (0.125 MHz typically) and detecting the time taken to reflect and return to the detector. This builds a “picture” of the metal wall thickness to identify any areas that may be corroded. This type of instrument is cheap and portable, so ideal for the analysis of multiple tanks or different areas in a large tank. The radiography technique is used to detect imperfections in small areas of interest and is often used in addition to UT testing. There are various versions available depending on the required output and data reporting. The meter contains a radioactive isotope that generates either gamma radiation or X-rays which can penetrate a wind range of materials and determine if corrosion is present. This technique is particularly useful in areas where there is tank insulation or protection with concrete or PVC that would stop sound waves from UT testing. Infra-red thermography (IRT) can create a thermal image of storage tanks by detecting emitted radiation (Fig. 7.4). Infra-red radiation is emitted by all objects

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FIGURE 7.4 Infra-red thermography image of tank sludge.

that have a recordable temperature. Thermographic cameras create a picture of variations in temperature, which makes them particularly useful in detecting different layers in crude tanks like sludge or biomass which will cause an increased corrosion risk. IRT is particularly useful as a tool to guide physical tank cleaning regimes, by indicating areas of concern with solid concentration.

7.3.2 Bacteria It has long been proposed and understood that the presence of bacteria in oil and gas production processes both decreases the quality of crude and increases the potential for damage to assets. The presence of bacteria in crude storage tanks is no different and should be considered as a root cause for some of the corrosion issues discussed above. Microbial-induced corrosion (MIC) is the generic term used to describe usually localized corrosion caused either directly or indirectly by bacterial activity. There are a variety of different pathways, with both anaerobic and aerobic species sometimes present in crude storage tanks, depending on the location. The presence of anaerobic sulfate-reducing bacterial strains is particularly of concern, as they produce acid species like H2S that contribute to the corrosiveness of the crude oil. H2S causes metals to crack and blister and also reacts with iron in unprotected tank structures which is solubilized at the corrosion cell anode and increases the overall corrosion reaction. All bacteria will form some kind of solid sludge as they reproduce and metabolize, known as biomass (Fig. 7.5), which will result in a localized environment suitable for under-deposit corrosion. Bacteria often have a significant impact on the storage of crudes because often the associated problems only manifest themselves well after the bacterial growth has reached a level of serious

7.3 Chemistry-related issues and solutions

FIGURE 7.5 An example of biomass tank sludge.

contamination, sometimes as an unexpected corrosion failure. Bacterial populations can reproduce at a very fast rate, with some strains able to double in numbers in a matter of minutes or hours. This means a few bacteria can become a colony of millions within a few days. It has been generally accepted that MIC results from the activities of a mixture of microbial communities, rather than a single type of bacteria [15]. Microbe communities present in tank bottom sludges, which often lead to MIC, are usually a mixture of aerobic and anaerobic types with both sessile and planktonic forms usually found. Sessile bacteria are known to cause more direct MIC issues and are usually found in nonuniform parts of the tank, like welds, scratches, rough surfaces, and attached to other solids [16]. The introduction of water into tanks, either as part of emulsion, distinct droplets, or sludge gathering at the bottom of the tanks, will provide potentially different environments for bacterial growth. Evaluation of bacterial colonies present in crude tank inlet fluids and sludge is key to developing a mitigation strategy. It is common to find bacteria with different oxygen requirements coexisting in the same system and the same deposits. In highly oxygenated systems, for example, anaerobic bacteria may survive in tiny crevices in pipe surfaces that are out of the direct flow of the oxygenated water. In addition, as populations of aerobic bacteria deposit on a system surface, oxygen diffusion to the surface is suppressed. This creates a reduced oxygen environment in which microaerophiles and anaerobes can thrive, shielded from the oxygen in the system by the aerobic bacteria. A sampling of tank fluids should be undertaken regularly and at various levels in each tank. The presence of anaerobic environments, particularly towards the bottom of storage tanks, means that care must be taken to sample with the least possible exposure to workers. Biomass sludges often contain entrapped hydrogen sulfide gas, which is extremely toxic. To address some of these potential

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concerns, there are a variety of analysis techniques available which range from the traditional, lost cost but less accurate techniques like serial dilution, to quicker and more accurate methods like nucleic acid, ATP, and APS which are bacteriaspecific. There is also a new technique recently introduced based on fluorescence which is not specific to any bacteria type, so suitable for the complete assay of all present bacteria [17].

7.3.2.1 Biocide treatment Current control of crude oil storage tank bacterial populations is almost exclusively conducted using biocide chemicals or by physical removal of the biomass sludge using scrapers or fluid jetting. There is some precedent of reducing the populations of anaerobic bacteria by recirculating aerated crude oil between lines and tanks; however, this would be very much dependent on the design of the vessels and the availability of adjoining lines. Biocides based on aldehyde chemistry are the most commonly used in the oilfield to treat both tanks and production vessels. The most widely used aldehyde is glutaraldehyde (glut), which is favored because it can kill a broad range of organisms by direct interruption of the bacteria membrane cell wall. Phosphorous-containing species like tetrakishydroxymethyl phosphonium sulfate (THPS) are also commonly used, with the choice dependent on the microbial populations and effects on biomass. Other biocides known for their effect on biomass are bisoxazolidine, 2,2-dibromo-3-nitrilopropionamide (DBNPA), and quaternary ammonium compound (quat). Selection of biocides should be undertaken based on both speed of kill and persistency. Table 7.1 shows a summary of some of the common commercially available biocides and general characteristics that can affect choice when considering tank treatment.

Table 7.1 Biocide chemicals properties and applications. Biocide

Kill speed

Protection time

DBNPA

5

1

Gluteraldehyde EDDM THPS Bronopol Quaternary ammonium Glutaraldehyde

5 4 5 3 1

2 2 2 2 5

Topside, water treatment, tank decontamination Water treatment, tank decontamination Tank decontamination Water treatment, tank decontamination Tank decontamination Water treatment, tank decontamination

1

5

Tank decontamination

Application

Speed of kill: 1 5 slow (days), 5 5 quick (,1 hour). Duration of protection: 1 5 few hours, 5 5 few weeks. DBNPA, 2,2-Dibromo-3-nitrilopropionamide; EDDM, (ethylenedioxy)dimethanol; THPS, tetrakishydroxymethyl phosphonium sulfate.

7.3 Chemistry-related issues and solutions

Batch injection of shock dosage of biocide is the usual mode of treatment for tanks and vessels, with quicker monitoring techniques being more useful in evaluating and optimizing these treatments. Biocide formulations containing surfactants are generally more effective at penetrating biofilm layers. For these nonoxidizing biocides, two key parameters, bacteriostatic level, and bactericidal level need to be considered for the most appropriate treatment levels. The bacteriostatic level is the level at which the biocide prevents further growth of the bacteria and is independent of the level of organisms in the system. The bactericidal level is the level at which the biocide kills all the bacteria in the system and is dependent upon the number of bacteria present. The most commonly used oilfield biocides, such as glutaraldehyde, THPS, and quaternary ammonium compounds will give an excellent kill with a relatively short contact time. In such cases, it is preferred to use a slug dose of 1000 ppm for 1 hour rather than 200 ppm for 6 hours, if the dosing equipment will permit. This treatment level is particularly effective for systems where there is evidence of significant biofilm growth, as these high treatment levels increase the likelihood of biofilm penetration. The level of bacterial activity will influence the frequency of slug doses in the system and the quality of tank fluids being treated.

7.3.2.2 Nonchemical treatment There are limited nonchemical options for dealing with bacteria in tanks, and those that are available usually require some additional chemical treatment to ensure effectiveness. Good housekeeping is essential and will reduce the demand for chemical applications. Depending on the setup of the crude tanks, it may also be possible to modify the environment so that conditions do not favor bacteria growth. It is difficult for bacteria to form thick biofilms on surfaces exposed to high velocity (greater than 1 m/s) water. Physical removal of sludge and deposits from tank bottoms eliminates some environments that are ideal for the growth of bacteria. Keeping tanks free of scale and other deposits will also help control bacteria. Dirty systems upstream of the crude storage usually require a substantial clean-up program involving pigging of lines, flushing with nonionic detergents, or acidizing, to minimize the solid carryover. Reduction of crude tank residence time before export will also reduce the potential for bacterial growth and associated biofilm formation. Changing the oxygen content of the environment is another technique that has been used to control sulfate-reducing bacteria (SRB) and other anaerobic populations. An increase in redox potential will prevent SRB from growing. A disadvantage of this technique is that the growth of other bacteria such as sulfur-oxidizing bacteria could be promoted. A further technique that has been applied in a limited number of systems is to artificially adjust the pH outside the favorable conditions (pH of 49). This will have an impact on either corrosion or scale control and is only practicable in a very limited number of situations.

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7.3.3 Emulsion The presence of emulsions is generally problematic in primary oil production systems. Emulsion in crude oil storage tanks is no different. It contributes to the formation of thick sludges which promote bacterial growth, induce corrosion, reduce tank capacity and cause issues with crude oil removal for export. Most production facilities have at least one separation phase upstream of crude storage tanks. The carryover of emulsion from the separation process into crude storage tanks can be caused by the presence of solids, fluctuations in temperatures, issues with separation equipment or chemical efficiency, or simply systems operating above their maximum designed capacity. Existing emulsion at tank inlets can be stabilized by any existing solids or chemical emulsifying agents. These agents can be present naturally, like small-sized waxes, asphaltenes, and scale, or the chemicals introduced earlier in the process. Corrosion inhibitors tend to cause emulsions as they are surface-acting agents. A sampling of suspected crude oil emulsion from storage tanks is a relatively simple process. The emulsion can often be seen visually, especially when compared to a known sample of dry crude oil. The main methods available to treat emulsions and associated sludges, other than reducing solids and fluid shear are heat, demulsification chemicals, and physical processing.

7.3.3.1 Heating The separation of oil and water by gravity in crude storage tanks can be increased by the addition of heat. This lowers the oil viscosity and helps the removal of gases entrained in the liquid mixture. It also helps increase the solubility of some natural emulsifying agents like waxes and other surfactants. It is common to see improved separation in oil production regions during summer, purely down to the increased ambient temperatures. The artificial addition of heat can be expensive and could also alter the crude oil characteristics by evaporating light ends, so it is usually confined just to the amount needed to keep storage tanks liquid in colder climates, rather than a single emulsion separation tool.

7.3.3.2 Demulsifier chemicals The following properties are important for determining demulsifier performance: • • • •

Surfactancy (oil/water interfacial properties) Ability to flocculate Ability to coalesce Wettability change

The demulsification process takes place at the oil/water interface of emulsion droplets, thus the active components in the demulsifier chemical must migrate rapidly to this interface to perform their function. The emulsifying agents are normally already

7.3 Chemistry-related issues and solutions

concentrated at the interface, and this creates an additional hindrance to the demulsifier. A good demulsifier must, therefore, not only rapidly migrate to the interface, but also successfully displace existing emulsifiers at that site. The droplets may now collide, rupturing the surrounding film and permitting coalescence. Solvent plays an important role, but some intermediates also assist in the migration of the active components. Most commercial demulsifier formulations are a mixture of different chemistries, each affecting some properties described above. The demulsifier can be injected continuously or applied as a batch treatment to remove interface “pads” that form intermittently. Demulsifier formulations are very crude and water specific, so often need prior testing on the emulsion to be treated. Fig. 7.6 shows the result of the demulsifier treatment of the emulsion sample from the tank interface.

7.3.3.3 Physical treatment The presence of emulsions and sludges that cannot be removed by periodical demulsifier injection may have to be treated using a mixture of heat, chemicals, and physical separation, especially where there are solids present. Fig. 7.7 illustrates the process of sludge remedial treatment in the crude oil tank. Initial laboratory testing involving heating, centrifuges, and chemicals can help design a suitable treatment scheme. Many of them share the same characteristics: • •

Large solids removal Homogenizing and processing, including the use of chemicals

FIGURE 7.6 Crude oil emulsion from tank interface (left) and separated into oil and free water (right).

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FIGURE 7.7 Schematic for crude tank sludge remedial treatment process.

• •

Removal of build solids Separation of oil, water, and fine solids

There are documented examples of large-scale treatment of hundreds of thousands of barrels of tank and pit sludge providing extra revenue and improving tank storage space. Fields in Kalimantan, Indonesia had accumulated .300,000 bbl of tank sludge which were characterized as containing mixtures of inorganic scale, wax, sand, and water in oil emulsion. The sludge was routed into a secondary containment/quarantine pit to reduce the effect on production system retention times. Pit storage is an environmental risk, and some deterioration of the pit walls was observed. A combination of chemical and mechanical treatment was designed and resulted in complete resolution of the sludge, providing .40,000 bbl that could be sold [18].

7.3.4 Carboxylate soaps Crude oil that contains high levels of carboxylic acids, particularly naphthenic acid derivates, can be susceptible to the formation of carboxylate soaps in production vessels and storage tanks. These issues are found most commonly in West Africa, Asia, and to a lesser extent in some North Sea fields. Carboxylate soaps tend to form in the presence of divalent ions, most commonly where the calcium concentration in produced water is high. Magnesium presents a lesser degree of concern, as the potential structure bonding is weaker. The terms “naphthenic” and “carboxylic” acids are sometimes used interchangeably, but naphthenic acids are a part of the carboxylic acid family based on cyclic groups connected to an alkyl chain. An example structure for a naphthenate molecule is shown Fig. 7.8. The Carbon chain number of these types of solids is high (C-70 1 ), and forms a sticky, soap-like structure that is difficult to remove. Fig. 7.9 presents these solids found in a crude storage tank. Remedial treatment of carboxylic acid solids in crude oil tanks usually requires some kind of acid, as the structure will dissolve directly in mineral and weak acids. Acids are usually added as a continuous injection stream into the

7.3 Chemistry-related issues and solutions

FIGURE 7.8 An example of naphthenate molecular structure.

FIGURE 7.9 Carboxylate solids from a crude storage tank.

tanks, or as a batch treatment. Acetic acid is the most popular choice. The formulation of acetic acids into demulsifier products can help control carboxylic acid formation upstream. There have also been some works conducted on the formulation and placement of nonacid demulsifier products which have shown promise in tackling this type of solids production [19].

7.3.5 Paraffin and asphaltene Solids accumulated during long-term storage include paraffins (waxes), asphaltenes, scale precipitates, corrosion products, as well as formation solids. Prolonged accumulation of these solids can present a plethora of problems such as reduced storage capacity, corrosion, out-of-specification fluids, increased maintenance, remediation costs, etc. The crude oil is depressurized and cooled after being produced from the reservoir. Once the temperature of the fluid near the pipe wall reaches below a critical temperature known as wax appearance temperature (WAT), wax crystals start to precipitate due to loss of solubility. As a result of this process, a radial wax concentration gradient is established. This radial concentration gradient leads to a

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mass flux of wax molecules toward the wall. The concentration gradient within the gel deposit leads to an internal diffusion of wax molecules inside the gel deposit, causing aging (as the solid wax content of the gel deposit increases with time). Wax deposition in the pipeline thus primarily occurs due to a molecular diffusion mechanism. The formation of wax crystals significantly increases the crude viscosity, as illustrated in Fig. 7.10. Asphaltenes are the n-pentane or n-heptane insoluble and toluene soluble fraction of crude oil, that remain soluble in a colloidal form under reservoir conditions. During the production process, asphaltenes become destabilized due to pressure, temperature, and compositional changes. factors that could bring about compositional changes include injection of carbon dioxide or natural gas, well stimulation, thermal recovery, chemical flooding, or other operations such as work over or sand control processes. Pressure is the most important factor in the asphaltene precipitation process. As the system is depressurized from the reservoir to pressure just above the bubble point, the light ends expand more than the heavier fractions. This increase in the molar volume of light ends continues up to the bubble point. At the bubble point, the light ends start leaving the liquid phase, which in turn reduces the molar volume of light ends. This, in turn, minimizes the solubility of asphaltenes due to changes in the density of the fluid causing destabilization. The destabilized asphaltenes then precipitate out of the crude oil

FIGURE 7.10 Changes of measured crude viscosity with the temperature at different shear rates [20]. Reproduced with permission from S. Kumar, V. Mahto, Emulsification of Indian heavy crude oil in water for its efficient transportation through offshore pipelines, Institute of Chemical Engineers, 115 (2016) 3443.

7.3 Chemistry-related issues and solutions

and may start to deposit in formations, flowlines, pumps, tubulars, safety valves, etc. causing numerous production issues. Once the crude oil enters a storage tank, it remains static under normal storage conditions. Under these conditions, the wax crystals collide with each other, combine, and adhere to the other solids, forming large-sized particles, which eventually settle under the action of gravity. As a result of this process, over a period the concentration of wax particles in the upper layer becomes lower than that in the lower layer, resulting in a possible back flux of the small particles. Therefore, after a certain degree of sedimentation, the settlement reaches an equilibrium state, and the sedimentary layer is no longer further thickened, which is called the settlement equilibrium state. The sedimentation process itself depends on the viscosity as well as the API gravity (heavy or light) of the crude oil. The resulting sludge is neither homogenous nor solid. The density, as well as the viscosity of the sludge, can change depending on the level of the sludge in the tank, external temperature (seasonal temperature variation), as well as crude API gravity. To prevent sludge formation, the minimum critical velocity for suspension (depending on whether the crude oil is light or heavy) must be maintained throughout the entire fluid volume. In most instances, it is impossible to maintain this velocity and hence the sludge formation cannot be avoided during crude storage. Conventionally, side-entry electric propeller mixers are installed on the tanks for agitation, but these mixers are often not enough to maintain the critical energy throughout the volume, resulting in sludge-free areas only surrounding the mixer. Substantial or severe sludge deposition occurs beyond a specific radius (69 m) from the mixer. The sedimentation process of asphaltene is mainly described by two thermodynamic models, the solution model, and the colloidal model. The solution model mainly considers that crude oil and asphaltene are a kind of homogeneous binary mixture, in which asphaltene is a solid solute and crude oil is a solvent. In this model, the amount of asphaltene dissolved in crude oil is determined by the solubility of asphaltene in crude oil, and whether asphaltene precipitates in crude oil is determined by whether the asphaltene content in dissolved crude oil reaches the saturation state. Kokal [21] and Singh [22] proposed a colloidal model. The model considers that asphaltene in crude oil is the core part, and the colloid adsorbed on asphaltene constitutes the dispersed phase, while other components constitute the continuous phase. In this model, a colloidal system under normal conditions is dynamically stable and the colloid is nonpolar. The dynamic equilibrium of the colloidal system may be broken due to the factors such as temperature, pressure, crude oil composition, precipitant, etc., and the colloidal molecules outside the asphaltene are destroyed, which makes the colloidal particles polar. These particles attract each other and collide with each other, and eventually settle under the action of gravity.

7.3.5.1 Thermal, mechanical and chemical methods To minimize the buildup of organic sludge (paraffin and asphaltenes), various chemical and nonchemical methods are used. Where possible, liquid in the tanks

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is maintained at higher temperatures. However, this is severely limited by the type of crude (light or heavy), WAT of the crude as well as the vapor pressure of the crude oil. In certain instances, where the gravity settling potential is significant, fluid is circulated between multiple storage tanks to introduce a minimum critical velocity for suspension. This slows down the process of settling precipitated wax particles to prevent the large sludge buildup in storage tanks. Chemicals are used to minimize the wax buildup in storage tanks. Two types of chemicals are used to achieve this goal: •

•

Inhibitors: Paraffin/asphaltene inhibitors are injected into the pipeline to delay the onset of precipitation in addition to reducing the gel strength of the resulting deposit. However, the effect of inhibitors is purely kinetic and diminishes if the crude oil is left in the storage tank for a significant length of time. Additionally, these inhibitors are not effective for paraffin crystals that precipitate in the bulk fluid below WAT. Asphaltene inhibitors are added to the flowline to reduce the asphaltene onset pressure. Dispersants: Dispersants prevent agglomeration and subsequent settling of precipitated paraffins. They are typically polymeric products and readily get adsorbed on paraffin crystals. The smaller-sized wax particles do not settle quickly under gravity, minimizing the sludge formed. In some applications, the dispersants are also added in a small dose during the fluid recirculation process between the tanks. Asphaltene dispersants are used to prevent the flocculation of asphaltenes and keep the particles dispersed in the fluid.

7.3.5.2 Remediation methods In case the sludge buildup has occurred over an extended period, remediation methods need to be used to remove and treat the sludge, recover crude oil, and isolate the hazardous waste for further processing and disposal. These methods include: •

•

•

Use of centrifuge systems: This is typically divided into three steps and includes pre-treatment of the sludge/oil from the tank, followed by desludging of solids, and then a further cleaning stage that involves centrifuges to maximize recovery of the crude oil. Use of jet mixers: The jet mixer introduces a high-pressure stream of diluent into the sludge, shearing and resuspending the wax molecules back into a useable state. This significantly reduces cleaning costs and outage time. This method can remove the sludge without removing the tank from service or disrupting operation. Use of heat: Tanks can be heat traced using piped production fluids as insulation or electric heating elements. There are also thermo-chemical treatments available that will create heat when injected together to allow the melting of the organic fractions as part of a cleaning process [23].

7.3 Chemistry-related issues and solutions

7.3.5.3 Selection of treatment method A key factor in selecting the right treatment methodology is to analyze crude oil properties such as crude oil wax/asphaltene percentage, WAT, asphaltene onset point, oil viscosity, as well as storage parameters such as temperature, flow rate, settling time, etc. One of the most critical aspects of sludge management is to estimate the amount of sludge likely to form under given storage conditions. Due to the different settling mechanisms of various substances in crude oil, the interactions between them and the influence on sedimentation are usually not well defined. Therefore, it is difficult to predict the amount of sludge deposition according to the type of components. Once the extent of sludge formation is known, a methodology (prevention or intervention) can then be selected. If using the chemical method for preventing/ minimizing sludge formation, inhibitors can be selected based on parameters such as pour point, API gravity, viscosity, etc. If chemical dispersants are used, the selection is typically carried out using a benchtop bottle test with optical methods such as Turbiscan or measuring the viscosity of the suspension at relevant temperatures. Many chemical companies have also developed proprietary methods for screening and selection of dispersants and inhibitors [24]. In case of desludging or remediation, it is important to conduct a sludge survey. This includes developing a sludge contour profile to estimate the amount and location of the sludge profile and to plan the desludging operation. With the associated software based on the models discussed above, it is possible to estimate the amount of sludge that may be present in the tank. Once this is known, a strategy for desludging can be safely formulated.

7.3.6 Inorganic solids Inorganic solids presented in crude oil and water production systems can often be carried over into storage tanks by free or emulsified water and end up as sludge, either on the bottom of the sides of the tank. The main types of inorganic solids usually encountered in crude oil tanks can be broadly categorized by mineral scales, corrosion products (due to either abiotic or microbiological reactions), and formation solids: • Mineral scales:

• Corrosion products:

• Formation solids:

Calcium carbonate, CaCO3 Barium sulfate, BaSO4 Strontium Sulfate, SrSO4 Calcium Sulfate, CaSO42H2O Iron oxides, FexOy Iron sulfides, FexSy Iron carbonate, FeCO3 Sand, SiO2 Limestone, CaCO3 Dolomite, CaMg(CO3)2 Clays, SixAlyOm(OH)n

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The types of corrosion products formed are primarily determined by the dissolved corrosive gases (O2, CO2, and H2S) in the stored fluids. With the oxygenated water, iron oxides are the predominant products even at a trace level of oxygen. They can be in various forms, such as magnetite (Fe3O4), hematite (Fe2O3), or goethite (FeOOH). In the absence of oxygen, the formation of iron carbonate (siderite) and iron sulfide is controlled by the relative concentrations of dissolved bicarbonate and bisulfide in the produced water. Iron sulfide has much lower solubility than iron carbonate; thus, it can form even at very low bisulfide levels. Similar to iron oxides, iron sulfides have different polymorphs, including iron-rich mackinawite (Fe11xS), iron-deficient pyrrhotite (Fe12xS), stoichiometric troilite (FeS), and iron disulfide pyrite and marcasite (FeS2).

7.3.6.1 Mineral scales Mineral scales accumulated in the storage tanks can result from the scale particles suspended in the incoming water or the entrapped particles at the interface of emulsion droplets. They can also form via precipitation reactions of the supersaturated water during storage, but the mass of scale formed by this mechanism is usually minor due to the small amount of produced water present in the tanks. The calcium carbonate scale is the most common mineral scale encountered in oilfield operations. Its formation is often related to changes in pressure or temperature. Decreasing partial pressures of CO2 and the associated increase of water pH shift the chemical equilibrium towards calcium carbonate formation. Increasing temperature also causes a drop in solubility and likely facilitates the scaling reaction. Crude storage tanks are usually at atmospheric pressure, the lowest pressure parts of oil and gas production systems. Thus, the formation of a calcium carbonate scale when water is present should always be considered a possibility. Identification of carbonate scale is relatively simple after removal of any oil layer that may be present. Reaction with hydrochloric acid (HCl) is vigorous and releases CO2 with visible effervescence. The formation of sulfate scales is mostly dependent on the concentrations of the scaling ions in water present in the tank. The scale usually forms as smaller crystals compared to the carbonate scale and can be distinguished by a lack of reaction with acids. Chelants like EDTA or similar will dissolve sulfate scales. Care must be taken when sampling for potential sulfate scales as there can often be co-precipitation of radioactive species with barium and strontium. This is often known as NORM (naturally occurring radioactive material). Most of the time radioactive sulfate scale will emit alpha or beta particles, so it can be detected using a standard Geiger counter. As these particles are stopped by sheet metal and the radiation is shielded by water and sludge, it is often not possible to detect from outside the crude storage tank. Removal and disposal of NORM must be handled according to local legislation covering radioactive materials but is usually a combination of a physical jetting, chemical dissolution, and either reinjection into disposal wells or encasement in safe materials like concrete.

Nomenclature

The most effective way of preventing scale formation is to reduce inlet water content for crude into any storage vessels. None of the scales listed above can be formed in dry crude oil.

7.3.6.2 Removal of inorganic solids The most common method for removal of inorganic solids from crude storage tanks is using high-pressure water jetting, sometimes alongside scale dissolver chemicals. The choice of dissolver will depend on the solid type, but broadly will be based on acid for calcium carbonate and iron compounds, and chelating agents blends for sulfate scales. Batch treatment of dissolvers in large crude storage tanks is usually not feasible due to the requirement of large volumes of chemicals. Sand and clays cannot be removed by chemical treatment. Inorganic solids are often removed as part of overall sludge treatments as discussed in Section 7.3.5.2.

7.4 Summary Crude oil storage tanks vary widely in terms of type and design depending on physical location, volumes of fluids stored and oilfield type. However, there are many similarities in terms of the issues faced in operating them successfully. Maintaining the integrity of the structures through protection from corrosion by the external conditions faced is key. Limiting the effects of solids formation, both organic and inorganic is important to ensure smooth operations and maximization of available residence time for the storage of crude oil before export. This chapter covered both the proactive measures that are often conducted, monitoring, as well as available techniques for remediation of potential operational issues.

Nomenclature API APS ATP bbl BS&W DBNPA EDDM EDTA ER ft GBS IRT LACT MIC

American Petroleum Institute adenosine 50 -phosphosulfate adenosine-50 -triphosphate barrels basic sediment and water 2,2-dibromo-3-nitrilopropionamide (ethylenedioxy)dimethanol ethylenediaminetetraacetic acid electrical resistance feet gravity-based structure infra-red thermography lease automated custody transfer microbial-induced corrosion

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CHAPTER 7 Crude oil storage

MMbbl NORM SRB THPS UT WAT

million barrels naturally occurring radioactive material sulfate-reducing bacteria tetrakishydroxymethyl phosphonium sulfate ultrasonic wax appearance temperature

References [1] API Specification 12B, Specification for Bolted Tanks for Storage of Production Liquids, American Petroleum Institute, Washington, DC, 2020. [2] API Specification 12F, Shop Welded Tanks for Storage of Production Liquids, American Petroleum Institute, Washington, DC, 2019. [3] API Specification 12D, Field Welded Tanks for Storage of Production Liquids, American Petroleum Institute, Washington, DC, 2017. [4] API Standard 650, Welded Tanks for Oil Storage, American Petroleum Institute, Washington, DC, 2020. [5] K. Ajay, J.S.N. Chaitanya, K. Chandramouli, et al., Study on rock caverns for storage of crude oil, International Journal for Modern Trends in Science and Technology 7 (2021) 262266. [6] T. Whited, X. Yu, R. Tems, Mitigating soil-side corrosion on crude oil tank bottoms using volatile corrosion inhibitors, NACE-20132242, NACE International Conference, March 1721, 2013, Orlando, Florida, USA. [7] K. Baker, T. Natale, A. Roytman, Mitigation of soil-side bottom corrosion of above ground storage tanks utilizing volatile corrosion inhibitors, NACE-20179462, NACE International Conference March 2630, 2017, New Orleans, Louisiana, USA. [8] M. Singer, F. Farelas, Z. Belarbi, et al., Review of volatile corrosion inhibitors evaluation methods and development of testing protocols, NACE-201913303, NACE International Conference, March 2428, 2019, Nashville, Tennessee, USA. [9] V. Carucci, J. Delahunt, Corrosion considerations for above ground atmosphere storage tanks, NACE-02487, NACE Corrosion Conference, April 711, 2002, Denver, Colorado, USA. [10] M. Childs, A.W. Sipkema, Hydrocarbon gas storage tank blanketing for FPSOs to eliminate VOC emissions, SPE-98763-MS, SPE International Conference on Health, Safety and Environment in Oil and Gas Exploration and Production, April 24, 2006, Abu Dhabi. [11] A. Ehi, Managing greenhouse gas emissions through minimizing the usage of blanketing gas on storage tanks - The Bonny oil and gas terminal case study, SPE178329-MS, SPE Nigeria Annual Conference and Exhibition, August 46, 2015, Lagos, Nigeria. [12] M. Stewart, Surface Production Operations: Volume 5: Pressure Vessels, Heat Exchangers, and Aboveground Storage Tanks, Elsevier, 2021. [13] API Standard 653, Tank Inspection, Repair, Alteration and Reconstruction, American Petroleum Institute, Washington, DC, 2014. [14] API Recommended Practice 575, Inspection Practices for Atmospheric or Low Pressure Tanks, American Petroleum Institute, Washington, DC, 2014.

References

[15] D. Pope, State of the art report on monitoring, prevention and mitigation of microbially influenced corrosion in the natural gas industry, Technical Report GRI-92/ 0382, Gas Research Institute, Chicago, 1992. [16] R.T. Huang, B.L. McFarland, R.Z. Hodgman, Microbial influenced corrosion in cargo oil tanks of crude oil tankers, NACE-97535, NACE International Conference, March 914, 1997, New Orleans, Louisiana, USA. [17] J. Fajt, A. Murphy, A. Jenkins, Development and field application of a new bacteria monitoring technique, NACE-201913158, NACE International Conference, March 2428, 2019, Nashville, Tennessee, USA. [18] D. Denney, Innovative chemical and mechanical processing for treatment of difficult production-waste sludge, Journal of Petroleum Technology 59 (2007) 7677. [19] D.L. Gallup, V. Denny, C.Y. Khandekar, Inhibition of sodium soap emulsions, West Seno, Indonesia field, SPE-130506-MS, SPE International Conference on Oilfield Scale, May 2627, 2010, Aberdeen, UK. [20] S. Kumar, V. Mahto, Emulsification of Indian heavy crude oil in water for its efficient transportation through offshore pipelines, Institute of Chemical Engineers 115 (2016) 3443. [21] S. Kokal, T. Tang, L. Schramm, et al., Elecrtokinetic and adsorption properties of asphaltenes, Colloids Surfaces A 94 (1995) 253265. [22] S. Peramanu, C. Singh, M. Agrawala, et al., Investigation on the reversibility of asphaltene precipitation, Energy Fuels 15 (2001) 910917. [23] O. Nelson, C. Khalil, L. Leite, et al., Thermochemical process to remove sludge from storage tanks, SPE-105765-MS, SPE International Symposium on Oilfield Chemistry, Feb. 28 - March 2, 2007, Houston, Texas, USA. [24] K. Leontartis, Quantification of asphaltene and wax sludge build-up in crude oil storage facilities, SPE-92958-MS, SPE International Symposium on Oilfield Chemistry, Feb, 24 2005, Houston, Texas, USA.

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CHAPTER

Geologic carbon storage: key components

8

Hakan Alkan1, Oleksandr Burachok2 and Patrick Kowollik2 1

TU Bergakademie Freiberg, Freiberg, Saxony, Germany 2 Wintershall Dea AG, Kassel, Hessen, Germany

Chapter Outline 8.1 Introduction ...............................................................................................326 8.2 Geologic carbon storage classifications, definitions, types ...........................330 8.2.1 Definitions ................................................................................331 8.2.2 Geologic carbon storage types ....................................................333 8.3 Key components of geologic carbon storage projects ...................................333 8.4 Surface components: capture, conditioning, and transport ............................335 8.4.1 Capture ....................................................................................335 8.4.2 Conditioning .............................................................................342 8.4.3 Transport ..................................................................................347 8.5 Subsurface components: exploration and reservoir .......................................353 8.5.1 Exploration and screening ..........................................................353 8.5.2 Storage capacity .......................................................................356 8.5.3 Injectivity .................................................................................373 8.5.4 Containment .............................................................................378 8.6 Risk assessment, monitoring, and validation ................................................383 8.7 Monitoring and validation ...........................................................................389 8.8 Regulations and certification .......................................................................391 8.9 Economics .................................................................................................396 8.10 Outlook ......................................................................................................401 Nomenclature ......................................................................................................405 References ..........................................................................................................408

Surface Process, Transportation, and Storage. DOI: https://doi.org/10.1016/B978-0-12-823891-2.00009-0 © 2023 Elsevier Inc. All rights reserved.

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8.1 Introduction It is most pronounced that Jean Fourier was the first who hypothesized the greenhouse effect expressing that the Earth’s atmosphere acts like an insulator from the radiation reaching the sun and making a warming effect [1]. Svante Arrhenius attempted to calculate the effect of carbon dioxide (CO2) on the earth’s ground temperature coming up with the result of reducing the amount of CO2 in the atmosphere by half could lower the temperature some 4
C5
C [2]. He also suggested nature and humans were responsible for the potential increase of CO2 in the earth’s atmosphere due to the “combustion and decay of organic bodies.” First attempts to measure the CO2 concentration in the earth’s atmosphere were started in 1900 and showed a tendency to group around 300 ppm [3]. Except for some rare news in popular journals and newsletters and few research publications, interest in the greenhouse effect and atmospheric CO2 levels was low. Starting from the second half of the 20th century, with the help of emerging analytics and increasing computing capacities, studies intensified on the greenhouse effect. The greenhouse effect and potential change in world climate reappeared in scientific and political discussions in the 1960s and, as a result, the use of chlorofluorocarbons (CFC) has been heavily regulated since the late 1970s, because of their destructive effects on the ozone layer. It should be noted that their effect is much stronger than CO2, one CFC molecule is about ten thousand times more influential than a molecule of CO2. The critical warnings about rising greenhouse gas (GHG) concentrations increased after the 1970s. With a better understanding of the role of the GHGs, efforts to reduce the anthropogenic emission of the gases have intensified in the last decades. Fig. 8.1 depicts the role and impact of the GHGs on world climate taking radioactive forcing (RF) as a measure of the balance of incoming and outgoing energy in the Earth-atmosphere system. This net gain of energy causes warming. In this figure, RF values are for changes relative to pre-industrial conditions defined at 1750 and are expressed in watts per square meter (W/m2). The annual greenhouse gas index (AGGI) introduced by NOAA is used to visualize the severity of climate change to policy makers and public opinion [4]. As a convention, the AGGI is normalized through the RF in the year 1990. It can be seen from Fig. 8.1 that the AGGI was 1.47 in 2020, which means that since 1990 the warming effect of the GHGs has increased by 47%. Fig. 8.1 also shows the relatively high and increasing share of the CO2 on climate. Approximately 2/3 of the total greenhouse effect was caused by CO2 in 2020. In terms of CO2 equivalents, the atmosphere contained 504 ppm of GHGs to which CO2 contributed with 412 ppm. The Intergovernmental Panel on Climate Change (IPCC) suggests that a concentration of CO2 alone at 550 ppm would result in an average increase of global temperature of around 3
C. According to IPCC ARG6 [5], the atmospheric CO2 concentrations in 2019 were higher than at any time in at least the past 2 million years (high confidence),

8.1 Introduction

3.5

1.6

other gases

Radioacve Forcing [W/m2]

2.5

N2 O

CFCs

1.2 1.0

2.0

CH4

0.8

1.5 0.6 1.0

0.4

CO2 0.5 0.0 1980

Annual Grennhause Gas Index

1.4

3.0

0.2

1985

1990

1995

2000 Year

2005

2010

2015

0.0 2020

FIGURE 8.1 Radioactive forcing and AGGI for various radioactively active gases. AGGI, Annual greenhouse gas index. Produced with data from NOAA/ESRL Global Monitoring Laboratory, The NOAA Annual Greenhouse Gas Index (AGGI), National Oceanic and Atmospheric Administration, Washington, DC, 2005.

and concentrations of CH4 and N2O were higher than at any time in at least 800,000 years. It is frequently reported that 2
C is a maximum temperature increase globally that can be tolerated for sustainable human life. Recent satellite data reveals that the Arctic is melting at an alarming rate. The end-of-season Arctic multilayer sea ice was roughly 50 cm thinner in 2021 than in 2019, showing a drop of around 16% in just 3 years. It is being replaced by less permanent seasonal sea ice that melts completely every summer [6]. The bar charts in Fig. 8.2 illustrate the cumulative CO2 emitted and to be emitted from 1850 until 2100 predicted in two different consumption scenarios based on multiple lines of evidence [5]. SSP12.6 (best case) predicted an approximate increase in terrestrial temperature of 1.8
C and SSP37.0 (mean worst case) predicted an approximate increase of 3.6
C. For the best-case scenario, a total of 3600 Gt of CO2 will be emitted by 2100 and, after the natural withdrawal by the ocean and land, approximately 1260 Gt will be released to the atmosphere which will contribute 1.4
C to the total global warming of 1.8
C. In the mean worst case scenario, the CO2 to be released into the atmosphere is estimated to be approximately 4500 Gt which will be responsible for 2.9
C of the total 3.6
C global warmings. With today’s measurements, it is obvious that global warming is rather approaching the worst cases scenario. To quantify the overall balance between GHG emitted and taken out of the atmosphere, the term “net zero emissions” is mostly used. The net negative CO2

327

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CHAPTER 8 Geologic carbon storage: key components

6° 0°

CO2, Gt

Only CO2 2.9°C

9000 8000

3.6°C

7000 6000

Atmosphere

6° 0°

5000

Only CO2 1.4°C

1.8°C

4000 3000

Atmosphere

2000

Ocean

1000

Land

Ocean

Land

0 SSP1-2.6 SSP3-7.0 Antropogenic CO2 Emissions aer Climate Scenarios of IPCC

FIGURE 8.2 Total amount of the CO2 emitted from 1850 until 2100 in two different scenarios. Reproduced with permission from V. Masson-Delmotte, P. Zhai, A. Pirani, et al. (Eds.), IPCC ARG 2021, Climate Change 2021: The Physical Science Basis, Cambridge University Press, Cambridge, In Press.

emissions occur when anthropogenic removals exceed anthropogenic emissions. It is a fact that even in case the transformation to renewables has been fully realized, the avoidance of CO2 emissions is impossible as significant process-bound industrial emissions (e.g., in the cement and chemical sector) will always persist. The only means to tackle this circumstance is anthropogenic removal of the CO2 (carbon dioxide removal, CDR). The latter gains in importance going one step ahead, i.e., pursuing the target of negative CO2 emissions. Several methods have been proposed for the removal of CO2 from the earth’s atmosphere. Fig. 8.3 gives an overview of principal removal methods and technologies. The concept of CO2 injection in the deep ocean to store it in dissolved or precipitated form was also supported by the idea of increasing their alkalinity by various measures. Reforestation is the most natural way of CO2 removal, and it was claimed that even without displacing agriculture and cities, the earth can sustain almost one billion hectares of new forests and that ideally would remove 25% of CO2 from the atmosphere [7]. Weathering and subsequent precipitation of Ca21 and Mg21 carbonates are the main processes that control the CO2 concentration in the atmosphere naturally. The enhanced weathering is affirmed to be a candidate for more CO2 removal and can be applied, e.g., by spreading finepowdered olivine on farmland or forestland to deploy the CO2 in the forms of carbonates. These enhanced natural forms of CO2 removal have unfortunately less

8.1 Introduction

OCEAN FERTILIZATION (OF)

DIRECT AIR CAPTURE (DAC)

BIOENERGY BIOMASS

METHANE REFORMATION

>20 Gta

CAPTURE ~2 Gta

(SS) SOIL SEQUESTRATION

STORAGE STORAGE ?

ENHANCED WEATHERING (EW)

Low (EW) – controversial (OF, SS) high cost (DOC) sequestraon

INDUSTRY CEMENT, ETC

FOSSIL FUEL POWER PLANTS

High emission, high storage capacity, high cost, lowemerging technical readiness levels

FIGURE 8.3 CO2 emission sources and carbon removal methods.

applicability because of economic and/or many side effects nowadays. The option of storing CO2 in ocean water has largely been abandoned since 2005 [8], because of high costs, low storage permanence, and considerable ecological impacts. Considering forestation, which is the cheapest one, land use is a major challenge regarding competing priorities, such as food production and conservation. In enhanced weathering, the negative side effects on the ecosystems are the main concerns. Direct air capture (DAC) is one of the most discussed methods in related literature [9]. However, on top of its high costs, the major obstacle is the need for a permanent storage solution for CO2 captured from the air. These CDR methods are estimated to capture and store a maximum of 2 Mta (Million tons per annum) of CO2. Power generation and various industrial sectors (e.g., steel and cement) belong to the major sources of anthropogenic CO2 which is needed to be captured. To avoid the release of this huge amount into the atmosphere, sequestration in geological formations seems to be the only feasible way with today’s technology. Approximately 62% of the global CO2 releases are emitted from large stationary point sources and are addressable with carbon capture and storage (CCS). The CCS technology can be applied to where a capture system can extract CO2 directly from an (exhaust) gas stream. The major amount of captured CO2 can only be removed permanently by injecting CO2 into geological formations, a process is also known as geologic carbon storage (GCS). From a theoretical

329

330

CHAPTER 8 Geologic carbon storage: key components

perspective, CO2 amounts higher than 20 Gt could be stored each year in geological formations. The idea of injecting CO2 into geological formations is not new. In the oil industry, CO2 is used as a solvent to mobilize the remaining oil in the reservoir. Following initiating works starting in the 1950s, a field application in the early 1970s was launched in the Kelly-Snyder field in Texas, USA [10]. Since then, CO2 flooding has been the most applied enhanced oil recovery (EOR) method for miscible and immiscible oil recovery processes. In other words, the oil and gas (O&G) industry possesses the required expertise and experience with injecting CO2 in geologic formations. The GCS was launched as part of CDR processes in which a relatively pure stream of CO2 from industrial and energy-related sources is separated (captured), conditioned (compressed and dried), transported, and finally injected into an appropriate geologic storage location for long-term isolation from the atmosphere. In 1989, the Massachusetts Institute of Technology initiated the Carbon Capture and Sequestration Technologies Program to investigate technologies to capture, utilize, and store CO2 from large stationary sources [11]. Koide et al. [12] assessed the CO2 storage capacity of sedimentary basins as 87 Gt with highly conservative assumptions. They also presented a preliminary design of injecting 5000 t/day of CO2 emitted from a thermal power plant into a selected geological formation over a period of 20 years. The scientific, as well as engineering studies, have been accelerated from this date on, accompanied by public and political support [13,14].

8.2 Geologic carbon storage classifications, definitions, types The techno-economic resourcereserve pyramid is mostly addressed in related literature to define the CO2 storage resources and capacities [15,16]. In this concept, the calculated capacity decreases gradually from the first-order theoretical capacity to the matched storage capacity (stored) at the top of the pyramid. Along with ongoing studies, researchers and authors from various disciplines proposed a wide range of terminology, definitions, and classification systems for the geological formations that can be used for CO2 storage [1721]. These mostly bear a significant resemblance to those of the petroleum industry, a fact which is expected given the industry’s long years’ experience with geologic definitions related to hydrocarbon resources. The petroleum resource management system (PRMS) issued in 2007 by Society of Petroleum Engineers (SPE) after a long preparation time was proposed as a model for a resource classification system to be used in GCS internationally [22]. Inspiring from the SPE PRMS and considering the initiatives of various international institutions such as the United Nations Economic Commission for

8.2 Geologic carbon storage classifications, definitions, types

Europe [23], SPE proposed the CO2 storage resources management system (SRMS). CO2 SRMS provides a measure for comparability and reduces the subjective nature of resource estimation with a consistent approach to assess storable quantities, evaluate development projects, and present results within a comprehensive classification framework [24]. In CO2 SRMS, the total storage resources are divided first as discovered and undiscovered. The discovered storage resources are then divided into stored, storage capacity, contingent storage, and inaccessible resources. Guidelines on the classification and categorization of CO2 storage resources are also defined. Fig. 8.4 illustrates the techno-economic resource-reserve pyramid (A) with the subdivision of GCS projects according to project maturity levels and (B) associated actions (i.e., business decisions) required to move a project toward commercial injection as proposed by SPE SRMS.

8.2.1 Definitions Relevant definitions of GCS partly based on SPE SRMS are discussed in the following to assure a systematic and standardized conceptualization. •

•

•

Resources: All quantities of naturally occurring pore volume (PV) potentially suitable for storage within the Earth’s crust, discovered and undiscovered (i.e., accessible, and inaccessible), plus those quantities already used for storage (i.e., stored). This definition draws the frame of the formations that can theoretically and practically be used for GCS. Stored: The quantity of discovered storage resources that have been exploited by a given date consisting of the cumulative quantity of CO2 injected and stored. Storage capacity: The quantities of total storage resources anticipated to be commercially accessible in the characterized geologic formation by application of development projects from a given date forward under defined conditions. Capacity must satisfy the following criteria: (1). It must be discovered and characterized (including containment), injectable, commercial, and not include CO2 previously-stored based on the development project(s) applied. (2). Proved capacity are those storable quantities that, by analysis of geoscience and engineering data, can be estimated with reasonable certainty to be commercially stored, from a given date forward, from known storable quantities and under defined economic conditions, operating methods, and government regulations. (3). If probabilistic methods are used, there should be at least a 90% probability that the quantities stored will equal or exceed the estimate. This amount is often referred to as 1P, also as proved. Additional capacities are defined as proved plus probable capacity (2P, with a probability of 50%) and proved plus probable plus possible (3P). 3P is equivalent to the high estimate scenario. When probabilistic methods are used, there should be at least a 10% probability that the actual quantities stored will equal or exceed the 3P estimate.

331

Theorecal Capacity

(A)

Commercial

CAPACITY

CONTINGENT RECOURCES Inaccessible

PROSPECTIVE RECOURCES Inaccessible

Uncertainty

Effecve Capacity

Sub-Commercial

Praccal Capacity

Undiscovered Resources

Matched Capacity

Discovered Recourses

Stored

Commerciality

CHAPTER 8 Geologic carbon storage: key components

Total Storage Resources

332

(B)

FIGURE 8.4 (A) The techno-economic resource pyramid for CO2 storage capacity and (B) SRMS classification of the GCS resources in terms of project maturity. GCS, Geologic carbon storage; SRMS, storage resources management system.

In addition, the following definitions are highly important for the development of GCS: •

•

Injectivity: In a general sense, the CO2 injectivity of a resource is the rate of the CO2 amount that a well can receive without causing any damage (e.g., fracturing) in the formation. This is a measure of the intake of a resource and a critical parameter for economics. To be evaluated in terms of capacity, there must be high confidence in the commercial injectability supported by actual injection or formation tests and confidence in the containment assessment. Containment: The ultimate objective of GCS projects is to permanently isolate the injected CO2 from the atmosphere. CO2 should be trapped in the geological formations by overlying impermeable rock strata formations and various physical mechanisms. In GCS jargon, the outflow of CO2 from storage formations into neighboring geologic formations and/or seabed (offshore) and/ or directly into the atmosphere through various pathways (e.g., wells, natural and induced geological features) is mostly called leakage. The leakage should be at an acceptable level and leak rates of 0.01% per year, equivalent to 99% retention of the stored CO2 after 100 years, is referred to by many stakeholders as adequate to ensure the effectiveness of CO2 storage. Moreover, the projected timeframe of safe containment ( . 1000 years) is one of the necessary criteria for estimating and identifying storable quantities.

8.3 Key components of geologic carbon storage projects

The capacity, injectivity, and containment (CIC) are the basic geological requirements of a GCS site and thereof screening parameters of higher relevance. •

Storage project: It represents the link between the storable quantities and the decision-making process, including budget allocation. In general, an individual project will represent a specific maturity level at which a decision is made on whether to proceed (i.e., spend money), and there should be an associated range of estimated storable quantities for that project.

A flowchart for the classification of storage resources based on the SRMS guidelines and terminology is provided by GCCSI [25]. Based on this classification and data from various countries, the stored capacity worldwide is given as 0.036 Gt whereas the total storage capacity is determined as 0.217 Gt. The resources including sub-commercial and undiscovered are added up to 13,000 Gt [25,26].

8.2.2 Geologic carbon storage types According to the original content of the geological formations, three types of GCS are most pronounced. These are storage in saline aquifers (SA), depleted/depleting hydrocarbon reservoirs (DHRs) (oil and gas), and enhanced coal bed methane (ECBM). The general characteristics of these GCS types in terms of storage as well as the estimated resources by various references [8,13,2528] are given as below. •

•

•

SA: Injecting CO2 into SA is known as the most potent way of GCS with high global capacities, however, with lower storage efficiencies per unit container volume when compared to DHRs. The global resources are estimated not lower than 10,000 Gt CO2, possibly up to 30,000 Gt. DHR: These are the oil (DOR) or gas (DGR) reservoirs that are currently in depletion or already depleted and can be used for GCS. The total worldwide resources are estimated to be up to 1000 Gt of CO2. CO2 can be injected into producing oil reservoirs to increase oil recovery (CO2 EOR). At the end of the lifetime of the reservoir, a part of the injected CO2 is produced back (in the best case recycled); however, the remaining CO2 after closing is considered stored. Injecting CO2 into producing gas reservoirs to enhance production is known to be inefficient (fast CO2 breakthrough) and not preferred. ECBM: This is the lowest capacity alternative for GCS. It relies on the desorption and displacement of methane by CO2 in ECBM formations. The global storage capacity is estimated to be no more than 200 Gt.

8.3 Key components of geologic carbon storage projects A GCS project consists of successive and complementary components as shown in Fig. 8.5. The terms CCS and CCUS are often used to summarize all GCS project components starting from the emission source and ending with encapsulating

333

Industrial/ Anthropogenic Emission CO2 Capture

CO2 Condioning (compression, pumping) Onshore Wellhead

Economics

CO2 Transport (Pipeline)

+

Offshore Plaorm

Cerficaon CO2 Transport (Ship)

Depleted Oil/Gas Reservoir Wellbore CO2 Injecon

Cap Rock (Containment)

CO2 Injecon

Saline Aquifer

Near wellbore

Monitoring

FIGURE 8.5 Schematic of the key components of GCS projects. GCS, Geologic carbon storage.

Well

8.4 Surface components: capture, conditioning, and transport

CO2 in the geological formation. With a few exceptions (e.g., ammonia/ethanol production and hydrogen generation plants), flue gas from various industrial sectors contains only small shares of CO2. To use the geological storage resource more efficiently and reduce the risks associated with injection, the CO2 should first be captured from the industrial flue gas using various methods. Subsequently, CO2 is conditioned for transport to the storage site by transforming it to the thermodynamic conditions selected for transport by pipeline or by ship (for offshore storage sites only). If transported via pipeline, highperformance (booster) pumps are needed to transfer the CO2 over long distances to the storage site. At its arrival at the wellhead, CO2 is injected via wellbore into the subsurface using the remaining energy at the pipeline outlet. If not sufficient, CO2 gets further pressurized before injection by intermediary pumps. CO2 flow-associated phenomena in the near-wellbore reservoir area might be challenging due to the thermodynamic contrast between arriving CO2 and native reservoir fluids. In the deep reservoir, CO2 is dispersed following the general rules of transport in porous media as a function of the dominating thermodynamics, petrophysics, and co-existing fluids. The storage of CO2 should be permanent. Therefore, the storage reservoir must be properly sealed through its confining formations. Special attention must be paid to the integrity of existing wells penetrating through the confining formations. A GCS project should have its validated economics throughout the full chain. The deployment always requires an approval—certification by a given authority confirming process design, storage safety, and capacity with a detailed monitoring strategy. The key components of GCS are described and discussed in the following sections in more detail.

8.4 Surface components: capture, conditioning, and transport 8.4.1 Capture In the full-chain industrial GCS operation, a high-purity CO2 stream ( . 95.5% concentration in the flue gas) is required as a byproduct of stationary industrial processes. As shown in Fig. 8.6, the most common sources of high-purity CO2 are hydrogen, ethanol and ammonia production as well as biofuel production through fermentation. Processes such as coal and cement production, which are major contributors to GHG emissions, release a much lower fraction of CO2 (#30%) in the total flue gas stream [29] where CO2 must be captured through capture and purification processes before being transported and then injected into geological formations considered for CO2 storage. The use of carbon separation in industrial processes dates to the 1930s when chemical solvents such as amines in aqueous solutions were used to capture CO2 from natural gas (NG). In the capture technology, various physicochemical principles are applied for the separation of CO2 from the whole gas stream. The

335

100

CO2 in gas stream

12

90

CO2 in Gas Stream [%wt]

70

10

8

60 50

6

40 30 20

4

2

CO2 Emission [Gta]

CO2 emission 80

10 0

0

FIGURE 8.6 Principal GHG emitting stationary industries, their CO2 emissions per year (average 2020), and the percentage of CO2 in the total emissions. Produced with data from Refs. [25,2831].

8.4 Surface components: capture, conditioning, and transport

definitions, as well as some drawbacks and application examples of mostly used and emerging separation processes, are summarized in Table 8.1. Absorption with chemical solvents (which use chemical bonds to capture CO2) or physical solvents (which use only the intermolecular Van der Waals force to capture CO2) is the oldest and most common method. Basically, the gas stream is dissolved in a liquid solvent and forms a solution. Subsequently, CO2 is released (solvent is regenerated) from the solution by changing pressure or temperature. Due to the different solubilities of the gas components in a particular solvent, the solvent can be used for selective separation. At lower CO2 partial pressure at the outlet, chemical solvents are more attractive for use as they have a higher absorption capacity. At higher partial pressure at the inlet, physical solvents are preferred, as the relationship between solvent capacity and partial pressure follows Henry’s Law (linear relation). In the solvent regeneration process, chemical solvents are usually regenerated by raising the temperature to release CO2. For physical solvents, the pressure is reduced [30,31,55]. These solvents operating at large-scale facilities can separate up to 4000 t/day of CO2 in synthetic gas (syngas) purification and NG processing. Chemical solvent-based systems that are commercially available or about to be introduced to the market are typically based on amine-based solvents such as those used in hydrogen manufacturing units (HMU) [32]. There are efforts to reduce the cost and energy requirements of chemical solvent technologies. In this framework, switching from aqueous to low-water solvents is also evaluated as a measure to increase mass transfer rates in normal chemical solvents with amines [56,57]. Adsorption is another CO2 capture process based on the physical and chemical bonding between CO2 and the surface of the adsorbent. CO2 from the gas stream is adsorbed onto a solid. Subsequently, CO2 is released by changing the pressure or temperature (the adsorbent is regenerated). Chemical bonding results in a strong interaction between the gas molecule and adsorbent and is mostly preferred for low-concentration gas streams. When the CO2 concentration is insignificant, temperature swing adsorption is often used in which the adsorbent is regenerated by raising its temperature to liberate the CO2 [42,58]. Physical adsorption has a weaker interaction between the gas molecule and sorbent and is typically applied to high CO2 concentration feed streams. Some well-known physical adsorbents are zeolite and amine sorbents [59]. A relatively new technology for CO2 capture is the use of membranes. CO2 dissolves in the selected membrane and diffuses at a rate proportional to its partial pressure. Two main membrane technologies are used currently. For the nonfacilitated technology, more energy is consumed as compression work to achieve the required capture efficiency. The capture efficiency also depends on the permeability of the membrane chosen. Its use is feasible in the capture from NG where the CO2 partial pressure is high. Although it has some advantages, such as low environmental impact, it could be challenging to integrate into existing power plants. The second type of membrane separation is the facilitated one, which consists of mobile or liquid phase carriers that facilitate the transport of CO2 as

337

Table 8.1 Gas separation technologies for CO2 capture from various industrial processes (Technology readiness level, TRL according to EU classification). TRL

Pros/challenges

Examples

The amine solvent reacts reversibly with CO2, forming water-soluble salts. (e.g., MEA, DEA). TRL 5 9

Mature technology for high rate, low PCO2 streams. Energy-intensive, high CAPEX and OPEX. Potential environmental impacts, equipment corrosion, and generation of volatile degradation compounds are other challenges [31] The selexol process operates at around ambient T, whereas the Rectisol process operates at T as low as 260
C. Preferred at high CO2 concentrations (25%70%) and under higher P (B100 bar) Higher absorption capacity, rate, selectivity, and degradation resistance over conventional amines. PCCCT [34] Regeneration can take place at P 5 30 bar, resulting in significant savings in energy, no thermal and oxidative degradation and high CO2 output, is used in post-C carbon-capture process [36,37] Use conventional infrastructure and have acceptable water tolerance and faster CO2 mass transfer than their aqueous counterparts. Higher solvent costs and higher intrinsic viscosities; to be tested for large-scale CO2 separations PSA is considered viable for separation of CO2 from flue gases containing about 5%15% v/v; is useful because of its short temporal need for regenerating the adsorbent [40]

CO2 separation from NG (Snovhit, Sleipner), coal combustion, and HMU (Shell Quest project [32])

Absorption with physical solvents (like Selexol, Rectisol); Industrially mature. TRL 5 9

SHA amino group attached to a tertiary carbon. TRL 5 79 CAP: Absorbing CO2 from the flue gas of a power plant at low temperature (0
C20
C) using ammonia as solvent. TRL 5 67

Absorption with water lean solvent: Chemical selectivity with step changes, efficiencies achieved by lower specific heats of organics. TRL 5 47 PSA/VSA: A cyclic process, allowing continuous separation. TRL 5 89

Large-scale facilities, H2 and NG production, (Exxon Shute Creek NG processing plant [33])

Petra Nova coal plant carbon capture project, USA [35] Mongstad, Norway using flue gas streams with high (16%) and low (3.6%) CO2 concentrations [38]

Commercial scale FEED studies: CHN Energy’s Jinjie pilot plant targeted for post-combustion separations [39]

Air Products Port Arthur SMR CCS gas separation in hydrogen production [41]

TSA: The saturated adsorbent is heated releasing adsorbed reactants. TRL 5 57

Gas separation membranes for NG processing with Knudsen diffusion principle. TRL 5 89 Polymeric membranes: Reactive and nonreactive membranes with various selectivity. TRL 5 67 ECM integrated with MCFCs. TRL 5 67

Cryogenic separation: Using distillation at very low T and high P. TRL 5 57 Polymeric membranes/Cryogenic separation hybrid. TRL 5 45 CaL: CO2-rich flue gas flows into the carbonator [52]. TRL 5 67

When the CO2 concentration is insignificant, TSA is often used but when the CO2 concentration is high PSA is preferred; CO2 purity 48%93%v [4244] Compact membrane system and high selectivity of amine-based absorption process; low environmental effect and degradation. The main challenge is the resistance on the membrane The modified membranes can reduce the processing cost and increase permeability with better stability at higher T New process to verify that the ECM can achieve at least 90% CO2 capture from the flue gas with no more than 35% increase in the cost of electricity Energy-intensive due to the operating T range. Formation of ice in pipeline system. Complex to operate Several compression applications at ambient T and P, suitable for producing liquid CO2 and ideal for high CO2 concentrations Suitable for retrofit, uses a cheap sorbent (limestone); CO2 purity is lower. The sorbent is vulnerable to decay and competing reactions causing chemical deactivation

Large pilot tests to FEED studies for commercial plants [28,45]

Petrobras Santos Basin Pre-Salt oil field [46]

FEED studies for large pilots [47]

Pilot at Plant Barr CEPP [48]

Pilot scale projects [49,50]

Large scale pilots [47,51]

Feasibility for using in CFPP [53], competitive in OFCP [54]

CaL, Calcium looping; CAP, chilled ammonia process; CEPP, combined electric power plant; CFPP, coal-fired power plant; DEA, di-ethanolamine; ECM, electrochemical membrane; FEED, front end engineering design; MCFCs, molten carbonate fuel cells; MEA, monoethanolamine; OFCP, oxyfuel cement plant; P, pressure; PCCCT, post-combustion carbon capture technology; PSA, pressure swing adsorption; SHA, sterically hindered amine; T, temperature; TRL, technology readiness level; TSA, temperature swing adsorption; VSA, vacuum swing adsorption.

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CHAPTER 8 Geologic carbon storage: key components

bicarbonate and support the permeability and selectivity of CO2 through the membrane. The mixed matrix membrane is a new type of technology that uses polymer membrane filters such as zeolite, mesoporous silica, and zeolitic imidazolate [51,60,61]. It is claimed that these modified membranes reduce processing costs with better stability at higher temperatures. The cryogenic separation approach involves several compression applications for separating the gas. CO2 is cooled until it becomes a solid that separates it from the gas stream. This technique is rather suitable for producing liquid CO2 and ideal for CO2 capture in high concentrations [51]. Its advantages include less water consumption, cheap chemical agents, corrosion-resistant and less impact on the environment. The main drawback is energy-intensive due to the operating temperature range [62]. Also, ice formation in the cryogenic approach often leads to clogging of the piping system, causing pressure drop and safety issues. In the calcium looping (CaL) process, CO2 reacts with calcium- or magnesium-bearing rocks to form magnesite or calcite at around 650
C. CO2 is converted to a solid substrate that can be reused as a building material or disposed of in surface facilities. Current challenges are the rate of reactions and mass of reactants (e.g., source of Mg, Ca), increased fossil-fuel consumption for required energy, disposal, and storage of materials, and impacts of mining for minerals used in the carbonation reactions [54]. The theoretical energy required for the capture of CO2 is significantly lower than the energy consumed in actual industrial capture. The energy efficiency of the capture technologies by eliminating the parasitic energy load is therefore one of the most relevant screening parameters for target use. Various studies have compared the different CO2 capture processes with their advantages and disadvantages for potential use in related industries and also provided improvement perspectives [8,30,63]. There are three types of deployment technologies for carbon capture from fossil fuel power plants: •

•

Pre-combustion (Pre-C): This is the process where CO2 is removed before the actual conversion of fuel to energy. This may involve reforming (gas) or gasifying (coal) the fuel to synthesize gas and using water gas shift to produce more hydrogen from water. The resulting gas stream contains more than 98% vol. CO2. The technology is mature and is being used commercially at the required scale in some industries. For coal-fired power plants, efficiency and cost are generally lower than for post-combustion capture. Some weaknesses of this technology are: (1) extensive supporting systems requirements, (2) temperature associated heat transfer problem, (3) efficiency decay issues associated with the use of hydrogen-rich gas turbine fuel, (4) high parasitic power requirement for sorbent regeneration, (5) inadequate experience due to few gasification plants currently operated in the market and (6) high capital and operating costs for sorption systems [64]. Post-combustion (Post-C): CO2 is captured from the flue gas released from industrial combustion or power stations. The resulting gas stream contains

8.4 Surface components: capture, conditioning, and transport

•

CO2 higher than 98 vol.% with high water levels (up to 5 wt.%) before compression. After compression, water levels are likely to be less than 2200 ppm. Further drying is mostly required. This is the most preferred option for retrofitting existing power (coal-NG) plants but can be integrated into new plants. Steam reforming is the dominant technology for hydrogen production today and a good example of post-combustion. Technology is more mature than other alternatives. Both adsorption and absorption, as well as membranebased technologies, can be applied. However, absorption processes based on chemical solvents are currently the preferred option. The major challenge is its large parasitic load [64]. Oxy-fuel combustion (OFC): This technology is based on burning fuel with pure oxygen instead of regular air and is preferred in coal-fired and gas-fired plants. The resulting CO2 purity is in general lower than pre- and post-combustion. Coalbased oxy-firing typically produces very high-water content (up to 20 wt.%) as well as higher SO2 and NOx contents. Large cryogenic oxygen production requirements may be cost-intensive. CO2 is recycled to the compressor to provide the expansion medium, instead of air. The potential for advanced oxygen separation membranes with lower energy consumption is reported [65].

The “chemical looping combustion” technology for coal-gasification plants is still under development. CO2 is the main combustion product in this technology, which remains unmixed with N2, thus avoiding energy-intensive air separation. The identification and analysis of promising CCU technologies, including their regulatory aspects for potential application, are investigated alongside the assessments on GCS [66]. The CO2 capture efficiency is defined as the ratio of the difference between the mass at the inlet and the outlet of the capture unit to the mass at the inlet. This depends obviously on the gas composition emitted from the source and on the efficient separation and capture technologies applied. The technology to be applied for the separation and capture of CO2 from the industrially applicable processes is selected based on various criteria. Some of the screening criteria for industrial applications can be listed as follows [28,31,32,64,67]: •

•

The CO2 source emission capacity and utilization: Both the CO2 source (CO2 concentration, the pressure of other gases in the stream, etc.) and the utilization of the captured CO2 (purity and thermodynamics) play an important role in the selection of the technology. Operability and reliability requirements focus on the impact of CO2 capture on the operational reliability of industrial equipment (e.g., power plants, cogeneration plants). Economic considerations: Capital (CAPEX) and operating expenditures (OPEX) should be minimized and optimized with ongoing operating costs over the 25 1 year life of the capture facilities especially if there are opportunities for synergies with existing treatment facilities in terms of utilities (power and steam) and chemical requirements.

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•

• • •

Footprint and overall size: This parameter can be critical especially if an expansion (upgrade) of the facility is in planning; this should be considered in the design, engineering, and construction of the capture facility. Maturity of the technology: Only the technologies with demonstrated proof of concept (PoC) are to be considered. Construction issues: These include space requirements, transportation and setups of components, and applicability of modularization. HSE risks: A broad perspective on health and safety issues should be considered in the technology screening process, including local and site-wide impacts

Example The ROAD project design applied post-C technology to capture the CO2 from the flue gases of a new 1069 MWe coal-fired power plant based on the conventional amine solvent MEA process. The capture unit has a design capacity of 250 MWe equivalent. During the operational phase of the project, approximately 1.1 Mta of CO2 would be captured and stored, with a full-load flow of 47 kg/s (169 t/h) of CO2 [68]. The industrial application of various capture technologies is reported in related reports [28]. Once CO2 is captured from the emitted gas it should be transported to the storage site. A reliable, safe and economically feasible CO2 logistic is a key feature of any GCS project. The CO2 logistic has normally two main pillars: transport and the physical and/or chemical processing of CO2-rich gas for transport and injection (conditioning). Using a power plant as an example, the logistical processes from the emission source to the injection site including conditioning and transport are shown in Fig. 8.7.

8.4.2 Conditioning Except for short distances and low flow rates, CO2 is generally transported in dense phase (liquid or supercritical) in pipelines and/or in ships. Therefore, the CO2, which is mostly in the gas phase at near atmospheric pressure (CO2-gas is usually available at pressures of 1.32.0 bar) by leaving the emitter, should be compressed to bring it into a dense phase. The capture and separation based on various processes increase the CO2 percentage in the range of 9598 vol.%, with impurities mainly water, O2, N2, H2S, and other gases as the by-product of the absorption/desorption process applied in the capture. However, purity levels above 95.5 vol.% CO2 and 99.7 vol.% are recommended for pipeline and ship transport, respectively. Thus, a further purification process following the capture should be applied in the compression stage. From capture to the injection potential components of conditioning the CO2 stream for transport and injection are: • •

Compression Purification

Mul-stage Compression

C1 Emission Source

C2

Pumping

Pipeline

C3

Capture Compression (C)

Cooling

Ship Liquefacon (Expansion)

Purificaon Removal of water & other impuries (SO2/H2S/etc) FIGURE 8.7 Schematic of conditioning and transport processes in a GCS project from the emitter to the wellhead.

Pumping and/ or heang

Injecon

CHAPTER 8 Geologic carbon storage: key components

• •

Liquefaction (in the case of ship transport) Pumping

The compression of CO2 is one of the most important, complex, and expensive components of the GCS. Its design requires high expertise and experience. The main design parameters are the inlet and outlet pressure as well as the inlet flow rate. Depending on these parameters, various types of compressors are available in the related industry. Fig. 8.8 shows a typical chart for the selection of fitto-purpose compressors. In GCS operations, centrifugal multi-stage and integrally geared compressors are preferred based on the flow rate and pressure ranges. The principal benefit of integrally geared compressors is the high compression ratio achievable in one machine, using multiple intercoolers to minimize power consumption with lower footprints. If the inlet pressure of CO2 into the compressor station is low (which is mostly the case), then the compression needs more than one stage. CO2 needs to be transformed from atmospheric conditions to conditions for pipeline transport in a dense phase, three or more stages are mostly necessary. Compression increases the temperature of the gas and thus the fluid must be cooled after the compression [69,70]. Without proper filtration of residual liquids and fine particles, compressor performance decreases. Before the CO2 enters the compressor, an intake scrubber is generally used to remove residual liquids in the stream. The separated CO2-rich gas is sent to compression and cooling, where water and other impurities such as 700

400 300 200 100

Rotary Screw

0 1.E+01

Reciprocang Mul-Stage

500

Centrifugal Single-Stage

600 Discharge Pressure [bar]

344

Centrifugal Mul-Stage

Integrally Geared

1.E+02 1.E+03 Inlet Volume [m3/min]

1.E+04

FIGURE 8.8 Main design parameters and available compressors for compression of CO2-rich gas streams.

8.4 Surface components: capture, conditioning, and transport

N2, O2, and H2S are removed by various physical processes, including adsorption and condensation. The cooling is usually achieved using an aerial cooler. After cooling, a liquid phase may form. This liquid is removed in intermediate scrubbers or knockout drums. If additional water drops out, several suitable technologies for CO2 dehydration like glycol-based systems utilizing triethylene glycol (TEG) and solid adsorption systems using the molecular sieve can be applied. This stage should bring the CO2 stream to the standards defined for transport, which is by pipeline for onshore plants and by pipeline and/or ship for offshore plants. The quality of the CO2, both in physical (temperature and pressure) and in chemical (the content of impurities) terms, must be defined in terms of the requirements for transport and storage (see the following section). The low suction pressure is not a major concern in the design of the compressor other than the need to be very conservative in the sizing of piping, valves, and vessels. The compressor performance should be examined for a range of pressures around the intended suction pressure. In this respect, speed changes (minimum and maximum), compressor internal temperatures with volume pockets in operation, and possible gas recycling are to be considered. The compression generates heat and discharge temperatures should be limited to about 150
C and should never exceed 180
C because of potential damage to the compressor. In addition, high temperatures can cause degradation of the compressor fluids as well as increase the corrosive damage to the compressors. A reliable evaluation of the phase behavior [pressure, volume and temperature, (PVT)] of CO2 including the impurities, is an essential prerequisite for the proper design of the compression technology and a number of the stages to be used. This is especially important at temperatures lower than 50
C where the compressibility factor (Z) of CO2 is extremely low, creating higher sensitivities to pressure and temperature. To prevent equipment damage due to the corrosion of pipes and the solidification of moisture during cooling, the concentrations of water, H2S, O2, SOx, and NOx are strictly controlled and included in the calculations. O&G industry relies on two main cubic equations of state (EoS): PengRobinson (PR) and SoaveRedlichKwong (SRK) with and without Peneloux correction [71]. Phase behavior of pure CO2 is best represented by the Span and Wagner EoS which is based on the explicit formulation in the Helmholtz free energy [72]. It was extended to gas mixtures, resulting in GERG-2008 EoS [73] and providing a better prediction of saturation pressures and more accurate density of the mixtures for most impurities regardless of mixture phases in comparison to PR and SRK EoS [74,75]. GERG-2008 is used as a reference model for the calculation of thermodynamic properties of NG by the NIST (The National Institute of Standards and Technology) reference material database. For practical engineering design applications, it is important to consider the impact of impurity type on phase behavior in comparison to single component CO2. Fig. 8.9 shows an example of phase envelops for binary mixtures of 95% CO2 with inert (5% N2) and hydrocarbon (5% C3H8) gases and compressibility factor (Z) of the same mixtures compared to pure CO2 near the critical temperature.

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CHAPTER 8 Geologic carbon storage: key components

100

1.0

90

0.9

T=31.5 °C

80

Compressibility Factor, Z

C

95% CO2 + 5% N2

60 50 40 95% CO2 + 5% C3H8

30

100% CO2 95% CO2 + 5%C3H8

100% CO2

70 Pressure [bar]

346

0.8

95% CO2 + 5%N2

0.7 0.6 0.5 0.4

20

0.3

10

0.2

0 -40

-32

-24

-16

-8

0

8

16

24

32

40

0

50

100

150

200

Pressure [bar]

Temperature [°C]

(A)

(B)

FIGURE 8.9 (A) Phase behavior and (B) compressibility factor, (Z) of CO2 in comparison with CO2 mixtures with N2 and C3H8 (propane) calculated with PengRobinson EoS near critical temperature of CO2.

In addition to reliable phase behavior prediction of the CO2 stream with impurities, the design of a compressor necessitates the application of the first law of thermodynamics, the conservation of energy. The work required to compress the gas between inlet (i) and outlet (o) is calculated as follows: W_ 5 m_ ðho 2 hi Þ 2 Q_

(8.1)

where W_ is the work flow (kW), h is enthalpy (kJ/kg), m_ is the mass flow rate (kg/ s), and Q_ is heat flow (transfer) (kW). As the heat transfer rate is small in comparison to the other terms in the equation, it is typical to assume that the compressor is adiabatic (Q_ 5 0). As the case with many equations in thermodynamics, the equation appears to be simple at first glance. However, in practice, it is quite demanding. The ideal case is that a compressor can operate isentropically resulting in the entropy of the stream remaining unchanged upon compression. The compression ratio for a compressor is defined as the outlet pressure divided by the inlet pressure. By definition, the compression ratio is always greater than one. CO2 compression uses mature technology usually applied in large-scale fertilizers manufacturing plants (i.e., production of urea) and similar technology is also used in NG pipeline transport worldwide. The main additional operating issues for CO2 compression are avoiding corrosion and hydrate formation [76]. To avoid potential issues with the heat exchanger, it should be preferred to use stainless steel throughout the compressor piping if H2S is present in the CO2-rich gas stream. Example Integrally geared compressors are in service for CO2 sequestration in Bismarck, North Dakota, USA (125 t/h, 1.15188 bar, 1.35 Mta) and in Norway with 1 Mta CO2 for injection into Sleipner saline reservoir. Suction volumes of integrally geared compressors are 10,000120,000 m3 /h. For the required energy, compressing CO2 from 1.3 bar to 150 bar is approximately 0.12 kWh/kg.

8.4 Surface components: capture, conditioning, and transport

8.4.3 Transport In GCS, CO2 is transported to the storage site via pipelines and/or ships. For shorter periods and lower volumes, road and rail tankers can be competitive. The tank trucks with trailers can operate under cryogenic conditions to transport up to 30 t under typically 17 bar and 230
C. Rail transport can carry larger volumes (up to 60 t each wagon) under a pressure of approximately 26 bars. Although these transport methods are used in the food industry, they cannot sustain a continuous and meaningful CO2 storage injection into a geological formation. O&G industry is familiar with both pipeline and ship transport for decades. Large-scale pipeline transport of CO2 is not a new technology; pipelines are used for EOR purposes in both the USA and Canada. Also, CO2 is separated from NG, transported with onshore or offshore pipelines, and injected for final storage, e.g., in Snohvit, Sleipner (Norway), and at In Salah (Algeria) sites [7779]. The main driver of CO2 transport design is density. Fig. 8.10 depicts the density of CO2 as a function of pressure for some relevant isotherms. The ranges of pipeline and ship transports are also highlighted in the figure. It can be concluded that the high density of CO2 in the dense (liquid and supercritical phase) favors the transport of high CO2 amounts by both ships and pipelines. It should be noted that despite its high density, the viscosity of CO2 is comparable with most of the gases being in the range of 0.040.15 mPa.s in dense phases, being slightly lower in the supercritical phase. In some limited cases where flow rates are low and/or, the low injection pressure is required at the injection wellhead, transport of CO2 in the gaseous phase via pipelines might be preferred.

1200 Liquid state 1000

Supercrical state

PIPELINE – LIQUID, SUPERCRITICAL CO2

Density, [kg/m3]

800

31°C

600

-30°C

400

-10°C

50°C

C; 30.9°C, 73.8 bar

10°C

70°C

90°C

110°C

Two-phase state

200 Gas state 0 0

20

40

60

80

100 120 Pressure, [bar]

140

FIGURE 8.10 Thermodynamics of CO2 preferred in pipeline and ship transport.

160

180

200

347

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CHAPTER 8 Geologic carbon storage: key components

For transporting large quantities of CO2 over long distances on land, pipelines are known to be the most viable method. Whether at dense (liquid or supercritical) or gas phases, to assure a single-phase flow along the pipeline from source to injection facility is extremely important, as a phase transition from dense to gas or vice versa in the pipeline would create operational and material challenges, especially in relation to pumping/compression. This requires maintaining a buffer zone between the operating region and the two-phase envelope. The typical (but not limited) range of pressure and temperature for a CO2 pipeline is between 85180 bar and 10
C44
C, to ensure a stable single-phase flow through the pipeline. The transport in the liquid phase has some advantages over supercritical transportation, such as higher density, lower compressibility, and less (pumping) energy requirements [80,81], yet design in the supercritical phase has become standard practice. The transition of CO2 from supercritical to liquid phase due to decreasing temperature in the pipeline is not seen as a challenge compared to liquidgas transition. For a given length and flow rate, the main design parameters are the diameter of the pipeline with selected material and pressure at the inlet and outlet which requires the definition of the compressions stations and booster pumps to maintain the required pressure. Larger diameter pipelines are more expensive; however, allow for lower pressure drop for the required flow rate also reducing the number of compressions—pumping stations. Pressure drop is calculated as a function of pipeline geometry (length and diameter) and roughness (friction factor), as well as flow rate, temperature, heat exchange with surroundings, static head resulting from positional differences (topology) on the pipeline path, and CO2 properties, especially compressibility. The maximum allowable operating pressure, minimum operating pressure, pipeline length, and operating temperatures are required in the initial design phases of a pipeline. The maximum pressure relates to the integrity of the pipeline and is usually determined by design specifications, regulations, and economic considerations. The minimum pressure should be set to avoid two-phase flow during flow in the pipeline. The pipeline length is governed by the route chosen between the CO2 source and the CO2 sink, which is normally dictated not only by topology and economics but also by regulations based on considerations of health and safety. According to the recommended practices for the design and operation of CO2 pipelines, the optimum wall thickness (tw, m) of the pipeline is calculated with the following equation [82]: tw 5

Pmo :d2 2SEF

(8.2)

where Pmo is the maximum operating (allowable) pressure (bar), d2 is the outside diameter of the pipeline (m), S is the specific yield stress of pipe material (bar), E is the longitudinal joint factor (51.0) and F is the design factor (50.72). The presence of impurities in the CO2 stream can have a significant effect not only on phase properties but in turn influencing CO2 density and viscosity. Its extent depends on the type, quantity, and combination of impurities present.

8.4 Surface components: capture, conditioning, and transport

Some combinations cause higher pressure and temperature drops for a given pipeline length than others, especially if hydrogen or nitrogen is present. This in turn has implications for the design of the pipeline, e.g., distances between booster pumping stations along the pipeline required to keep the pressure sufficiently high to maintain a dense phase. The pipeline CAPEX and OPEX increase with the incidence of intermediate booster stations which, in any event, are not viable for subsea pipelines. Therefore, as in the case of compression design, reliable prediction of the phase behavior is a necessity in pipeline design as well [8385]. Particularly, H2O and O2 can have excessive negative impacts on pipelines including fracture propagation, corrosion, non-component deterioration, and the formation of hydrates and clathrates. The presence of H2O concentration above 50 ppm may lead to the formation of carbonic acid inside the pipeline and cause corrosion. Hydrates may also affect the operation of valves and pumps. The estimated values of corrosion on the carbon steel commonly used for pipeline construction can be up to 1 mm/year, which necessitates the consideration of corrosion allowance in the pipeline design. In case of potential leakage from a pipeline, the content of H2S and CO may have to be below 200 and 2000 ppm, respectively, due to health and safety considerations [86,87]. These amounts should be below the threshold for the short-term exposure limits (STEL) that gives the maximum amount of a compound that one can be exposed to for 15 minutes without adverse health effects. Several CO2 specifications and recommendations for maximum impurity concentrations have been published [7887] and an excerpt showing the tentatively most aggressive impurities (H2O, H2S, O2, NOx, SOx, CO) is shown in Table 8.2. In addition, to ensure proper transportation and storage of the CO2, it is necessary to consider the requirements in that respect. However, since the concentration of a single impurity depends on the interplay with the concentrations of other impurities, it is not possible, due to lack of data and current understanding, to state a fixed maximum concentration of a single impurity when other impurities are, or may be, present. All common impurities (CO, N2, CH4, NO, O2 and H2S) studied are found to increase pressure loss, while H2O and SO2 cause a decrease in pressure loss [90]. The common impurities increase the critical pressure of the CO2 fluid above 73.8 bar. The HC impurities (such as CH4, C2H6, C3H8 as well as H2S and SO2) increase the critical temperature, while all common impurities decrease the critical temperature. In general, no impurity is desirable because they may create a two-phase region and additional corrosion concerns [91]. For a long time, no standardized materials and procedures have been applied in the transport of CO2. ISO 27913 (2016) specifies requirements and recommendations not covered in existing pipeline standards for the transportation of CO2 streams from the capture site to the storage facility [89]. The system boundary between capture and transportation is the point at the inlet valve of the pipeline, where the composition, temperature, and pressure of the CO2 stream are within a certain specified range determined by the capture or capture-compression processes to meet the requirements for transportation. The boundary between

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CHAPTER 8 Geologic carbon storage: key components

Table 8.2 Standards that are mostly applied in pipeline and ship transport of CO2. For pipeline transport Impurity H2O H2S CO CH4 N2 O2 Ar H2 CO2

Specification 100 ppm 100 ppm 35 ppm

10 ppm

Max allowable 500 ppm 200 ppm 2000 ppm Aquifer ,4 vol.% EOR ,4 vol.% ,4 vol.% Not well defined ,4 vol.% ,4 vol.% .95.5 vol.%

Reason Corrosion, hydrate Health and safety Health and safety Costs, energy Costs, energy, performance Costs Corrosion, challenges in storage Costs Costs, energy Economy, operational efficiency

50 ppm 200 ppm 2000 ppm 0.3 vol.% 0.3 vol.% Not well defined 0.3 vol.% 0.3 vol.% .95.7 vol.%

Corrosion, icing on heat exchanger Health and safety Health and safety Costs for liquefaction, hydrate Costs for liquefaction, hydrate Costs for liquefaction, hydrate Costs for liquefaction, hydrate Costs for liquefaction, hydrate Dry ice formation (hydrate)

For ship transport H2O H2S CO CH4 N2 O2 Ar H2 CO2

30 ppm 9 ppm 100 ppm

5 ppm

EOR, Enhanced oil recovery. Produced with data from Refs. [86,88,89].

transportation and storage is the point where the CO2 stream leaves the transportation pipeline infrastructure and enters the storage infrastructure. ISO 27913 also includes aspects of CO2 stream quality assurance, as well as converging CO2 streams from different sources. HSE aspects specific to CO2 transport and monitoring are considered. An essential requirement for the efficient and safe design and operation of a CO2 pipeline is the accurate transient flow modeling of the fluid phase and composition of the CO2-rich mixture. Routine analyses are required to verify that the CO2 stream compositions comply with the approved CO2 specifications for the pipeline transportation network and the storage site. The measurement of CO2 flow in GCS pipelines is more challenging than metering the oil, gas, or multiphase flow in the air and gas industry, due to the readily varying physical properties of CO2. Metering orifice and venturi meters are used as state-of-the-art measurement technologies [92].

8.4 Surface components: capture, conditioning, and transport

Examples For storage in an offshore DGR in the Netherlands, the pipeline operating at 80120 bar and up to 80
C requires an 8-stage compressor for conditioning onshore. For storage in another DGR in the Netherlands, the pipeline at 1922 bar and ambient temperature (10
C30
C) requires a 4-stage compressor [68]. In the Porthos project (Netherlands), the shore-based pipeline has the capacity of 5 Mta CO2 at 40 bar (gas CO2) [93]. A 150 km seabed pipeline transports dense CO2 from the onshore capture-compression station to the offshore Snohvit (Norway) GCS site. A list of CO2 pipelines is summarized by Noothout et al. [94]. There is more than 46,500 km of CO2 pipelines worldwide both onshore and offshore [95]. In the ships, due to material properties, the CO2 should be a liquid phase in a lower pressure range. Ship transport is considered at low (68 bar), medium (15 bar), and high-pressure (4560 bar). The practice of transporting liquefied gas, such as LPG (liquified petroleum gas) and LNG (liquified NG), is well established in the O&G industry and they can be carried in fully refrigerated, semipressurized, or fully pressurized ships [96]. Recently, there is a consensus on the low-pressure-based transport approach being the techno-economic optimal transport condition [9799]. In fully refrigerated LPG and LNG ships, the cargo is maintained in the liquid phase at near atmospheric pressure solely by refrigeration (low-pressure transport). Semi-pressurized ships operate under conditions approaching the triple point (where the density is the highest), while the fully pressurized ships operate closer to the critical point (high-pressure transport). For a fixed volume vessel, almost twice as much CO2 can be transported at low pressures near the triple point than at high pressures near the critical point. For smallscale ship transport, the CO2 is liquefied and transported in semi-pressurized vessels at 1420 bar. The liquefaction process is needed for ship transportation. This is achieved by expansion after the compression to 20 bar and CO2 is condensed in a closed refrigeration cycle (propane or NH3) at temperatures of 225
C to 230
C. This process is used in several plants up to capacities of 250,000 tons per year. The power consumption is approximately 150 kWh/t of liquid CO2. For large-scale liquefaction, an open cycle using CO2 as a refrigerant can be applied. In addition, water and volatile gases are removed by means of gas scrubbers and an adsorption gas drier and in a column, respectively. The process may deliver CO2 at 6.5 bar and 252
C to the storage tanks before the transport in ships [100,101]. Ship transport is subjected to stricter specifications than pipeline transport. The water content should be lower due to possible water freeze-out during the liquefaction process. The main difference is that liquid CO2 cannot contain more than roughly 0.3 mole% of volatiles such as N2, as a higher content will increase the liquefaction costs and decrease the temperature for the same pressure, thereby increasing the risk for dry-ice formation. Similar to pipeline transport, the reliable prediction of phase behavior is important in designing ship transport [102].

351

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During normal filling and unloading of the ships, the pressure in the storage vessels is maintained by gas unloading/injection conditions. This implies that the net transport is defined from the liquidgas density difference at transport conditions. For the food-grade conventional liquid CO2 transport, it is common practice to fill the vessels using subcooled liquid CO2 to prevent pressure creep during transport. This practice is more complicated for low-pressure transport. Controlled or accidental depressurization might be accompanied by the formation of solid CO2. Special consideration is therefore required for the vent system used. It is concluded that the current design codes for LPG and LNG tankers do not cover the transport of liquid CO2, as the pressure and density ranges are different [89]. Such codes need to be developed. Functional requirements and optimization of the CO2 shipping chain, with a focus on the offshore offloading system using a range of typical North Sea reservoirs (SAs or DHRs), is discussed by Neele et al. [103]. The results provide insight into the requirements for offshore offloading from a ship into an injection well for a range of potential North Sea storage reservoirs. The results of the analyses are presented in terms of pumping and heating requirements (to bring the CO2 from the conditions in the ship to conditions acceptable for the injection well) and the required investment cost and operational cost of shipping CO2. Liquid CO2 can be loaded onto ships using a conventional articulated loading arm developed for other cryogenic liquids such as LPG and LNG. An alternative would be to use a flexible cryogenic hose, however, this is less reliable with a higher risk of failure and leakage. Cryogenic pumps located near the intermediate storage transfer the liquid CO2 via an insulated pipeline, specified for the liquid storage conditions, to the loading arm and the ship. A second line returns gas from the empty ship’s tank, and any boil-off gas produced on loading, to the compressors of the liquefaction plant. Loading rates can be quite high allowing ships to be loaded within a day, for example, Vermeulen proposes to fill at a rate of 2875 t/h using two loading arms, allowing a 30,000 m3 ship to be filled within 12 h [104]. Examples The ammonia producer Yara International trades much of its CO2 byproduct and transports it by sea from production sites in Norway and the Netherlands to seven import and distribution terminals around western European coasts using three ships with capacities of 9001200 t of CO2 at 1520 bar and around 230
C. Beyond this, the shipping company IM Skaugen has six 10,000 m3 ships in their fleet which are rated to 7 bar, 2104
C, and are registered for carrying liquid CO2, however, their normal cargo is LPG. Large-scale shipbased transport of CO2 is relatively new. As part of the first phase of its CO2 transport and storage infrastructure development to be used in Northern Lights GCS project in Norway, two dedicated CO2 carriers, each with a cargo size of 7500 m3 (approximate transport mass 8000 t per cargo) and a length of 130 m were ordered [105].

8.5 Subsurface components: exploration and reservoir

In some cases, the use of a pipeline or ship as a transport option could not be a priori; a decision would need a benchmarking. As several options (diameter or ship size) are possible for each technology, the optimum scenario of each option should first be identified before the two cost-optimized supply chains are compared [106]. The reuse of the O&G pipeline for CO2 logistics is a viable option, provided that it is suitable. The equipment must have a pressure rating and material specification appropriate for the project and the remaining lifetime must be sufficient. The industry has enough practices, procedures, and equipment to assess the suitability of existing infrastructure for re-use. IEAGHG provides a general approach to assess the reuse together with a re-use index tool and examples from the GCS applications [107].

8.5 Subsurface components: exploration and reservoir 8.5.1 Exploration and screening Screening and selection of a geological formation for CO2 storage is usually the starting point of GCS projects and involves overlap/interaction with various disciplines. The goal is to define a GCS site with the required storage characteristics— capacity, injectivity, and confinement (CIC)—that also fits emission rate with operational and economic background. Dedicated projects investigated screening criteria for GCS sites [108,109]. Various studies proposed the principles for selecting SAs and DHRs to be used in GCS [110114]. In general, a sedimentary formation with sufficient PV to store and permeability to inject and being confined with sealing rocks without tectonic disturbance is a natural candidate for GCS. However, applicability may vary regionally due to hurdles, especially regarding legislation and regulations as well as CO2 logistic (remote areas). Depth is one of the most important criteria to assure the required thermodynamic conditions for storing CO2 in the dense phase. The mineralogy of the host rock as well as of the cap rock should also be assessed before a decision is made as they may provide important hints on the long-term mineralization (capacity) and dissolution (confinement). The distance of the potential GCS site to the CO2 source as well as, regulatory, environmental, and social concerns should be evaluated within the screening frame. A process flowchart is proposed by National Energy Technology Laboratory (NETL) that can be used for screening GCS sites [109]. The process begins with the definition of potential subregions and includes a step-by-step algorithm with input from the project management. First, regional geologic data is analyzed, which leads to the regional proximity analysis as the second step. This step addresses the question of whether the areas are attractive for storage, considering the environment, society, and infrastructure (pipeline, CO2 resources, etc.). The third step requires answering

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the question of whether social context analysis indicates a receptive community. These three steps are iterative, and the result is a ranked list of potential sites. To facilitate the preliminary screening process the NETL introduces an Excel-based public screening tool (CO2-SCREEN) for calculating potential CO2 storage resources for saline formations based on a probabilistic assessment of influencing parameters such as P, T, formation volume, thickness, porosity, and depth, all linked to CO2 properties [109]. The results are estimates for discovered resources that should be further analyzed for commercial capacity. Example In Table 8.3, geologic screening criteria applied for selecting GCS sites in the Nordic region are presented with modifications to generalize the requirements. It should be noted that some reservoir properties are GCS type dependent; e.g., the preferred ranges for permeability should usually be higher for SAs if compared to DHRs or brine salinity is less important for DGR. The first step on the development path as a GCS site is the accomplishment of the screening process with positive results. The following steps include detailed analysis of available data and acquisition of those data to be required yet, modeling and front-end-engineering studies followed by the operation and monitoring. Fig. 8.11 provides a simplified workflow that can be used for GCS selection and development. Table 8.3 Geologic screening criteria for geologic carbon storage sites. Reservoir properes Depth Porosity Permeability Heterogeneity Pore pressure Temperature Brine salinity Thickness (net sand) Seal properes Thickness Tectonic/Fault Lateral extent Mulple seals Lithology of the primary seal Safety/risk Seismicity Groundwater contaminaon Maturity/Available data Wells Seismic survey

N/G, Net to gross ratio.

Best 800-2500 m >20% >100 mD Low N/G=0.4 Hydrostac or lower Low T gradient Low > 50 m

Challenging 600-800 m 10-20% 10-100 mD Moderate N/G=0.1-0.4

>50 m No tectonic/no faults Connous More than one Homogeneous clay, mud or evaporates

20-50 m Moderate

None/low No risk

Moderate Unsure

High Risk

Wells through the actual trap 3D

Wells through analogue geology Younger 2D

No well data

15-50 m

Only one Chalk

Not preferred