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Advances in clean hydrocarbon fuel processing
© Woodhead Publishing Limited, 2011
Related titles: Advanced membrane science and technology for sustainable energy and environmental applications (ISBN 978-1-84569-969-7) Progress in membrane materials, selective membrane design, and computer modelling and simulation have contributed greatly to the application of advanced membranes in conventional and alternative power sectors, as well as to clean industry applications. Research and development of membrane systems continues apace towards the goal of simple, efficient, and easily integrated systems that offer low-cost, reliable processing and operation. This book presents a comprehensive review of membrane science and technology. Handbook of biofuels production (ISBN 978-1-84569-679-5) The global increase in the use of biofuels as substitute transportation fuels has prompted research and development into advanced chemical, biochemical and thermochemical biofuels production. This research and development is aimed at improving the quality and environmental impact of biofuels production, as well as the overall efficiency and output of biofuels production plants. The range of biofuels has also increased, with market developments leading to the increased production and utilisation of advanced biofuels such as biosyngas, biohydrogen and biobutanol. Advanced power plant materials, design and technology (ISBN 978-1-84569-515-6) Fossil-fuel power plants generate the majority of the world’s power, but many plants are ageing and cannot meet rising global energy demands and increasingly stringent emissions criteria. To ensure security and economy of supply, utilities are building a new generation of advanced power plants with increased output and environmental performance. This book initially reviews improved plant designs for efficiency and fuel flexibility, including combined-cycle technology and utilisation of lower-grade feedstocks. Coverage extends to advanced material and component use, and the incorporation of alternative energy conversion technology. Details of these and other Woodhead Publishing materials books can be obtained by:
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© Woodhead Publishing Limited, 2011
Woodhead Publishing Series in Energy: Number 19
Advances in clean hydrocarbon fuel processing Science and technology Edited by M. Rashid Khan
Oxford
Cambridge
Philadelphia
New Delhi
© Woodhead Publishing Limited, 2011
Published by Woodhead Publishing Limited 80 High Street, Sawston, Cambridge CB22 3HJ, UK www.woodheadpublishing.com Woodhead Publishing, 1518 Walnut Street, Suite 1100, Philadelphia PA 19102-3406, USA Woodhead Publishing India Private Limited, G-2, Vardaan House, 7/28 Ansari Road, Daryaganj, New Delhi – 110002, India www.woodheadpublishingindia.com First published 2011, Woodhead Publishing Limited © Woodhead Publishing Limited, 2011 The authors have asserted their moral rights. This book contains information obtained from authentic and highly regarded sources. Reprinted material is quoted with permission, and sources are indicated. Reasonable efforts have been made to publish reliable data and information, but the authors and the publisher cannot assume responsibility for the validity of all materials. Neither the authors nor the publisher, nor anyone else associated with this publication, shall be liable for any loss, damage or liability directly or indirectly caused or alleged to be caused by this book. Neither this book nor any part may be reproduced or transmitted in any form or by any means, electronic or mechanical, including photocopying, microfilming and recording, or by any information storage or retrieval system, without permission in writing from Woodhead Publishing Limited. The consent of Woodhead Publishing Limited does not extend to copying for general distribution, for promotion, for creating new works, or for resale. Specific permission must be obtained in writing from Woodhead Publishing Limited for such copying. Trademark notice: Product or corporate names may be trademarks or registered trademarks, and are used only for identification and explanation, without intent to infringe. British Library Cataloguing in Publication Data A catalogue record for this book is available from the British Library. Library of Congress Control Number: 2011935448 ISBN 978-1-84569-727-3 (print) ISBN 978-0-85709-378-3 (online) ISSN 2044-9364 Woodhead Publishing Series in Energy (print) ISSN 2044-9372 Woodhead Publishing Series in Energy (online) The publisher’s policy is to use permanent paper from mills that operate a sustainable forestry policy, and which has been manufactured from pulp which is processed using acidfree and elemental chlorine-free practices. Furthermore, the publisher ensures that the text paper and cover board used have met acceptable environmental accreditation standards. Typeset by RefineCatch Limited, Bungay, Suffolk Printed by TJI Digital, Padstow, Cornwall, UK
© Woodhead Publishing Limited, 2011
Contents
Contributor contact details Woodhead Publishing Series in Energy Part I Overview and assessment of hydrocarbon fuel conversion processes 1
Characterization and preparation of biomass, oil shale and coal-based feedstocks
xi xv
1 3
O. TRASS, University of Toronto, Canada
1.1 1.2 1.3 1.4 1.5 1.6 1.7 1.8
Introduction Types and properties of feedstock Coal feedstock characterization and requirements Coal cleaning and preparation techniques Coal slurry fuels Future trends Sources of further information and advice References
3 4 15 22 28 43 45 48
2
Production, properties and environmental impact of hydrocarbon fuel conversion
54
J. G. SPEIGHT, CD & W Inc., USA
2.1 2.2 2.3 2.4 2.5 2.6 2.7 2.8
Introduction Production of hydrocarbon fuels Properties of hydrocarbon fuels Use and energy efficiency Environmental impact Toxicity hazards Future trends in fuels production and properties References
© Woodhead Publishing Limited, 2011
54 58 63 72 72 77 80 82
vi
Contents
3
Life cycle assessment (LCA) of alternative hydrocarbon fuel conversion
83
J. M. ANDRESEN and S. LI, University of Nottingham, UK
3.1 3.2
Introduction Life cycle assessment: environmental, energetic and techno-economic issues Life cycle assessment of fuel conversion routes and alternative feedstock utilisation Conclusions and future trends Sources of further information and advice References
86 97 98 98
Part II Solid hydrocarbon fuel processing and technology
103
3.3 3.4 3.5 3.6
4
Direct liquefaction (DCL) processes and technology for coal and biomass conversion
83 84
105
C. BURGESS CLIFFORD and C. SONG, Penn State University, USA
4.1 4.2 4.3 4.4 4.5 4.6 4.7 4.8 4.9 4.10 4.11 5
Introduction Feedstocks for direct liquefaction Basics of coal and biomass/lignin reaction chemistry Process variables: coal rank, solvent, catalyst, temperature, pressure and residence time in direct liquefaction (DCL) Known process technologies Product output and quality issues Process control and modeling techniques Advantages and limitations Future trends in direct coal liquefaction Sources of further information and advice References
105 109 115
Gasification process technology
155
119 129 142 145 145 147 148 149
C. HIGMAN, Higman Consulting GmbH, Germany
5.1 5.2 5.3 5.4 5.5 5.6 5.7 5.8 5.9
Introduction Gasification in the refinery environment Basic principles Building blocks for complete systems Hydrogen and power plant as an example of a complete system Advantages and limitations Future trends Sources of further information and advice References
© Woodhead Publishing Limited, 2011
155 155 158 173 179 182 182 183 184
Contents
6
Pyrolysis processes and technology for the conversion of hydrocarbons and biomass
vii
186
J. M. ANDRESEN and XIAO Y. LIM, University of Nottingham, UK
6.1 6.2 6.3 6.4 6.5 6.6 6.7 6.8 7
Introduction Applicable feedstocks Process technology Basic reactions Thermodynamics/reaction kinetics Catalyst and solvent utilization Conclusion and future trends References
186 187 189 191 193 194 196 197
Biomass catalysis in conventional refineries
199
J. A. MELERO, A. GARCÍA and J. IGLESIAS, Universidad Rey Juan Carlos, Spain
7.1 7.2 7.3 7.4 7.5 7.6 7.7
Introduction Biomass feedstock: availability and diversity Catalytic cracking of biomass feedstock Hydrotreating of biomass feedstock Production of conventional liquid fuels from sugars Future trends References
199 202 204 215 227 231 233
Part III Liquid hydrocarbon fuel processing and technology
241
8
Sulfur removal from heavy and light petroleum hydrocarbon by selective oxidation
243
M. RASHID KHAN, Saudi Aramco, Saudi Arabia and E. SAYED, Imperial College, UK
8.1 8.2 8.3 8.4 8.5
Introduction Background Oxidative desulfurization chemistry Conclusions References
243 243 248 259 260
9
Partial oxidation (POX) processes and technology for clean fuel and chemical production
262
R. L. KEISKI, S. OJALA, M. HUUHTANEN, T. KOLLI and K. LEIVISKÄ, University of Oulu, Finland
9.1 9.2 9.3 9.4 9.5
Introduction Process technology and methods of partial oxidation (POX) Basic partial oxidation reactions Catalysts utilized Process control and modelling techniques
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262 264 268 275 275
viii
Contents
9.6 9.7 9.8
Advantages, limitations and optimization Future trends References
277 278 280
10
Hydroconversion processes and technology for clean fuel and chemical production
287
P. R. ROBINSON, PQ Optimization Services, USA
10.1 10.2 10.3 10.4 10.5 10.6 10.7 10.8 10.9
Introduction to petroleum refining Environmental protection Hydroconversion overview Economics of hydroconversion Chemistry of hydroconversion Supported-metal hydroconversion catalysts Commercial hydroconversion units Future trends in hydroconversion References
Part IV Gaseous hydrocarbon fuel processing and technology 11
Middle distillate fuel production from synthesis gas via the Fischer-Tropsch process
287 290 291 292 292 307 312 324 325
327
329
F. G. BOTES, L.P. DANCUART, H. G. NEL, A. P. STEYNBERG and A.P. VOGEL, Sasol Technology, South Africa, B. B. BREMAN, Sasol Technology Netherlands, The Netherlands, and J.H.M. FONT FREIDE, BP, UK
11.1 11.2 11.3 11.4 11.5 11.6 11.7 11.8 11.9 11.10 12
Introduction Process technology Basic principles of the reaction process Catalyst utilisation Product upgrading and quality issues Process modelling and control Advantages, limitations and optimisation for synthetic middle distillate fuels Future trends Sources of further information and advice References Methanol and dimethyl ether (DME) production from synthesis gas
329 332 339 343 346 350 353 355 357 358 363
D. SEDDON, Duncan Seddon & Associates Pty. Ltd, Australia
12.1 12.2 12.3
Introduction Process technology and new innovations Basic principles of methanol synthesis
© Woodhead Publishing Limited, 2011
363 364 371
Contents
ix
12.4 12.5 12.6 12.7 12.8 12.9
Catalysts Product quality Estimation of production costs Future trends Sources of further information and advice References
374 376 377 381 382 382
13
Advances in water-gas shift technology: modern catalysts and improved reactor concepts
387
E. V. REBROV, Queen’s University Belfast, UK
13.1 13.2 13.3 13.4 13.5 14
Introduction Modern reactor concepts Advanced catalytic systems Conclusions and future trends References
387 388 397 406 408
Natural gas hydrate conversion processes
413
M. MAX and A. H. JOHNSON, Hydrate Energy International (HEI), USA
14.1 14.2 14.3 14.4 14.5 14.6 14.7 14.8
Introduction Factors important for hydrate conversion Resource potential Conversion processes Advantages, limitations and optimization Future trends Sources of further information and advice References
Part V Operational issues and process improvement in hydrocarbon fuel processing plant 15
Environmental degradation in hydrocarbon fuel processing plant: issues and mitigation
413 416 423 424 430 430 431 432
435
437
F. ROPITAL, IFP Energies Nouvelles, France
15.1 15.2 15.3 15.4 15.5 15.6 15.7
Introduction Types of degradation and their main locations Protection and mitigation technologies Plant management techniques Future trends Sources of further information and advice References
© Woodhead Publishing Limited, 2011
437 438 450 454 460 460 461
x
Contents
16
Automation technology in hydrocarbon fuel processing plant
463
S. EL FERIK, King Fahd University of Petroleum and Minerals, Saudi Arabia
16.1 16.2 16.3 16.4 16.5 16.6 16.7 16.8 17
Introduction Automation technology survey: from exploration to processing Fundamentals of process control Control design Future trends in automation technology Working towards a broader integration of control and operation Conclusions References Advanced process control for clean fuel production: smart plant of the future
463 464 470 481 489 490 493 494 496
O. TAHA and M. RASHID KHAN, Saudi Aramco, Saudi Arabia
17.1 17.2 17.3 17.4 17.5 17.6 17.7 17.8 17.9 18
Introduction Incentives for smart process control technologies Smart instrumentation of the future Advanced process control (APC) and optimization solutions Model predictive control technology (MPC) Real-time optimization (RTO) technology Control performance monitoring (CPM) Driving future innovation, sustainability and performance in process control technologies References
496 496 497 497 498 499 502
Process modeling for hydrocarbon fuel conversion
509
505 507
E. VASQUEZ, Maryland, USA and T. ELDREDGE, Connecticut, USA
18.1 18.2 18.3 18.4 18.5
Computational fluid dynamics (CFD) modeling techniques Empirical modeling techniques Process flow models Chemical kinetic modeling References
509 519 523 530 544
Index
546
© Woodhead Publishing Limited, 2011
Contributor contact details
(* = main contact)
Editor
Chapter 3
Dr M Rashid Khan Leader, Corporate Intellectual Assets Management Technology Management Division PO Box 5337 Saudi Aramco Dhahran 31311 Saudi Arabia
Professor John M Andresen* and Dr Sujing Li Energy & Sustainability Division Faculty of Engineering University of Nottingham NG7 2RD UK
E-mail: [email protected] [email protected] [email protected]
E-mail: [email protected] [email protected]
Chapter 4
Chapter 1 Professor Olev Trass Professor Emeritus Chemical Engineering and Applied Chemistry University of Toronto 200 College St Toronto, ON M5S 3E5 Canada E-mail: [email protected]
Chapter 2
Dr Caroline Burgess Clifford* and Dr Chunshan Song EMS Energy Institute Penn State University C211 Coal Utilization Laboratory University Park PA 16802 USA E-mail: [email protected] [email protected]
Chapter 5
Dr James G Speight CD&W Inc 2476 Overland Road Laramie WY 82070 USA
Mr Chris Higman Higman Consulting GmbH Sachsenstrasse 19 D-65824, Schwalbach Germany
E-mail: [email protected]
E-mail: [email protected]
© Woodhead Publishing Limited, 2011
xii
Contributor contact details
Chapter 6 Professor John Andresen and Xiao Y Lim* Energy & Sustainability Division Faculty of Engineering University of Nottingham NG7 2RD UK E-mail: [email protected] [email protected]
Chapter 7 Professor Juan Antonio Melero* and Alicia García Department of Chemical and Environmental Technology Universidad Rey Juan Carlos C/ Tulipán s/n E–28933 Móstoles Madrid Spain E-mail: [email protected]
José Iglesias Department of Chemical and Energy Technology Universidad Rey Juan Carlos C/ Tulipán s/n E–28933 Móstoles Madrid Spain
E-mail: [email protected] [email protected] [email protected]
E Sayed Imperial College London UK
Chapter 9 Professor Riitta L Keiski*, Dr Satu Ojala, Dr Mika Huuhtanen, Dr Tanja Kolli and Professor Kauko Leiviskä University of Oulu Department of Process and Environmental Engineering Mass and Heat Transfer Process Laboratory and Control Engineering Laboratory PO Box 4300 FI-90014 University of Oulu Finland E-mail: [email protected]
Chapter 10 Dr Paul R Robinson PQ Optimization Services 3418 Clear Water Park Drive Katy TX 77450 USA E-mail: [email protected]
Chapter 8 Dr M Rashid Khan Leader, Corporate Intellectual Assets Management Technology Management Division PO Box 5337 Saudi Aramco Dhahran 31311 Saudi Arabia
Chapter 11 F G Botes*, L P Dancuart, H G Nel and A P Vogel Sasol Technology 1 Klasie Havenga Road Sasolburg, 1947 South Africa E-mail: [email protected]
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Contributor contact details
xiii
Kenner LA 70065-1126 USA
A P Steynberg Sasol Technology 33 Baker Street Rosebank Johannesburg, 2196 South Africa
E-mail: [email protected] [email protected]
B B Breman Sasol Technology Netherlands Vlierstraat 111 NL-7544 GG, Enschede The Netherlands J H M Font Freide BP UK
Chapter 15 Dr Francois Ropital IFP Energies nouvelles Rond Point de l’Echangeur de Solaize 69360 Solaize France E-mail: [email protected]
Chapter 16
Chapter 12 Duncan Seddon Duncan Seddon & Associates Pty Ltd 116 Koornalla Crescent Mount Eliza 3930 Australia E-mail: [email protected]
Dr Sami El Ferik Associate Professor Systems Engineering Department King Fahd University of Petroleum and Minerals Dhahran 31261 Saudi Arabia E-mail: [email protected]
Chapter 13
Chapter 17
Professor Evgeny V Rebrov School of Chemistry and Chemical Engineering Queen’s University Belfast Stranmillis Road BT9 5AG Belfast UK
Dr Othman Taha PO Box 12543 Saudi Aramco Dhahran, 31311 Saudi Arabia
E-mail: [email protected]
Chapter 14 Dr Michael Max* and Arthur H Johnson Hydrate Energy International (HEI) 612 Petit Berdot Drive
E-mail: [email protected] [email protected]
Dr M Rashid Khan PO Box 5337 Saudi Aramco Dhahran 31311 Saudi Arabia E-mail: [email protected] [email protected] [email protected]
© Woodhead Publishing Limited, 2011
xiv
Contributor contact details
Chapter 18 Dr Edmundo R. Vasquez Combustion Specialist 8320 Park Crest Drive 20910 Silver Spring, Maryland USA
Dr Thomas V. Eldredge* 28 Blakeman Drive 06468 Monroe, Connecticut USA E-mail: [email protected]
E-mail: [email protected]
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Woodhead Publishing Series in Energy
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Generating power at high efficiency: Combined cycle technology for sustainable energy production Eric Jeffs Advanced separation techniques for nuclear fuel reprocessing and radioactive waste treatment Edited by Kenneth L. Nash and Gregg J. Lumetta Bioalcohol production: Biochemical conversion of lignocellulosic biomass Edited by K.W. Waldron Understanding and mitigating ageing in nuclear power plants: Materials and operational aspects of plant life management (PLiM) Edited by Philip G. Tipping Advanced power plant materials, design and technology Edited by Dermot Roddy Stand-alone and hybrid wind energy systems: Technology, energy storage and applications Edited by J.K. Kaldellis Biodiesel science and technology: From soil to oil Jan C.J. Bart, Natale Palmeri and Stefano Cavallaro Developments and innovation in carbon dioxide (CO2) capture and storage technology Volume 1: Carbon dioxide (CO2) capture, transport and industrial applications Edited by M. Mercedes Maroto-Valer Geological repository systems for safe disposal of spent nuclear fuels and radioactive waste Edited by Joonhong Ahn and Michael J. Apted Wind energy systems: Optimising design and construction for safe and reliable operation Edited by John D. Sørensen and Jens N. Sørensen Solid oxide fuel cell technology: Principles, performance and operations Kevin Huang and John Bannister Goodenough Handbook of advanced radioactive waste conditioning technologies Edited by Michael I. Ojovan Nuclear safety systems Edited by Dan Gabriel Cacuci
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Woodhead Publishing Series in Energy Materials for energy efficiency and thermal comfort in buildings Edited by Matthew R. Hall Handbook of biofuels production: Processes and technologies Edited by Rafael Luque, Juan Campelo and James Clark Developments and innovation in carbon dioxide (CO2) capture and storage technology Volume 2: Carbon dioxide (CO2) storage and utilisation Edited by M. Mercedes Maroto-Valer Oxy-fuel combustion for power generation and carbon dioxide (CO2) capture Edited by Ligang Zheng Small and micro combined heat and power (CHP) systems: Advanced design, performance, materials and applications Edited by Robert Beith Advances in clean hydrocarbon fuel processing: Science and technology Edited by M. Rashid Khan Modern gas turbine systems: High efficiency, low emission, fuel flexible power generation Edited by Peter Jansohn Concentrating solar power (CSP) technology: Developments and applications Edited by Keith Lovegrove and Wes Stein Nuclear corrosion science and engineering Edited by Damien Féron Power plant life management and performance improvement Edited by John Oakey Direct-drive renewable energy systems Edited by Markus Mueller and Henk Polinder Advanced membrane science and technology for sustainable energy and environmental applications Edited by Angelo Basile and Suzana Pereira Nunes Irradiation embrittlement of reactor pressure vessels (RPVs) in nuclear power plants Edited by Naoki Soneda High temperature superconductors (HTS) for energy applications Edited by Ziad Melhem Infrastructure and methodologies for the justification of nuclear power programmes Edited by Agustín Alonso Santos Waste to energy (WtE) conversion technology Edited by Marco Castaldi Polymer electrolyte membrane and direct methanol fuel cell technology Volume 1: Fundamentals and performance Edited by Christoph Hartnig and Christina Roth Polymer electrolyte membrane and direct methanol fuel cell technology Volume 2: In situ characterisation techniques Edited by Christoph Hartnig and Christina Roth Combined cycle systems for near-zero emission power generation Edited by Ashok Rao Modern earth buildings: Materials, engineering, construction and applications Edited by Matthew R. Hall, Rick Lindsay and Meror Krayenhoff
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Handbook of metropolitan sustainability: Understanding and improving the urban environment Edited by Frank Zeman Functional materials for energy applications Edited by John Kilner, Stephen Skinner, Stuart Irvine and Peter Edwards Nuclear decommissioning: Planning, execution and international experience Edited by Michele Laraia Nuclear fuel cycle science and engineering Edited by Ian Crossland Electricity transmission, distribution and storage systems Edited by Ziad Melhem Advances in biodiesel preparation: Second generation processes and technologies Edited by Rafael Luque and Juan Antonio Melero Biomass combustion science, technology and engineering Edited by Lasse Rosendahl Ultra-supercritical coal power plant: Materials, technologies and optimisation Edited by Dongke Zhang Radionuclide behaviour in the natural environment: Science, impacts and lessons for the nuclear industry Edited by Horst Geckeis and Christophe Poinssot Calcium and chemical looping technology for power generation and carbon dioxide (CO2) capture: Solid oxygen- and CO2-carriers P. Fennell and E.J. Anthony Materials ageing and degradation in light water reactors: Mechanisms, modelling and mitigation Edited by K.L. Murty Structural alloys for power plants: Operational challenges and high-temperature materials Edited by Amir Shirzadi, Rob Wallach and Susan Jackson Biolubricants: Science and technology Jan C.J. Bart, Emanuele Gucciardi and Stefano Cavallaro Wind turbine blade design and materials: Improving reliability, cost and performance Edited by Povl Brøndsted and Rogier Nijssen Radioactive waste management and contaminated site clean-up: Processes, technologies and international experience Edited by William E. Lee, Michael I. Ojovan, Carol M. Jantzen Probabilistic methods of strength reliability and their application for optimum nuclear power plant life management (PLiM) Gennadij V. Arkadov, Alexander F. Getman and Anderi N. Rodionov Coal utilization in industry: Towards cleaner production Edited by D.G. Osborne Coal power plant materials and life assessment: Developments and applications Edited by Ahmed Shibli
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1 Characterization and preparation of biomass, oil shale and coal-based feedstocks O. TRASS, University of Toronto, Canada Abstract: Aspects of feedstock preparation for biomass, oil shale and coal conversion are presented. Conversion processes include combustion, pyrolysis, gasification, liquefaction and various combined upgrading operations to obtain energy, gaseous and liquid fuels, and other valuable products. Crushing and grinding are required to obtain easy-to-handle feedstocks with desirable particle sizes and shapes. Depending on downstream processing, various feedstock requirements must be met, including levels of cleaning. Wet-ground, beneficiated coal slurries provide a valuable feedstock which may be economically produced from fine waste coal. Slurries are typically water-based for combustion and gasification and oil-based for direct coal liquefaction. Key words: biomass feedstock preparation, oil shales, coal characteristics, coal cleaning techniques, beneficiated coal slurry fuels.
1.1
Introduction
Over the years, the image of coal has undergone major shifts in public opinion. It was an important driving force of the industrial revolution and displaced biomass, mainly wood, wind and water, as the major source of energy and power. A century ago, oil started to displace coal in some applications and, by the 1950s, coal bunkers had given way to oil tanks on industrial sites and also in many homes. The current perception of coal is that of a ‘dirty’ fuel and an environmental culprit, at least in the public mind, in media and especially among the Greens, concerned environmentalists who carry a lot of weight, also with politicians. One might argue that into the 1970s coal was, indeed, not only black but ‘dirty’, as much of the coal used in power plants was not even washed . . . and the United States’ Clean Coal Technologies (CCT) program was still some way off. The ‘Oil Shock’ changed all that in 1973. Cheap, easy-to-handle oil suddenly became more expensive, courtesy of the oil embargo and the Organization of Petroleum Exporting Countries (OPEC). Not only industry, but also a number of power plants, especially on the US East Coast, were using oil as a fuel and felt the pinch. The potential salvation was to shift back to coal, possibly in the form of coal–oil or coal–water slurries. The impressive and productive CCT program did get going, and rapid progress lasted until 1986. Then the oil price crashed! That oil price drop, from around US $30 to below $10/bbl, stopped development virtually completely in the industry, and research support was also cut soon thereafter. Implementation of the CCT results suffered as well, and efforts to 3 © Woodhead Publishing Limited, 2011
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Advances in clean hydrocarbon fuel processing
promote alternative energy sources were set back. Coal liquefaction and gasification plants built to demonstrate newer technologies were simply shut down. Twenty years later, and we were again facing high oil prices, with the peak reaching $147 per barrel. This time, the four-fold price drop which soon followed was less devastating as it did not last long enough. At the time of writing, half of that drop has been retraced, and the new incentives are stronger; not only for wind, biomass and solar energy, but also for producing clean, easy-to-handle liquid and gaseous fuels from King Coal. In the early 1980s there was some discussion about the concept of the ‘coal refinery’. In a manner analogous to the refining of petroleum, it was presumed also that the various coal constituents should each be used in the most effective way (Oder, 1981; Trass, 1984). The idea, in its simple conception, was not only to beneficiate the coal by various ash removal techniques, but also to separate the different coal macerals. The reactive ones, such as vitrinite, would then be directed to a liquefaction plant, and the more inert fusinite or inertinite should go to a power plant for combustion. Just as one now talks about the potential ‘biomass refinery’ (PAPTAC, 2007; Pinatti et al., 2010) to be implemented, particularly in pulp and paper mills, so it pays to resurrect the idea of coal refining. It is hoped that this volume will provide some guidance and help in that regard. In the meantime, a different concept of the coal refinery has emerged, where it means production of a range of products for different uses rather than using subcomponents of the raw material. That, really, is a combination of different processes used on the same site rather than refining in the usual sense. In this chapter we will deal with key requirements of feedstocks for different processing techniques. As the range of potential raw materials is wide, some will be dealt with in a rather cursory fashion and others more thoroughly, with at least some emphasis on specific downstream processing steps beyond the broad, general requirements.
1.2
Types and properties of feedstock
When considering the range of processes covered in this volume, one must cover also the appropriate range of potential feedstocks. Starting with King Coal, we have coal of several ranks ranging from anthracite, through bituminous, subbituminous, down to lignites. Then, starting with peat, we have surely reached the biomass feedstock category (even though peat is rather aged compared with the current harvest), wood, agricultural products and residues, and grasses. A common requirement of all these feedstocks is the need to get them in a condition which allows for relative ease of handling. They should also be relatively clean; that is, free of stones, soil and a wide range of undesirable impurities. This, minimally, implies cleaning, cutting, crushing and grinding to bring the feedstock
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to a usable form. Beyond that, we need to look at particular types of material, and the processing they need to undergo. Let us start with a brief look at the ‘youngest’, harvested biomass, and then proceed to the mined varieties, oil shale and coal. Both biomass and oil shale will be covered in this section; coal, being the most important feedstock, will be dealt with more thoroughly thereafter.
1.2.1 Biomass feedstocks Biomass is the largest source of renewable energy, supplying about 13% of the world’s total energy demand (Demirbas et al., 2009). The use of biomass for energy is attractive because supplies are large, widespread and renewable, and its use has the potential to be carbon neutral, i.e., the CO2 released can be balanced by the CO2 consumed by new biomass growth. The use of biomass for generating electricity, for heating, in particular for buildings, and in ‘co-generation’ or ‘combined heat and power’ operations that produce both electricity and useful thermal energy, is increasing rapidly in Europe where the annual rate of increase in the use of biomass energy is now about 25% per annum (Magelli et al., 2009). Direct combustion of biomass residues to produce electricity is expected to be the major application for biomass in the foreseeable future. Nevertheless, biomass processing to get liquid biofuels and gas is also growing in importance, along with the production of biomass-based chemicals. There are immense variations in biomass depending on its source. The major group is undoubtedly forest biomass, substantially residues from lumber and pulp and paper operations. Others include industrial and agricultural wastes as well as peat. Considering the major group further, there are important distinctions between hardwoods and softwoods, i.e. deciduous and coniferous tree species, individual species differences in each category, as well as differences due to tree age or site of growth. Bark and foliage, including twigs and branches, add further to both structural and chemical composition variations of woody biomass. Other biomass sources such as animal waste or aquatic plants and algae exhibit even greater differences in their properties and will not be considered here. Feedstock requirements For combustion (as one common form of biomass conversion), particle size, shape and density are important parameters for fuel handling. Thus, pre-treatment of woody biomass such as primary and secondary size reduction, drying and densification are required before it can be used. The essential size reduction step is energy intensive and should be optimized. Esteban and Carrasco (2006) reviewed the literature on particle size requirements for combustion of woody biomass in pulverized fuel burners and concluded that the wood particles need to be below 1000 µm, as compared with below 100 µm for coal. Such particle sizes allow complete combustion in about half a second, the typical
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residence time in a burner. The proportion of very fine wood particles, below 100 µm, is also important because these particles are responsible for early ignition. Often, combustion of biomass is carried out in combination with coal or coke so as to have a higher energy density while still benefiting from the ‘renewable’ aspect for more sustainable fuel sources. Matching of particle sizes is then important. Grinding of biomass to the sizes required for suspension firing is typically done with hammer mills. Esteban and Carrasco (2006) give detailed data from an experimental study of grinding wood chips and bark in two stages of hammer milling. Their reported energy requirements are 120, 150 and 35 kWh/t (oven dry) for poplar chips, pine chips and pine bark, respectively. These energy requirements are below 3% of the fuel heating value. Similar data for grinding of switchgrass, wheat straw and corn stover with a hammer mill are given by Bitra et al. (2009). An alternative for this type of application is the Szego Mill™, a ring-roller mill developed at the University of Toronto in conjunction with General Comminution Inc., which has been used to grind various woody materials as well as minerals such as coal. Grinding may be done dry or wet, depending upon the desired product. Gravelsins and Trass (2011) report on the performance of the Szego Mill for the grinding of Energex™ industrial pelletized wood waste, consisting of a mixture of bark and wood from coniferous species, over a wide range of operating conditions. A brief description of grinding equipment including hammer mills and the Szego Mill is given under Section 1.5.3, Preparation of coal slurry fuels. Turning our attention to microbiological and thermochemical processing, rather than combustion, important requirements include access to the internal spaces of particles, i.e. for heat and mass transfer between the particle and its surroundings. The latter is here interpreted rather broadly as surface or diffusional access for microorganisms and enzymes on the one hand as well as for heat and typically smaller reactant and product molecules throughout the particle structure, at least over the reaction time. Thus, particle size is again vitally important. Lignocellulosic materials, composed mostly of cellulose, hemicellulose, lignin and, of course, water, form complex structures. These vary significantly from species to species and are not evenly broken down by micro-organisms. While we might wish to emphasize the requirements for thermal and chemical biomass processing, the range of processes is still broad and, hence, generalization is difficult. Some issues to be kept in mind include (Lipinski, 1985):
• • • • • • •
Specific biomass type and its resultant properties. Anatomical parts, e.g. wood chips or bark. Particle size and its distribution, i.e. uniformity of particle sizes; also particle shape. Moisture content (very important also for pre-processing). Homo- or heterogeneity of the feedstock materials. Ease of handling. Bulk density.
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Grossly simplifying a complex topic, it is essential to have relatively small, uniformly-sized particles as feed material to the eventual processing steps; hence, pre-treatment is important. Pre-treatment considerations For forest products, we can think of the sequel of steps going from the tree to the log, then to chips, and possibly sawdust. Separately from chips, we have bark and foliage and here it pays to keep in mind the idea of a ‘forest products refinery’ as a first step of that segregation. Mechanical pre-treatment has many routes, starting with anatomical separation. The roll splitter described by Jones (1982) is an example of suitable equipment when comminution to smaller sizes is vital. For many processes, wood chips from a conventional chipper are the preferred, easyto-handle, convenient feedstock. Typical chip sizes, with the chip thickness in the 3 mm range, will inevitably lead to relatively slow processes. Therefore, mechanical comminution to the sawdust or even finer powder size is often desirable. Considering the desired particle size, not only is the absolute size important, but also the size range, i.e. the particle size distribution (PSD). If particles are too small in a particular downstream process, they may not be retained long enough in the reaction zone. Those that are too large may not react completely, or secondary reactions may occur within those particles. Hence, a narrow PSD, or good uniformity, is usually desired and allows for optimal heat and mass transfer rates. For the fine powder processor, this typically means closed circuit grinding with recycle. There is a range of comminution equipment available, with the most commonly used mills of the hammer mill type. Often it is desirable to combine particle size reduction with other processes. An example is the steam explosion process, which will convert chips to fibers or powder (formerly Stake Technologies, now SunOpta BioProcess Inc., a division of SunOpta Foods). It also disrupts the lignocellulose structure by means of hydrolysis. The product has a relatively high moisture content, however, so steam explosion is suitable only for processes where this is not a disadvantage, e.g. for typical wet microbial processing. Torrefaction is another semi-processing step that should be mentioned. It is a mild pyrolysis pretreatment that causes changes in the chemical and physical properties of biomass, making it hydrophobic and more brittle, with an increased energy density (on a mass basis). Essentially, biomass progresses along the coalification band and its properties start to resemble those of lignite coals. This may be more important for moist, non-woody biomass. Bridgeman et al. (2008) provide a description and examples. The idea of an integrated biomass refinery has recently become popular, for a variety of reasons. A new journal, Biorefining Magazine, presented its inaugural issue in September, 2010 and a number of biomass-related conferences offer symposia reflecting the idea of biomass refining. The basic idea is still to disassemble
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wood into its constituent parts and to obtain maximum value from each such component. The paper by Pinatti et al. (2010) describes a holistic, integrated sequence of processing steps, multiple products, energy conversion, and recycling of wastes. The abstract of that paper is reproduced below as an illustration of that potential. Biomass Refinery is a sequential of eleven thermochemical processes and one biological process with two initial basic treatments: prehydrolysis for lignocelluosics and low temperature conversion for biomass with medium-tohigh content of lipids and proteins. The other ten processes are: effluent treatment plant, furfural plant, biodiesel plant, cellulignin dryer, calcination, fluidized bed boiler, autothermal reforming of cellulignin for synthesis gas (syngas) production, combined cycle of two-stroke low-speed engine or syngas turbine with fluidized bed boiler heat recovery, GTL technologies and ethanol from cellulose, prehydrolysate and syngas. Any kind of biomass such as wood, agricultural residues, municipal solid waste, seeds, cakes, sludges, excrements and used tires can be processed at the Biomass Refinery. Twelve basic products are generated such as cellulignin, animal feed, electric energy, fuels (ethanol, crude oil, biodiesel, char), petrochemical substitutes, some materials (ash, gypsum, fertilizers, silica, carbon black) and hydrogen. The technology is clean with recovery of energy and reuse of water, acid and effluents. Based on a holistic integration of various disciplines, Biomass Refinery maximizes the simultaneous production of food, electric energy, liquid fuels and chemical products and some materials, achieving a competitive position with conventional and fossil fuel technologies as well as payment capacity for biomass production. Biomass Refinery has a technical economical capability to complement the depletion of the conventional petroleum sources and to capture its GHGs resulting a biomass + petroleum ‘green’ combination.
Subsequent processing may well involve depolymerization and fractionation of the lignocellulose. The question, then, of what is pre-treatment and what is subsequent processing becomes rather diffuse. Indeed, pre-treatment should be integrated with the development of specific conversion processes. We will not return to biomass but deal subsequently with oil shales and coal-based feedstocks.
1.2.2 Oil shales The next category of feedstocks to be mentioned is really based on oil. Here, one may include oil shale, oil sands and various heavy oils, but we will restrict our attention to oil shales. Oil sands are found only in the Province of Alberta, Canada and therefore are not of as broad an interest as the other feedstocks. Heavy oils are really feedstocks ready for processing, so minimal preparation is needed. Oil shales are plentiful in many parts of the world. Nevertheless, the typical reader knows very little about their occurrence, properties or utilization. Hence, it seems appropriate to provide more background than may otherwise be needed for
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either biomass or coal, both being well known, even though the latter is the most important feedstock here. Extensive research has been carried out on oil shale during the last century, much of it in Estonia. Most of the recent efforts have been described in the journal Oil Shale published by the Estonian Academy Publishers in Tallinn since 1983. That journal has just initiated a new series called ‘Historical Reviews’. The first of these, by Harvey (2010), describes the shale-oil industry in Scotland where the first commercially successful production of oil from shale was undertaken, in the county of West Lothian, between Glasgow and Edinburgh. Scotland, apparently, was the global pioneer of the modern oil industry and, for a few decades in the second half of the nineteenth century, the leading oil producer in the world. All that was based on oil shale as feedstock! It is time to get back to that valuable resource now that we appear to be past Hubble’s peak as far as far as free-flowing oil supplies are concerned. The first patent in Britain for extracting crude oil from shale was registered in 1694, entitled ‘Method for Production of Large Quantities of Pitch, Tar and Oil from a Special Rock’. A small commercial plant was built in France in 1838, but the major effort started after another patent was registered in 1850 by James Young which covered both retorting and refining the shale oil and purifying paraffin wax from it. By 1865, when Young’s son took out another patent for an early version of the cracking process, more than 100 million litres (approximately 600 000 barrels) of oil were being produced annually. Oil production increased rapidly thereafter, peaking at 27 million barrels per annum in 1913. A significant number of by-products were also made, for example mothballs and candles. Young’s patented process was used in Pennsylvania to refine the free-flowing oil discovered there in 1859, a technology obviously learned from the advances made in Scotland refining the kerogen from the oil shale. When British Petroleum struck oil in the North Sea in 1970, a new era in Scotland’s oil history began, further North, near Aberdeen, and at a rate an order of magnitude higher. While oil shale is mined and utilized in several countries, the oil shale industry in Estonia is currently the most developed in the world. This is particularly the case in power plants but also in other respects, such as retorting. Several types of shale are found in Estonia. Nearly one-quarter of the country is lying on a rather thick layer of dictyonema shale, which has a relatively low organic matter content of about 13–14%, and hence of low heating value. The only oil shale being mined and utilized is the kukersite shale, often referred to simply as ‘Estonian oil shale’. Urov and Sumberg (1999) offer an overview of oil shales in Estonia and compare these with deposits elsewhere in the world. Raukas and Punning (2009) point out that over one billion tonnes of oil shale has been produced in Estonia, the largest producer and consumer in the world since the 1960s. They also point out the serious environmental problems that have arisen from such intensive exploitation of the resource, and describe some remedial steps being taken.
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Oil shale is sometimes referred to as ‘bituminous shale’ to distinguish it from other shales containing little organic matter, which might be just trapped gas. Some such fossilized sedimentary rocks do provide ‘shale gas’ which is about to change significantly the world’s energy scene, with exploitation having started just recently. Hydraulic fracturing or ‘fracking’ of the shale beds is used to liberate the gas. This chapter will review oil shale properties and requirements for some of its different uses. Composition of oil shale kerogen The organic matter in oil shale is called kerogen. It has been formed from protista, bacteria, algae and plankton; mostly unicellular marine organisms. As a result, there are no intercellular bonds and the main consolidating substance in the fuel is fat from the original material. As kerogen is of sapropelic origin – from original marine sources – it has a low solubility in solvents and is characterised by high hydrogen and oxygen content. The atomic H/C ratio is about 1.5, approximately the same as in crude oil. While the mineral matter in shales may vary greatly both in composition and in the amounts present, the kerogen part, originating largely from similar organisms, is reasonably consistent and does not depend on the total amount of kerogen present. This is in contrast to the humic fuels where many different plant species have gone into forming the peat, lignite or coal. In addition to carbon, hydrogen and oxygen, kerogen contains also some sulfur. Nitrogen content is low. One characteristic is the rather high chlorine content in the organic matter. Its presence can be explained by the marine origin of oil shale. The average chemical composition of the kerogen may be represented by the empirical formula C10 H15 OCl0.03S0.07. Volatiles matter content of oil shale is very high: 85–90% of the total organic matter. The kerogen is structured as a high-molecular-weight homogeneous compound insoluble in organic solvents and resistant to many reagents and acids (Urov and Sumberg, 1999). This organic matter is located in nests of shale mineral matter as colonies of small particles. Nest dimensions vary over a wider range, reaching 140 microns. The insoluble kerogen is transformed into partially soluble thermobitumen in the temperature range 250–450 °C. Beyond that thermal degradation starts and shale oil and char are formed. The kerogen content in the Estonian oil shale as received, depending on heating value, is in the range 23–26%. Composition of oil shale mineral matter There are two large groups of minerals in oil shale: a carbonate part and a sandyclay or terrigenous part. The latter is intensely intertwined with organic matter and is considered an inherent mineral impurity. In contrast, carbonate minerals in an
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oil shale deposit usually occur as separate layers and can, at least to some extent, be considered extraneous mineral matter. Yet there may also be individual calcite crystals densely bound to organic matter and distributed quite uniformly, in the 5–50 µm size range. For most fuels there is a close relationship between mineral matter and ash quantities and compositions. As oil shale contains carbonate minerals, however, the amount of mineral CO2 must be established simultaneously with ash determinations. For that, a fuel sample may be treated with a solution of hydrochloric acid and the amount of released CO2, e.g. from CaCO3, measured and allowance made. There may be significant variations in mineral matter proportions depending on location. Particularly, the carbonate part can vary in as-received oil shale, also depending on fuel quality, i.e. some enrichment at the mine site, which will give different carbonate/sandy-clay ratios. The main component of the oil shale carbonate part is calcium oxide followed by magnesium oxide. In the sandy-clay part, SiO2, Al2O3 and K2O are the main components. The corresponding minerals are calcite and dolomite and, in the sandy-clay group, quartz, feldspar and various micas. Tables 1.1 and 1.2 give respective chemical and mineral compositions (Ots, 2006). Much additional material is available (Dyni, 2003; Lille, 2003; Veiderma, 2003). Table 1.1 Chemical composition of oil shale mineral matter Component
Content (%)
Sandy-clay part SiO2 CaO A12O3 Fe2O3 TiO2 MgO Na2O K2O FeS2 SO3 H2O Total Carbonate part CaO MgO FeO CO2 Total
59.8 0.7 16.1 2.8 0.7 0.4 0.8 6.3 9.3 0.5 2.6 100.0 48.1 6.6 0.2 45.1 100.0
Source: Ots, 2006
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Mineral
Formula
Content (%)
Carbonate minerals Calcite
CaCO3
69.1
Dolomite
CaMg(CO3)2
30.6
Siderite
FeCO3 Total
0.3 100.0
Sandy-clay minerals Quartz
SiO2
23.9
Rutile
TiO2
0.7
Orthoclase
K2O Al2O3-6SiO2
Albite
Na2O-Al2O3-2SiO2
Anorthite
CaO-Al2O3-6SiO2
Muskovite
[K1-x(H2O)x-Al3Si3O10(OH)2]2
Amphipole
[NaCa2Mg4(FeAl)Si8O22(OH)2]
2.1
Markasite
FeS2
9.3
Limonite
Fe2O3-H2O
2.9
Gypsum
CaSO4-2H2O Total
29.0 6.0 1.4 23.7
1.0 100.0
Source: Ots, 2006
Pretreatment requirements The requirements vary greatly depending on the subsequent process to be used. For combustion there are at least three types of boilers to be considered: those with grate firing, pulverized firing and those utilizing fluidized beds. Pulverized combustion (PC) requires the finest grind, with a median size between 35–60 µm. Grate firing can accept large lumps but its use has virtually ceased. Fluidized bed combustion (FBC) is the newest technology for oil shale combustion and still under development. A special issue of Oil Shale, with 14 papers, edited by Ots (1997) makes the case for using circulating fluidized bed combustion (CFBC) technology as the most appropriate for combustion of Estonian oil shale. Two CFBC units have been built at the Narva Power Plant by Foster Wheeler Energia Oy, Finland, and are performing well (Hotta et al., 2005). Fuel particle size requirements are simple, as the FBC is able to accept particles up to 25 mm in size (D50 ≈ 1 mm). The particles are broken down in the combustor, with the exiting bottom ash median size D50 ≈ 0.75 mm and the fly ash (after inside separation) D50 ≈ 15 µm.
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The main advantage of the CFBC units is their ability to capture sulfur almost quantitatively, by making the huge excess of calcium in the shale available. Also, at the lower operating temperatures and with coarser shale particles compared to PC, there is less decomposition of the CaCO3 with the concomitant reduction of mineral CO2 emissions (Arro et al., 2001). The other important utilization for which oil shale feedstock must be prepared is thermal processing, i.e. retorting for oil production. For this, in the Estonian experience, several types of equipment have been used: tunnel kilns, rotating retorts and vertical retorts of different designs. Early equipment, in the 1920s, had capacities of 20–40 t/d. The latest ones, in the 1980s, have oil shale throughputs of 1000–3000 t/d. When the heat carrier was gas, large chunks of shale could be used, up to 125 mm in size, for apparently slow heating/reaction rates. When solid heat carrier technology was utilized, the top size had to be reduced to 25 mm, still in a size range that could be attained with crushers, with no grinding required. Grinding characteristics of oil shale are such that each resultant particle size has a different composition. The coarse particles are mainly carbonates, i.e. calcite and dolomite, while the sandy-clay constituents concentrate into the finest particle range. The most kerogen-containing, organic parts remain in the middle. In view of the geological description of oil shale particles and layers, this is not surprising. The carbonates, substantially limestone, are harder to crush and remain in the coarser fractions, the already finely divided sandy-clay material will be liberated into the finest range, and the somewhat sticky organic matter will occupy the middle ground. Heating value (Hv) variations are observed accordingly. The highest, in the test results shown by Ots (2006) are in the range 15–19 MJ/kg, in the size range 40–80 µm. At 20 µm (mostly sandy-clay minerals), Hv ≈ 9–11 MJ/kg and in the size range 150–300 µm, Hv ≈ 4–8 MJ/kg. Density variations are the reverse. Kerogen has the lowest density, reported as 980–1100 kg/m3, with all minerals well above 2200 kg/m3. Aside from a modest level of enrichment at the mine, it is apparently not the practice to concentrate the organic matter prior to use. Some suggestions in that direction have been made, however. Plamus et al. (2011) show significant benefits attainable in CFBC boilers when the oil shale is ground finer and enriched to a higher heating value (Hv from 8.5 up to 11.1 MJ/kg). In view of the different grinding characteristics as well as the large density variation, further separation can readily be accomplished, if warranted, by potentially novel further processing techniques. Indeed, well-enriched and much finer oil shale particles (D50 = 14 µm) have been used in various pyrolysis/kinetics studies (Volkov et al., 2010). A rather complete description of the kukersite kerogen and its geological formation is also given in that paper. The first information and data about scientific research on Estonian oil shale date from 1771 and some activity has continued ever since. Industrial utilization started during World War I, in 1916, initially to make gas for St. Petersburg and soon in the cement industry. The main use of oil shale in Estonia, starting in 1924, has been for power generation where it slowly replaced wood, peat and imported
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coal (Ots, 2006). Oil production from retorting the kukersite shale started seriously in the interwar years and has continued at varying levels of intensity since that time. Extraction of various chemicals has been a third important area of activity. Much of the extensive published literature is available in Russian only, often as Transactions of the Tallinn Technical University. It may be appropriate here to add another such publication which emphasizes the breadth of research and development activities, by Zaretskii et al. (2001). Similar activities for the production of oil and chemicals have more recently been commenced also elsewhere, for example in Australia, Brazil, China and the USA. For either use, it is necessary to crush the mined shale as an essential preparation step. For more rapid thermal processing the oil shale should also be ground more finely, to a size suitable for the equipment used, and enrichment should be considered, certainly to reduce carbonates content, and possibly even the fine sandy-clay components. With greatly enhanced interest in such processing expected in the near future, undoubtedly new processing technologies will emerge and likely require more extensive feedstock preparation. A brand-new book from China, Oil Shale – Petroleum Alternative by Qian (2010) covers everything in its 625 pages, from history and geology to combustion, processing, design and development, and an outlook upon the world’s oil shale industry.
1.2.3 Coal-based feedstocks Before turning to the material in the heading, it will be of some relevance to point out that co-generation, co-processing, etc., are gaining some ground. Combustion of coal in a power plant or industrial furnace along with some biomass is already quite common. The idea is to add the biomass as a ‘sustainable component’, i.e. helping with the issue of greenhouse gases, mainly CO2, and perhaps also getting a reduction in the sulfur content of the fuel, among other potential benefits. Co-processing of coal and biomass, e.g. for gasification, is also gaining some ground, as a small addition of coal or coke to the biomass feed will give greater energy density to the fuel, thus adding more energy to the system. For those cases it is of course important to balance the particle size requirements, as perhaps the most important issue. It is reasonable to include under the ‘coal-based’ heading petroleum coke, which has many similar properties, as well as the youngest of the coals, lignite. Typically, lignites have a higher moisture level and less sulfur. Ash, as always, ranges according to the site and how the material was deposited. Petroleum coke, or pet coke, has usually the opposite special properties: minimal retained moisture, low ash, but a high level of sulfur, typically well above that of coal. As far as coal contaminants are concerned, both ash and sulfur levels vary widely. Pyritic sulfur really fits both categories and can be removed quite readily unless it is finely distributed. The level of organic sulfur in coal is a much more serious issue. Different coal regions and
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seams show rather extreme differences, with Western North American coals having low sulfur levels, whereas those in the East have typically high levels, in the range 3–6% total sulfur, of which up to one-half may be organically bound. Another, totally different case of co-processing is that of simultaneous heavy oil upgrading and coal liquefaction. Both individually studied operations reap significant benefits from the simultaneous process, yet in rapid sequel, (1) this was deemed to be ‘the most valuable development’ in the fossil energy field (U.S. Department of Energy) and (2) promptly terminated by Government of Canada, supposedly at the suggestion of the oil industry. It is time to resurrect this processing technology, with its own coal preparation and catalystis features.
1.3
Coal feedstock characterization and requirements
Information on coal classification and detailed characterization is available in several sources, e.g. handbooks edited by Wen and Lee (1979) and Cooper and Ellingson (1984), or the new Handbook of Combustion (Lackner et al., 2010). Coverage in this section is therefore limited to a brief overview along with an emphasis on those properties that are important in coal conversion, be it liquefaction or gasification. It is also relevant to recall that conversion includes combustion, i.e. conversion directly to thermal energy rather than to other combustible forms. Further, as the coal-based feedstocks are the most important in this context, neither biomass nor oil shales will be covered in any further detail. As has been emphasized by Berkowitz (1994), in the second edition of his book An Introduction to Coal Technology, the section on coal and its properties, Part I, did not have to be renewed to any great extent from the first edition in 1979. Nothing had to be substantially updated. That statement still rings true. Part II, Processing of Coal, in contrast, had to be changed considerably, largely taking account of the work done under the US Clean Coal Technologies Program (U.S. Department of Energy, DOE) and other coal processing activities.
1.3.1 Coal classification and characterization Coal is a complex, heterogeneous material which is abundantly available in many parts of the world. Hence, its great importance, but not without the many processing problems that arise from its complexities, both its chemical composition and physical structure, all influenced by its origin, location and past history over millennia. Origins and formation It is well known that coal is a product of plant matter that has undergone dramatic developments, both biochemical or diagenetic, at the early stages of
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its development, and metamorphic once imbedded in the earth crust and exposed to high temperatures and pressures. The time scale has been of the order of tens or hundreds of millions of years, through many tectonic disturbances and geological ages. As a result, dramatic variations in coal properties should not be unexpected, arising from both the different types of original plant material as well as the local conditions over those millions of years. Combinations of pressure, temperature and time determine the rank of the coal, in a series stretching from lignite through subbituminous and bituminous coals to anthracite and graphite. Distribution Coal resources are, indeed, broadly distributed, as coal is found on all inhabited continents and is also likely on Antarctica. Total world coal resources are immense, measured in millions of megatonnes. In many countries, the proven reserves amount to 100 years or more of usage at current rates, with, undoubtedly, much more yet to be found. Hence, King Coal is of vital importance on the world energy scene. Coal composition Here we will briefly cover both the chemical composition and the petrographic structure of the coals. For chemical composition we use either the ‘proximate’ or the ‘ultimate’ (elemental) analysis. Each has it appropriate uses. The parameters of proximate analysis are fixed carbon, volatile matter, ash or, more properly, ash-forming inorganic mineral matter, and finally moisture. Calorific value is also measured at the same time. The higher the fixed carbon, the higher the coal rank. Thus, anthracite has a fixed carbon content equal to or greater than 85% on a moisture-free basis. For ultimate analysis, the elemental composition of the organic material is reported, usually on a moisture-free basis, as percentage carbon, hydrogen, nitrogen, sulfur (organic) and oxygen. Mineral matter is then handled separately and includes pyritic sulfur. Mineral matter constituents cover a broad range. Some, such as silica, originate from the initial plant material and are, then, very finely distributed. Others have seeped into the coal seams at different times during coal development or have simply been mixed into the coal seams by geological forces. Yet others arise from the way coal is mined, originating from the various soil layers in the vicinity of the coal seams. The last one is particularly important when coal seams and the intervening soil layers are relatively thin and modern mining equipment mixes them up. The most common constituents are clay minerals, with a lot of aluminum; silica, predominantly quartz; various carbonates, mostly from calcium; iron and magnesium, and sulfides, primarily pyrite and marcasite. Other metals found frequently are manganese, sodium and lithium. The minor, or trace constituents of mineral matter are many. Berkowitz (1994, p.48) lists over twenty of these along with their maximum and typical concentrations
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in the ashes. Their influence on downstream processing may or may not be important. A greater concern is often environmental. For example, mercury has become an issue in coal combustion and must be captured before flue gases enter the atmosphere. Petrographic structure As with most other aspects of coal science, the petrographic coal structure is also complicated. Multiple subdivisions can be observed, mostly from reflectance measurements. The first subdivision of structure is termed the lithotype, four groupings of which are defined. These in turn are subdivided into microlithotypes along with their dominant maceral groups, and, of course, their many subdivisions. Let us again refer the reader to Berkowitz (1994) for details and simply list the important maceral groups along with their key properties:
• • •
Vitrinites, or huminites in low-rank coals, are the most reactive of the maceral groups. Exinites comprise another, somewhat less reactive, maceral group. Inertinites, particularly fusinite, are the least reactive of the maceral groups.
The macerals appear in various sizes, have different densities and can be separated by density differences after adequately fine grinding. This is practical for research purposes, in small quantities. Industrial maceral separation is not yet practiced and, instead, different coals can be selected for different purposes, depending on their major maceral compositions. Classification There are many ways to classify coals. Berkowitz (1994) reviews the various systems. Here, let us disregard the differences between humic and sapropelic coals by just stating that the humic ones are the common coals and the sapropelic ones arise mainly from plankton, algae and anaerobically putrefied vegetation associated with marine environments, akin to oil shale kerogen. Various scientific schemes rely on the carbon/hydrogen ratio, oxygen content as well as the amounts of fixed carbon and volatile matter, all related also to calorific value as well as to moisture and ash content. Berkowitz (1994) concludes that no classification scheme can be useful for all processes and operations and, hence, cannot stand in lieu of coal specifications. The simplest practical system disregards chemical and quality details and classifies the coals by age, fixed carbon and volatiles content as well as heating value. The ASTM system (1984), used mainly in North America, distinguishes among four coal classes, each further subdivided, as shown in Table 1.3. We observe the familiar groups: anthracite, bituminous, subbituminous and lignite coals, their volatility levels and their mineral matter-free heating values. It should
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1. Metaanthracite 2. Anthracite 3. Semianthracite
1. Low volatile bituminous 2. Medium volatile bituminous 3. High volatile A bituminous 4. High volatile B bituminous 5. High volatile C bituminous
1. Subbituminous A 2. Subbituminous B 3. Subbituminous C
1. Lignite A 2. Lignite B
I. Anthracite
II. Bituminous
III.Subbituminous
IV. Lignitic
dmmf = dry mineral matter-free; mmf = mineral matter-free
Source: ASTM, 1984
Group
Class
Table 1.3 ASTM classification of coal by rank
LigA LigB
subA subB subc
Lvb Mvb hvAb hvBb hvCb
ma an sa
Abbreviation
Volatile matter (% dmmf) phenol > diaryl ethers = o-alkyl-substituted phenols = alkylfurans > benzofurans > dibenzofurans. In view of available biomass feedstocks, the reactivities of phenols, acids and esters are the most relevant. Phenols, which may account for up to 25 wt% of liquids obtained by pyrolysis of lignocellulosic materials, are refractory oxygenates. The overall mechanism for the hydrodeoxygenation of the o-substituted phenols shown in Fig. 7.7 includes two main hydrodeoxygenation reactions, direct hydrodeoxygenation and hydrodeoxygenation via hydrogenated phenol, occurring in parallel (Furimsky et al., 1986). In the latter case, H2O elimination may result in the formation of the intermediate methylcyclohexene species, which are hydrogenated rapidly. The formation of cyclohexene, alkylcyclohexenes and methylcyclopentanes is also shown in Fig. 7.7, although these are only minor products. Guaicacol (GUA) and substituted GUAs have attracted much attention because of their relatively high content in bio-oils and low stability. The hydrotreatment of guaicacol in the presence of CoMo and NiMo catalysts was studied in a batch reactor by Laurent and Delmon (1994a, 1994b). They proposed the mechanism shown in Fig. 7.8, which considers the hydrogenolysis of the
7.7 Mechanism of hydrodeoxygenation of 2–methylphenols (adapted from Furimsky et al., 1986).
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7.8 Mechanism of hydrodeoxygenation of guaiacol (adapted from Laurent and Delmon, 1994a).
methoxy group of guaiacol to catechol and methane as the first stage, followed by the elimination of one OH group from catechol in the second stage to produce phenol. Coke was formed from both guaicacol and the primary product catechol. Elliott and co-workers developed a two-step hydrotreating process for the upgrading of bio-oils using sulphided Co-Mo/Al2O3 or sulphided Ni-Mo/Al2O3 catalysts (Elliot et al., 1988; Elliot and Neuenschwander, 1996). The first step involves a low-temperature (270 °C, 136 bar H2) catalytic treatment that hydrogenates the thermally unstable bio-oil compounds, which would otherwise undergo thermal decomposition to form coke and plug the reactor. The second step involves catalytic hydrogenation at higher temperature (400 °C, 136 bar H2). During this process, 20–30% of the carbon in the bio-oil is converted into gasphase carbon, decreasing the overall yield. Catalyst stability and formation of gums in the lines were identified as points of major uncertainty of the process, and future work is needed to develop improved hydrotreating catalysts. Bio-oil was hydrotreated at high pressures (138–172 bar) and low space velocities (0.1–0.2 LHSV) by Marinangeli et al. (2006). At these high pressures and low space velocities, hydrodeoxygenation predominates. Large quantities of hydrogen are required to generate water during hydrodeoxygenation because of the high level of oxygen (46%) in bio-oil. The resulting hydrotreated oil was then cracked in an FCC or hydrocracker to produce gasoline. This approach is unlikely to be commercially viable because of the high hydrogen requirement and the high capital cost of the hydrotreatment step. The challenges of feeding hydrodeoxygenated pyrolysis oil in standard refineries have been also studied by de Miguel Mercader et al. (2010). Different HDO reaction end temperatures (230–340 °C) were evaluated in a 5 L autoclave, keeping the other process conditions constant (290 bar, 5 wt% Ru/C catalyst), in order to find the required oil product properties necessary for successful FCC co-processing (miscibility with FCC feed and good yield structure: little gas/coke production and good boiling range liquid yields). After hydrodeoxygenation, the upgraded pyrolysis oil underwent phase separation resulting in an aqueous phase, some gases (mainly CO2 and CH4) and an oil phase. Although the oil and the aqueous phase yields remained approximately constant when the HDO reaction
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temperature was increased, a net transfer of organic components (probably hydrodeoxygenated sugars) from the aqueous phase to the oil phase was observed, increasing the carbon recovery in the oil product (up to 70 wt% of the carbon in pyrolysis oil). The upgraded oils were subsequently tested in a lab-scale catalytic cracking unit (MAT reactor), assessing the suitability of HDO oils to be used as FCC feed. In spite of the relatively high oxygen content (from 17–28 wt%, dry basis) and the different properties of the HDO oils, they all could be successfully dissolved in and co-processed (20 wt%) with a long residue, yielding near normal FCC gasoline (44–46 wt%) and light cycle oil (23–25 wt%) products without an excessive increase of undesired coke and dry gas formation, as compared to the base feed. Samolada et al. (1998) reported a two-step process of thermal hydrotreatment and catalytic cracking of biomass flash pyrolysis liquids (BFPLs). Thermal hydrotreatment of BFPLs can be effectively operated producing liquid products which can be upgraded in a refinery. The heavy liquid product of this process (HBFPL), mixed with light cycle oil (LCO) (15/85 wt%/wt%), was considered as a potential FCC feedstock. Commercially available cracking catalysts were found to have an acceptable performance. The obtained bio-gasoline quality is comparable with that of the VGO cracking but with low yields ∼20 wt%. The upgrading of BFPL to transportation fuels by using a hydrotreating technology was also studied in the Chemical Process Engineering Research Institute (CPERI) in collaboration with Veba Oil (Lappas et al., 2009). With the thermal hydrogenation of bio-oil a deoxygenation conversion of 85 wt% was achieved, producing hydrotreated oil with oxygen content of about 6.5 wt%. Due to the low-oxygen content of the thermally hydrotreated bio-oil, it can be separated by distillation in a light and a heavy fraction. The light fraction comprised components mainly in the gasoline and diesel range, and thus it could be directly blended with the corresponding petroleum fractions. The heavy fraction has similar characteristics to conventional VGO. This heavy fraction could be used as co-feed with vacuum gasoils (VGO). The bio-oil co-processing technology proposed by Lappas et al. (2009) is shown in Fig. 7.9. The co-processing of VGO with the heavy fraction showed that the presence of the bio-oil fraction favours the gasoline and diesel production but increases the coke yield. However, depending on the concentration of biomass liquids, it was shown that this option is technically viable for FCC units running with good quality feedstocks, that is, the FCC unit with excess coke burning capacity. Fogassy et al. (2010) have recently evaluated the impact of adding 20 wt% HDO-oil to a conventional FCC feedstock. The VGO and bio-oil mixtures were co-fed into a fixed-bed reactor simulating FCC conditions using an equilibrated industrial FCC catalyst. Co-processing of 20 wt% HDO-oil with VGO gave comparable yields for the gasoline fraction to that of the pure VGO cracking. However, during co-processing, oxygen removal from HDO-oil oxygenated components consumes hydrogen coming from the hydrocarbon feedstock
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7.9 Bio-oil co-processing technology investigated by Lappas et al. (2009).
7.10 Reaction pathways for the catalytic cracking of HDO-oil oxygenates (adapted from Fogassy et al., 2010).
(Fig. 7.10). As a result the final product composition is poor in hydrogen and contains more coke, aromatics and olefins. Moreover, the phenolic fraction was not completely converted.
7.4.2 Hydrotreating of triglyceride-based feedstock: green diesel Biodiesel is usually prepared by the transformation of vegetable oils and fats through triglyceride transesterification towards fatty acid methyl esters (FAME). This well-established conversion method allows upgrading of these renewable energy resources to get diesel-type fuels, showing high energy content with low
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amounts of oxygen (Huber et al., 2007). However, the final products achieved through this methodology display several properties that do not match the specifications for automotive fuels – for instance high density and boiling point and poor stability against oxidation (Tang et al., 2008), leading to several technical problems, like filters between the fuel tank and engine becoming blocked (Donnis et al., 2009). Besides, the production of FAME leads to the formation of glycerol as a by-product which needs to be marketed – a difficult task because of the abundance of this chemical – to make the process profitable. These and other problems do not make attractive the transesterification of triglyceride oils to conventional petroleum refineries. As an alternative, the transformation of triglycerides in a refinery hydrotreating unit to form lineal alkanes allows production of diesel-like fuels. Though hydrotreating, because of the needing for hydrogen, is much more expensive than other refinery alternatives, such as catalytic cracking, the products achieved through this pathway – green diesel – are pure hydrocarbons indistinguishable from their petroleum counterparts. In fact, green diesel displays a high cetane number, showing good potential to meet current specifications for petroleum-derived diesel-like fuels. In this sense, several studies reveal a meaningful decrease of hydrocarbons, carbon monoxide and nitrogen oxides in diesel engine emissions when using green and conventional diesel mixtures instead of petroleum-derived diesel as fuels (Soveran et al., 1992; Stumborg et al., 1996; Aatola et al., 2008). Apart from these beneficial features, the green diesel option to process triglyceride-oils displays an important advantage over the transesterification production process in hydrotreating units that are actually present in refineries, where high vacuum gasoil (VGO) is already treated with hydrogen. Chemical compatibility allows processing of these feedstocks together with crude-derived fractions in the current infrastructure (Holmgren et al., 2007), so that there is no need to build new plants (Huber and Corma, 2007). From a chemical point of view, the total hydrogenation of triglycerides leads to n-alkanes and propane as the main reaction products and water, CO and CO2 as by-products. This process involves several reactions that can be lumped into two reaction pathways (Donnis et al., 2009; Semejkal et al., 2009): hydrodeoxygenation (HDO) and hydrodecarboxylation (HDC). The former involves the formation of long n-alkanes showing the same number of carbon atoms than the original fatty acid alkyl chain as well as propane, coming from the hydrodeoxygenation of glycerine. On the contrary, hydrodecarboxylation involves the loss of a carbon atom – the atom economy suffers some decrease – coming from the carboxylic group at the fatty acid chain. In this way, the main n-alkanes obtained as products display one less carbon atom than the original fatty acid alkyl chain, this carbon being lost as carbon monoxide or dioxide. An overview of these reaction pathways is depicted in Fig. 7.11. Considering the hydrodeoxygenation mechanism, the mass balance indicates the need for 12 moles of hydrogen per mole of triglyceride plus an additional mole of H2 per double bond present in the fatty acid alkyl chains. This is the faster
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7.11 Proposed scheme for triglyceride hydrotreating main reaction mechanisms. Major products have background tint.
transformation, and thus occurs in the first term, when hydrotreating triglycerides (Huber and Corma, 2007). Thus, for instance, the total hydrogenation of rapeseed oil (4 double bonds per mole) by HDO needs of 16 moles of hydrogen per mole of triglyceride, leading to a mixture of water (six moles), propane (one mole) and a mixture of n-C18 and n-C22 (three moles) – rapeseed oil fatty acid profile is mainly composed by oleic, linoleic and linolenic (C18 ∼35 wt%) acids and erucic acid (C22 ∼ 40 wt%). On the other hand, if considering the hydrodecarboxylation mechanism, only 3 moles of hydrogen are needed to process a mole of triglyceride plus the additional H2 to reduce each double bond (7 moles of H2 per mole of rapeseed oil). This comparison suggests hydrodecarboxylation should be favoured over the hydrodeoxygenation to reduce hydrogen consumption, a critical issue in the profitability of hydrotreating units, but apart from the main reactions proposed in Fig. 7.11, several other chemical gas-phase transformations have to be considered in the HDC mechanism. The mere presence of CO2 in the reaction system leads to the existence of methanation, or at least the partial reduction of the same, as well as the water gas shift reaction (Snåre et al., 2009), though to a minor extent (Fig. 7.12). Thus, it is important to consider these reactions when evaluating the process economy since they lead to the consumption/production of hydrogen.
7.12 Secondary reactions taking place during triglycerides hydrotreating.
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Thus, when considering the methanation of carbon dioxide – or its transformation into carbon monoxide and subsequent methanation – together with the hydrodecarboxylation of starting triglycerides, twelve additional hydrogen moles should be added to the three already considered from the HDC transformation, plus the hydrogen consumed in reducing hydrocarbon bonds. Hydrodecarboxylating rapeseed oil would need 19 moles of hydrogen per mole of triglyceride, leading to n-C17 and n-C21 as main products (three moles), water (six moles), propane (one mole) and methane (three moles). If considering the hydrogen consumption, it seems that the hydrodeoxygenation pathway is more attractive than hydrodecarboxylation, but such a low difference between H2 consumption by both mechanisms and the similarities observed on the final hydrocarbon products do not allow determination of which is the best option, being dependent on the process and on the catalyst used in the hydrotreating unit. The most important challenge when using triglyceride-containing vegetable oils in conventional refinery hydrotreating units seems to be the development of an adequate catalytic technology for treating these new biomass feedstock, but the design of the catalyst depends on the desired reaction pathway. As previously stated, the hydrotreating of triglycerides can follow two reaction pathways, hydrodeoxygenation and hydrodecarboxylation, and both alternatives have been a matter of intensive research looking for suitable catalytic technology for driving these reactions. However, in both cases, the reported research has followed different paths (Krár et al., 2010): current hydrodesulphuration technology in petroleum industry has been investigated for the oxygen removal of triglycerides, because hydrodesulphuration and hydrodeoxygenation are, a priori, rather similar reactions. On the contrary, triglyceride decarboxylation has been investigated through the development and use of a different catalytic technology: supported noble metals (Snåre et al., 2006). The hydrodecarboxylation route in hydrocarbons production from free fatty acids has been investigated through the use of different conventional hydrogenation catalysts, employing a huge variety of active metal species supported on few catalytic supports. Reaction experiments indicate that supported metal carbonbased catalysts display a much higher selectivity towards hydrodecarboxylated products than analogue catalysts based on different supports (Snåre et al., 2006). With regards to the active metal species, palladium and platinum display a much better catalytic performance than other metals like ruthenium, rhodium or iridium. However, in terms of the efficient use of the hydrogen, which is the key factor in hydrotreating units, palladium is superior to platinum because the former mainly drives decarboxylation reactions, whereas platinum produces hydrodecarbonylation – which is the removal of the carboxylic group of free fatty acids leading to carbon monoxide and water, a transformation which involves a higher hydrogen consumption. Nevertheless, the preparation conditions used for these Pd/C materials exert a dramatic influence on their catalytic behaviour, the particle size of the final supported metal species being one of the most important variables determining the catalytic activity (Simakova et al., 2009).
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The use of these catalysts for treating different feedstocks has also been reported. Thus, free fatty acids, alkyl esters and triglycerides have been assayed, leading in every case to the same major products, which are the hydrocarbons – both saturated and unsaturated – formed by an alkyl or alkenyl chain with one carbon atom less than the original fatty acid chain (Kubicˇková et al., 2005). However, the hydrodecarboxylation reaction rate is rather low in comparison with double-bond hydrogenation rate. Besides, hydrotreating free fatty acids seems to be rather effective and fast, being more difficult than the hydrodecarboxylation of alkyl esters (Snåre et al., 2008) and even more complicated triglycerides. These differences seem to be caused by a different reaction mechanism. Thus, whereas the deoxygenation of free fatty acids follows a hydrodecarboxylating pathway, in the case of fatty acid alkyl esters it mostly proceeds via hydrodecarbonylation (Mäki-Arvela et al., 2007), though both types of reactions coexists in both cases (Snåre et al., 2009), and the dominant one can be tuned depending on the catalyst and the reaction conditions; for example, increasing hydrogen concentration enhances the decarboxylating activity. Bearing in mind that hydrodesulphuration and hydrodeoxygenation involve analogous reactions, triglyceride oxygen removal by hydrodeoxygenation seems to be a rather easy task to implement on conventional refinery hydrotreating units used for hydrodesulphuration of petroleum-derived streams. Hydrodesulphurization (HDS) is a widespread and mature technology conventionally used in refineries. HDS usually coexists with HDO and hydrodenitrification (HDN) for the removal of sulphur, oxygen and nitrogen, respectively. The most important industrial catalysts used for HDS conventionally involve alumina-supported molybdenum and tungsten sulphides as main catalytic species, usually promoted with cobalt and/or nickel (Kubicˇka, 2008). Nevertheless, due to the higher catalytic activity, Co-Mo and Ni-Mo are the most widespread catalysts in hydrotreating units (Babich and Moulijn, 2003). These catalysts are employed because of their high resistance against sulphur poisoning, in contrast with noble metals which display a much higher hydrogenating activity but a lower resistance against deactivation with sulphur. Despite the chemical analogy between hydrodeoxygenation and hydrodesulphuration, important different behaviours are found for the HDS catalysts when hydrotreating triglyceride-containing feedstocks in comparison with petroleum-derived streams. These differences are associated with the feedstock nature and the final form of the heteroatoms to be removed (oxygen and sulphur), even though the reaction pathways and mechanisms are rather similar in HDS and HDO. Thus, hydrodesulphurization leads to the formation of sulphidric acid (H2S) whereas hydrodeoxygenation leads to the formation of water (H2O), both of them interacting with the surface of sulphided catalysts (Ferrari et al., 2001). It is well known that sulphidric acid promotes the acid catalytic activity of these catalysts (Laurent and Delmon, 1994b), enhancing the reaction rate of the acid-driven transformations. On the contrary, water seems to exert a negative
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influence on the catalytic activity of HDS catalysts, but this is only because of the presence of H2S, since H2S and H2O compete for their coordination to metal sites (Laurent and Delmon, 1994b). In case of hydrotreating sulphur-free feedstocks, like triglyceride containing feedstocks, H2S is absent and thus water is free to interact with hydrogenating sites, leading to the inhibition of the hydrogenation reactions (S¸enol et al., 2005a). The work from S¸enol et al. (2005b) described the hydrogenation of model aliphatic methyl esters in presence of alumina-supported sulphided CoMo and NiMo catalysts as a sequence of three reaction pathways: the formation of alcohols which evolve towards hydrocarbons by dehydration, the deesterification between the alcohol and the carboxylic acid functionality and the hydrogenation of the carboxylic acid towards hydrocarbon, either passing through the alcohol or not. Thus, dehydration, hydrolysis and hydrogenation reactions are present at the same time when hydrotreating these triglyceride-containing feedstocks, the two first reactions being promoted by acid catalysis whereas the last one is driven by the hydrogenating activity. In fact, some decarboxylating activity (S¸enol et al., 2005b) has also been found to be present, which could not be avoided, at least under the employed reactions conditions. This last reaction is also driven by the acid sites of these catalysts, whose presence is associated to sulphydryl acid groups (Ferrari et al., 2001). A proof of this behaviour is the low hydrodecarboxylating activity of the same catalysts when used as non-sulphided metal oxides (S¸enol et al., 2005b). One possibility to avoid this deactivation is to drive these HDO reactions in the presence of sulphidric acid, but the enhancement of acid activity leads to a meaningful increase of the extent of hydrodecarboxylation reactions (Laurent and Delmon, 1994b; Ferrari et al., 2001). On the contrary, hydrogenolysis and hydrogenation reactions are consequences of the presence of sulphur vacancies associated with molybdenum atoms (Leliveld et al., 1998). When treating more realistic feedstocks like sunflower (Krár et al., 2010) or waste cooking oils (Bezergianni et al., 2010a, 2010b), the observed reaction pathways were in essence the same as those previously described for model compounds, suggesting that the performance of conventional hydrodesulphuration catalysts was almost the same. However, several interesting results are found when treating triglyceride feedstocks. Increasing the reaction temperature leads to a much higher extent in the removal of oxygen, though this enhancement on triglyceride conversion is accompanied by a much higher rate in hydrodecarboxylation, isomerization and even hydrocracking reactions, which are more important when using NiMo/γ-Al2O3 catalysts (Gusmão et al., 1989; da Rocha Filho et al., 1993; Šimácˇek et al., 2009). Thus, though the triglycerides conversion is higher insofar as the operation temperature increases, the yield towards green diesel decreases – a proof of this fact is the decreasing of the cetane number in the final product or the increasing of the bromine index when operating at high temperature (∼400 °C). In this way, it seems preferable to be operating at
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lower temperatures and recirculating the heavy fraction and the residue to maximize the yield towards biodiesel. Some modifications of conventional hydrodesulphuration catalysts have been reported, looking for better catalytic activity of the same in triglyceride HDO treatments. Most of these improvements lay on the modification of the catalytic support, mainly tackled through the enhancement of the support surface area (Kubicˇka et al., 2009) or the assay of different supports like silica or silica alumina mixed oxides (Liu et al., 2009; Kubicˇka et al., 2010). Nevertheless, though the performance of these modified HDS catalysts is good there are still scarce examples of the adaptation of the same for hydrotreating triglyceride-containing feedstocks, and much more effort needs to be made. As an alternative, the use of conventional HDS catalyst for treating mixed blends composed by triglyceridecontaining oils and petroleum-derived oils is now becoming of more and more interest to petroleum refineries. This strategy is readily applicable in conventional refineries without the need to implement large modifications to existing hydrotreating units. The coprocessing of vegetable oils with petroleum-derived streams has been tackled through the use of conventional hydrodesulphuration catalysts instead of supported noble metals because of the sulphur resistance of the former. Thus, Huber et al. (2007) reported the use of sunflower oil together with heavy vacuum oil (HVO) as feedstock to be treated in a hydrotreating unit in presence of a conventional sulphided NiMo/γ-Al2O3 catalyst. Interestingly, this treatment option produced, under certain conditions, much better results than treating both feedstocks separately. Thus, treating both feeds together led to a higher amount of straight alkyl chains in the range C15 to C18 than if treating pure sunflower oil – the given reason for this improvement is that dilution of free fatty acids (FFA) inhibits polymerization and hydrocracking reactions (Lappas et al., 2009). Besides, since hydrodesulphuration is a much slower reaction than alkane production from the vegetable oils, the use of feedstocks mixtures does not affect the rate of desulphuration. Similar findings to this pioneer work have also been found using CoMo/ γ-Al2O3 catalyst in the hydrotreatment of cottonseed oil (Sebos et al., 2009). In addition, quality enhancement of several properties of the final products were also found, like a cetane index (Bezergianni et al., 2009) which showed higher values compared to that achieved when treating only petroleum fractions. The direct application of the existing infrastructure in petroleum refineries for treating petroleum streams–vegetable oils mixtures makes possible extensive industrial application in the near future. Much important industrial experience has been gained in the use of vegetable oils and animal fats in existing industrial hydrotreating units, and much research effort has been carried out by oil refining companies. Most of these developments consist of modified gasoil hydrotreating processes to which a blend of gasoil and triglyceride-containing mixtures is fed (Mayeur et al., 2009; O’Connor et al.,
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7.13 Flow diagram and reaction steps of the UOP/ENI Ecofining process.
2008). However, Neste Oil OYJ, a Finnish oil refining company, has taken a step forward and has licensed a new process, NExBTL, for the production of green diesel from pure vegetable oils and fats (Snåre et al., 2009; Markkanen et al., 2010). The first two production plants (170 000 tons per year) for this process were built in Porvoo refinery. In addition, two new plants, one in Rotterdam and another in Singapore, are now under construction (800 000 tons per year). A similar process is that licensed by UOP/ENI, the Ecofining process (Kalnes et al., 2007; Baldiraghi et al., 2009), which also involves the same two reaction stages as the NExBTL process (Fig. 7.13). Thus, the two distinct stages comprise the hydrodeoxygenation step into which triglyceride-based biomass and hydrogen are fed. Obviously, as previously stated, not only HDO reactions occur during this step, but HDC and subsequent methanation also takes place in this first reaction step. The second stage involves the hydroisomerization of the deoxygenated product to improve its cold properties. Light ends can be used to produce hydrogen, which is then recycled to the reaction stages in order to increase the profitability of the processes. In essence, both processes, NExBTL and Ecofining, are rather similar and the final products coming from the process are the same: light naphtha, propane and green diesel. This last is the major component, comprising more than 80 wt% of the final products. Final properties of this green diesel are rather similar to those achieved through the production of Fischer–Tropsch liquid fuel from syngas. Table 7.3 lists the properties calculated for diesel-like fuels, including green diesel and its blending with conventional diesel fuel. These fuels are featured by a high cetane number, which involves very good engine efficiency, low oxygen content,
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Table 7.3 Properties for different diesel-like fuels obtained through different techniques Property
Diesel fuel1
Biodiesel2
Cetane number
53
50
70–90
Oxygen content (wt%)
0
11
0
Sulphur content (wt ppm)
diphenyl sulfide > 4,6-dimethyldibenzothiophene > 4-methydibenzothiophene > dibenzothiophene > benzothiophene > thiophenes. This trend confirms that the refractory sulfur compounds in HDS are the most reactive in the oxidation. The reactivity of the compounds seems to correlate well with their electron density except for the dibenzothiophenes with methyl substitutes at 4 and 6 positions.3,6,11,13,21,22,48 Two main catalysts used for selective desulfurization are organic acids and polyoxometalates. Organic acids include formic acid, acetic acid, etc. Polyoxometalates have long been studied for oxidation reactions, particularly the polyoxometalate/ hydrogen peroxide system for organic substrate oxidations.3,6,11,21,22,48 The tungsten and molybdenum polyoxoperoxo species in the presence of hydrogen peroxide were also studied. Mure Te et al. have carried out a comparative study of the dibenzothiophenes oxidations using a series of polyoxometalates as catalyst precursors to better understand the oxidation reactivity of the typical refractory sulfur compounds in diesel fuels18. Using model compounds of diesel, experiments were carried out to compare the reactivity of the different dibenzothiophenes in oxidation reactions using polyoxometalate/H2O2 systems. Oxidation reactivity decreased in the order of dibenzothiophene > 4-methyldibenzothiophene > 4,6-dimethyldibenzothiophene, and the same reactivity trend that exists in HDS (Fig. 8.4). Oxidation of the
8.4 4,6 DMBT conversions as function of reaction time at various reaction temperatures. Data points represent experimental results. Lines are predictions from a kinetics model.
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dibenzothiophenes was achieved under low reaction conditions but the reaction rate increased with temperature or reaction time even for the least reactive 4,6-dimethyldibenzothiophene. There is a decrease in reactivity of dibenzothiophenes as methyl substitutes increased at the 4 and 6 positions on dibenzothiophene rings. Interestingly, in a formic acid/H2O2 system, the oxidation reactivity of the dibenzothiophenes showed the reverse trend, suggesting that stereoscopic hindrance might play a role when bulky polyoxoperoxo species, which likely form in a hydrogen peroxide solution, act as catalysts.18 The polarity of the hydrocarbon fuels increases with increasing aromatic concentration, such that the solubility of the sulfones in the oil also increases with increasing aromatic concentration. As a result, the extractability of the sulfones from high-aromatic-content oil is relatively low. The results show that the highly substituted sulfones of DBTs and BTs are removed more easily compared to those sulfones having alkyl substituents of low carbon number. The extractability of sulfones depends on the dipole moment values for the compounds. Highly alkyl-substituted DBT-O2 and BT-O2, of high dipole moment values, is therefore extracted easily from hydrocarbon fuels into acetonitrile solution. The aromatics in hydrocarbon fuels are oxidized to their corresponding carbonyl compounds, which impacts fuel quality.50
8.3.5 Review of inventions and recent innovations Some of the granted patents in oxidative desulfurization technology are reviewed in Table 8.3. Guth and Diaz (1974) and Guth et al. (1975) used nitrogen dioxides followed by extraction with methanol to remove both sulfur and nitrogen compounds from petroleum. Tam and Kittrell (1984) removed sulfur and heteroatom nitrogen compounds from shale oils, by reacting the oil with an oxidizing gas containing nitrogen oxides and then extracting the oxidized oil with solvents. Patrick et al. (1990) used nitric acid or nitrogen oxides as oxidants. The oxidized products are composed of a liquid phase and a byproduct that is a semisolid-like residue with high sulfur content. US Patent No. 3,847,800, Guth and Diaz proposed a process for treating diesel fuel that used oxides of nitrogen as the oxidant. Nitrogen oxides have several disadvantages. In the presence of oxygen, nitrogen oxides initiate a very non-selective form of oxidation, termed auto-oxidation. Several side reactions also take place including the creation of nitro-aromatic compounds, oxides of alkanes and arylalkanes, and auto-oxidation products. Nitrogen oxide based oxidants do not yield the appropriately oxidized sulfur compounds in distillate hydrocarbon fuels without creating many undesirable byproducts. The Guth and Diaz patent proposes the use of methanol, ethanol, a combination of the two and mixtures of these and water as an extraction solvent for polar molecules. Although these have been proved to be acceptable extraction solvents for this system, they do not perform as well as others.32
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Table 8.3 Some of the granted patents in oxidative desulfurization technology31–46 Technology type
Patent no.
Main features
Oxidation + solvent extraction.
6,406,616 (Unipure) • Hydrocarbons are contacted at and report from elevated temperatures with an BP, PetroStar oxidizing/extracting solution of and others formic acid, a small amount of hydrogen peroxide, and nearly 25 wt % water. • H2O2 used as oxidizing agent followed by silica adsorption of oxidized sulfur compounds. • Pyroxyacetic acid is used also to oxidize sulfur compounds followed by extraction process using methanol, dimethyl formamide and methyl pyrrolidone or acetone. • Disadvantage: in some cases, much of the usable oil (30 %) is removed along with the sulfur compounds. Large volume of solvent is required.
Peroxometal
5,958,224 (Exxon)
• Oxidizing sulfur compounds to sulfoxides and sulfones using peroxometal complex as oxidizing agent. • Producing diesel oil < 50 ppm sulfur.
Different types of acid
5,087,350 (Laboratorios Paris)
• Removing S and Fe from coal or oil and their derivatives or from minerals by oxidative extraction performed with hypochlorous acid (HClO). • Phosphotungstic acid + H2O2 system can be as oxidizing agent.
HDS catalyst
6,277,271 (UOP LLC)
• Hydrocarbonaceous oil + recycle stream containing sulfur-oxidized compounds contacted with a HDS catalyst. • Resulting hydrocarbonaceous stream + oxidizing agent give sulfur-oxidated compounds. • Hydrocarbonaceous oil with less sulfur-oxidated compounds. • Portion of the sulfur-oxidated compounds is recycled to the HDS reaction.
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H 2O 2
6,881,325 (BP)
• Mixing the feed with ammonium salt + acid • Producing less sulfur and/or less nitrogen fuel. • Immiscible acid phase is recycled to oxidation zone.
NOx and methanol
3,847,800 and 4,746,420 (REI Tech.)
• NOx used as oxidant agent to oxidize sulfur compounds. (Mainly producing sulfones). • Then, methanol or/and ethanol or/ and mixture of both with water are used as solvent for polar molecules. • Disadvantage: several side chain reactions including the creation of nitro-aromatic compounds, oxides of alkanes and arylalkanes.
Thermal oxidation
3,847,798 and 5,288,390 (Sun Company)
• Oxidation of high boiling petroleum fractions above 550 °F. • Treating the oxidized compounds at about 1300 °F to produce H2S. • Disadvantage: only small degree of desulfurization and loss of valuable products due to cracking and coke formation.
Biodesulfurization
6,130,081 • Decomposing heterocyclic sulfur (PEC, Tokyo) compounds by cleaving C–S and 5,968,812 (EBC, bonds under high-temperature Woodlands) conditions in isolated microorganisms. • Improving the removal of sulfinate group from an organosulfinate compound using an effective amount of Lewis acid.
Ultrasound
6,827,844 (SulphCo) and 6,402,939 (SulphCo)
• Fossil fuel mixed with aqueous liquid + dialkyl ether all subjected to ultrasound. • Producing an organic phase as desulfurized fuel. Also • Fossil fuel mixed with H2O2 in aqueous organic subjected to ultrasound. • Oxidizing sulfur compounds to sulfones. • Removing sulfones by polar extraction.
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One oxidative desulfurization method involves the oxidation of high boiling petroleum fractions (above about 288 °C) followed by treating the oxidized compounds at elevated temperatures (to about 700 °C) to form hydrogen sulfide and/or hydro-processing to reduce the sulfur content of the hydrocarbon material. See, for example, US Pat. No. 3,847,798 (Sun Yoo)31 and US Pat. No. 5,288,390 (Vincent A Durante).37 Such methods have proven to be of only limited utility since only a small degree of desulfurization is achieved and there is substantial loss of valuable products that may result due to cracking and/or coke formation. Therefore, there was a need to develop a process which gives an increased degree of desulfurization while decreasing cracking or coke formation.31,37 Several different oxygenation methods for improving fuels have been described in the past. For example, US Pat. No. 2,521,69829 describes a partial oxidation of hydrocarbon fuels as improving cetane number. This patent suggests that the fuel should have relatively low aromatic ring content and high paraffin content. US Pat. No. 2,912,31330 states that an increase in cetane number is obtained by adding both a peroxide and a dihalo compound to middle distillate fuels. US Pat. No. 2,472,15228 describes a method for improving the cetane number of middle distillate fractions by the oxidation of saturated cyclic hydrocarbon or naphthenic hydrocarbons in such fractions to form naphthenic peroxides. This patent suggests that the oxidation may be accelerated in the presence of an oil-soluble metal salt as an initiator, but carried out in the presence of an inorganic base. The naphthenic peroxides formed are undesirable initiators for gum formation. Consequently, gum inhibitors such as phenols, cresols and cresyic acids must be added to the oxidized material to minimize or prevent gum formation. These latter compounds can be toxic and carcinogenic.28–30 US Pat. No. 4,494,961 (Chaya Venkat and Dennis E. Walsh)33 relates to improving the cetane number of raw, untreated, highly aromatic, middle distillate fractions having a low hydrogen content by contacting the fraction at a temperature of from 50–350 °C and under mild oxidizing conditions in the presence of a catalyst which is either (1) an alkaline earth metal permanganate, (2) an oxide of a metal of Groups IB, IIB, IIIB, IVB, VB, VIB, VIIB or VIIIB of the periodic table, or a mixture of (1) and (2). European Patent Application 0,252,606 A2 also relates to improving cetane number of a middle distillate fuel fraction which may be hydro-refined by contacting the fraction with oxygen or oxidant, in the presence of catalytic metals such as tin, antimony, lead, bismuth and transition metals of Groups IB, IIB, VB, VIB, VIIB and VIIIB of the periodic table, preferably as an oil-soluble metal salt. The application states that the catalyst selectively oxidizes benzylic carbon atoms in the fuel to ketones.33 US Pat. No. 4,723,963 to William F. Taylor34 suggests that cetane number is improved by including at least 3 wt% oxygenated aromatic compounds in middle distillate hydrocarbon fuel boiling in the range of 160–400 °C. This patent states that the oxygenated alkyl aromatics and/or oxygenated hydroaromatics are preferably oxygenated at the benzylic carbon proton.34
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Table 8.4 shows some of the latest filed patent applications in oxidative desulfurization47 Oxidative desulfurization of middle distillates by reaction with aqueous hydrogen peroxide catalyzed by phosphotungstic acid and tri-noctylmethylammonium chloride as phase transfer reagent followed by silica adsorption of oxidized sulfur compounds has been described by Collins et al.7 This described the oxidative desulfurization of winter-grade diesel oil which had not undergone hydrotreating. While Collins et al. suggest that the sulfur species resistant to hydrodesulfurization should be susceptible to oxidative desulfurization, the concentrations of such resistant sulfur components in hydrodesulfurized diesel may already be relatively low compared with the diesel oils treated by Collins et al.7 Table 8.4 Some of the latest filed patent applications in oxidative desulfurization47 Technology type
Patent application no. Main features
O2 gas
20060021913 (BP)
• Contacting the feedstock with an O2 gas in an oxidation zone in the presence of an oxidation catalyst comprising a zeolitic material. • Reducing the sulfur and/or nitrogen content of a distillate feedstock.
Acetic acid
20040118750 (BP)
• Contacting a hydrocarbon feedstock with an immiscible phase containing hydrogen peroxide and acetic acid in an oxidation zone to selectively oxidize the sulfur and nitrogen species. • Then the hydrocarbon phase is passed to an extraction zone wherein aqueous acetic acid is used to extract a portion of any remaining oxidized impurities. • The acetic acid phase effluent from the oxidation and the extraction zone is recovered for optional recycle back to the oxidation and extraction zones.
HDS catalyst
20050040078 (UOP)
• Hydrocarbonaceous oil + recycle stream containing sulfur-oxidized compounds contacted with a HDS catalyst. • Resulting hydrocarbonaceous stream + oxidizing agent give sulfur-oxidated compounds. • Hydrocarbonaceous oil with less sulfur-oxidated compounds. • Portion of the sulfur-oxidated compounds is recycled to the HDS reaction. (Continued overleaf)
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Table 8.4 (Continued) Technology type
Patent application no. Main features
H2SO4
20050218038
• Contacting polyene compounds with sulfuric acid at ambient temperature to form two phases: • Hydrocarbon phase with no loss in octane number and a • Spent acid phase. • Hydrocarbon phase then can be sent to desulfurization process
O2 gas + metal oxide catalyst
20050150819
• Converting the heterocyclic sulfur compounds in the gas phase oxygenated products as well as sulfur-deficient hydrocarbons by contacting it with metal oxide catalyst in the presence of O2 gas. • Recovering the oxygenated products separately from a hydrocarbon stream with substantially reduced sulfur.
Sonic Energy
20050182285
• Exposing the crude or organic liquid to a sonic energy at different frequency help to reduce the sulfur and nitrogen level. • Advantages: • Can be performed with or without oxidizing agent. • With or without elevated temperature or pressure. • Can be performed as either continuous process or batch process.
Ultrasound
20040222131
• The sulfur species present in the crude oil or other refinery streams is oxidized to sulfoxides and/or sulfones by ultrasound with or without oxidizing agent. • The oxidized sulfur species are converted to hydrogen sulfide gas by hydrodesulfurization processes.
US Pat. No. 4,746,420, issued to Darian and Sayed-Hamid, also proposes the use of nitrogen oxides to oxidize sulfur- and nitrogen-containing compounds followed by extraction using two solvents – a primary solvent followed by a cosolvent that is different from the primary.19 The sulfur and nitrogen results published in this patent are consistent with those expected from incomplete oxidation of these compounds followed by extraction.35
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Tetsuo claims many oxidants as being essentially equal in their ability to oxidize sulfur- and nitrogen-containing compounds.10 Many of these oxidants are not selective as found later. Oxidizers that proceed by an auto-oxidation mechanism involving a free radical tend not to be selective for the sulfur- and nitrogen-containing compounds of interest, producing numerous side reactions and, hence, various undesirable byproducts. The author does not provide the level of sulfur removal that his method achieves. Zannikos et al. describe an oxidation and solvent extraction technique for the removal of sulfur-containing compounds.51 Peroxyacetic acid was used in an inefficient manner to oxidize the sulfur compounds in a diesel fuel. Methanol, dimethyl formamide and N-methyl pyrrolidone were used as simple one-stage extraction solvents at different ratios. The results of their work show these solvents removed much of the usable oil along with the oxidized sulfur compounds. To get sulfur levels of approximately 500 ppm with these solvents, the researchers reported a loss of 30% or more of the overall fuel. Such a loss is completely unacceptable on a commercial basis. In a 1993 European Patent, Funakoshi and Aida claim a method of recovering organic sulfur compounds from liquid oil using oxidizing agents, followed by distillation and solvent extraction or adsorption.9 The organic sulfur is recovered as sulfones or sulfoxides. Organic sulfur compounds in fuels could be effectively recovered by a simple solvent extraction process.35 Using acetone, dimethylformamide or other solvents, more than 90% sulfur removal from various hydrocarbon fuels (ranging from gasoline to straight-run bottoms) could be achieved through six to eight stages of extractions with a solvent to oil ratio of 1:1. When an oxidation step is applied before the extraction, an even higher degree of sulfur removal is obtained. Earlier work at Alberta Research Council by McFarlane and Hawkins has shown that organic sulfur in bitumen and synthetic crude oil could be converted to sulfones by hydrogen peroxide or performic acid although these researchers found that the extraction of sulfur compounds from bitumen was ineffective. Researchers from BP Chemical reported that dibenzothiophene could be 100% converted to sulfones by using a phosphotungstic acid/hydrogen peroxide system under mild conditions.16 Treatment of gas oils with the phosphotungstic acid/ hydrogen peroxide system shows that all the sulfur compounds present are oxidized. The results also suggest that highly substituted dibenzothiophenes are the most readily oxidized species containing a thiophenic nucleus. Zannikos et al. report that a combination of oxidation with solvent extraction is capable of removing up to 90% of the sulfur compounds in petroleum fractions at acceptable liquid yield.51 The oxidation process itself leads to substantial sulfur removal without affecting the boiling point distribution. Dolbear and co-worker report that the more refractory sulfur compounds could be removed effectively using appropriate oxidants and catalysts at near-ambient temperature and pressure.4,5,12 PetroStar Inc. is one of the companies that are seriously pursuing this approach.20,27,49 Because of their leading work in this direction, PetroStar was selected by the US Department of Energy as
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one of the three teams to lead the development of ultra-clean fuels by developing new refining processes that remove sulfur pollutants from crude oil.27 The desulfurization reactivity of actual hydrocarbon fuels, such as straight-run light gas oil (LGO), commercial hydrocarbon fuels (CLO) and light cycle oil (LCO), of differing sulfur and aromatic concentrations, was studied using oxidant system H2O2 and AcOH, by Yasuhiro Shiraishi et al.51 The desulfurization efficiency for hydrocarbon fuels fell in the order LGO > CLO > LCO. This order is the same as that of the aromatic concentration in hydrocarbon fuels, and demonstrates that high-aromatic-content hydrocarbon fuels are difficult to desulfurize. The sulfones formed by oxidization are not only removed into the aqueous phase but also form an insoluble precipitate and remain in the hydrocarbon fuels. The low desulfurization efficiency for hydrocarbon fuels is caused by the accumulation of sulfones in the resulting oil. UniPure ASR-2 is a two-step process. The sulfur species are oxidized into sulfones using an oxidant carried in the aqueous phase with a liquid catalyst. This reaction consumes ’small’ amounts of oxidant; conversion of sulfur compounds into sulfones occurs quickly with a 5 min reactor residence time. After separation from the oil, the aqueous phase – containing spent catalyst and some sulfones – is sent to a recovery section. At recovery, the sulfones are removed and the catalyst is regenerated. The oil phase, which also carries some sulfones, is forwarded to an extraction step that uses a solid adsorbent. Methanol is used to regenerate this Adsorbent. The final product is diesel with 5 ppm of sulfur. Recovered sulfones can be disposed off in the refinery coker. One ton per day of sulfone is extracted for each 1000 b/d of diesel (500 ppm sulfur) processed with present equipment configurations for low sulfur specifications.
8.3.6 Process development A typical oxidative desulfurization process looks relatively simple. The process is simply composed of a hydrodesulfurization process to reach about 500 ppm of sulfur and followed by oxidative desulfurization process to reduce the sulfur level to about 10 ppm. Several major problems are impediments for commercial development. First, the oxidants chosen do not always perform selectively. Many oxidants engage in unwanted side reactions that reduce the quantity and quality of the hydrocarbon fuels. The second problem is the selection of a suitable solvent for the extraction of the sulfur compounds. Using improper solvent may result in removing desirable compounds from the fuel or extracting less than a desired amount of the sulfur compounds from the fuel. Finally, further work is needed to identify the most effective catalyst and economical means to use or dispose of the byproducts.
8.3.7 Process economics Lyondell Chemical developed a semi-commercial process at a 1 b/d in-house pilot unit using t-butyl hydroperoxide oxidant which has the advantage of being fuel
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Table 8.5 A comparison between conventional hydrotreating and oxidative desulfurization process14 Parameter
Base operation
Conventional Oxidative hydrotreating desulfurization
Diesel sulfur, (wt ppm) Diesel yield, Iv-% Capital cost ($MM) Hydrogen cost ($MM/yr) Utilities cost ($MM/yr) Catalyst cost ($MM/yr) Total operating cost ($MM/yr) Delta product value ($MM/yr) Total operating cost ($bbl)
400 98.5
> benzoquinolines, carbazoles Oxygen compounds Oxygen compounds include carboxylic acids, phenols, naphthenic acids, furans and compounds containing both oxygen and sulphur or both oxygen and nitrogen. Some are formed during transportation and storage due to exposure to air. In biocrudes, much of the oxygen occurs in polyhydroxy compounds such as glycerol. Oxygen compounds are very reactive in hydroconversion units, where the oxygen is converted into water. A rough rule of thumb is that the oxygen content of a straight-run crude is similar to the nitrogen content. A major exception comes from biocrudes, which contain almost no sulphur and nitrogen but can include more than 5 wt% oxygen. Consequently, biocrudes can be (and are) processed in conventional petroleum refineries, but the amounts must be limited. Trace contaminants Trace contaminants poison hydroconversion catalysts. Nickel and vanadium are present as porphyrins. The key feature of porphyrins is a planar arrangement of four nitrogen atoms, which form coordination compounds with cationic metals. Porphyrins containing iron and magnesium are the active components of haemoglobin and chlorophyll, respectively. The presence of porphyrins is consistent with the biological origin of petroleum. Soluble iron is associated with naphthenic acids. Salt water dissolves in petroleum and brings with it sodium, calcium, and other alkali- and alkaline earth metals. Such metals can also be present as organometallics. Organoarsenic compounds are found in certain conventional crudes, but also in synthetic crudes from oil sands and oil shale. Arsenic and mercury are especially severe catalyst poisons. Not all contaminants are natural. Organosilicon compounds are added to improve flow in pipelines and control foaming in delayed coking units. Opening the wrong valve can send the wrong substance to the wrong place. Harmful substances include:
• • •
Phosphorous and/or boron are present in motor-oil additives and used motor oil. Lead and manganese are no longer added to gasoline to improve octane, but they still turn up now and then in old, idle storage tanks. PERC (perchloroethylene) is used to generate hydrogen chloride in catalytic reformers. PERC can induce severe corrosion.
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10.5.2 Hydroconversion versus thermal conversion Conversion processes that break carbon–carbon bonds can be lumped into three categories: 1 2 3
Thermal cracking – visbreaking, delayed coking and fluid coking, Catalytic cracking – fluid catalytic cracking (FCC), Catalytic cracking in the presence of hydrogen – hydrotreating, hydrocracking.
It is often said that the processes in Categories 1 and 2 reduce H/C by removing carbon. In fact, they produce a combination light and heavy material. In delayed coking, the H/C of the heavy product (coke) is substantially lower than the average H/C of the other products – liquids and gaseous products. FCC makes high-quality gasoline, but it also makes coke and dense, low-quality +205 °C ‘cycle oils.’ The coke accumulates on the circulating catalyst. Burning that coke provides the heat required for cracking, which in the absence of hydrogen is endothermic. Category 3 includes both hydrotreating and hydrocracking, which increase the H/C ratio by adding hydrogen. Both in FCC units and hydrocrackers, catalytic conversion requires catalysts with acidic sites, which generate carbenium ions. In thermal conversion, free radicals are the reactive intermediates. Table 10.3 illustrates the practical differences.
Table 10.3 Comparison of conversion processes Process
Type
Products
Hydrocracking
Catalytic cracking in the presence of hydrogen
• • • • •
Minimal production of C1 and C2 Major products: C3+ Significant product branching Minimal olefin formation Minimal coke formation
FCC
Catalytic cracking
• • • • • • •
Minimal production of C1 and C2 Significant production of C3= Major products: C3+ Significant product branching Significant aromatics production Moderate olefin production Significant coke formation
Coking
Thermal cracking
• Significant C1, C2, and C2= • Significant olefin production • Significant coke formation
C1 is methane, C2 is ethane, C2= is ethylene, C3= is propylene
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10.5.3 Hydroconversion reactions Hydroconversion reactions fall into the following categories:
• • • • • •
Saturation of aromatics (ASAT), saturation of olefins (OSAT) and saturation of diolefins. HDx processes – hydrodesulphurization (HDS), hydrodenitrogenation (HDN), hydrodeoxygenation (HDO) and hydrodemellation (HDM). Hydrocracking. Isomerization. Catalytic reforming. Polymerization, oligomerization and alkylation.
Table 10.4 gives a summary of hydroprocessing reactions, showing their relative amounts of hydrogen consumption and heat release. Hydrotreating reactions Saturation. The saturation of olefins is rapid and (essentially) irreversible. It releases considerable heat. To control this heat release, refiners load low-activity catalysts at the top of reactors. Figure 10.9 provides an example of the reversible saturation of aromatics. Figure 10.10 (Robinson, 2006) summarizes thermodynamic calculations for the saturation
Table 10.4 Summary of hydroprocessing reactions Reaction type
Illustration
∆HR*
Minimal C-C bond breaking Hydrodesulphurization (HDS) † Hydrodenitrogenation (HDN) Hydrodeoxygenation (HDO) Hydrodemetallization (HDM) Saturation of aromatics Saturation of olefins Isomerization
R-S-R + 2 H2 → RH + R2H + H2S R = N-R + 3 H2 → RH + RH + NH3 R-O-R + 2 H2 → RH + RH + H2O R-M + ½ H2 + A → RH + M-A C10H8 + 2 H2 → C10H12 R = R + H2 → HR-RH n-RH → i-RH
2.5–3.0 2.5–3.0 2.5–3.0 3 3 5.5 n/a
Significant C–C Bond breaking Dealkylation of aromatic rings Opening of naphthene rings Hydrocracking of paraffins
Φ-CH2-R + H2 → Φ-CH3 + RH Cyclo-C6H12 → C6H14 R-R + H2 → RH + RH
−1.3–1.7 −1.3–1.7 −1.3–1.7
Other reactions Coke formation Mercaptan formation
2 ΦH → ΦΦ + 2 H2 R = R + H2S → HS-R-RH
+3 3
Source: Hsu and Robinson, 2006 * Kilojoules per standard m3 of H2 consumed. For exothermic reactions, ∆HR is negative. † R = alkyl; Φ = aromatic; M = Fe, Ni or V; A = metals-adsorbing material
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10.9 An example of a reversible saturation/dehydrogenation reaction.
of naphthalene to tetralin and then to decalin. At high pressures and low temperatures, equilibrium favours saturation. At low pressures and high temperatures, equilibrium favours dehydrogenation. The reference also shows that high temperatures favour the condensation of polyaromatics into larger polyaromatics. Coincidentally, hydroconversion units operate in the ‘crossover’ region between 315 °C and 425 °C, where both saturation and dehydrogenation can occur. Removing ‘difficult’ sulphur and nitrogen from polyaromatic molecules requires prior saturation of one or more aromatic rings, which becomes more difficult at higher temperatures. HDS. Figures 10.11 and 10.12 show important HDS reactions. Figure 10.11 shows the mechanism for the HDS of 4,6-dimethyldibenzothiophene (4,6DMDBT), one of the least-reactive sulphur compounds in petroleum. The reaction proceeds via both routes, but the overall rate for the indirect route is faster. After deep desulphurization to make ULSD, the only remaining sulphur compounds in the diesel product are dimethyl- and trimethyldibenzothiophenes. HDN. Figures 10.13 and 10.14 show HDN reactions. Figure 10.14 illustrates the mechanism for the HDN of quinoline. As with sulphur removal from hindered DMDBTs, removing nitrogen from quinoline requires prior saturation of an aromatic ring. © Woodhead Publishing Limited, 2011
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10.10 Results of free-energy calculations on a seven-component system: system naphthalene-tetralin decalin-xylene-hydrogen-chrysene (hydrogen not shown).
10.11 Mechanism for hydrodesulphurization by the ‘direct’ mechanism.
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10.12 Two mechanisms for the hydrodesulphurization of 4,6-dimethyldibenzothiophene.
10.13 Hydrodenitrogenation reactions.
Hydrocracking reactions Hydrocracking reactions break carbon–carbon bonds. In both catalytic cracking (FCC) and hydrocracking, the rupture of C–C bonds is catalyzed by strong solid acids. The most widely-used hydrocracking catalysts contain synthetic zeolites, such as H-Y, or amorphous silica/alumina (ASA). Hydrocracking catalysts also contain metals, which saturate the fragments produced by cracking. The most
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10.14 Mechanism for the hydrodenitrogenation of quinoline.
commonly used noble metal is palladium. The most commonly used base metal catalysts contain either MoS2 promoted by NiS, or WS2 promoted by NiS. Mechanism. Figure 10.15 outlines the mechanism for the hydrocracking of n-heptane. Dehydrogenation over a metal at a metal site generates 2-heptane, which is converted into a carbenium ion (also known as a carbonium ion or carbocation) via proton addition at protic acid sites. Different carbenium ions are generated by hydride abstraction at Lewis acid sites. The carbenium ions can undergo a variety of reactions. They can isomerize into branchedchain intermediates via alkyl shift. They can ‘crack’ into smaller molecules via β-scission. Methane is not produced because the smallest possible olefin and the smallest possible carbenium ion are C2 entities. They can be hydrogenated into stable end-products; in the example, these are butanes, ethane and isohexane. They can even react with olefins to make larger molecules. (Pine, 1987). Nitrogen slip. Acid sites are poisoned by organic nitrogen. Therefore, in commercial hydrocrackers the feed is hydrotreated to reduce the nitrogen content. ‘N slip’ is the concentration of nitrogen in the feed to the hydrocracking catalyst. (Note that some refiners define N slip as the concentration of nitrogen in the pretreater effluent. The two are not necessarily the same.) A typical N slip target is 50 ppmw, but for some units the target is much lower. Noble metal catalysts are sensitive to sulphur, so for these there is also an ‘S slip’ target, typically 95 wt%. Every C–C bond is a potential reactant, so while a system might run out of sulphur, it ‘never’ runs out of C–C bonds. For a feed with an average chain length of C20, there are 18 C–C bonds per molecule. On this basis, the weight ratio of potential reactants (sulphur versus hydrocarbon) is not 50, it is 900. Of course, reaction rates are just as important, if not more so. A rule of thumb is that raising the temperature 10 °C doubles the reaction rate. The hotter it gets, the faster it goes. Consequently, controlling heat release is the most important aspect of designing and operating hydrocrackers. Excessively high temperatures cause runaway reactions, some of which have been deadly. Recombination. In hydrotreaters, light mercaptans are easy to desulphurize. But in hydrocrackers, conditions favour the production of light mercaptans by the ‘recombination’ of hydrogen sulphide with the olefin intermediates. If not
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removed, the mercaptans can poison downstream catalytic reforming catalysts. Refiners mitigate this problem with hydrotreating (‘post-treat’) catalysts. Catalytic reforming reactions The primary purpose of a catalytic reforming (CatRef) unit (CRU) is to produce C6–C10 aromatics and hydrogen from naphthenes and other saturated molecules in heavy naphtha. Aromatics have high octane numbers, and C7-plus reformate is an excellent blend stock for gasoline; due to the toxicity of benzene, regulations limit the amount of benzene in gasoline. Alternatively, the aromatics can be used to produce petrochemicals. CatRef catalysts contain highly dispersed platinum (Pt), platinum/rhenium (Pt/Re), platinum/rhenium/tin (Pt/Re/Sn) supported on high-surface-area gamma alumina. Hydrogen chloride is added to provide acidity, usually by injecting tetrachloroethylene or perchloroethylene. The catalytic activity of platinum is poisoned by sulphur, so feeds to CRUs must be very clean. The upper limit on sulphur is 1 ppmw. Hydrotreating is used to achieve this limit. Figure 10.17 shows some typical catalytic reforming reactions. Reaction 4 – the dehydrogenation of naphthenes to form aromatics – predominates. CRUs also make coke and C4-minus light ends. Heavy naphtha from a hydrocracker is rich in naphthenes, which makes it an excellent CatRef feed.
10.17 Catalytic reforming reactions.
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Hydroisomerization reactions One might argue that hydroisomerization is not a hydroconversion process because it does not consume hydrogen. We cover it here because, like hydrocracking, the mechanism involves alkyl shift in carbenium ions (see Fig. 10.15, Reaction 6). Hydroisomerization converts n-paraffins into isoparaffins. Isopentane and isohexanes have much higher octanes than n-pentane and n-hexane, so they are better for gasoline blending. Isobutane reacts with olefins in alkylation units to make high-octane C7 and C8 products. (The alkylation mechanism involves carbenium ions generated either by hydrofluoric or sulphuric acid.) The most common catalyst for isomerising n-butane is highly dispersed Pt on gamma alumina. As with catalytic reforming, chloride is used as a promoter. The catalyst is highly active. This allows operation at low temperatures, which favours isobutane. For isomerising n-pentane and n-hexane, the catalyst is either chloridepromoted Pt on alumina or Pt on zeolite. Zeolite catalysts require higher temperatures but are less susceptible to water. As with n-butane isomerization, operation at lower temperatures favours branched isomers.
10.6
Supported-metal hydroconversion catalyst
Most hydroconversion catalysts comprise active metals dispersed on a solid, high-surface-area support (Scherzer and Gruia, 1996; Magee and Dolbear, 1998; Ancheyta and Speight, 2007); exceptions include bulk sulphuric acid and hydrofluoric acid, which are used for commercial alkylation processes, and solid phosphoric acid (SPA), which is used for the oligomerization of propylene. Common supports include bohemite (γ-alumina), amorphous silica-alumina and synthetic zeolites. The active constituents of commercial hydrotreating catalysts are molybdenum sulphide (MoS2) promoted by sulphides of nickel, cobalt or both. For both NiMo and CoMo catalysts, the most common support is γ-alumina. Fresh catalysts are very strong and have high surface areas – typically 250–400 m2/g. Guard material protects main-bed hydrotreating catalysts by capturing particulates and removing trace contaminants. Hydrocracking catalysts are bifunctional. They contain both strong solid acids and metals. The active metals can be noble (Pd or Pt) or non-noble (nickel-promoted MoS2 or WS2). The acid function is provided either by amorphous silica alumina (ASA) or synthetic zeolites in a binder matrix. For zeolite catalysts, the zeolite content can be as low as 0.5 wt% or as high as 80 wt%. The gross dimensions of hydrocracking catalysts are similar to those for hydrotreating catalysts. In most hydrocrackers, the first few catalyst beds contain a high-activity hydrotreating catalyst, which removes nitrogen and sulphur. It also drives the saturation of olefins and aromatics. Catalytic reforming catalysts comprise combinations of noble (and other) metals on γ-alumina supports. The metal combinations include Pt/Re and Pt/Re/ Sn. Acidity is provided by chlorination.
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10.6.1 Catalyst manufacturing The following steps are used to prepare the supported metal catalysts for hydroconversion processes:
• • • • • •
precipitation filtration (or centrifugation), washing and drying forming calcining impregnation and co-mulling activation.
Other steps, such as kneading, mulling, grinding and/or sieving also may be used. Precipitation In the precipitation step, two solutions are combined to form a desired solid. Mixing a solution of aluminum nitrate (Al(NO3)3) with a solution of sodium aluminate (Na2Al2O4) yields a gel of aluminum hydroxide (Al(OH)3). As the gel ages, tiny crystals grow larger and a pore structure starts to develop. The zeolites used in hydrocracking catalysts also are prepared by precipitation. Zeolites occur naturally, but the ones used for catalysis are synthetic. Figure 10.18 outlines a common procedure for synthesizing Na-Y and H-Y zeolites. Synthetic zeolites are used as drying agents, ion-exchangers and molecular sieves for gas separation. Their microporosity provides them with high surface area, and they can be converted into solid acids with superb catalytic activity. The Al(III) atoms in zeolites replace Si(IV) atoms in a SiO2 superstructure. To maintain a neutral charge, every aluminum atom must be accompanied by a counter-ion such as Na+, K+, H+, NH4+, etc. Counter ions can be swapped via ion exchange. When Na-Y zeolite is exchanged with an ammonium salt, the Na+ ion is replaced by NH4+. When NH4-Y is heated to the right temperature, the ammonium ion decomposes, releasing NH3(gas) and leaving behind highly acidic H-Y zeolite. The synthetic zeolites used in catalysts for hydrocracking include X, Y, and ZSM-5. ZSM-5 is a shape-selective zeolite made by including a soluble organic template, such as a quaternary ammonium salt, in the mix of raw materials. ZSM-5 is used for catalytic dewaxing. Due to its unique pore structure, it selectively cracks waxy n-paraffins into lighter molecules. Filtration, washing and drying Filtration, washing and drying remove undesired impurities. For example, after the precipitation of Al(OH)3 from aluminum nitrate and sodium aluminate, the
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10.18 A procedure for synthesizing Na-Y and H-Y zeolites.
co-produced sodium nitrate is washed away with water. Adding ammonium hydroxide expedites sodium removal. Subsequent drying removes excess water and initiates the transformation of Al(OH)3 into alumina (Al2O3). Forming Catalysts can be produced as cylinders, shaped extrudates, spheres or pellets. A shaped-extrudate cross-section can look like a three- or four-leaf clover without the stem (Fig. 10.19). Compared to cylindrical extrudates, shaped extrudates have a higher surface-to-volume ratio, and the distance from the outside of a particle to the center is shorter. To make extrudates, a paste is formed and forced through a die. The resulting spaghetti-like strands are dried and broken into short pieces with a length/diameter ratio of 2:4. The particles are calcined, which hardens them and removes additional water and volatile molecules such as ammonia. The particles have diameters ranging from 1.3–4.8 mm. Spherical catalysts are made (1) by spray-drying slurries of catalyst precursors, (2) by spraying liquid onto powders in a tilted rotating pan or (3) by dripping
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10.19 Catalyst shape examples.
a silica-alumina slurry into hot oil. Pellets are made by compressing powders in a dye. Impregnation and co-mulling Impregnation distributes active metals within the pores of a catalyst support. Like sponges, calcined supports are especially porous. When they are exposed to aqueous solutions containing active metals, capillary action pulls the aqueous phase into the pores. After drying, additional solution may be added to increase loading of the same or a different active metal. Catalysts can also be made by co-mulling the active metal oxides with the support. Co-mulling tends to cost less because it requires fewer steps. It also produces materials with different activities – sometimes higher, sometimes lower – than impregnation. Activation Prior to use, supported hydroconversion catalysts must be activated. After the steps described above, the active metals are present as oxides. Catalysts containing noble-metal oxides such as platinum or palladium are activated by direct reduction with high-pressure hydrogen at 350 °C. Catalytic reforming catalysts must be chlorinated. Non-noble-metal hydrotreating and hydrocracking catalysts are activated by conversion of the oxides to sulphides, again under flowing hydrogen. The sulphur can come from H2S, from the organic sulphur compounds in start-up oil, or from the decomposition of a sulphiding agent. The most common sulphiding agent is dimethyl disulfide (CH3-S-S-CH3). Some suppliers offer pre-sulphided or pre-sulphurized catalysts.
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10.6.2 Catalyst cycle life In fixed-bed units, catalyst life has a major impact on the economics. Cycles can be as short as 12 months and (in rare cases) as long as 60 months. Two-year cycles are typical. At the start of a cycle, average reactor temperatures are low, ranging from 288–349 °C. As the cycle proceeds, the catalyst deactivates, so refiners raise temperatures to maintain HDS, HDN or conversion. A catalyst cycle ends for one of the following reasons:
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•
The temperature required to achieve the unit’s main process objectives hits a metallurgical limit. For a typical hot-wall reactor operating at 100–170 barg, the maximum peak temperature is 441 °C. To ensure safe operation, the maximum average reactor temperature for hydrotreating is about 427 °C. For hydrocracking, the maximum average temperature is lower. Depending on the activity of the catalyst, it can be 404–415 °C. The main process objective can be met only at reduced feed rate. Pressure drop across the unit reaches a limit. The overall pressure drop is the difference in pressure between the recycle compressor suction and the recycle compressor discharge. At start-of-run, the pressure drop across the catalyst is low (0.2–0.7 bar) for each bed – but it increases as the run proceeds. Usually, the increase is largest in the first catalyst bed, which is most susceptible to fouling. Attempts to continue running a unit despite very high pressure drop can deform the quench-deck support beams inside a reactor, or collapse the ‘elephant stool’ at the bottom. A typical upper limit across the high-pressure loop is 13.6 bar. Selectivity decreases. If the production of light gases exceeds the capacity of one or more towers in the downstream gas plant, operators must decrease feed rate or reduce conversion. Both options are expensive. Excess gas production can render a unit inoperable by overwhelming downstream equipment, such as strippers, fractionators or offgas compressors. It also consumes expensive hydrogen and converts it into low-value RFG and LPG. High temperatures also decrease the amount of middle distillates and increase their aromatics content. A related unit shuts down. Related units might include an upstream FCC or vacuum distillation unit, an upstream hydrogen source or a downstream catalytic reformer. Feed upsets. For a fixed-bed hydrotreater, a slug of residue can poison part of the catalyst with trace contaminants or foul it with particulates, asphaltenes and/or refractory carbon. In fixed-bed units, poisoning and fouling usually are confined to the top few feet of the first catalyst bed. If so, the ruined catalyst can be skimmed and replaced during a brief, scheduled shutdown that does not require a cycle-ending catalyst change-out. Equipment failure. Hardware problems occur most frequently in rotating equipment – pumps and compressors. Such upsets can trigger temperature excursions requiring EDP.
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10.7
Commercial hydroconversion units
10.7.1 Fixed-bed hydrotreating and hydrocracking For both hydrotreating and hydrocracking, the required hydrogen can come from catalytic reformers, steam/methane reformers, olefin plants or some combination of the three. Liquid feeds come from atmospheric and/or vacuum distillation units, from thermal cracking units (delayed cokers, fluid cokers, visbreakers) or FCC units. Feeds from primary fractionation are called ‘straightrun.’ Those that come from FCC or thermal cracking units are called ‘cracked stocks.’ Middle distillates from a hydrocracker usually meet or exceed finished product specifications, but the heavy naphtha from a hydrocracker usually goes to a catalytic reformer for octane improvement or production of petrochemical aromatics. The fractionator bottoms can be recycled or sent to an FCC unit, an olefins plant or a lube base-stock plant. According to a tally prepared from the survey of refining process units published by the Oil & Gas Journal (Koottungal, 2008), there are roughly 1350 fixed-bed hydrotreaters and 150 fixed-bed hydrocrackers in the world, plus 17 ebullated bed hydrocrackers. The word ‘roughly’ is used because some so-called ‘mild’ hydrocrackers use nothing but hydrotreating catalysts, and others have been converted into ULSD hydrotreaters. Including demonstration plants, at least three slurry-phase units are processing heavy oil. Prior to 1950, 12 VCC slurry-phase units were producing 100 000 barrels per day of hydrocarbon liquids from coal (Motaghi et al., 2010). Fixed-bed reactors can be modelled with trickle-bed kinetics, as described in a classic article by Satterfield (1975). In trickle-bed reactors, a gas phase and a liquid phase flow downward concurrently over a packed bed of catalyst particles. Commercial trickle-bed reactors operate adiabatically. Fixed-bed process flow schemes Figure 10.20 presents a process flow scheme for a typical single-reactor fixed-bed hydrotreater. Reaction conditions (Table 10.5) depend on feed quality and process objectives. Oil and hydrogen-rich gas flow down through reactors loaded with catalysts. Make-up gas comes in to replace consumed hydrogen. Gas flow can be once-through in naphtha hydrotreaters, but in distillate and VGO hydrotreaters, unconsumed hydrogen is recycled. Hydrotreating is exothermic, and at constant pressure trickle-bed reactors are adiabatic, so the release of heat raises temperatures. To divide the heat release into smaller, safer portions, commercial reactors have multiple catalyst beds with cooling (‘quench’) in between. In a quench section (see below) hot process fluids from the preceding bed are combined with relatively cold hydrogen-rich quench gas before the mixture passes into the next bed.
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10.20 Process flow sketch for a single-reactor fixed-bed hydrotreater. Table 10.5 Typical operating conditions for fixed-bed hydrotreaters and hydrocrackers Process, feedstock types
Approx. H2 Partial Pressure Approx. Maximum psig kPa Conversion (wt%)
Hydrotreating Naphtha LGO (kerosene) HGO (Diesel), LCO VGO, VBGO, DAO, CGO, HCO Residual oil Mild hydrocracking VGO, VBGO, DAO, CGO, LCO, HCO Once-through hydrocracking VGO, VBGO, DAO, CGO, LCO, HCO Residual oil Recycle hydrocracking VGO, VBGO, DAO, CGO, LCO, HCO Ebullated-bed hydrocracking VGO, VBGO, DAO, HCO Vacuum residual oil
250–450 250–600 600–800 800–2000 2000–3000
1800–3200 1800–4250 4250–5600 5600–14 000 14 000–21 000
5 5 5 10 10
800–1200
5600–8400
40
1500–2000 2000–3000
10 000–14 000 14 000–21 000
90 40
1500–2000
10 000–14 000
99
2000 2000–3000
14 000 14 000–21 000
99 70
In all cases, operating temperatures range from 315°C to 440°C (600 to 825°F) LGO = light gas oil, HGO = heavy gas oil, LCO = FCC light cycle oil, HCO = FCC heavy cycle oil, VGO = vacuum gas oil, VBGO = visbreaker gas oil, DAO = deasphalted oil, CGO = coker gas oil
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Hydrotreating produces both H2S and NH3. Under reaction conditions, these remain in the gas phase. But at lower temperatures, they combine to form solid ammonium bisulfide (NH4SH). Ammonia also reacts with chlorides to form NH4Cl; chloride can come with make-up gas, feed, or wash water. These salts deposit in air coolers and heat exchangers, blocking flow and – even worse – inducing corrosion. Fortunately, they are water soluble, so they can be controlled by injecting wash water into the reactor effluent. The reactor effluent is cooled by heat exchange with cold feed then cooled further with air or water (or both). It then flows into a series of flash drums. The cold high-pressure separator (CHPS) separates recycle gas from the other process fluids. Water containing ammonium salts comes out the bottom. The remaining material goes to a low-pressure separator (LPS). In the LPS, the condensation of liquid products separates them from methane, ethane and any remaining hydrogen. The C3-plus liquids flow to a stream stripper or fractionator. Figure 10.21 shows a two-reactor once-through hydrocracker, along with some typical properties of feed and product streams. Figure 10.22 shows a two-stage hydrocracker, in which unconverted oil from the first two reactors goes to a third reactor for additional conversion. With respect to equipment and process flow, fixed-bed hydrocrackers are similar to fixed-bed hydrotreaters. Hydrotreating is an integral part of hydrocracking, because hydrotreating is used to remove nitrogen from the feed.
Fixed-bed hydrocracking yields and product selectivity As shown in Table 10.6, a fixed-bed hydrocracking unit can have significant product flexibility, producing either large amounts of C4-plus naphtha or large amounts of middle distillates. In petroleum refining, this kind of flexibility is unique.
Fixed-bed hydroconversion equipment As shown in Fig. 10.21 and 10.22, for fixed-bed hydrocrackers, process flows can differ substantially. All designs include pumps, heaters, compressors, catalysts, flash drums (separators), product recovery equipment, emergency depressuring systems and of course reactors. The most common differences are mentioned below.
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Heaters and heat exchange. Many heaters are combined phase, which means that hydrogen and oil are mixed prior to heating. Others heat only the recycle gas or only the oil. In all units, much of the heat generated by the process is recovered by heat exchange against incoming oil. Compressors. Most hydrotreaters and hydrocrackers have one recycle gas compressor (RGC), which can be either centrifugal or reciprocating, and at
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10.21 Two-reactor once-through hydrocracker with some typical feed and product streams.
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10.22 Two-stage hydrocracker with recycling of unconverted oil to a third reactor.
• •
least one reciprocating make-up gas compressor. Make-up gas can enter the unit either before or after the RGC. Two-stage hydrocrackers, especially those with a sweet second stage, may have a separate sweet-gas RGC for the second stage. In this instance, ‘sweet’ means that the H2S content of the recycled gas is less than 20–50 ppmv. Amine treating. In many hydrotreaters and some hydrocrackers, H2S is scrubbed from the recycle gas with a high-pressure amine treater. H2S inhibits HDS reactions, so its removal enhances desulphurization. Amine treaters also increases recycle gas purity. Flash drums. After primary heat exchange against cold feed, the reactor effluent goes through a series of flash drums. Older hydrocracking units have just a cold high-pressure separator (CHPS), which separates gases from liquids, and a cold low-pressure separator (CLPS), which removes additional light ends ahead of the fractionator. The reactor effluent (CHPS feed) is cooled either with air, water or both. Many newer units with heavy feed have four separators – HHPS, CHPS, HLPS, and a CLPS. The hot high-pressure separator (HHPS) feeds a CHPS and a hot low-pressure separator (HLPS); the
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Table 10.6 Flexibility of catalytic hydrocracking Feed
Straight-run vacuum gas oil
Boiling range (°C) API gravity Specific gravity Nitrogen (ppmw) Sulphur (wt%)
340–550 22.0 0.9218 950 2.5
Primary product objective Weighted average reactor temp (°C) Product yields (vol% fresh feed) Butanes Light naphtha Heavy naphtha Kerosene or gas oil Total C4–plus Chemical H2 consumption (Nm3/m3) (Scf/bbl) Product qualities Light naphtha (C5–82 °C) RON clear Heavy naphtha P/N/A RON clear End point (°C) Kerosene Flash point (°C) Freeze point (°C) Smoke point, mm FIA aromatics (vol%) End point (°C) Gas oil Cloud point (°C) API gravity Cetane number Flash point (°C) End point (°C)
•
Naphtha base
Kerosene −6
Gas Oil −12
126
8 18 29 69 124
7 16 21 77 121
345 2050
315 1870
292 1730
79
79
80
45/50/5 41 216
44/52/4 63 121
67 118
11 25 90
38 −48 34 7 282 −15 44 55 52 349
CHPS bottom goes to a CLPS. Both of the low-pressure bottom streams go to the fractionator. Product recovery. For a distillate hydrotreater, it is common to use a steam stripper to separate H2S and light gases from the liquid product. Most hydrocrackers have full-blown fractionation sections comprising a main tower with at least one side draw. Many include a pre-flash between the LP separator and main tower. Some have a naphtha splitter to separate light and heavy naphtha, and some use vacuum distillation to lift incremental gas oil out of
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the fractionator bottoms. In most once-through hydrocrackers, the bottom product stream goes to another unit – FCC, lube plant or olefin plant. If the bottom stream is light enough, it can be blended directly into diesel. In highconversion units, the fractionator bottom stream is recycled, either to the pretreat section, to the first-stage cracking reactor or to a separate second-stage reactor. Emergency depressuring. In 1997, the American Petroleum Institute published the 4th edition of API 521, Guide for Pressure Relieving and Emergency Depressuring. The most recent version is equivalent to ISO 23251. Even so, one can still find differences in emergency depressuring (EDP) systems. For more information, please review the Guide. Key concepts include the following facts: ○ Just like combustion requires heat, fuel and oxygen, hydrocracking requires heat, oil and hydrogen. Removing any one of the three will stop the reaction. ○ The flow of process fluids is the only thing that removes heat. Stagnation is dangerous. ○ Depressuring is the only way to quickly remove hydrogen. The recommended practice is to have two systems, one which depressures the unit at an initial rate of 6.8 bar/min, and another which depressures the unit at an initial rate of 13.6 bar/min. Some licensors design for initial rates of 10.3 and 20.7 bar, respectively. Depressuring is doubly effective, because it also removes oil and heat. ○ Other actions – stopping the heater, stopping the flow of oil or by-passing heat exchangers – can be effective, but they take a relatively long time. Reactors. Some hydrocrackers have multiple one-bed reactors, but most reactors have multiple beds with quench zones in between. Quench zone designs are proprietary. Figure 10.23 shows generic components, but it doesn’t necessarily show the ‘latest’ or ‘best.’ Additional information is available from process licensors. In a quench section, hot process fluids from the preceding bed are combined with relatively cold hydrogen-rich quench gas before the mixture passes into the next bed. In addition to lowering the temperature of the reacting fluids, quench sections improve flow distribution with radial mixing and redistribution of the reactants to the next bed. The major constituents of a quench deck are the gas nozzle, liquid collector, gas/liquid mixer and re-distributor. The gas nozzle brings in cool, hydrogenrich quench gas. Old ones were very simple – just a tube with one or more holes in it. In newer units, quench gas comes in through the side of the reactor. In some old ones, it came down through the top. New nozzles deliver gas at multiple points through spokes or rings. The liquid collector provides radial (horizontal) liquid–liquid mixing. In the gas/liquid mixer, a valve tray or bubble-cap tray provides intimate contact between gases and liquids. The re-distributor sends an even spray of fluids down to the catalyst bed below. In addition to cooling, quench zones improve radial flow distribution. We can think of a catalyst bed as a stack of thin, horizontal discs. Ideally, the top disc is
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10.23 Quench zone for a fixed bed hydroprocessing unit.
the coolest, the bottom disc is the hottest. But in real units, the downward flow of reactants is never perfectly uniform. To one extent or another, temperatures in a given horizontal plane are different. The differences are larger near the bottom. The flow of reaction fluids moves heat through the unit, so uneven flow leads to uneven temperatures. The difference between the highest and lowest temperature at the bottom of a catalyst bed is called the ‘radial temperature difference’ (RTD). One can never know the actual highest and lowest temperature, because thermocouples cannot be placed everywhere. But if the measured RTD is small – less than ∆3 °C – one can assume that the actual RTD also is small, and that flow through the bed is nearly uniform. But if the measured RTD is large, the actual RTD is almost certainly larger, and should raise concerns about hot spots and other potentially dangerous symptoms of mal-distribution.
10.7.2 Ebullated bed hydrocracking In contrast to fixed-bed hydrocrackers, ebullated bed (e-bed) units can process large amounts of residual oils. Catalyst life does not limit these units, because fresh catalyst is added and spent catalyst is removed continuously. In e-bed units
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(Fig. 10.24), hydrogen-rich recycle gas is bubbled up through a mixture of oil and catalyst particles. This provides three-phase turbulent mixing, which is needed to ensure a uniform temperature distribution. At the top of the reactor, catalyst is disengaged from the process fluids, which are separated in downstream flash drums. Most of the catalyst is returned to the reactor. Some is withdrawn and replaced with fresh catalyst. According to Ancheyta and Speight (2007) and the references cited therein, when compared to fixed-bed processes e-bed technology offers the following advantages:
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The ability to achieve more than 50 wt% conversion of atmospheric residue due to: ○ improved heat distribution and heat transfer. E-bed reactors are nearly isothermal, which enables higher operating temperatures and hence higher conversion ○ use of catalysts with smaller diameters. This reduces diffusion effects, allowing faster specific reaction rates.
10.24 Reactor internals for an ebullated bed hydrocracker.
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Ample free space between catalyst particles, which allows entrained solids to pass through the reactor without accumulation, plugging or build-up of pressure drop. Better liquid-product quality than delayed coking.
Disadvantages versus fixed-bed processes include high catalyst attrition, which leads to high rates of catalyst consumption; higher installation costs due to larger reactor volume and higher operating temperatures; and sediment formation. Recent improvements include second-generation catalysts with lower attrition; catalyst rejuvenation, which allows the reuse of spent catalysts; improved reactor design leading to higher single-train feed rates; and two-reactor layouts with inter-stage separation.
10.7.3 Slurry-phase hydrocracking Slurry-phase hydrocracking converts vacuum residue – or even coal tar – in the presence of hydrogen under severe process conditions – more than 450 °C and 14 000–20 800 kPa. The catalysts are non-acidic, comprising divided carbon or iron sulphate. The cracking is thermal, not catalytic, so the reaction intermediates are free radicals, not carbenium ions. The additives prevent excessive coking. Inside the reactor, the liquid/powder mixture behaves as a single phase due to the small size of the additive particles. The advantages of slurry-phase processes include:
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The ability to achieve more than 95 wt% conversion of vacuum residue and other feedstocks, including coal (Motaghi et al., 2010). Good product quality in two-stage designs, which incorporate a fixed-bed hydrotreating reactor – see Table 10.7 (Motaghi et al., 2010). Feedstock flexibility. Low-cost, micron-size disposable catalysts. Good external heat transfer in a nearly isothermal reaction regime. For a given volume of residue feed, lower reactor volume than e-bed processes.
Disadvantages include the co-production of high-metals and high-sulphur refractory pitch.
10.7.4 Catalytic reforming The three major process flows for catalytic reforming are:
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semi-regenerative cyclic continuous catalyst regeneration (CCR).
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Table 10.7 Two-stage VCC process yields and product selectivity Property
Feed
Flow rate (barrels/day) Product selectivity (wt% SCO) (vol% SCO) Specific gravity API gravity Sulphur (wt%) Nitrogen (wt%) Carbon residue (wt%) Nickel (ppmw) Vanadium (ppmw)
92 309 100 000
1.012 8.33 5.04 0.45 14.48 81 207
Synthetic LPG crude oil
100 100 0.8451 35.9 0.004 0.033 0.016