Membrane Engineering for the Treatment of Gases: Volume 2: Gas-separation Problems Combined with Membrane Reactors [1 ed.] 1849732396, 9781849732390

Citation preview

Membrane Engineering for the Treatment of Gases

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP001

Volume 2: Gas-separation Problems Combined with Membrane Reactors

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP001

View Online

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP001

Membrane Engineering for the Treatment of Gases Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli Department of Chemical Engineering and Materials, The University of Calabria, Italy

and Giuseppe Barbieri National Research Council - Institute for Membrane Technology, Italy

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP001

View Online

ISBN: 978-1-84973-239-0 A catalogue record for this book is available from the British Library r Royal Society of Chemistry 2011 All rights reserved Apart from fair dealing for the purposes of research for non-commercial purposes or for private study, criticism or review, as permitted under the Copyright, Designs and Patents Act 1988 and the Copyright and Related Rights Regulations 2003, this publication may not be reproduced, stored or transmitted, in any form or by any means, without the prior permission in writing of The Royal Society of Chemistry, or in the case of reproduction in accordance with the terms of licences issued by the Copyright Licensing Agency in the UK, or in accordance with the terms of the licences issued by the appropriate Reproduction Rights Organization outside the UK. Enquiries concerning reproduction outside the terms stated here should be sent to The Royal Society of Chemistry at the address printed on this page. The RSC is not responsible for individual opinions expressed in this work. Published by The Royal Society of Chemistry, Thomas Graham House, Science Park, Milton Road, Cambridge CB4 0WF, UK Registered Charity Number 207890 For further information see our web site at www.rsc.org

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP005

Preface Separations of downstreams and treatment of upstreams in any type of process require most of the energy necessary in the production cycle. The concentration and/or purification of the downstreams of a production unit on the basis also of the market specifications, the ratio adjustment of reactor feed streams to fit the best operating condition, the conditioning or pretreatment of the feed streams for contaminants removal, etc. are some examples of typical separations that occur in the process industry every day. Membrane engineering has been growing significantly in the last few years and membrane operations are the dominant technology in various areas today, e.g. in seawater desalination, in waste-water treatment and reuse, in artificial organs, in food juice treatment, etc. The intrinsic properties of membranes, such as molecular separations, the possibility of coupling reaction and separation in the same unit, etc. help to confirm membrane engineering as a powerful tool for realizing the process intensification strategy, which is the best answer today to sustainable industrial growth. For instance, reverse osmosis was demonstrated as requiring an energy load about 10 times lower than that of a thermal process. Therefore, in such fields of application, membrane separations are recognized today, among the different technologies, as the ‘best available technology’. The use of membranes in the separation of gases is also a fast-growing field and in various cases membrane technology competes with traditional operations. The separation of air components or oxygen enrichment by means of membranes has been growing substantially during the past 10 years. The oxygen-enriched air produced has been used in various fields, including chemical and related industries, medical fields, food packaging, etc. In industrial furnaces and burners, for example, the injection of oxygenenriched air (25–35% oxygen) leads to higher flame temperatures and reduces the volume of ‘parasite’ nitrogen to be heated. Mixtures containing more than 40% of O2 or 95% by volume of N2 from the air can be obtained together with Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

v

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP005

vi

Preface

membrane systems which dominate the fraction of the nitrogen market today for applications of less than 50 tonnes day 1. Membrane gas separation, however, is far from covering its real potential and further progress will be a challenge in the coming years in many applications, such as CO2 capture and utilization, H2 separation and purification, dehydration of gaseous streams, etc. Currently, problems related to the pre-treatment of the streams, membrane life-time and their selectivity and permeability, still slow down the growth of large scale industrial applications. Membrane engineering will pursue the design and development of new polymeric, inorganic and hybrid materials with tailored and improved mass transport properties and which are able to withstand more aggressive environments and a wider range of operating conditions. The development of new knowledge for the better utilization of these unit operations in integrated membrane systems, combing various membrane operations in the industrial process will also be part of the work for a sustainable industrial growth. The goal of this book is to present the main aspects and challenges related to membrane engineering for the treatment of gases starting from the fundamentals to the industrial application, focusing on polymeric, metallic and other inorganic membranes such as zeolites, perovskites and also carbon membranes. It is intended to provide a wide and critical state of the art on membrane technology for the treatment of gases, with emphasis for the application of membrane engineering in various fields of gas separation. Analysis by molecular design and of aging phenomena show how to improve our understanding of the fundamentals of mass transport of gas molecules through thin and ultra-thin glassy polymer membranes. A state of the art of macro-scale simulation studies is provided, particularly for CO2 post-combustion capture by means of membrane gas separation to identify the most relevant and efficient processes which fit the separation targets, including also an analysis of the target materials and cost performances. A very critical state of the art on current available polymers and recent progress on high performance polymers for next-generation gas separation applications is also presented. Membranes have to be assembled in modules before their utilization; their design depends also on the membrane configuration, such as flat sheet or hollow fiber. All the aspects concerning different module types and packing density, manufacturing costs and range of application are also analyzed and presented. An overview is provided about the applications of membranes in natural and biogas treatment, petroleum refining and petrochemicals production, presenting the technical challenges and the market for which the membranes are intended, giving a detailed presentation of existing commercial membrane technologies and considering the future trends for research in each addressed application. The potential of simulations in membrane engineering are shown in a techno-economic analysis of multi-stage membrane processes. This provides an insight into economics and related energy consumption and recovery avoidance costs with reference to the application of membranes CO2 capture from coal-fired power plants. The commercial applications of membranes in gas separations are also widely described and the critical needs in the

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP005

Preface

vii

development of this technology and some of the factors which are impeding the growth of gas separation membranes are addressed together with a perspective on further commercial growth. The use of hybrid process solutions involving synergistic membranes and fully established non-permeating technologies, such as adsorption and absorption, for improving the separation performance with marginal costs is described in depth, highlighting the cooperative way of operating. The main trends in Pd-based membrane development and criteria for their scale-up are presented in Chapters 10–14, specifically for very thin composite membranes. The basic features of membrane reactors are widely discussed for this application of this innovative technology, also considering membrane costs analysis on different membrane reactor architectures, by means of case studies. The packed bed and fluidized bed configurations most often used for membrane reactors for hydrogen production and purification studies are proposed, discussing the performances of the membrane reactor by means of modeling and simulations. An interesting strategy for the redesigning of more compact and efficient processes for pure hydrogen production than conventional ones is proposed, considering also the effect of the permeation reduction owing to concentration polarization and inhibition. This reduction is quantified by means of the concentration polarization and inhibition coefficients already included in the Sievert’s law equation that can be used, in its integrated form, to evaluate the hydrogen permeating flux. Relevant applications of carbon molecular sieve membranes are discussed for selected industrial separations such as CO2/CH4 in biogas upgrading, H2/CH4 wherever relevant, CO2 capture from flue gases, air separation, petrochemical and high-temperature applications. The advantages and drawbacks offered by perovskite membranes are analyzed for high temperature oxygen separation, also by introducing engineering and scale-up issues as well as an economic evaluation. An overview on the current and potential applications of zeolite membranes in the treatment of gases is provided showing the possibility of zeolite use for separating light gases such as CO2/N2 and CO2/CH4 as well as in the deep purification of H2 rich streams when these membranes are catalytically active. The application of mixed ionic–electronic conduction for oxygen separation is discussed, highlighting all the advantages over other methods, such as the large membrane area per unit packing volume, the reduced resistance to oxygen permeation, and the easy assembly into membrane modules offered by the hollow fiber configuration as well as the new strategies for improving the membrane properties, thereby making them ideal also for practical application in oxygen production. The last chapter gives a more comprehensive approach and discusses the role of membrane gas separation and membrane engineering in the re-designing of industrial applications in terms of new, recently introduced metrics. It provides an analysis of some processes for hydrogen production/separation that can be easily extended in other separation processes. This is a useful tool for the evaluation of pros and cons during the design phase of a new plant, where the membrane operations would replace traditional ones to pursue the strategy of process intensification.

View Online

viii

Preface

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP005

We wish to thank Dr Adele Brunetti for her collaboration in the preparation of this book, and for giving us the benefit of her knowledge in the field of gas separation and membrane reactors; she has been very useful for coordinating our activities during the various aspects of the final editing. Enrico Drioli and Giuseppe Barbieri The University of Calabria and National Research Council Institute on Membrane Technology, Italy

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

Contents Volume 1 Chapter 1

Multi-scale Molecular Modeling Approaches for Designing/ Selecting Polymers used for Developing Novel Membranes Elena Tocci and Pluton Pullumbi 1.1 1.2

Introduction Computational Methods 1.2.1 Atomistic Simulation Methods 1.3 Numerical Simulation of Polymer Membranes 1.3.1 Force Field and Choice of Ensemble 1.3.2 Generation of Amorphous Cell Packings 1.3.3 Realistic Amorphous Cell Selection 1.3.4 Estimation of Gas Transport Properties through Amorphous Cells 1.4 Concluding Remarks Acknowledgements References Chapter 2

Simulation of Polymeric Membrane Systems for CO2 Capture Eric Favre 2.1

2.2

Introduction 2.1.1 Global Warming and Carbon Capture 2.1.2 Membrane Processes and Carbon Capture Membrane Module Simulation Framework 2.2.1 Identifying Capture Step Boundary Conditions 2.2.2 Membrane Module Design: a Simplified Framework

Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

ix

1

1 5 8 16 16 17 18 19 20 23 23 29

29 29 31 32 32 34

View Online

x

Contents

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

2.2.3

Membrane Module Design: Classical Methodology 2.3 Simulation Studies for Post-combustion CO2 Capture by a Membrane Gas Separation Module 2.3.1 Addressing the Separation Problem: Selectivity Challenge 2.3.2 Tackling the Energy Requirement Issue 2.3.3 The Energy Requirement/Membrane Area Trade-off 2.3.4 Towards Multi-stage Processes 2.4 Scientific and Technological Challenges 2.4.1 Improved Materials: Selectivity and Productivity 2.4.2 Beyond Model Mixtures 2.4.3 Alternative Approaches and Prospects 2.5 Concluding Remarks 2.6 List of Symbols Acknowledgements References Chapter 3

40 42 44 45 47 48 48 51 52 53 54 54 54

Physical Aging of Membranes for Gas Separations B.W. Rowe, B.D. Freeman and D.R. Paul

58

3.1 3.2 3.3

58 60

Introduction Aging Behavior in Thin and Ultra-thin Films Additional Experimental Methods used to Study Physical Aging 3.4 Influence of Previous History and Experimental Conditions on Aging 3.5 Modeling Physical Aging Behavior 3.6 Concluding Remarks References Chapter 4

38

67 72 75 78 80

Recent High Performance Polymer Membranes for CO2 Separation S.H. Han and Y.M. Lee

84

4.1 4.2

84 86

Introduction Background 4.2.1 Solution-diffusion Mechanism for Gas Permeation 4.2.2 Trade-off Relationship in Gas Separation 4.2.3 High Performance Polymer Membranes for Gas Separation

86 88 89

View Online

xi

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

Contents

Chapter 5

4.3

Sorption-enhanced Polymer Membranes 4.3.1 Poly(Ethylene Oxide) Membranes 4.3.2 PEO-based Block Copolymer Membranes 4.3.3 Dendrimer Membranes 4.4 Diffusion-enhanced Membranes 4.4.1 Substituted Polyacetylene-based Membranes 4.4.2 Amorphous Fluoropolymer Membranes 4.4.3 Polymers with Intrinsic Microporosity 4.4.4 Thermally Rearranged Polymer Membranes 4.5 Concluding Remarks References

92 93 97 103 105 105 107 110 113 119 119

Design of Membrane Modules for Gas Separations M. Scholz, M. Wessling and J. Balster

125

5.1 5.2

125 127 127 129 131

5.3

5.4

5.5

Introduction Membrane Modules 5.2.1 Plate-and-frame Modules 5.2.2 Spiral-wound Modules 5.2.3 Hollow Fiber Modules 5.2.4 Comparison of the Different Module Configurations Operation of Gas Separation Hollow Fiber Membrane Modules 5.3.1 Flow within the Fiber (Lumen-side Feed, Shell-side Feed) 5.3.2 Operational Modes 5.3.3 Flow Patterns Mathematical Description of the Performance of a Gas Separation Module 5.4.1 Characteristic Numbers 5.4.2 Description of Concentration, Pressure and Temperature Profiles 5.4.3 Energy Balance 5.4.4 Pressure Losses Non-ideal Construction of Membrane Modules and the Influence of Non-idealities of Defect-free Dense Hollow Fiber Membranes 5.5.1 Influence of Fiber Diameter Variation 5.5.2 Influence of Variation in Membrane Thickness 5.5.3 Influence of Variation in Fiber Length 5.5.4 Influence of Membrane Defects 5.5.5 Influence of Blocked Fibers

132 133 133 135 136 138 138 140 141 142

142 143 143 144 144 146

View Online

xii

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

Chapter 6

Contents

5.6 Concluding Remarks 5.7 List of Symbols References

147 148 149

Gas/Vapor Permeation Applications in the Hydrocarbonprocessing Industry Arnaud Baudot

150

6.1

Natural and Biogas Membrane Processing 6.1.1 Membrane Suppliers 6.1.2 Membrane-based Acid Gas Removal 6.1.3 Removal of Hydrogen Sulfide 6.1.4 Other Membrane-based Natural Gas Treatments 6.2 Petroleum Refining 6.2.1 Hydrogen Purification 6.2.2 Gasoline Isomerate Fractionation 6.3 Petrochemicals 6.3.1 Separation of Light Olefins/Paraffins 6.3.2 Separation of Xylene Isomers 6.3.3 Recovery of Monomers 6.4 Concluding Remarks 6.5 List of Abbreviations References Chapter 7

Membrane Gas Separation Processes for Post-combustion CO2 Capture Peter Michael Follmann, Christoph Bayer, Matthias Wessling and Thomas Melin 7.1 7.2

7.3 7.4

Introduction Boundary Conditions 7.2.1 Upstream Boundary Conditions: the Power Plant 7.2.2 Downstream Boundary Conditions: CO2 Transport 7.2.3 Downstream Boundary Conditions: CO2 Storage 7.2.4 Summary of Boundary Conditions Membranes and Membrane Model Driving Force 7.4.1 Feed Compression 7.4.2 Suction at the Permeate Side 7.4.3 Feed Compression and Suction at the Permeate Side 7.4.4 Sweep Operation

150 150 152 160 163 168 168 170 173 173 184 185 185 186 187

196

196 197 197 198 199 200 200 201 201 202 202 203

View Online

xiii

Contents

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

7.5

Chapter 8

Techno-economic Analysis 7.5.1 Process Configurations 7.5.2 Key Performance Indicators and Economics 7.5.3 Process without Retentate Recycling 7.5.4 Process with Retentate Recycling 7.6 Competing Technologies 7.7 Concluding Remarks Acknowledgements References

203 203 204 206 208 211 211 212 212

Commercial Applications of Membranes in Gas Separations Pushpinder S. Puri

215

8.1

Gas Separation Membrane Systems 8.1.1 Gas Separation Processes 8.1.2 Polymeric Gas Separation Membrane Systems 8.2 Major Gas Separation Membrane Producers 8.3 Gas Separation Membrane Applications 8.3.1 Air Separation Membranes 8.3.2 Air Drying 8.3.3 Hydrogen Separation Membrane Systems 8.3.4 Natural Gas Upgrading Systems 8.3.5 Carbon Dioxide Separation Membrane Systems: CO2 Capture from Flue Gases 8.3.6 Organic Vapor Separation Systems 8.4 Concluding Remarks References Chapter 9

Novel Hybrid Membrane/Pressure Swing Adsorption Processes for Gas Separation Applications Isabel A.A.C. Esteves and Jose´ P.B. Mota 9.1

Gas Separation Technologies 9.1.1 Introduction 9.1.2 Pressure Swing Adsorption 9.2 Hybrid Membrane/PSA Processes for Gas Separation 9.2.1 Scheme A: the More Permeable Component is the Least Adsorbed 9.2.2 Scheme B: the More Permeable Component is the More Adsorbed 9.3 Concluding Remarks References Subject Index

215 216 217 225 226 226 230 232 234 239 240 242 243 245

245 245 247 251 260 268 272 273 276

View Online

xiv

Contents

Volume 2

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

Chapter 10

Modeling of Membrane Reactors for Hydrogen Production and Purification F. Gallucci, M. van Sint Annaland and J.A.M. Kuipers 10.1 10.2 10.3

Chapter 11

1

Introduction Limit Conversion in Membrane Reactors Packed Bed Membrane Reactors 10.3.1 One-dimensional Models 10.3.2 Two-dimensional Models 10.4 Fluidized Bed Membrane Reactors 10.4.1 Modeling of Fluidized Bed Membrane Reactors 10.4.2 Multi-Scale Modeling of Dense Gas–Solid Systems 10.5 Appendix A: Constitutive Equations used in Packed Bed Modeling 10.6 Appendix B: Constitutive Equations used in Fluidized Bed Modeling 10.7 List of Symbols References

1 3 6 7 13 17

Palladium-based Membranes in Hydrogen Production Rune Bredesen, Thijs A. Peters, Marit Stange, Nicla Vicinanza and Hilde J. Venvik

40

11.1 11.2

40

Introduction Conventional Hydrogen Production and Applications 11.2.1 Hydrogen Production 11.2.2 Hydrogen Application 11.3 Development of Palladium-based Membranes and Stability Issues in Hydrogen Production 11.3.1 Membrane Development 11.3.2 Membrane Fabrication Methods 11.3.3 Palladium-alloys and their Implications for Membrane Stability 11.3.4 Structural Stability of Composite Palladium-based Membranes 11.4 Integration of Palladium-based Membranes in Hydrogen Production 11.4.1 Methane Reforming 11.4.2 Water Gas Shift

18 25 29 32 34 35

41 41 44 45 45 45 47 49 50 50 55

View Online

xv

Contents

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

11.4.3 11.4.4

Reforming of Alcohols Dehydrogenation and Coupled Endothermic and Exothermic Reactions 11.4.5 Decomposition of Ammonia 11.5 Demonstration of Up-scaled Hydrogen Production by Palladium-based Membrane Reactors 11.6 Examples of Up-scaled State-of-the-Art Palladium-based Membrane Technology 11.6.1 CRI-Criterion 11.6.2 Pall Corporation 11.6.3 Energy Centre of the Netherlands 11.6.4 Membrane Reactor Technologies 11.7 Concluding Remarks Acknowledgements References Chapter 12

60 62 63 64 64 65 69 70 70 71 71

Membrane Reactors in Hydrogen Production A. Brunetti, G. Barbieri and E. Drioli

87

12.1 12.2

87

Introduction Membranes and Membrane Reactors for Hydrogen Production 12.3 Current and Potential Applications of Membrane Reactors for Hydrogen Production 12.3.1 Steam Reforming of Methane and other Light Hydrocarbons 12.3.2 Water Gas Shift 12.3.3 Carbon Monoxide Clean-up 12.4 New Indexes for the Comparison of Membrane and Traditional Reactors 12.4.1 Case Study: Water Gas Shift Reaction in a Membrane Reactor 12.5 Concluding Remarks 12.6 List of Symbols, Abbreviations and Dimensionless Numbers Acknowledgements References Chapter 13

57

Palladium-based Selective Membranes for Hydrogen Production G. Iaquaniello, M. De Falco and A. Salladini 13.1

Basic Features of Membrane Reactors 13.1.1 Selective Membranes 13.1.2 Membrane Fabrication Methods

90 92 93 95 99 100 100 107 107 108 108

110

110 111 114

View Online

xvi

Contents

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

13.1.3

Palladium-based Membranes Available on the Market 13.1.4 Membrane Cost Analysis 13.2 Membrane Reactor Architectures 13.2.1 Configuration Layouts 13.2.2 Benefits and Drawbacks 13.3 Case Studies 13.3.1 Natural Gas Steam Reforming 13.3.2 Water Gas Shift Reactor 13.3.3 Propane Dehydrogenation 13.3.4 Catalytic Partial Oxidation 13.3.5 Catalytic Decomposition of Hydrogen Sulfide 13.4 Concluding Remarks References Chapter 14

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes Giuseppe Barbieri, Alessio Caravella and Enrico Drioli Palladium-based Membranes: Overview and Potential for Hydrogen Purification 14.2 Objectives 14.3 Gas–surface Interactions for Palladium-based Membranes 14.4 Concentration Polarization in Gas Separation 14.5 Inhibition by Carbon Monoxide in Palladium-based Membranes 14.6 Coupled Effect of Concentration Polarization and Inhibition by Carbon Monoxide 14.6.1 Concentration Polarization Coefficient 14.6.2 Inhibition Coefficient 14.6.3 Overall Permeation Reduction Coefficient 14.6.4 Main Results of Analysis 14.7 Concluding Remarks 14.8 List of Symbols and Abbreviations Acknowledgement References

116 117 118 119 121 124 124 127 128 129 130 132 133

137

14.1

Chapter 15

137 139 139 140 142 142 143 148 149 151 157 158 158 158

Carbon Molecular Sieve Membranes for Gas Separation May-Britt Ha¨gg and Xuezhong He

162

15.1 15.2

162 164

Introduction Production of Carbon Molecular Sieve Membranes

View Online

xvii

Contents

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

15.2.1 Material Selection 15.2.2 Material Functionalization 15.2.3 Precursor Preparation 15.2.4 Pretreatment 15.2.5 Carbonization 15.2.6 Post-treatment 15.3 Characterization for Carbon Molecular Sieve Membranes 15.3.1 General Characterization Techniques 15.3.2 Gas Sorption 15.3.3 Gas Permeation 15.3.4 Aging and Regeneration 15.4 Theory on Transport Mechanisms for Carbon Molecular Sieve Membranes 15.4.1 Knudsen Diffusion 15.4.2 Selective Surface Flow 15.4.3 Molecular Sieving 15.5 Module Construction 15.6 Potential Industrial Applications for Carbon Molecular Sieve Membranes 15.6.1 Biogas 15.6.2 Natural Gas 15.6.3 Flue Gas 15.6.4 Air Separation 15.6.5 Petrochemical Industry 15.6.6 High-temperature Applications 15.7 Concluding Remarks Acknowledgement References Chapter 16

Perovskite Membranes for High Temperature Oxygen Separation F. Liang and J. Caro 16.1 16.2

Introduction Materials Aspects of Oxygen Transporting Membranes 16.3 Oxygen Separation by Oxygen Transporting Membranes 16.3.1 Using Sweep Gases 16.3.2 With Evacuation on the Permeate Side 16.3.3 Applying Elevated Pressure on the Permeate Side 16.3.4 Combining Evacuation of the Permeate Side and Elevated Pressure on the Feed Air Side

164 165 166 166 168 169 170 170 173 174 177 178 179 179 180 180 181 181 182 183 184 185 186 187 188 188 192

192 194 196 196 201 202

203

View Online

xviii

Contents

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

16.4

Oxygen Separation from Air with its Immediate Consumption in a Partial Oxidation 16.4.1 Partial Oxidation of Methane to Syngas 16.4.2 Oxidative Coupling of Methane 16.4.3 Oxi-dehydrogenation of Alkanes to the Corresponding Olefins 16.5 Oxygen Separation from Oxygen-containing Gases and its in situ Consumption in a Partial Oxidation 16.5.1 Water as an Oxygen Source for Hydrogen Production Coupled with Synthesis Gas or Ethylene Production 16.5.2 Decomposition of N2O and NO into Nitrogen and using the Abstracted Oxygen for Synthesis Gas Production 16.6 Engineering and Scale-up Aspects 16.7 Comparing Cryogenic Air Distillation, Pressure Swing and Permeation with Organic and Inorganic Membranes: Economic Evaluation 16.8 Concluding Remarks Acknowledgements References

Chapter 17

Chapter 18

204 204 205 207 209

210

212 214

215 217 218 218

Zeolite Membranes for Gas Separations C. Algieri, G. Barbieri and E. Drioli

223

17.1 17.2 17.3 17.4

Introduction to Zeolite Membranes Preparation of Zeolite Membranes Mass Transport in Zeolite Membranes Zeolite Membranes and Gas Separations 17.4.1 Carbon Dioxide Separation 17.4.2 Hydrogen Separation 17.5 Concluding Remarks 17.6 List of Symbols Acknowledgements References

223 225 229 231 231 240 248 249 249 249

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production X. Tan and K. Li

253

18.1 18.2

Introduction Oxygen Permeation in MIEC Ceramic Membranes 18.2.1 Oxygen Permeation Mechanism 18.2.2 Permeation Flux 18.2.3 Stability

253 254 254 255 258

View Online

xix

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

Contents

Chapter 19

18.3

Development of MIEC Hollow Fiber Membranes 18.3.1 Preparation 18.3.2 Surface Modification 18.3.3 Mechanical Strength 18.4 Design of Hollow Fiber Membrane Systems 18.4.1 Operation Mode 18.4.2 Design Equation 18.4.3 Hollow Fiber Membrane Systems 18.5 Energy Consumption and Cost Analysis 18.6 Concluding Remarks 18.7 List of Symbols References

260 260 263 264 265 265 266 269 270 274 275 276

New Metrics in Membrane Gas Separation A. Brunetti, G. Barbieri and E. Drioli

279

19.1 19.2

279

Introduction Current Applications of Membranes in Gas Separation 19.2.1 Case Study: Hydrogen Recovery 19.3 Comparison of Membrane Gas Separation and the Other Separation Technologies: Engineering Evaluation 19.3.1 Technologies for Gas Separation 19.3.2 Selection Guidelines for Gas Separation 19.3.3 Case Study: Selection Guidelines for the Separation and Recovery of Hydrogen in Refineries 19.4 New Metrics for Gas Separation 19.4.1 Case study: H2 Separation from H2/N2 and H2/CO Mixtures with co-polyimide Hollow Fiber Modules 19.5 Further Evaluations in Membrane Process Design: The Exergetic Aspects 19.6 Concluding Remarks Acknowledgement References

Subject Index

280 281

284 284 286

291 294

296 298 300 300 300

302

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-FP009

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

CHAPTER 10

Modeling of Membrane Reactors for Hydrogen Production and Purification F. GALLUCCI*, M. VAN SINT ANNALAND AND J.A.M. KUIPERS Chemical Process Intensification, Multiphase Reactors, Eindhoven University of Technology, Eindhoven, the Netherlands

10.1 Introduction Membrane reactors (often multiphase reactors) integrate a catalytic reaction (generally reforming reactions in case of hydrogen production) and a separation through a membrane in a single unit. This combination of process steps results in a high degree of process integration/intensification. When compared with a conventional configuration in which a reactor is followed by a separation unit, the use of membrane reactors can bring various potential advantages such as reduced capital costs (due to the reduction in the number of process units), improved yields and selectivities (due to the shift in the reaction equilibrium in case of selective removal of one of the products) and reduced downstream separation costs (the separation is integrated). The success of membrane reactors for hydrogen production is basically associated with (i) the advances in membrane production methods; and (ii) the design of innovative reactor concepts, which allow the integration of separation and energy exchange, the reduction of mass and heat transfer resistances and simplification of the housing of the membranes. Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

1

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

2

Chapter 10

As a result of continuous improvements in membrane science, ultra-thin, highly permeable and highly selective H2 membranes have recently become available, which has triggered the development of novel, improved membrane reactor concepts for the production of ultra-pure hydrogen. Different types of membrane reactors for hydrogen production have been proposed in the literature. Most of the previous work has been performed in packed bed membrane reactors (PBMRs); however, there is an increasing interest in novel configurations such as fluidized bed membrane reactors (FBMRs) and membrane micro-reactors (MMRs), especially because better heat management and decreased mass transfer limitations can be obtained in these novel reactor configurations. The aim of this chapter is to show the design features of different types of membrane reactors (ranging from packed bed to fluidized bed reactors), the simulation of which is an important step towards the scale-up and industrial exploitation of membrane reactors. The packed bed membrane reactor configuration is the first and most studied configuration for hydrogen production in membrane reactors. In fact, the first studies on membrane reactors mainly focussed on the effect of the hydrogen permeation through the membranes on the performance (in terms of conversion) of the reaction system. Thus, it was relatively straightforward to compare (both experimentally and theoretically) the performance of two packed bed reactors in one of which the tubular wall was replaced by a membrane. Packed bed membrane reactors have been used for producing hydrogen via reforming of methane, reforming of alcohols, autothermal reforming, partial oxidation of methane, water gas shift, etc. In the packed bed the catalyst is in a fixed configuration and in contact with a hydrogen permselective membrane. The most used packed bed configuration is the tubular one where the catalyst may be packed either inside the membrane tube or in the shell side, while the permeation stream is collected on the other side of the membrane (in case of hydrogen selective membranes) or one reactant is fed at the other side of the membrane (in case of oxygen selective membrane). The models used for PBMRs are essentially the same models available for fixed bed reactors in which the permeation through the membrane is added as a source/sink term in one-dimensional (1D) models and incorporated in the boundary conditions at the membrane wall in two-dimensional (2D) models. In this chapter both 1D and 2D models (pseudo-homogeneous and heterogeneous) will be described and applied to hydrogen production for a standard reaction system (methane reforming). One-dimensional models have been extensively used in the literature to simulate membrane reactors and to compare the reactor performance with the conventional systems (without membranes). This comparison has been, so far, fair enough because thick membranes (i.e. low flux membranes) were generally considered in those works. Even 40–100 mm thick self-supported membranes have been considered. At these conditions (unfortunately too far away to be

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

3

considered interesting for industrial exploitation) the flux through the membrane is generally the limiting step in the reactor performance, and external mass transfer limitations (bed-to-wall mass transfer limitations, also known as concentration polarization) can be neglected and 1D models can well describe the laboratory scale experimental results. Moreover, often steam reforming reactions carried out in a small reactor placed in a furnace, have been considered. In these cases even a 1D isothermal model was sufficient to model the system. The application of 1D non-isothermal models can show some limitations of PBMRs, especially in case of autothermal reforming (where an oxidation reaction is used to supply the energy required for the reforming). On the other hand, with ultra-thin (high permeation flux) membranes, which have recently become available, it has been experimentally shown that the extent of bed-to-wall mass transfer limitations in case of hydrogen purification/ production become prominent, which greatly influences the reactor performance. When these limitations prevail, the hypotheses behind the 1D model are no longer valid and more sophisticated 2D models need to be used. In this chapter it will be shown how 2D models can be used to predict the extent of external mass transfer limitations and their effect on the reactor performance. Also the effect of a radial porosity profile (important where the ratio of the tube diameter over the particle diameter is smaller than about 10) can be included in the model. The main drawbacks of the packed bed membrane reactors are related to the temperature profiles occurring in these reactors (with possible hot spot formation) which are detrimental for the membrane stability, the bed-to-wall mass transfer limitations and, to some extent, also the intra-particle mass transfer limitations because relatively large particle sizes are often applied to prevent large pressure drops. All these detrimental phenomena can be circumvented by using a fluidized bed membrane reactor. In this case the membranes are immersed in a fluidized bed of small catalytic particles. The fluidization regime results in a virtually uniform temperature throughout the reactor even in highly exothermic reactions. At the same time the bed-to-wall mass transfer limitations are strongly reduced while the small particle size also results in negligible intra-particle mass transfer limitations. Possible bubble-to-emulsion mass transfer limitations can be reduced by optimal positioning of the membranes in the fluidized bed. In this chapter the fluidized bed membrane reactors will be simulated with a two-phase phenomenological model. The extent of bubble-toemulsion mass transfer limitations will also be discussed. The results in terms of recovery and conversion in fluidized bed membrane reactors will be compared with the results of packed bed membrane reactors. Finally, the multi-scale approach to gas–solid fluidized beds will be shortly introduced to indicate how the modeling can be improved in the future.

10.2 Limit Conversion in Membrane Reactors Membrane reactors are often used to circumvent the equilibrium constraints which limit the conversion in conventional reactor systems. Although a

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

4

Chapter 10

membrane reactor is, in principle, able to give higher conversion than a traditional system, it is a good exercise to first compute the limit conversion1,2 attainable in a membrane reactor before proceeding in a more detailed simulation and experimentation of such a system. From simplified calculations it can be found that this limit conversion exists only in some cases, being 100% in the other cases.2 Let us now consider the main differences between a traditional system and the corresponding membrane system: whatever the reactor type considered (packed bed, stirred tank, fluidized bed, etc.) the difference between the two systems is the selective permeation through the membrane of one or more species which occurs in the membrane system. We can use this knowledge to compute the limit conversion in a membrane reactor by simply adding the equilibrium of permeating species through the membrane to the condition of chemical equilibrium of the traditional system. This means that, the pseudo-equilibrium state of the membrane system is achieved if the two following conditions are simultaneously satisfied:  Chemical equilibrium inside the reaction zone (as for the traditional system)  Partial pressures equilibrium between the reaction zone and permeation zone (valid for the membrane reactor). If a non-isothermal system is considered, the energy balance has to be satisfied along with the above-mentioned conditions. The term pseudoequilibrium is used here to indicate the limit conversion and to compare it with the equilibrium conversion of a traditional system. Since the membrane reactor is an open system, it is indeed not fully correct to use the term ‘‘equilibrium’’. By using this definition of pseudo-equilibrium, it is straightforward that the limit conversion of a membrane system can be as high as 100% in particular cases. Let us consider two similar cases. The first is a system where the total pressure at the permeation side is zero and only (one or more) products can permeate membrane, while the second case occurs when, at the permeation side, an infinite amount of inert (and not permeating) gas is used and only products can permeate through the membrane. In both cases, the partial pressure of the products in the permeation side is always equal to (or close to) zero, so that the second condition (equilibrium of the partial pressures) is reached only if the reaction conversion and products permeation is complete: limit conversion is 100%.2 Let us now consider a simple reaction in a membrane reactor where only a product is permeating through the membrane (and where the pressure in the permeate side is different than zero). For this calculation consider the typical steam reforming of methane which is an equilibrium reaction producing hydrogen, carried out in a Pd-based dense membrane which allows only the product hydrogen to permeate and leave the reaction zone.

View Online

5

Modeling of Membrane Reactors for Hydrogen Production and Purification

For simplicity we will consider only one reaction as: CH4 þ H2 O , CO þ 3H2

ð10:1Þ

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Applying the conditions for the limit conversion results in the following:  Local chemical equilibrium inside the reaction zone  Hydrogen permeation equilibrium through the membrane (partial pressure of hydrogen is the same in the two zones of the membrane reactor). The first condition can be written by using the stoichiometric method with the ideal gas law as: Keq ¼

YH3 2 YCO 2 P ¼ f ðTÞ YCH 4 YH2 O

ð10:2Þ

The second condition leads to the following equality:3 PH2;reac ¼ PH2;perm

ð10:3Þ

Let us solve this system with initial concentrations CH4/H2O ¼ 1/3, no initial products and a permeation pressure equal to 1 bar. The calculation show the results reported in Figure 10.1.2 From the figure it can be seen that the membrane reactor is able to overcome the equilibrium conversion of a conventional system; however, this calculation also shows that the membrane reactor has a limit conversion which is well below 100% for a wide range of temperatures and pressures. For instance, by working at temperatures lower than 900 K and 20 bar it will be impossible to

Figure 10.1

Equilibrium conversion for conventional system (TR) and limit conversion for membrane system (MR) as a function of temperature and reaction pressure. Reprinted from Gallucci et al.2 with permission of Wiley.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

6

Chapter 10

reach 100% conversion. This implies that it is not worth installing more membrane area or thinner membranes in a real reactor when the conversion has reached the limit conversion. The same calculation can be done for different reaction systems, with more reactions taking place simultaneously and with more compounds permeating the membrane, and it is irrespective of the membrane reactor considered. Once the limit conversion for the reactor has been computed, the detailed simulation of the membrane reactor can be carried out by using the models described in the following sections.

10.3 Packed Bed Membrane Reactors A packed bed membrane reactor is an assembly of usually uniformly sized catalytic particles, which are randomly arranged and firmly held in position within a vessel or tube. A permeable membrane (generally tubular) is immersed within the particles or represents the tube wall of the fixed bed. The PBMR could look, for example, like a tube-in-shell or a multi-tubular reactor. Zooming in on the reaction zone the different phenomena occurring in the reactor can be described as follows:  The reactants are transported first from the bulk of the fluid to the catalyst surface.  The reactants permeate through the pores of the catalyst.  The reactants adsorb on the surface of the pores.  The chemical transformation takes place.  The formed products desorb from the surface.  They are transported back into fluid bulk.  The desired product is transported from the bulk to the membrane surface.  The product is transported through the membrane and separated from the reaction zone. It has to be noted that the last phenomenon is, in general, a combination of different contributions (elementary steps) depending on the type of membrane (porous or dense, organic or inorganic, etc.) and it is often represented by a phenomenological permeation equation (see below). Along with these general steps, the convection of the bulk fluid is tied in with heat and mass dispersion. Dispersion effects are largely caused by the complex flow patterns in the reactor induced by the presence of the packing, by transport phenomena like molecular diffusion, thermal conduction in fluid and solid phases and radiation and by the presence of the membrane itself. Last, but not least, chemical reactions are generally accompanied with heat generation or consumption to be taken into account when modeling the process. The above-mentioned phenomena make an exact description of a packed bed membrane reactor either impossible or lead to very complex mathematical problems. The more detailed the mathematical model, the more parameters it will contain.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

7

Most of the elementary steps described above can hardly be individually and independently investigated and for this reason the more detailed models suffer from a lack of accurate parameter estimations. Therefore, the description of PBMRs is often carried out via simplified models capturing the most crucial and salient features of the problem at hand. The best model is selected on the basis of the properties of the particular system under consideration, the features of the system one is interested in and the availability of the parameters included in the model. The most commonly used class of PBMR models is continuum models. In this type of models the heterogeneous system is treated as a one-phase or multi-phase continuum. To simulate a PBMR, appropriate reaction rate expressions are required and the transport phenomena occurring in the catalyst pellet, bulk fluid and their interfaces, as well as through the membrane need to be modeled. These phenomena can be classified into the following categories:       

Intra-particle diffusion of heat and mass Heat and mass exchange between catalyst pellet and bulk fluid Convection of the fluid Heat and mass dispersion in the fluid phase Thermal conduction in the solid phase Heat exchange with the confining walls Heat and mass exchange through the membrane walls.

The degree of sophistication of the model is determined by the accepted assumptions and, consequently, by the way the aforementioned phenomena are incorporated in the model. According to the classification given by Froment and Bishoff,4 the continuum models can be divided in two categories: pseudo-homogeneous and heterogeneous models. In pseudo-homogeneous models it is assumed that the catalyst surface is totally exposed to the bulk fluid conditions, i.e. that there are no fluid-toparticle heat and mass transfer resistances. On the other side, heterogeneous models take conservation equations for both phases into account separately. A general schematic classification of continuum models is given in Figure 10.2 and the following sections (Iordanidis)5:

10.3.1 10.3.1.1

One-dimensional Models One-dimensional Pseudo-homogeneous Model

The 1D pseudo-homogeneous model is the most used model to describe packed bed membrane reactors, especially for laboratory-scale applications. In its simplest form, namely the plug flow steady state model, the model describes only axial profiles of radially averaged temperatures and concentrations.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

8

Chapter 10

Figure 10.2

Schematic classification of classical continuum models.

The equations read as:6 Continuity equation @ @ ðerg Þ þ ðerg uÞ ¼ 0 @t @z

ð10:4Þ

Total momentum balance equation @ @ @p @ ðerg uÞ þ ðerg u2 Þ ¼ e  berg u  ðetg Þ þ erg g @t @z @z @z

ð10:5Þ

Friction coefficient b ¼ 150

ð1  eÞ2 mg ð1  eÞ eu þ 1:75 3 2 e3 dr e rg dr

ð10:6Þ

where tg is calculated depending on the nature of fluid. Component mass balance   @ @ @ @oi ðerg oi Þ ¼  ðerg uoi Þ þ rg Dax;i þ S r;i  Ji @t @z @z @z

ð10:7Þ

where the source term, Sr,i, and the trans-membrane flux term, Ji, depend on the reaction system considered and membrane used, respectively. In particular, the source term has the following general formula: X nr Sr;i ¼ ð1  eÞrg Mi ð10:8Þ j¼1 gij rj

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

9

The flux term depends on the membrane used. For hydrogen production and purification dense hydrogen permselective membranes are often used, which often exhibit infinite selectivity towards hydrogen (in the case of Pd-based dense membranes). In this particular case, the flux term reduces to the flux of hydrogen though the membrane and it is equal to zero for all the other components. To compute the hydrogen flux through such a membrane, the Richardson equation is generally applied, which reads:   Ea pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi PH2 ;upstream  PH2 ;downstream Pe0 exp ð10:9Þ RT JH2 ¼ dm This formula is, in general, valid for permeation through thick Pd-based membranes and it has been verified at different times especially with pure gas permeation tests. However, hydrogen production is often obtained through the water gas shift reaction or steam reforming reactions. These reactions proceed with production of carbon monoxide that can poison the Pd-based membranes and reduce their hydrogen flux (depending on Pd alloy, temperatures and CO content). To incorporate this poisoning effect a different equation can be used (Sieverts–Langmuir’s model) as described by Barbieri et al.:7   Ea pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi pffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi 0   Pe exp PH2 ;upstream  PH2 ;downstream KCO PCO RT JH2 ¼ 1  WðTÞ 1 þ KCO PCO dm

ð10:10Þ

The energy balance is: 

erg Cp;g þ ð1  eÞrs Cp;s

   @T  @ @ @T ¼ Cp;g erg uT þ lax þ Sh ð10:11Þ @t @z @z @z

were the source term Sh reads: Sh ¼ ð1  eÞrs

nr X

rj DHj

j¼1

ð10:12Þ

This model can be reduced if isobaric conditions are considered and/or axial dispersion is neglected. Most of the times, simulations are carried out in steady state conditions, so that the whole model is reduces neglecting the time derivative. Generally, standard Danckwerts’ boundary conditions at the reactor inlet and outlet can be assumed for the gas phase balance.8

10.3.1.2

One-dimensional Heterogeneous Model

The heterogeneous model is interesting to study a.o. the effect of internal (in the catalyst) mass transfer limitations on the performances of the membrane

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

10

Chapter 10

reactor. It is clear that the equations related to the catalyst particles and the inter-phase exchange has to be taken on board in this model. To simplify the description, let’s assume the catalyst particles to be of spherical shape of constant particle diameter dp, so that a 1D model can be used to describe the profiles inside the catalyst particle. The catalyst structure is assumed to be macro-porous, so that transport mechanisms like viscous transport or Knudsen diffusion can be neglected. It is assumed that the component mass transport inside the particle is described by Fick’s law of diffusion (due to the relatively low concentrations of the relevant components). The gas phase component mass balance now reads: Component mass balance     @ @ @ @oi erg oi ¼  erg uoi þ rg Dax;i  ans;i  Ji @t @z @z @z

ð10:13Þ

Catalyst phase mass balance   @ 1 @ 2 @oi r Deff ;i r ðroi Þ ¼ 2 þ Sr;i @t r @r @r

ð10:14Þ

where Sr;i ¼ rs Mi

nr X

gij rj

j¼1

ð10:15Þ

The boundary conditions for the catalyst phase are:  @oi  Deff ;i r ¼ ns;i @r r¼R  @oi  ¼0 @r r¼0

ð10:16Þ ð10:17Þ

The inter-phase mass transfer is written as: ns;i ¼ rks;i ðoi  os;i Þ

ð10:18Þ

where ks,i is the interphase mass transfer coefficient. The energy balance for the gas phase is then:    @ @T b @ @T b b erg uT þ lax ¼ Cp;g erg Cp;g þ aqjr¼R @z @z @t @z

ð10:19Þ

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

11

and the energy balance for the solid phase is:

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001



erg Cp;g þ ð1  eÞrs Cp;s

 @T 1 @2  r q þ Sh ¼ 2 @t r @r

ð10:20Þ

where q ¼ le f f ;i

and

Sh ¼ ð1  eÞrs

@T @r

nr X

ð10:21Þ

rj DHj

j¼1

ð10:22Þ

The boundary conditions for the solid phase are:  @T  ¼0 @r r¼0

and

 T r¼R ¼ T b

ð10:23Þ

The remaining of the constitutive equations are reported in Section 10.5. The application of these reactor models is quite straightforward. For example, let us consider again the hydrogen production in a membrane reactor applied for methane steam reforming. Suppose that the membrane used is a dense defect free Pd-based membrane which obeys the Richardson equation (10.9). The reactions taking place in the reactor are: CH4 þ H2 O , CO þ 3H2

ð10:24Þ

CO þ H2 O , CO2 þ H20

ð10:25Þ

where the rate expressions are the following:9   k1 PCH4 pH2O  P3H2 pCO =Keq;1 r1 ¼ P1:596 H2O

ð10:26Þ

  k2 PCO pH2O  PH2 pCO2 =Keq;2 PH2O

ð10:27Þ

r2 ¼

The hydrogen permeation rate through the palladium membranes follows Richardson’s equation, where the values of the apparent activation energy Ea and pre-exponential factor Pe0 are 12 540 J mol1 and 2.21  1003 mol s1 m2 Pa0.5, respectively (experimentally determined). By using a 1D heterogeneous model, in non-isothermal conditions (by solving the energy balances mentioned above) the effect of heat profiles can be studied in the reactor. A typical result is depicted in Figure 10.3 where in the packed bed membrane reactor a temperature drop in the first part of the reactor

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

12

Chapter 10

Figure 10.3 Axial temperature profile in a packed bed membrane reactor for methane

steam reforming. Reprinted from Gallucci et al.10 with permission of Professor T. Nejat Veziroglu.

is observed (and similar results are obtained by changing the heat flux through the walls, the flow rates etc).10 The results show a temperature drop of 80–100 K in the first part of the reactor, which can give stability and sealing problems for the membrane. In fact, the membrane material should stand at a great axial temperature gradient which can cause the detachment of the Pd-based layer from the support with consequent loss in permselectivity. Moreover, the first part of the membrane is not effectively used since it is working at low temperature which, following the Richardson equation, results in a lower hydrogen flux. The decrease of temperature at the beginning of the reactor also gives a decrease of the reaction rate. The result is an increase of the membrane area needed for the required conversion. In particular, the membrane area required increases by around 21% in comparison to isothermal operation. This suggests that the assumption of isothermal conditions should be used with great care when dealing with methane reforming in packed bed membrane reactors. The application of 1D models is quite useful in describing the axial temperature and concentration profiles inside the reactor, while accounting for the effects of intra-particle mass transfer limitations, and membrane characteristics on the reactor performances and is useful for a first design of the membrane reactor. A big limitation of such a model is that radial profiles are ignored. If one considers that in packed bed membrane reactors for hydrogen production, the hydrogen should ‘migrate’ from the catalytic bed (where it is produced) to the membrane surface (where it permeates), it is clear that radial profiles can be

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

13

of great importance. In fact, the hydrogen partial pressures used in the Richardson equation to compute the hydrogen flux correspond to the partial pressures at the two faces of the membrane; the partial pressure upstream at the surface of the membrane is (much) lower than the bulk partial pressure, if radial profiles are present, in which case a 1D model over-estimates the flux (or under-estimates the membrane area required for a given hydrogen recovery). To study the effects of bed-to-wall mass transfer limitations (also called concentration polarization) a 2D model is thus required as described in the following sections.

10.3.2

Two-dimensional Models

In the following section a two-dimensional model will be described that is used for the computation of temperature and concentration profiles inside a packed bed membrane reactor for hydrogen production. For simplicity, only a pseudohomogeneous model will be described. The extension of the heterogeneous model is analogous to the 1D model.

10.3.2.1

Two-dimensional Pseudo-homogeneous Model

A pseudo-homogeneous, two-dimensional reactor model for membrane reactors consists of the total gas-phase continuity and Navier–Stokes equations augmented with gas-phase component mass balances and the overall energy balance. The model is based on the standard dispersion model that describes the gas phase mass and energy transport as convective flow with superimposed radial and axial dispersion. The model equations in two-dimensional axi-symmetrical cylindrical coordinates and the boundary conditions are listed below. The following assumptions have been made in this model (although the model could be extended to include these phenomena):  The particle size is sufficiently small so that both intra-particle mass and heat transfer limitations and external mass and heat transfer limitations from the gas bulk to the catalyst surface can be neglected.  Homogeneous gas phase reactions are neglected in view of the relatively low temperatures.  The gas bulk can be described as an ideal Newtonian fluid. The most important constitutive equations for the reaction kinetics, membrane flux are described above while axial and radial dispersion coefficients are discussed in Section 10.5. The main equations read: Continuity equation   @  erg þ r  erg u ¼ 0 @t

ð10:28Þ

View Online

14

Chapter 10

Total momentum balance equation

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

     @ erg u þ r  erg uu ¼ erp  berg u  r  etg þ erg g @t

ð10:29Þ

Friction coefficient

b ¼ 150

ð1  eÞ2 mg ð1  eÞ ejuj þ 1:75 dp e3 e3 rg dp2

ð10:30Þ

where tg is calculated as (10.Newtonian fluid):   h i ¼ 2 tg ¼  lg  mg ðr  uÞ I  mg ðr uÞ  ðr uÞ T 3

where

juj ¼

qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi u2r þ u2z and

rg ¼

Mg p RTg

ideal gas

ð10:31Þ

ð10:32Þ

Boundary conditions for a 2D model for packed bed membrane reactor @uz ¼0 @r uz ¼ 0

Center (r=0) Wall (r=R)

ur ¼ 0 ur ¼

F00m erg

Inlet (z=0)

uz ¼

Outlet (z=L)

p ¼ p0

JH2 erH2

@ur ¼0 @z @ur ¼0 @z

Component mass balance      @  @oi erg uoi ¼ r  erg uoi þ r  rg Di þ Sr;i @t @z  Dr;i 0 where Di ¼ 0 Dz;i

ð10:33Þ ð10:34Þ

where the source term Sr,i equals: Sr;i ¼ ð1  eÞrs Mw;i

nr X j¼1

gij rj

for i ¼ 1; 2 :: nc

ð10:35Þ

Energy balance 

erg Cp;g þ ð1  eÞrs Cp;s

    @T ¼ Cp;g r  erg uT þ r  l þ Sh @t @z

 @T

ð10:36Þ

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

where Di ¼



lr 0 0 lz



15

ð10:37Þ

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

where the source term Sh reads: Sh ¼ ð1  eÞrs

nr X j¼1

rj rHj

for j ¼ 1;2 ::nr

ð10:38Þ

Boundary conditions Position

Mass balance

Energy balance

Center (r=0)

@oi ¼0 @r

@T ¼ 0 @r

Wall (r ¼ R) Inlet (z ¼ 0) Outlet (z ¼ L)

JH2 ¼ ur rH2 e @oi ¼ 0 i 6¼ H2 @r 00 @oi Fm;i þ erg uz oi ¼ rg Dz;i @z Areactor @oi ¼0 @z

@T ¼ 0 adiabatic @r T ¼ Twall heated wall l

Cp;g T0 F00m @T þ Cp;g erg uz T ¼ @z Areactor

@T ¼0 @z

This model, applied to a packed bed membrane reactor for hydrogen production through methane reforming can give in particular indications on the extent of mass transfer limitations. Let us compute the radial H2 concentration profiles at different axial positions at isothermal conditions. As can be seen in Figure 10.4, radial concentration profiles are present but not very pronounced. It can be concluded that for the membranes used and for small membrane diameters (1 cm in the simulation shown in the figure), the bed-to-wall mass transfer limitations have a negligible influence on the required membrane area. With the actual developments and optimization of Pd-based membranes, higher membrane fluxes have become available. Whether concentration polarization will occur with increased permeability was investigated numerically. Membrane research has produced thinner membranes with higher hydrogen permeability which will result (Figure 10.5) in higher mass transfer limitations in packed bed membrane reactors (even with small reactor diameters) and consequently in larger membrane area required for the same hydrogen separation (with respect to a case without mass transfer limitations). It has been found that, in the worst case the packed bed membrane reactor requires almost double the membrane area with respect a membrane reactor without mass transfer limitations.11 Of course the same model can be used for evaluating the temperature profiles inside the reactor as was demonstrated by Tiemersma et al.12

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

16

Chapter 10

Figure 10.4

Radial profile of the H2 weight fraction for the isothermal operation mode. Reprinted from Gallucci et al.10 with permission of Professor T. Nejat Veziroglu.

Figure 10.5

Relative H2 weight fraction for the isothermal packed bed for different membranes. Reprinted from Gallucci et al.11

The extent of mass and heat transfer limitations in packed bed membrane reactors have forced researchers to investigate other solutions to circumvent those limitations. In this respect, membrane assisted fluidized bed reactors have

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

17

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

shown great advantages compared to packed beds. This brings us to the modeling of fluidized bed reactors discussed in the next session.

10.4 Fluidized Bed Membrane Reactors The integration of membranes (dense or porous, generally non-catalytic) inside a fluidized bed of catalyst, allows the combination of the benefits of both separation through membrane and benefits derived from the fluidization regime. It is well known (as described above) that packed bed membrane reactors suffer from the same disadvantages as conventional packed bed reactors; that is to say: (i) relatively high pressure drop, (ii) possible mass transfer limitations owing to the relatively large particle size to be used, (iii) radial temperature and concentration profiles, (iv) difficulties in reaction heat removal or heat supply, and (v) low specific membrane surface area per reactor volume. On the other hand, as summarized in the review presented by Deshmukh et al.,17 the main advantages of the fluidized bed membrane reactors are:  Negligible pressure drop; no internal mass and heat transfer limitations because of the small particle sizes that can be employed  Virtual isothermal operation (even in case of highly exothermic reactions)  Flexibility in membrane and heat transfer surface area and arrangement of the membrane bundles  Improved fluidization behavior as a result of:  Compartmentalization, i.e. reduced axial gas back-mixing  Reduced average bubble size due to enhanced bubble breakage, resulting in improved bubble to emulsion mass transfer. Some disadvantages are of course foreseen such as:  Difficulties in reactor construction and membrane sealing at the wall  Erosion of reactor internals and catalyst attrition. The last disadvantage can be really critical if high selective thin layer membrane is used inside the fluidized bed. Any erosion on the membrane surface can result in a decreased permselectivity and a decrease in overall membrane reactor performance. For this reason, membranes to be used in fluidized membrane reactors should be protected by erosion, perhaps by using a porous media between the membrane layer and the fluidized bed or using selective layer inside the tube. Fluidized bed membrane reactors for pure hydrogen production have been studied by different research groups (see Rahimpour,18 Patil et al.,19 Gallucci et al.,20 Mahecha-Botero et al.,21 and Abashar and Elnashaie).22 In this case, as discussed in the first part of the review, Pd-based membranes are inserted in fluidized bed reactors where reforming of hydrocarbons takes place.

View Online

18

Chapter 10

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

10.4.1

Modeling of Fluidized Bed Membrane Reactors

A complete overview of the characteristics and related models for fluidized beds is reported in the excellent book by Kunii and Levenspiel.23 It is quite accepted that the most difficult fluidized bed reactor to be simulated is a bubbling fluidized bed, where the description of bubble behavior should be taken into account along with the description of solid movement and reactions occurring on the solid surface. A frequently used phenomenological description of the two-phase flow phenomena in fluidized bed reactors is based on the bubble assemblage model, originally proposed by Kato and Wen.24 A typical 1D two-phase model for a membrane assisted fluidized bed reactor can be used for the simulation of the fluidized bed membrane reactor for hydrogen production via methane reforming. A schematic representation of the gas flows between the compartments of the bubble and emulsion phases is depicted in Figure 10.6. The model main assumptions are:  Hydrogen permselective membranes are immersed in the reactor.  The reactor consists of two phases, viz. the bubble and emulsion phase.

Figure 10.6

A schematic representation of the two-phase fluidized bed reactor model (FBMR) (E ¼ emulsion phase, B ¼ bubble phase).

View Online

19

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

 The gas flowing through the emulsion phase is considered to be completely mixed in each section and at incipient fluidization conditions.  The bubble phase gas is assumed to be in plug flow (i.e. a large number of continuously stirred tank reactors), where the bubble size and the bubble rise velocity changes for each section.  The heterogeneous reactions (methane steam reforming and water gas shift reactions) take place only in the emulsion phase, assuming that the bubble phase is free of catalyst particles.  Gas removed from the fluidized bed via membranes is assumed to be extracted from both the emulsion phase and bubble phase, distributed according to the local bubble fraction. The gas extracted from the emulsion phase is subsequently instantaneously replenished via exchange from the bubble phase (to maintain the emulsion phase at minimum fluidization conditions) (following Deshmukh et al.).25,26  A uniform temperature is assumed throughout an entire section of the fluidized bed, assuming no heat losses to the surroundings (adiabatic conditions) and no heat transfer limitations between the bubble and emulsion phase.27,28 The mass and heat balance equations are as follows: Total mass balance usb;n1 AT rb;n1  usb;n AT rb;n þ use;n1 AT re;n1  use;n AT re;n nc n

o X 00 membrane 00 membrane þ fi;mol Mw;i Amembrane eb;n þ fi;mol Mw;i Amembrane 1  eb;n ¼0 i¼1

ð10:39Þ

Bubble phase component mass balances* usb;n1 AT rb;n1 usb;n AT rb;n  þ

nc X i¼1

nc X i¼1

Kbe;i;n Vb;n rb;n ðob;i;n oe;i;n Þ

00

membrane ji;mol Mw;i Amembrane eb;n þ ½oe;i;n SF ðQÞob;i;n SF ðQÞ ¼ 0

ð10:40Þ

Emulsion phase component mass balances* use;n1 AT re;n1  use;n AT re;n  

nc X i¼1

00 membrane

ji;mol

nc X i¼1

Kbe;i;n Vb;n rb;n ðob;i;n  oe;i;n Þ

 Mw;i Amembrane 1  eb;n 

 ½oe;i;n SF ðQÞ  ob;i;n SF ðQÞ ¼ 0

nrxn X j¼1

!

n j;i rj Ve;n rp;n ð1  ee Þ

ð10:41Þ

View Online

20

Chapter 10

Transfer term

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Q ¼ use;n1 AT re;n1  use;n AT re;n  þ

nc X i¼1

nc X i¼1

 00 membrane ji;mol Amembrane 1  eb;n

Kbe;i;n Vb;n rb;n ðob;i;n  oe;i;n Þ

ð10:42Þ

where use;n AT ¼ ue;n AT ð1  eb;n Þ

usb;0 AT ¼ utot AT eb;0 use;0 AT

*Note that:

ð10:43Þ

¼ utot AT ð1  eb;0 Þ

SFðxÞ ¼



x if x > 0 0 if x  0

ð10:44Þ

Energy balance (in the case of the energy supply inside the reactor) nc X

Tfeed

Hi

i¼1





nc  X usb;n¼0 AT rb;i;n¼0 þuse;n¼0 AT re;i;n¼0  HiTout i¼1

usb;n¼N AT rb;i;n¼N þuse;n¼N AT re;i;n¼N



(

nc X i¼1

HiTout



00 membrane

fi;mole



00 membrane

Mw;i AT eb;n þfi;mole

Mw;i AT ð1eb;n Þ

) 

þE ¼ 0 ð10:45Þ

where E depends on the kind of energy supply used (see e.g. Gallucci et al.20). All the parameters used are described in Section 10.6. As can be seen in Figure 10.6 the model allows to change the number of continuously stirred tank reactors in both bubble and emulsion phases. These parameters can be used to investigate the effect of the degree of gas back-mixing in the bubble and emulsion phases. Moreover these are adjusting parameters to be evaluated through a model validation with experimental data. For example, during steam reforming in a 60 cm high FBMR with inserted 10 dead-end membranes, Gallucci et al.20 concluded that for the predicted membrane fluxes matched reasonably well with the experimental measured fluxes when both the bubble and emulsion phases are considered in plug flow. The model can be used to simulate reactors as big as industrial scale reactors. For example, the membrane area required for a given conversion or a given hydrogen production can be evaluated. Gallucci et al.20 quantitatively evaluated the reactor performance in terms of CH4 conversion, CO selectivity and H2 recovery as a function of the load-tosurface ratio (reciprocal to the required membrane area). Figure 10.7 clearly

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

21

Figure 10.7 CH4 conversion as a function of L/S (load-to-surface area) at different pressures and temperatures H2O/CH4 ¼ 2. Reprinted from Gallucci et al.20 with permission from Wiley.

Figure 10.8

CO selectivity and H2 recovery as a function of L/S (load-to-surface area) at different pressures and temperatures H2O/CH4 ¼ 2. Reprinted from Gallucci et al.20 with permission from Wiley.

shows that for each temperature and pressure investigated, complete CH4 conversion can be achieved with a load-to-surface ratio below 1, while with a loadto-surface ratio above 100 the equilibrium CH4 conversion is obtained and the extent of hydrogen extraction via the membranes is too small to affect the steam reforming equilibrium. With a load-to-surface ratio below 0.6 complete CH4 conversion, maximal H2 recovery and minimal CO selectivity can be realized (Figures 10.7 and 10.8). However, a good compromise between reactor

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

22

Chapter 10

performance and membrane investment costs is probably achieved with a loadto-surface ratio of 1:6, with which about 80% H2 recovery and over 90% CH4 conversion is obtained. However, for this case the CO selectivity might easily exceed 20%, so that a post-treatment of the retentate might be necessary. Alternatively, higher H2O/CH4 ratios could be used. As already stated above, a fluidized bed membrane reactor can be used for circumventing the bed-to-wall mass transfer limitations. However, especially if a bubbling fluidized bed is considered, an important transfer limitation affecting its performance is the mass transfer limitation between the bubble phase and the emulsion phase. In fact, the gas transported inside the bubbles should be exchanged with the emulsion phase to react. A high mass transfer limitation (low mass transfer coefficient) between bubble phase and emulsion phase results in a larger gas slip via the bubble phase and a lower conversion degree. In our fluidized bed membrane reactor model the bubble-to-emulsion phases mass transfer coefficient is calculated with the equations derived for a fluidized bed without internals. Although the internals (solid membranes) should enhance the mass transfer characteristics of the bed, at the moment, reliable equations for bubble-to-emulsion phases mass transfer coefficient for fluidized bed with inserts are not available. As a matter of fact, the bubble-to-emulsion phase mass transfer limitation increases with increasing bubble diameter, which itself increases by increasing the reactor length as schematically indicated in Figure 10.9.

Figure 10.9

Schematic representation of the membrane reactor concept with bubble increasing in size.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

Figure 10.10

23

Effects of bubble to emulsion phase mass transfer. Reprinted from Gallucci et al.10 with permission of Professor T. Nejat Veziroglu.

As a result of this bubble size increase, the methane conversion decreases as indicated in Figure 10.10. The figure shows that the methane conversion decreases by increasing the mass transfer limitations. In case of mass transfer limitations calculated as a fluidized bed reactor without internals (worst case) the methane conversion decreases tremendously as compared with the case without mass transfer limitations (which can be indicated as the ideal condition for fluidized bed reactor). That also shows that by improving the mass transfer by a factor of 10 results in a conversion close to the ideal case without mass transfer limitations. Simulations show that, in order to achieve the same conversion degree of a fluidized bed membrane reactor without mass transfer limitations, the membrane area installed in the reactor needs to be increased 2.4 times with respect to the case without limitations as reported in the figure. Figure 10.10 shows that a decrease of 10 times in the mass transfer limitations is enough to reach the limit conversion required. However, even considering the worst case (bubble to emulsion phase mass transfer coefficient equal to a fluidized bed without internals) the mass transfer problem in the fluidized bed can be easily circumvented. In fact, the mass transfer resistance is higher when the bubble diameter becomes larger, and the bubble diameter increases with the increasing of the bed height, so that we can reduce the bubble diameter by inserting stagers such as meshing wires at different reactor heights (i.e. staging the fluidized bed reactor). In Figure 10.11 the conversion in a fluidized bed with mass transfer limitations is shown for different numbers of stages. For these simulations the axial position of the stagers has not been optimized. This means that the distance

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

24

Figure 10.11

Chapter 10

Conversion reached for a given area in case of mass transfer limitations for different stages (FBMR). Reprinted from Gallucci et al.10 with permission of Professor T. Nejat Veziroglu.

between two stages is constant for one simulation and it is given by dividing the total length of the reactor by the number of stages. The area used in this simulation was kept the same as was needed in the case of no mass transfer limitations. From the figure it can be seen that the conversion required can be achieved already with three to four stages. Thus, dividing the reactor in different stages completely circumvents the problems of mass transfer limitation for the fluidized bed membrane reactor. Similar models, although with different simplifications or more phenomena considered, have been used for simulation fluidized bed membrane reactors by different authors.29–32 Recently, Mahecha-Botero and co-workers presented a generalized comprehensive model which characterizes multiple phases and regions (low-density phase, high-density phase, staged membranes, freeboard region) with the possibility to include new features or simplifications in order to simulate different fluidized bed (membrane) reactors. For a more detailed description of the model and assumptions an interested reader is referred to.33 All the models proposed for fluidized bed membrane reactors have the same limitations. These models are phenomenological models that make use of closure equations originally derived for fluidized bed without internals. Although it is known that the presence of internals (membranes and permeation of gas through them) may vary the behavior of bubbles, mass transfer and solid circulation inside the bed, reliable closure equations for membrane assisted fluidized bed are not yet available. A way to solve this problem is to use a so-called multi-scale modeling of dense gas–solid systems, where low level

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

25

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

models are used for the detailed description of the system and for deriving closure equations for the higher level models. In the following section the multiscale modeling applied to fluidized bed membrane reactors is described.

10.4.2

Multi-Scale Modeling of Dense Gas–Solid Systems

Understanding, and hence predicting, mass, momentum and heat transfer in large-scale dense gas–solid fluidized beds has historically proven to be difficult. The difficulty increases even further when membranes are immersed into the solid suspension and gas is extracted (or added) through the membranes. A possibility to better understand the macroscopic behavior of these systems is the study of all prevailing phenomena in reactive gas–solid suspensions at all relevant scales, ranging from the microscopic to the macroscopic level, using computational fluid dynamics (CFD). Especially the link between the most elementary interactions that take place in the system, i.e. the exchange of mass, heat and momentum at the level of surface of the individual particles, to the macroscopic circulation pattern, extent of gas back mixing and solid mixing and their effect on the reactor performance, are of great importance. To achieve this goal, a so-called multi-scale approach (Figure 10.12) can be adopted in which the gas–solid flows are considered at different distinct levels of

Figure 10.12

Multi-scale approach to dense gas–solid flow. In the DNS and DPM models, the solid phase is represented by the actual particles. At the TFM level, the solid phase is considered as a continuum. In DBM the fluid is considered as discrete phase. Phenomenological models are based on the assumptions described above.35

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

26

Chapter 10

modeling, ranging from a detailed description at small scales (micrometer range), to an effective description at large scales (order of meters).34 In this approach, the smaller scale models are used to determine the closure equations to be used in larger scale models. The final aim is to obtain better and more general closure equations for heat, momentum and mass transfer that can be applied in phenomenological models and account for the presence of and permeation of gas through membranes in membrane assisted fluidized bed reactors instead of the previous described (empirical) closure equations obtained for reactors without membranes. At the most fundamental level (corresponding to direct numerical simulations), the gas flow field is modeled at scales smaller than the size of the particles. The interaction of the gas phase with the particles is considered by imposing suitable boundary conditions at the surface of the solids. The flow field between the spheres can be solved for instance by the lattice Boltzmann method. The detailed gas–particle interaction (in the form of drag force) obtained from this microscopic level lattice Boltzmann method simulation will be applied to higher level models instead of the conventional empirical correlations, which are only valid for spatially homogeneous flows. At a higher level, the flow field is modeled at a scale much larger than the size of the particles, and the fluid velocity and pressure are obtained by solving the volume-averaged Navier–Stokes equations. The particle–particle interactions (particle–wall as well) are formulated with the so-called discrete particle models (DPMs), which are based on the schemes that are traditionally used in molecular dynamics simulations, with the addition of dissipation of mechanical energy. At the macroscopic scale, a two-fluid model (TFM) is used where the continuum description is employed for both the particle phase and the fluid phase.36,37 The information obtained in the two lower-level models is then included in the continuum models via the kinetic theory of granular flow (KTGF). The advantage of this model is that it can predict the flow behavior of gas–solid flows at life-size scales and these models are therefore widely used in commercial fluid flow simulators of industrial scale equipment. At a larger scale, the discrete bubble model (DBM) is especially interesting for studying the macro-scale emulsion phase circulation patterns induced by bubble–bubble interactions and bubble coalescence. In the DBM, the bubbles are modelled as discrete elements and are tracked individually during their rise through the emulsion phase, which is now considered as a continuum. The advantage is that DBM fully accounts for the two-way coupling between the bubbles and the emulsion phase, which enables direct computation of the emulsion phase velocity profiles.38 The interested reader is referred to the reviews by van der Hoef et al.39,40 for more details on the multi-scale approach. In the following an application of the multi-scale modeling for membrane assisted fluidized beds will be given. As already mentioned, an important feature of the membrane assisted fluidized bed to be elucidated is the effect of membranes (and permeation through them) on the bubble behavior (and hence on mass/heat transfer). A DPM can be used in combination with an immersed boundary method to study the behavior of the gas in the immediate vicinity of the membrane. With this

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

27

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

approach, bubble size distribution, particle and gas mixing as a function of both particle properties and operating conditions can be studied in great detail. In the DPM, all particles are tracked individually using Newton’s second law, where vi is the velocity, ri the position, mi the mass and Vi the volume of particle i: mi

 dvi d 2 ri Vi b  ¼ mi 2 ¼ Vi rP þ ug  vi þ mi g þ Fcontact;i ss dt dt

ð10:46Þ

The forces on the right-hand side represent the pressure gradient, drag force, gravity and collision forces respectively. Particle–particle and particle–wall collisions are calculated by taking a soft sphere approach.41 Gas-phase hydrodynamics are described with the continuum model and volume-averaged Navier–Stokes equations. For a more detailed description of the DPM the interested reader is referred to a review by Deen et al.42 In the immerse boundary method the interaction between the Eulerian grid and the immersed object occurs through Lagrangian force points equally distributed over the object’s surface (see Figure 10.13). Each force point exerts a force on the gas phase such that the interpolated gas phase velocity equals the specified gas velocity at the position of that force point.34 Collisions between particles and the immersed object is treated in a similar way as any other collision between particles and walls. The model can be used

Figure 10.13

Schematic illustration of the cylindrical membrane with force points inside a computational domain.34

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

28

Figure 10.14

Chapter 10

Typical simulation result of a DPM-IBM simulation with a horizontal membrane inserted in the fluidized bed reactor.34

to investigate the influence of addition and extraction of gas on particle mixing and bubble size distribution inside the bed. A snapshot of a typical result is reported in Figure 10.14. By changing the imposed flux through the membrane, the effect of gas extraction on the bubble diameter and bubble size distribution can be investigated. CFD models have proven to be very accurate in predicting the behavior of some gas–solid systems (e.g. bubble formation and bubble size distribution),43 but also fail in describing other systems (e.g. the TFM-KTGF strongly overestimates the extent of solid mixing due to neglect of particle–particle friction).44 An important ingredient of the multi-scale modeling approach is that the models need to be thoroughly validated before performing a detailed simulation. The validation can be performed either via experiments or via simulation experiments performed with a lower level validated model. Experimental techniques for validating such models should preferentially be non-invasive techniques in order to avoid any external interference on the (reactive) flow. An example of a state-of-the-art experimental technique for small scale validation is based on high speed cameras which can capture the particle movements inside the fluidized bed through accessible windows. An example is reported in Figure 10.15.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

Figure 10.15

29

Experimental set-up. The local flow structure is measured with the PIV method. The local porosity and CO2 concentration are obtained from local\through beam IR absorbance measurements.

In the particle image velocimetry (PIV) method, a CMOS high speed camera is used to record two subsequent images of the flow, separated by a short time delay, from which an instantaneous velocity field can be constructed.45 The images recorded during the PIV measurements are also analyzed by means of digital image analysis (DIA) to compute the instantaneous particle volume fraction from the solids intensity distribution in the image.43 The particle volume fraction distribution provides insight in the heterogeneity of the flow. Moreover, the particle volume fraction maps obtained from the DIA can be combined with the PIV velocity data to obtain the particle volume flux. These data are afterwards directly compared with the model results for the direct validation of CFD model.46 Only when accurately validated, can the CFD models be used for developing the closure equations required for the higher level and phenomenological models.

10.5 Appendix A: Constitutive Equations used in Packed Bed Modeling In this appendix the constitutive equations used in the modeling of packed bed membrane reactors (either 1D or 2D) are listed. Components properties have be evaluated by making use of book The Properties of Gases and Liquids.53

View Online

30

Chapter 10

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

The definition of porosity profile and the way it was computed in the 2D model is also reported.

Effective Radial and Axial Dispersion Coefficients for Mass and Energy:14–16 Effective dispersion of mass

Effective dispersion of energy

Radial

pffiffiffiffiffiffiffiffiffiffi Dr;j ¼ 1 1  e Dm i þ

udp   PeN f Dt =dp

pffiffiffiffiffiffiffiffiffiffi udp ¼ 1  1  e Dm i þ 8

lr lbed;0 Pex l Pe   ¼ bed;0 þ x ¼ þ lg lg lg 8 KN f Dt =dp

Axial

 pffiffiffiffiffiffiffiffiffiffi udp lr lbed;0 Pex l Pe Dz;i ¼ 1  1  e Dm   ¼ bed;0 þ x ¼ þ i þ 2 lg lg lg 2 KN f Dt =dp  

 p ffiffiffiffiffiffiffiffiffiffi lbed;0 lrad ¼ 1 1e 1þ lg lg 9 8 2 3 > > lg > > > > B = 6 1 pffiffiffiffiffiffiffiffiffiffi< 2 lcat B þ 1 Bþ1 7 1 lcat 6 7   þ 1e ln þ 6 7  2 lg 4 lg 5 lg lg > > lg B 2 lg > > > > 1 B þ :1  lcat B 1 B lcat lrad lcat ; lcat with

  0:23 T 2 lrad ¼ dp 2 100 1 erad   1  e 10=9 C ¼ 1:4 B¼C e

Pex ¼

usup rg Cp;g XF lg

with XF ¼ 1:15 for spherical particles

Porosity Profile When catalytic particles (spheres) are packed in tubes as in the packed bed membrane reactors, a porosity profile occurs which influences the performance of the reactor. The porosity profile in a packed bed of uniform spheres was studied by several research groups as a function of the distance from the wall (Benenati and Brosilow,47 and Schuster and Vortmeyer48). The data presented by Benenati and Brosilow formed the basis of many approaches to develop a correlation for the radial porosity distribution. Ridgway and Tarbuck49 started with the geometric deviation of the porosity function for a bed of regular close-packed spheres at a straight wall, which they subsequently fitted to the data of Benenati and Brosilow introducing three

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

31

correction factors to account for the random character of the packing (two randomizing factors) and the cylindrical form of the bed. Martin50 proposed the following empirical representation: 8 2 > 1z0 < emin þ ð1  emin Þz

p  e ð zÞ ¼ > : e0 þ ðemin  e0 Þez=4 cos z z 0 C where

Rr z

dp=2

and

C¼



0:816 0:876

dt =dp ¼ N dt =dp ¼ 20:3

The minimum porosity is in the range of emin 0.20–0.26 and e0 is the bulk porosity of the bed undisturbed by wall effects. Several authors present approximate porosity functions, which involve only the porosity increase towards the wall without any minima in the profile or including only the first minimum. Vortmeyer and Schuster51 give the following equation for a circular tube:   Rr eðrÞ ¼ e0 þ ð1  e0 Þ exp 2 dp Hunt and Tien13 suggest a similar correlation replacing the constant 2 by 6 for perfect spheres and 8 for slightly irregular or non-uniform particles. In Figure 10.A1(a) the different porosity profiles are compared with the data points of Benenati and Brosilow, while Figure 10.A1(b) shows the resulting dimensionless axial velocity profiles, calculated with the PBMR model with the membrane flux set to zero.

Figure 10.A1

(a) Porosity distribution near the wall in a bed of spherical particles. Data points by Benenati and Brosilow (1962) for dt/dp ¼ 20.3; e0 ¼ 0.38. (b) Axial velocity profiles resulting from the different porosity distributions.52

View Online

32

Chapter 10

Table 10.A1

Comparison of the predictions for the average porosity, pressure drop and maximum axial velocity by the different porosity distribution models (dt/dp ¼ 20.3, ee ¼ 0.38)

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Porosity model Martin Hunt and Tien Vortmeyer and Schuster

hei

0.399 0.409 0.447

DP[Pa]

umax/hui

504 470 215

4.31 4.96 6.99

To evaluate the proposed radial porosity correlations several criteria have to be considered:  Average porosity  Pressure drop  Influence on velocity and concentration profiles. From Figure 10.A1(b) it can be concluded that the porosity function of Vortmeyer and Schuster results in an overestimation of the by-pass flow, and as an consequence thereof the gas velocity in the core of the bed is predicted significantly lower than the velocities resulting from the profile of Martin, opposite to the correlation proposed by Hunt and Tien.13 Furthermore, a comparison of the average porosity and pressure drop (Table 10.A1) clarifies that the correlation of Hunt and Tien is to be preferred and it is the correlation used in the 2D simulations proposed in this chapter.

10.6 Appendix B: Constitutive Equations used in Fluidized Bed Modeling In this appendix the constitutive equations used in the modeling of fluidized bed membrane reactors are listed. Components properties have be evaluated by making use of book The Properties of Gases and Liquids.53 Parameter

Equation

Archimedes number23

Ar ¼

Minimum fluidization velocity54

umf

Bed voidage at minimum fluidization velocity54 Projected tube area for a square bed

emf

 dp3 rg rp  rg g

m2g !qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi  mg ¼ ð27:2Þ2 þ 0:0408Ar  27:2 rg dp !0:021 rg 0:029 ¼ 0:586Ar rp

AT ¼ D2T

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

(Continued ) Parameter

Equation

Rise velocity of a single bubble Velocity of rise of swarm of bubbles23 Initial bubble diameter (porous plate distributor)23

ubr ¼ 0:711ðgdb Þ1=2

Maximum bubble diameter Superficial bubble gas velocity

db;max ¼ DT   usb;max  usb 0:55z ¼ exp hmf DT usb;max  usb;0

Maximum superficial bubble gas velocity

usb;max ¼ u0  umf

Initial superficial bubble gas velocity

usb;0 ¼ ubr;0 db0

 where db0 ¼ 1  hmf hf

Superficial emulsion gas velocity

ub ¼ u0  umf þ 0:711ðgdb Þ1=2  2 db0 ¼ 0:376 u0  umf

use ¼ u0  usb

usb ub den ¼ 1  dbn db ¼

Bubble phase fraction Emulsion phase fraction24

hf Nb h ¼ AT Nfb db;n

Volume of emulsion phase in the n-th compartment

Ve;n ¼ AT

Volume of bubble in the n-th compartment

Vb;n



Bubble diameter

db ¼ db;max  ðdb;max  db;0 Þe

Height of bed expansion

hf ¼ hmf

0:3z DT



C1 C1  C2

where;   ub;0 0:275 exp  DT ub;avg    us 0:275 C2 ¼ b 1  exp  DT ub;avg  1=2 ub;avg ¼ u0  umf þ 0:711 gdb;avg C1 ¼ 1 

Average bubble rise velocity23 Gas exchange coefficient

23

Kbc Kce 1 Kbe

! 1=2 1=4   Dg g umf ¼ 4:5 þ 5:85 5=4 dp db   Dg emf ub 1=2 ¼ 6:77 db3 1 1 ¼ þ Kbc Kce

33

View Online

34

Chapter 10

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

10.7 List of Symbols Symbol

Description

Units

Ar AT Amembrane,n dp Cp D Dg eb ee Ea g Hj T Hi,x

Archimedes number Area of bed cross section Membrane surface area per cell, n Particle diameter Heat capacity Dispersion coefficient Gas diffusivity Bubble phase fraction Emulsion phase fraction Activation energy for hydrogen permeation Gravitational acceleration (=9.81) Enthalpy of species j Enthalpy of component i at temperature T at position x Permeation flux through membrane Reaction rate constant for i-th reaction Bubble-to-cloud phase mass transfer coefficient for component I in cell n Bubble-to-emulsion phase mass transfer coefficient for component i in cell n Cloud-to-emulsion phase mass transfer coefficient for component i in cell n Adsorption constant for CO Equilibrium constant for j-th reaction

— m2 m2 m J (kg K)1 m2 s1 m2 s1 — — J mol1 m s2 J mol1 J mol1

J ki Kbc,i,n Kbe,i,n Kce,i,n KCO Keq,i ks,i Mw hMi Pi Pe0 r R rj R Sh Sr,i t T u us V vj,i Y z a b d r

Interphase mass transfer coefficient Molar mass for component i Average molar mass Partial pressure for component i Pre-exponentional factor for permeation of Pd membrane Radial coordinate Universal gas constant (=8.3145) Reaction rate for j-th reaction Inner tube radius Source/sink term for heat balance Source/sink term for mass balance Time Temperature Mixture velocity Superficial gas velocity Volume Stoichiometric coefficient for j-th reaction and i-th component Mole fraction Axial coordinate Heat transfer coefficient Friction factor Membrane thickness Density

mol (m2 s) — s1

–1

s1 s1 Pa1 (Depending on the reaction) s1 kg mol1 kg mol1 Pa mol (s m2 Pa0.5)1 m J (mol K)1 mol (kgcat s)1 m J (m3 s)1 kg(m3 s)1 s K m s1 m s1 m3 — — m J (m2 K s)1 — m kg m3

View Online

Modeling of Membrane Reactors for Hydrogen Production and Purification

35

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

(Continued ) Symbol

Description

Units

e ee l mg tg W(T)

Porosity Emulsion phase porosity Thermal conductivity Viscosity of gas Stress tensor Correction factor for the Sievert– Langmuir model Molar flux component i through the membrane per cell Weight fraction

— — J(m K s)1 Pa  s kg (m s)1 —

f0membrane i,mol o Subscripts 0 ax b cat e eff g i j n r s z Superscripts b

mol (m2  s)1 —

Reactor inlet Axial Bubble phase Catalyst Emulsion phase Effective Gas phase Component i Number of reaction Number of continuously stirred tank reactors for emulsion or bubble phase Radial co-ordinate Solid phase Axial co-ordinate

— — — — — — — — — — — — —

Bulk condition

—

References 1. S. Hara, G. Barbieri and E. Drioli, Limit conversion of a palladium membrane reactor using counter-current sweep gas on methane steam reforming, Desalination, 2006, 200, 708–709. 2. F. Gallucci, M. De Falco and A. Basile, A simplified method for limit conversion calculation in membrane reactors, Asia-Pacific J. Chem. Eng., 2010, 5, 226–234. 3. G. Marigliano, G. Barbieri and E. Drioli, Equilibrium conversion for a Pdbased membrane reactor. Dependence on the temperature and pressure, Chem. Eng. Proc., 2003, 42, 231–236. 4. G. F. Froment and K. B. Bischoff, Chemical Reaction Analysis and Design, John Wiley, 1979. 5. A. A. Iordanidis, Mathematical Modeling of Catalytic Fixed Bed Reactors, PhD thesis, University of Twente, 2002. 6. J. R. Welty, C. E. Wicks and R. E. Wilson, Fundamentals of Momentum, Heat and Mass Trasfer, John Wiley and Sons, New York, 1969.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

36

Chapter 10

7. G. Barbieri, F. Scura, F. Lentini, G. De Luca and E. Drioli, A novel model equation for the permeation of hydrogen in mixture with carbon monoxide through Pd-Ag membranes, Sep. Purif. Technol., 2008, 61, 217–224. 8. P. V. Danckwerts, Continuous flow systems. Distribution of residence times, Chem. Eng. Sci., 1953, 1, 1–13. 9. T. Numaguchi and K. Kikuchi, Intrinsic kinetics and design simulation in a complex reaction network: steam reforming, Chem. Eng. Sci., 1988, 43, 2295–2301. 10. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, Theoretical comparison of packed bed and fluidized bed membrane reactors for methane reforming, Int. J. Hydrogen Energy, 2010, 35, 7142–7150. 11. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, High performance hydrogen membranes deserve optimal reactor design, NPT, 2010, 2, 14–15. 12. T. P. Tiemersma, C. S. Patil, M. V. Sint Annaland and J. A. M. Kuipers, Modelling of packed bed membrane reactors for autothermal production of ultrapure hydrogen, Chem. Eng. Sci., 2006, 61, 1602–1616. 13. M. L. Hunt and C. L. Tien, Non-darcian flow, heat and mass transfer in catalytic packed-bed reactors, Chem. Eng. Sci., 1990, 45, 55–63. 14. P. Zehner and E. U. Schlu¨nder, Wa¨rmeleitfa¨higkeit von schu¨ttungen bei ma¨ssigen temperaturen, Chem-Ing-Tech., 1970, 42, 41. 15. E. U. Schlu¨nder and E. Tsotsas, Warmeubertragung in Festbetten, Durchmischten Schuttgutern und Wirbelschichten. G.T.Verlag, Stuttgart, 1988. 16. U. Ku¨rten, M. van Sint Annaland and J. A. M. Kuipers, Oxygen distribution in packed-bed membrane reactors for partial oxidations: Effect of the radial porosity profiles on the product selectivity, Ind. Eng. Chem. Res., 2004, 43, 4753–4760. 17. S. A. R. K. Deshmukh, S. Heinrich, L. Mo¨rl, M. van Sint Annaland and J. A. M. Kuipers, Membrane assisted fluidized bed reactors: Potentials and hurdles, Chem. Eng. Sci., 2007, 62, 416–436. 18. M. R. Rahimpour, Enhancement of hydrogen production in a novel fluidized-bed membrane reactor for naphtha reforming, Int. J. Hydrogen Energy, 2009, 34, 2235–2251. 19. C. S. Patil, M. Van Sint Annaland and J. A. M. Kuipers, Design of a novel autothermal membrane-assisted fluidized-bed reactor for the production of ultrapure hydrogen from methane, Ind. Eng. Chem. Res., 2005, 44, 9502–9512. 20. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, Autothermal reforming of methane with integrated CO2 capture in a novel fluidized bed membrane reactor. Part 2: Comparison of reactor configurations, Top. Catal., 2008, 51, 146–157. 21. A. Mahecha-Botero, T. Boyd, A. Gulamhusein, N. Comyn, C. J. Lim, J. R. Grace, Y. Shirasaki and I. Yasuda, Pure hydrogen generation in a fluidized-bed membrane reactor: Experimental findings, Chem. Eng. Sci., 2008, 63, 2752–2762.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

37

22. M. E. E. Abashar and S. S. E. H. Elnashaie, Feeding of oxygen along the height of a circulating fast fluidized bed membrane reactor for efficient production of hydrogen, Chem. Eng. Res. Des., 2007, 85, 1529–1538. 23. D. Kunii and O. Levenspiel, Fluidization Engineering, Series in Chemical Engineering, Butterworth-Heinemann, Newton, 1991. 24. K. Kato and C. Wen, Bubble assemblage model for fluidized bed catalytic reactors, Chem. Eng. Sci., 1969, 24, 1351–1369. 25. S. A. R. K. Deshmukh, J. A. Laverman, A. H. G. Cents, M. Van Sint Annaland and J. A. M. Kuipers, Development of a membrane assisted fluidized bed reactor. 1. Gas phase back-mixing and bubble-to-emulsion phase mass transfer using tracer injection and ultrasound experiments, Ind. Eng. Chem. Res., 2005, 44, 5955–5965. 26. S. A. R. K. Deshmukh, J. A. Laverman, M. Van Sint Annaland and J. A. M. Kuipers, Development of a membrane assisted fluidized bed reactor. 2. Experimental demonstration and modeling for the partial oxidation of methanol, Ind. Eng. Chem. Res., 44, 5966–5976. 27. S. A. R. K. Deshmukh, M. Van Sint Annaland and J. A. M. Kuipers, Heat transfer in a membrane assisted fluidised bed with immersed horizontal tubes, Int. J. Chem. React. Eng, 2005, 3, A1. 28. C. S. Patil, M. Van Sint Annaland and J. A. M. Kuipers, Experimental study of a membrane assisted fluidized bed reactor for H2 production by steam reforming of CH2, Chem. Eng. Res. Des., 2006, 84, 399–404. 29. A. M. Adris, C. J. Lim and J. R. Grace, The fluidized bed membrane reactor for steam methane reforming: model verification and parametric study, Chem. Eng. Sci., 1997, 52, 1609–1622. 30. Z. Chen, Y. Yan and S. S. E. H. Elnashaie, Modeling and optimization of a novel membrane reformer for higher hydrocarbons, AIChE J., 2003, 49, 1250–1265. 31. K. Johnsen, H. J. Ryu, J. R. Grace and C. J. Lim, Sorption-enhanced steamre forming of methane in a fluidized bed reactor with dolomite as CO2-acceptor, Chem. Eng. Sci., 2006, 61, 1195–1202. 32. I. A. Abba, J. R. Grace and H. T. Bi, Application of the generic fluidizedbed reactor model to the fluidized-bed membrane reactor process for steam methane reforming with oxygen input, Ind. Eng. Chem. Res., 2003, 42, 2736–2745. 33. A. Mahecha-Botero, J. R. Grace, C. Jim Lim, S. S. E. H. Elnashaie, T. Boyd and A. Gulamhusein, Pure hydrogen generation in a fluidized bed membrane reactor: Application of the generalized comprehensive reactor model, Chem. Eng. Sci., 2009, 64, 3826–3846. 34. J. F. De Jong, H. J. van Gerner, M. van Sint Annaland and J. A. M. Kuipers, Development of a Novel Hybrid Discrete Particle – Immersed Boundary Model for Fluidized Bed Membrane Reactors, Seventh International Conference on CFD in the Minerals and Process Industries. CSIRO, Melbourne, 2009. 35. http://fcre.tnw.utwente.nl/cfd.php?menuid¼2

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

38

Chapter 10

36. J. A. M. Kuipers, K. J. van Duin, F. P. H. van Beckum and W. P. M. van Swaaij, A numerical model of gas-fluidized beds, Chem. Eng. Sci., 1992, 47, 1910. 37. D. Gidaspow, Multiphase Flow and Fluidization Continuum and Kinetic Theory Descriptions, Academic Press, Boston, 1994. 38. G. A. Bokkers, J. A. Laverman, M. van Sint Annaland and J. A. M. Kuipers, Modelling of large-scale dense gas–solid bubbling fluidised beds using a novel discrete bubble model, Chem. Eng. Sci., 2006, 61, 5590–5602. 39. M. A. van der Hoef, M. Ye, M. van Sint Annaland, A. T. Andrews IV, S. Sundaresan and J. A. M. Kuipers, Multi-scale modelling of gas-fluidized beds, Adv. Chem. Eng., 2006, 31, 65. 40. M. A. van der Hoef, M. van Sint Annaland, N. G. Deen and J. A. M. Kuipers, Numerical simulation of dense gas-solid fluidized beds: A multiscale modeling strategy, Annu. Rev. Fluid Mech., 2008, 40, 47. 41. M. Ye, Multi-level modeling of dense gas-solid two-phase flows, PhD dissertation, University of Twente, 2005. 42. N. G. Deen, M. van Sint Annaland, M. A. van der Hoef and J. A. M. Kuipers, Review of discrete particle modeling of fluidized beds, Chem. Eng. Sci., 62, 28–44. 43. D. J. Patil, M. van Sint Annaland and J. A. M. Kuipers, Critical comparison of hydrodynamic models for gas-solid fluidized beds - Part I: Bubbling gas-solid fluidized beds operated with a jet, Chem. Eng. Sci., 2005, 60, 57–72. 44. M. van Sint Annaland, G. A. Bokkers, M. J. V. Goldschmidt, O. O. Olaofe, M. A. van der Hoef and J. A. M. Kuipers, Development of a multi-fluid model for poly-disperse dense gas–solid fluidised beds, Part II: Segregation in binary particle mixtures, Chem. Eng. Sci, 2009, 64, 4237–4246. 45. J. Westerweel, Fundamentals of digital particle image velocimetry, Meas. Sci. Technol., 1997, 8, 1379. 46. J. M. Link, C. Zeilstra, N. G. Deen and J. A. M. Kuipers, Validation of a discrete particle model in a 2D spout-fluid bed using non-intrusive optical measuring techniques, Can. J. Chem. Eng., 2004, 82, 30. 47. R. F. Benenati and C. B. Brosilow, Void fraction Distribution in Beds of Spheres, AIChE J., 1962, 8, 359. 48. J. Schuster and D. Vortmeyer, Ein einfaches Verfahren zur na¨herungsweisen Bestimmung der Porosita¨t in Schu¨ttungen als Funktion des Wandabstandes, Chem.-Ing.-Tech., 1980, 52, 848. 49. K. Ridgway and K. J. Tarbuck, Voidage fluctuations in randomly-packed beds of spheres adjacent to a containing wall, Chem. Eng. Sci., 1968, 23, 1147–1155. 50. H. Martin, Low Peclet number particle-to-fluid heat and mass transfer in packed beds, Chem. Eng. Sci., 1977, 33, 913.

View Online

Downloaded on 30 March 2012 Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00001

Modeling of Membrane Reactors for Hydrogen Production and Purification

39

51. D. Vortmeyer and J. Schuster, Evaluation of steady flow profiles in rectangular packed beds by a variational method, Chem. Eng. Sci., 1983, 38, 1691. 52. U. Ku¨rten, Modeling of packed bed membrane reactors: Impact of oxygen distribution on conversion and selectivity in partial oxidation systems, PhD dissertation, University of Twente, 2003. 53. R. C. Reid, J. M. Prausnitz and B. E. Poling, The Properties of Gases and Liquids, McGraw-Hill, New York, 1988. 54. C.-Y. Shiau and C.-J. Lin, An improved bubble assemblage model for fluidized-bed catalytic reactors, Chem. Eng. Sci., 1993, 48, 1299–1308.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

CHAPTER 11

Palladium-based Membranes in Hydrogen Production RUNE BREDESEN,a THIJS A. PETERS,a MARIT STANGE,a NICLA VICINANZAb AND HILDE J. VENVIKb a

SINTEF Materials and Chemistry, P.O. Box 124, Blindern, N-0314 Oslo, Norway; b Department of Chemical Engineering, Norwegian University of Science and Technology, N-7491 Trondheim, Norway

11.1 Introduction Hydrogen is one of the most important chemicals used in industry today. In addition, hydrogen is prospected as an energy carrier for the future, since the energy contained in the molecule can be efficiently converted to electric energy in fuel cells. Different requirements to hydrogen purity or hydrogen gas mixture composition exist depending on the application, therefore different production and separation technologies may be applied. Membranes with high hydrogen permeation and selectivity are identified as a promising technology for efficiency improvement and cost reduction for hydrogen production. Palladium and certain Pd-alloy compositions are known to selectively absorb and diffuse hydrogen in the solid. Therefore, membranes made from these materials give 100% separation selectivity, and thus, direct production of high purity hydrogen provided a defect-free membrane. The large number of experimental and modeling activities presented in this chapter show that Pd-based membrane reactors can convert many chemicals more efficiently to hydrogen and other products than conventional reactors. Thus, a legitimate question is: Why is the technology not in widespread use by Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

40

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

41

industry already? The main reason for this is that Pd-based membrane technology was regarded as too expensive until recent years due to the cost of the noble metal. But, particularly in the two last decades, concerns about the environment and new governmental policies have driven research on hydrogen as an energy carrier for greenhouse gas mitigation and reduction of local air pollution. The great potential of the Pd-based membrane technology, manifested by the economic support from public and industrial sectors, can be read from the rapidly increasing number of publications on the topic in international scientific journals.1 We will show that this research effort has resulted in great progress in the technology, which will result in significant cost reduction. This is mainly due to the development of composite membranes consisting of thin Pd or Pd-based layers (o10 mm) on mechanically strong porous supports. Also progress in sealing technology, module design and understanding of operational issues are important factors supporting the likelihood of successful commercialization. The chapter will show several companies’ up-scaling efforts encouraged by both the technical progress and economical potential. Successful implementation of these efforts will place Pd-based membrane reactor technology in the forefront of promising key enabling technologies for improving energy efficiency and reducing the environmental impact of human activities. Many excellent reviews on Pd-based membranes, membrane reactors and applications exist.1–9 Our intention is to give the reader insight about the status and to pin-point some trends and main challenges related to Pd-based membranes for hydrogen production.

11.2 Conventional Hydrogen Production and Applications 11.2.1

Hydrogen Production

The most cost and energy efficient way to produce hydrogen in large quantities is by conversion of natural gas. Depending on scale and the application of the product different combinations of reactions, reactor types and process schemes are applied.10–13 Similarities also exist between this technology, and the generation of hydrogen from gaseous, liquid and solid carbon-containing fuels. Table 11.1 gives an overview over the main reactions for hydrogen production from fossil and renewable fuels. Hydrogen is produced from natural gas by the endothermic steam methane reforming (SMR) reaction (eqn (11.1)) at high temperature. Heat is conventionally supplied by burning methane externally at the reactor wall. The steam reforming reaction may be completely or partly replaced by the exothermic partial oxidation of methane (eqn (11.2)), where CO and H2 are thermodynamically favored at high temperature. Whereas steam reforming (eqn 11.1)) always is a heterogeneously catalyzed reaction, the partial oxidation (eqn (11.2)) may or may not be catalytic. The reaction proceeds by first

View Online

42

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Table 11.1

Chapter 11

Main chemical reactions the production of hydrogen from natural gas, liquid fuels, coal and biomass

Reaction

DH0298 (kJ  mol–1)

Equation number

CH4 þ H2 O $ CO þ 3H2 2CH4 þ O2 ! 2CO þ 4H2 CH4 þ 2O2 ! 2CO2 þ 2H2 O CH4 þ CO2 $ 2CO þ 2H2 CO þ H2 O ! H2 þ CO2 Cx Hy Oz þ nO2 þ mH2 O ¼ CO þ H2 C2 H5 OH þ 3H2 O $ 2CO2 þ 6H2 C þ H2 O $ CO þ H2 2C þ O2 $ 2CO

206 –36 –803 247 –41 — 347.4 131 –283

(11.1) (11.2) (11.3) (11.4) (11.5) (11.6) (11.7) (11.8) (11.9)

consuming most of the oxygen in fast, complete oxidation reactions such as methane combustion (eqn (11.3)), followed by endothermic steam (eqn (11.1)) and dry reforming (eqn (11.4)) reactions downstream. In autothermal reforming (ATR), steam and oxygen are both added and the oxygen is consumed in a homogeneous zone followed by a heterogeneous (catalytic) reforming section. Under reforming conditions (700–1100 1C), the weakly exothermic water gas shift (WGS) reaction is also relatively fast and usually at equilibrium (eqn (11.5)). The product effluent is hence always a mixture of hydrogen, steam, carbon monoxide and carbon dioxide, while nitrogen and other inerts (Ar) may be present. Reactions (11.1)–(11.4) are all thermodynamically favored by low pressure, but due to investment cost (i.e. reactor size) and further processing of the product mixture, elevated pressure (20–50 bar) is often preferred. The pressure increase must be compensated by increase in temperature to maintain complete methane conversion. The technology is often referred to as synthesis gas (syngas) technology, for which the main principles are the same whether the product mixture (H2/CO/CO2) is to be applied in the methanol or Fischer– Tropsch syntheses, or if hydrogen production is targeted for ammonia production or energy applications. The choice of process lay-out, reactor type and auxiliary process units will vary with the application of the syngas, as well as with the scale. Autothermal reforming is, for example, preferred in very large scale, with oxygen supplied from an air separation unit (ASU), while the cost of air separation becomes less favorable as the scale decreases. In general, any liquid hydrocarbon or oxygenate may be subjected to similar schemes, i.e. reactions at high temperature with steam and/or oxygen, to produce H2 and CO (eqn (11.6)).10,12,13 The oxygen needs to be restricted in order to limit the complete oxidation reactions, and the oxygen/steam ratio will determine the overall exo-/endothermicity. The oxygen and steam levels also affect the potential side reactions that lead to the formation of solid carbon. Since this can cause severe reactor problems (plugging, hot spots), there is a lower limit to the so-called steam-to-carbon (S/C) ratio. If only oxygen is fed, the balance between carbon formation and complete oxidation is a delicate one. Propane,14,15 methanol, ethanol,16,17 glycerol,18 gasoline and diesel19,20 are examples of (liquid) energy carriers that have been studied for hydrogen

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

43

production by (eqn (11.6)), whereas heavy hydrocarbons have been subjected to gas phase partial oxidations to produce syngas. The carbon formation potential generally increases with the C-chain length as well as with the number of C–C double and triple bonds. Synthesis gas may also be produced by gasification of coal or biomass,12,13,21–26 enabling production of hydrogen from these important sources. Although the biomass will contain both H and O atoms, the reactions may be represented by reactions with carbon (eqns (11.8) and (11.9)). Downstream processing of gasification effluents, i.e. separation or further reaction, is however more challenging than in the case of natural gas because of the wide range of possible impurities and by-products associated with the coal or biomass. Common impurities include sulfur and nitrogen compounds, tars, various metals, and chlorine, in concentrations varying with the source of the feedstock and the gasification technology applied. To optimize for hydrogen production and minimize (unwanted) CO, the forward WGS reaction (eqn 11.5)) needs to be promoted. Conventionally, e.g. hydrogen production for ammonia synthesis, this is done through stepwise reductions in temperature while passing the mixture over two or more catalytic reactors downstream. This will usually bring the concentration of CO down to 1–3%. Alternatively, the reforming (eqns (11.1) and (11.4)) and WGS equilibriums can be manipulated through continuously removing H2 and/or CO2 from the steam reforming/partial oxidation by a membrane or a sorbent. The use of Pd-based membranes for this purpose is well covered in the next sections, but other hydrogen-selective membranes may also be considered.2–4,8 For the purpose of H2 production only, the use of a H2-sorbent has to our knowledge not been reported, but the concept has been patented.27 And whereas current CO2 selective membranes can not work under reaction conditions (T,P),28 or are currently very immature,29–31 several research groups32–35 have investigated the use of inorganic CO2 sorbents and proved that until saturation of the sorbent, complete conversion of the fuel can be obtained jointly with very low levels of effluent CO and CO2. This technology, however, needs further research in terms of sorbent stability and development of reactors and process lay-out for the complete sorption-regeneration scheme. Product extraction by sorption or membrane separation may alternatively enable more advantageous process conditions, such as lower temperature or higher pressure. Much work in this field has been undertaken by Air Products and collaborators to produce H2 for power generation with CO2 capture.34 Further processing may include separations as well as reactions, depending on scale and application. H2O may be condensed, and CO2 may be extracted by liquid absorption. Separation to give high hydrogen purity (Z99.999%) can also be done in a pressure swing adsorption unit. If removal of trace CO (and CO2) is required, reaction with H2 (methanation) or O2 (preferential oxidation) can be applied. In general, purification by CO2 liquid absorption and pressure swing, or catalytic reaction steps to remove trace levels of CO add cost, complexity and reduces the energy efficiency, particularly in the small scale. The use of hydrogen selective membranes in separation units or membrane reactors is

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

44

Chapter 11

therefore to a large degree pursued to obtain simplified, compact and efficient hydrogen production. In addition, a pressurized stream enriched in CO2 can be targeted. It could also be mentioned that efficient oxygen separation technology that carries reasonable cost in a wide range of scale is another ‘enabling’ technology. Application of oxygen membrane technology for air separation has been investigated by many researchers,36 and Air Products has announced a plan for commercialization of this technology.37 Many other hydrogen production methods exist. With the exception of electrolysis, these are all at R&D level. Direct water splitting occurs at very high temperatures (water spontaneously dissociates at around 2500 1C) and is not used as such for hydrogen production. With electrolysis, electro-catalytic water splitting can be done well below 100 1C for production of hydrogen and oxygen. This technology which directly incorporates H2/O2 separation, is expensive and consumes electric energy at, so far, inacceptable energy efficiencies for application in energy systems. It is therefore not further covered here. Hydrogen may also be produced as a by-product during the production of chlorine by electrolysis from brine and in dehydrogenation of hydrocarbons. Photocatalytic water splitting and microbiological hydrogen production (algae, bacteria) are other options being heavily researched with a long-term perspective.

11.2.2

Hydrogen Application

The largest application of H2 is for the processing of fossil fuels in the petroleum and chemical industry. Here, most of the hydrogen produced worldwide is directly converted in a downstream process, such as the production of ammonia and fertilizers, or refinery processes that require hydrogen, e.g. hydrogenation and desulfurization processes. In addition, hydrogen has been regarded during the last decades as a possible future energy carrier in the socalled hydrogen economy. An important application is fuel for fuel cell vehicles and small combined heat and power systems, which could reduce the small distributed emissions of CO2 generated by the fossil fuel based alternatives. Future large scale hydrogen production with CO2 capture and sequestration (CCS) may also be combined with heat and power generation in pre-combustion decarbonization processes (PCDC). As mentioned, large scale hydrogen production from natural gas is currently preferred from a cost and efficiency perspective. The demand for hydrogen production technology for smaller scale is growing, however, driven by the development of fuel cell technology on one hand and the distributed nature of renewable sources (biomass/biogas) on the other hand. Hydrogen vehicle filling stations, electricity distribution sub-station, distributed ammonia production, low emission auxiliary power units running on diesel, and computers and mobile phones running on methanol are examples of the wide range of scale. Reduced cost of reforming technology through innovation and technological progress are needed to avoid that the hydrogen produced becomes prohibitively expensive.

View Online

Palladium-based Membranes in Hydrogen Production

45

11.3 Development of Palladium-based Membranes and Stability Issues in Hydrogen Production Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

11.3.1

Membrane Development

Commercial H2 selective Pd-based membranes are available in the form of relatively thick (20 mm or more) tubes or foils manufactured by cold-working techniques. The H2 flux, being in many cases inversely proportional to the thickness of the membranes, is too low for most applications to give a favorable cost-performance combination. Thus, development of membranes with reduced Pd-alloy layer thickness is necessary. Research in recent years has therefore focused on the development of composite membranes consisting of a thin Pd-based separation layer on a mechanically strong support. The typical stateof-the-art membrane consists of a separation layer of less than about 10 mm thickness on a ceramic or metallic support. Examples of commercial development of composite membranes are given in Section 11.6.

11.3.2

Membrane Fabrication Methods

A large number of different membrane fabrication methods have been applied including chemical vapor deposition,38–41 physical vapor deposition,42 such as sputtering,43–48 but the most common is electroless plating.49–52 The latter method has low cost and is applicable for large-scale production, but becomes increasingly more complex with the number of alloying elements. In the case of Pd-Au membranes, for example, at the temperature required for alloying (550 1C), the composite membrane integrity is found to slowly decline.53 This is an important limitation in the membrane production as more advanced ternary or quaternary alloys may be needed to improve performance and stability. The less common sputtering method enables production of complex alloys with controlled stoichiometry in an easily controlled manner,45,54 but the method is currently less developed for large-scale membrane production. Generally, for these fabrication methods, the thin Pd or Pd-alloy layer is prepared directly on the surface or inside the pores of the support.55 Different materials such as ceramics, glass and metals, have been employed as porous supports. For metallic supports and high operation temperatures (>400 1C) a barrier layer made from porous ceramics, like zirconia, zeolite Sil-1, yttriastabilized zirconia, titania and silica, or TiN56–64 to limit interdiffusion of metals between support and separation layer is needed. Moreover, Ma and co-workers developed a controlled in situ oxidation of the porous stainless steel (PSS) prior to plating in order to produce an oxide layer acting as a diffusion barrier between Pd and the PSS support.65 Several membranes prepared by this technique have been stable under hydrogen permeation conditions for over 1400 h at temperatures exceeding 500 1C.66 The use of ceramic supports as, e.g. pure alumina is not believed to cause significant inter-diffusion problems. A recent study, though, has reported intrusion of aluminium from the alumina support into the palladium layer at

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

46

Chapter 11

temperatures of 650 1C resulting in a significant decrease in the hydrogen permeability.67 Concerning the preparation of thin membranes directly on porous supports, a lower thickness limit seemingly exists for which a dense metal layer can be obtained. This thickness limit increases with increasing surface roughness and pore size in the support’s top layer.44,68 Clearly, this relation puts strong demands on the support quality in terms of narrow pore size distribution, and the amount of surface defects.69,70 Therefore both pore size and roughness of the support surface are often reduced by the application of meso-porous intermediate layers prior to deposition of the permselective metal layer.71 This procedure facilitates the preparation of thin defect-free membranes because it is relatively easier to cover small pores by filling them with metal.69 It is therefore conceivable that for a certain low Pd-alloy thickness and support pore size, the H2 flux becomes limited by the support resistance.2 A method to circumvent support quality problems is to fill a hydroxide gel in the substrate pores and defects, preventing the formation of pinholes during subsequent electroless plating of the palladium layer.72 The increase in mass transfer limitations is claimed to be limited because the volume of the hydroxide decreases greatly after thermal treatment.73 Additionally, this method prevents infiltration of palladium into the substrate. Another option to circumvent strong demands on support quality is to simply seal defects present after membrane manufacturing by selective deposition of Pd within them. Successfully sealing of defects in Pd and Pd-Ag membranes by directed electroless plating has been demonstrated.74 This was achieved by feeding the metal source and the reducing agent from opposite directions to the defect zone.74 Also an osmotic concentration gradient75 has been used to direct the deposition to open voids. Membranes prepared by the former point plating method practically maintain their original high H2 permeability, which suggests that the metal layer thickness did not increase and metal deposition was essentially restricted to defect sites. Point plating is also suited for restoring the selectivity of spent Pd membranes, which could lower the operation costs of technical membrane separators substantially.74 A two-step membrane manufacturing process76 has been reported where a defect free Pd-alloy membrane is first prepared by sputtering deposition onto the ‘perfect surface’ of a silicon wafer, for example. In a second step the membrane is removed from the wafer and transferred to a porous stainless steel support (see Figure 11.1). This allows the preparation of very thin (B1–2 mm) defect-free membranes supported on macroporous substrates (pore size equals B2 mm). By this technique, the ratio of the membrane thickness over the pore size of the support may become less than 1, which is two orders of magnitude smaller than obtained by more conventional membrane preparation techniques.2 Tubular-supported palladium membranes prepared by the two-step method77 show a H2/N2 permselectivity equal to B2600 at 26 bars and hydrogen flux of 2477 mL(STP) min1 cm2. Since the method enables the combination of macro-porous stainless steel supports and thin membrane layers, the support resistance is negligible.77,78

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

47

Figure 11.1 Two-step membrane manufacturing method developed at SINTEF (a) Pd-23%Ag film prior to removal from the silicon wafer; (b) unsupported Pd-23%Ag film; (c) Pd-23%Ag film during wrapping on the tubular porous stainless steel (PSSs) support; (d) Four long (50 cm) Pd-23%Ag/ stainless steel composite membranes.77

Energy Centre of the Netherlands (ECN) has published a report79 that compares technologies for production of thin Pd-alloy films onto porous tubes on an industrial scale. Both chemical and physical industrial film deposition technologies were considered. According to the requirements, the up-scaled technology should be applicable to porous tubes, give defect-free membrane films with homogeneous thickness on a large area, and give a controlled formation of a Pd-film with one or two alloying elements. In addition, microstructure, hydrogen permeance, thermal cycling behavior and robustness of the Pd-alloy membranes produced by different methods were compared. The conclusion from this study was that the vacuum (low pressure) techniques in general, and the sputtering technique in particular, were the most versatile techniques for the direct deposition of binary, ternary, or multi-component Pdalloy films on a porous support. Technologies such as atomic layer deposition (ALD), chemical vapor deposition (CVD) and electroless deposition were less favorable due to very low deposition rate, highly toxic precursors and difficulties in controlling the alloy composition and thickness homogeneity, respectively.

11.3.3

Palladium-alloys and their Implications for Membrane Stability

In pure Pd, an a-to-b hydride phase transition may occur in hydrogen below about 290 1C,80 and only a few cycles through this transition makes the material brittle and must be avoided. By alloying Pd with different elements, the phase

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

48

Chapter 11

transition can be suppressed, and in particular alloys with 20–30 wt% Ag and 40 wt% Cu have been frequently investigated as membranes.69,81–85 The Ag alloys have higher hydrogen permeability than pure Pd (up to about 70%), while the Cu alloy has about the same.81,84 Palladium has also been alloyed with other metals, including Ru, Au, Y, Ni, Fe to obtain increased permeability, greater mechanical strength, inhibition of undesired grain growth, or enhancement of the resistance against poisoning of CO or sulfur-containing species.86–92 The majority of the work related to Pd-based membranes is on the highly permeable Pd-Ag alloy. The drawback is, however, that these alloys are prone to poisoning by CO and sulfur-containing gases69,84,93–95 leading to reduced permeability, or even to complete deterioration of the membrane.94,96,97 The extent to which this affects stability and flux depends on many parameters, and the reported literature appears only partly consistent in this respect. The latter is probably mainly due to the insufficient understanding of the reactions mechanisms and how the surface properties, which is usually not well described on the atomic level, varies with synthesis methods, thermal history, surface composition, morphology, etc. CO reduces the flux due to competitive adsorption on the membrane surface, thereby displacing hydrogen. As an adsorptive effect, its importance increases with decreasing temperature and membrane thickness. Encouraging results have been reported for Pd-23%Ag membranes where the inhibiting effect of CO was compared before and after thermal treatment in air.98 In absolute numbers, the flux in presence of the same amount of CO was approximately two times higher after air treatment than before.98 It has been suggested that the air treatment is merely a cleaning effect that removes much of adsorbed surface contamination,99,100 but it is difficult to apprehend how this explains the observed relative flux improvement in the presence of CO. Other research has pointed to more complex effects of air treatment on flux related to microstructural changes and changes in surface compositions by segregation effects.101 Since the relative flux reduction appears to depend on the surface composition and morphology, research to optimize these parameters may lead to membranes that show a more robust performance in the presence of CO. A similar effect has been observed for many other hydrocarbons, i.e. adsorption and subsequent hydrogen flux reduction. Post-inspection of the surface often reveals carbonaceous surface contamination, like carbon,102 CH4103,104 or propylene.105 Secondary effects of CO adsorption is the catalytic decomposition and formation of carbon, carbide (Pd1–xCx)102,106 or carbonate phases.107 While reversible CO adsorption decreases with increasing temperature, catalytic decomposition appears to increase with temperature.102 The tendency of deposit forming due to catalytic decomposition is reported to be counteracted by the presence of steam108–110 mitigating the flux reduction. Production of hydrogen from coal and other sulfur-containing sources is challenging for Pd-based membranes due to their limited stability towards sulfur. Even a few ppm of sulfur in the gas reduces the flux drastically due to strong surface adsorption, and may further lead to degradation of the material

View Online

Palladium-based Membranes in Hydrogen Production

49

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

96,97

caused by formation of bulk Pd4S. Membranes consisting of alloys of Pd-Cu and Pd-Au have received attention69,84,89,96,111–113 because these alloy systems offer some sulfur resistance. Hydrogen permeation results for a Pd-Au membrane show that this alloy exhibited better resistance to bulk sulfidation as compared to the Pd-Cu alloy system.96,112 Although the hydrogen flux was reduced up to 80% in the presence of H2S, the permeance could nearly be fully recovered after a temperature increase, indicating that the poisoning of the Pd-Au alloy membrane was caused only by reversible surface site blocking. The absolute value of the permeability for these alloys, however, are low compared to that found in, for example, Pd-23%Ag in the absence of H2S, meaning that the flux penalty due to H2S is significant in absolute terms. Recent publications by Sholl et al. provide guidelines, based on density functional theory calculations, to identify ternary alloys which retain the favorable surface chemistry of Pd-Cu or Pd-Au binary alloys but are predicted to yield higher H2 fluxes.114,115 Experimental data on chemically more robust ternary alloys, however, has so far been limited by the large resources needed to test multiple materials. Coating as a protection has also been suggested and demonstrated.116,117 Recently, it was reported that flux reduction in the presence of H2S could not be observed in a Pd-based membrane having a surface protective film.117 The coating route, probably in combination with alloy optimization, holds promise for the use of Pd-based membranes in gas mixtures containing reactive components.

11.3.4

Structural Stability of Composite Palladium-based Membranes

Although low membrane thickness is important to increase the H2 flux, the need for a thin Pd layer challenges the stability of the whole composite membrane structure. Instability of the interface between the membrane layer and the support is mainly caused by two effects: a badly matched expansion coefficient between the support material and membrane layer, and interfacial diffusion between the support and membrane layer causing the membrane layer to fail. The latter issue was previously discussed. Stress generated due to different thermal expansion coefficients, as well as the volume change due to hydrogen solubility, may cause detachment, cracking and wrinkling of the membrane layer. The interfacial stress also increases with membrane thickness.118 Seemingly, fabrication methods where the membrane partly or fully forms in the pores give some increased stability.55 The thermal expansion coefficient of stainless steel is close to that of palladium, which reduces problems with the thermal expansion coefficient mismatch. Guazzone et al. concluded after an extensive analysis of microstrain and stress in electroless deposited thin Pd films that the magnitude of stress only played a minor role in leak formation using porous stainless steel as support.119 Leak formation and growth, however, seem to be related to inhomogeneous sintering. It has been shown the formation of pinholes is initiated above a certain temperature120,121 which appears to depend on the membrane thickness.120

View Online

50

Chapter 11

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

11.4 Integration of Palladium-based Membranes in Hydrogen Production A number of studies have verified the beneficial potential of Pd-based membrane reactors for hydrogen production by investigating critical parameters such as pressure, temperature, catalyst, gas composition and reactor design on the overall performance, as will be discussed below. As the thermodynamics, kinetics and by-product formation vary with the type of fuel/feedstock applied, both the Pd-based membrane and operating conditions need to be tailored in order to guarantee a cost-effective process and sufficient membrane stability.

11.4.1 11.4.1.1

Methane Reforming Steam Reforming of Methane

As mentioned, steam reforming of natural gas requires high temperature (>850 1C), because it is highly endothermic and equilibrium limited. Hydrogen extraction through a membrane may therefore shift the equilibrium to obtain complete conversion and higher yields at lower temperature (or higher pressure). Pd-based membranes also offer the additional advantage of high selectivity. Pure hydrogen (>99.9%) can be produced without further downstream purification by pressure swing adsorption or similar, which is economically favorable compared to conventional technology.122 Table 11.2 shows experimentally reported results for methane steam reforming with Pd-based membrane reactors. Temperatures lower than B500 1C significantly reduces the H2 partial pressure in the reaction zone, and most studies are therefore performed at 500 1C or above. Reports suggest, however, that for membrane temperatures higher than about 550 1C, the thickness should not be below B10 mm since morphological changes generating membrane defects accelerate.120,121 This challenges the economics, and should guide further research on thin membranes in the direction of membrane stability and process optimisation combination. An important factor for the operation is the catalyst activity and stability. The catalyst activity and/or loading must be sufficiently high to sustain chemical equilibrium adjacent to the membrane. Tong and Matsumura129 compared two Ni-reforming catalysts and demonstrated the importance of having sufficient activity throughout the catalyst bed, i.e. under changing gas composition. Steam reformers are normally not limited by the rate of reaction, but rather the rate of heat transfer to the reaction zone. Ni, as well as Ru or other precious metal-based catalysts, should hence be able to match the membrane transport provided efficient heat transfer. In addition, the catalyst must be sufficiently stable and selective under the conditions created by the hydrogen extraction. Reactions leading to the formation of carbon species (‘coke’) on the catalyst surface, and possibly also on reactor and membrane materials, may be favored by the reduced hydrogen partial pressure.130 Again, the above-mentioned alternatives to Ni may offer some advantages, but the expensive metals will also add significantly to the cost.

P (bar)

3 9 20 3 1

1 9

9

3

450 500 500 500 500

500 500

550

550

3

3

3 3

2 3 3 3 3

S/C (mol  mol1)

6

4

Pd/Al2O3

Pd/MPSS

20 4

50 20 20 6 13 90 88.7

Ni/Al2O3 Ni-La/ Mg-Al Ni-La/ Mg-Al Ni/Al2O3 100

98.8

50 40 85 98 85

CH4 conversion (%)

Ni/Al2O3 Ni-based Ni-based Ni/Al2O3 —

Thickness (mm) Catalyst

Pd23wt%Ag Pd/SS Pd/SS Pd/MPSS Pd/CeO2/ MPSS Pd/PG Pd/Al2O3

Membrane

—

95

— 91.8

70 30 90 — —

H2 recovery (%)

6

11.9

o2.1 10.3

3.07 1.18 2.5 2.6 10.25

H2 flux (N m3  m2  h1)

Experimentally reported results for methane steam reforming with Pd-based membrane reactors

T (1C)

Table 11.2

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

125

127, 128

126 127

123 124 124 125 60

Ref.

View Online

Palladium-based Membranes in Hydrogen Production 51

View Online

52

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

11.4.1.2

Chapter 11

Autothermal Reforming and Combined Reforming of Methane

Autothermal reforming may lead to high temperature and possible hot spots in the first part of the catalyst bed since the exothermic oxidation reactions are faster than the reforming reaction.131 The process therefore needs careful control of the operation parameters. High conversion at fairly low temperature is reported for autothermal reforming, e.g. 95% CH4 conversion at 470 1C and 7 bar which was compared to 37% for a traditional reactor. The 24 mm thick Pd-Ag membrane used132 showed hydrogen flux of 0.093 mol m2 s1.

11.4.1.3

Dry Reforming of Methane

The dry (CO2) reforming of methane (eqn (11.4)) gives a lower H2/CO2 ratio than steam reforming and is more endothermic.130,133–135 The equilibria (eqns (11.4) and (11.5)) can be favored towards product formation by applying a hydrogen selective membrane. The coke formation potential, however, is greater than in steam reforming, and likely to be enhanced by the hydrogen extraction.130 Ni-based reforming catalysts on supports with high oxygen mobility have, nevertheless, been reported to suppress the carbon deposition.130 Ru catalysts demonstrated higher activity than Ni-based systems in dry reforming.133 Recently, using an improved Ru catalyst and oxygen co-feeding, Mu´nera et al.136 showed that the catalyst at 550 1C could sustain the equilibrium in the membrane reactor for the permeate side sweep flow range employed. A commercial Pd-Ag membrane was applied, but no membrane thickness was given.

11.4.1.4

Membrane Reactor Design in Methane Reforming

Optimisation of the reactor design is important to avoid mass and heat transfer limitations. As mentioned, the reforming reaction requires efficient heat exchange to avoid temperature decrease and gradients. An estimation of steam reforming in a packed-bed tubular membrane reformer (Figure 11.2), operating at 600 1C for production of 2.4 million Nm3 H2 day1, shows that for a membrane permeance of 0.4 mol m2 s1 bar1 the necessary membrane surface area corresponds fairly well with the required heat exchange area.122

Figure 11.2

Schematic drawing of a packed bed membrane reactor, after Patil.137

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

Figure 11.3

53

Scheme of the permeative-stage membrane reactor (PSMR) considered, after Caravella et al.140

For conventional reformers operating at high temperature, the heat is provided by burners directly to the tube walls. The lower operation temperature of membrane reactors allows other solutions as heated gas or flameless heat distributors to be used.138,139 Caravella et al.140 suggested a so-called permeativestage membrane reactor in steam reforming of methane, see Figure 11.3. The reactor consists of several subsequent reactive sections (containing the catalyst) and permeative sections (containing the membrane). Simulation of catalyst distribution and process parameters showed possibility of 21% improvement in hydrogen recovery compared to a conventional packed bed membrane reactor with the same catalyst volume and membrane area. Membrane Reactor Technologies Ltd (MRT) has experimentally verified the permeative-stage membrane reactor concept.141 With the membranes outside the reactor, operation at more favorable conditions for both reaction (750 1C) and membrane separation (450 1C or lower) is possible. A decrease in the metal cost of palladium-based membranes by 86.5% and membrane area by 470% to achieve equal hydrogen production capacity was reported. The volume of reformer decreases accordingly, thus, the costs of both the reactor and membrane module are reduced. As mentioned the entrance temperature in autothermal reformers can be much higher than further down the packed bed reactor. To avoid damage of the membrane an alternative could be to increase the separation between the homogeneous oxidation reaction and the downstream heterogeneous reforming reaction, possibly in two unit operations142,143 and allow the temperature of the latter6,144–151 to decrease somewhat. This idea is embodied in the suggested fluidized bed reactor design, which operates close to isothermal conditions. Two fluidized bed reactor concepts have been suggested by the same group147 to provide the heat required to sustain the steam reforming reaction (Figure 11.4). The first integrates an oxygen transport membrane that partially oxidises methane upstream to generate hot gas. The second integrates a separate Pd-based membrane in the same reactor volume where air is fed to the permeate side and heat is generated by oxygen reacting with hydrogen diffusing through the membrane. An important result of the simulations in the study by

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

54

Figure 11.4

Chapter 11

Schematic representation of the two fluidized membrane reactor concepts for autothermal methane reforming with integrated CO2 capture (a) Methane combustion configuration; (b) Hydrogen combustion configuration, after Patil et al.148

Gallucci et al.144 was the identification of mass transfer limitation between the bubble phase and the emulsion phase (where the catalytic reaction takes place) in the fluidised bed reactor. This may potentially result in a large degree of slip of gas out of the reactor through the bubble phase. The authors showed that to break up the bubbles by inserting certain stagers as meshing wires at different heights in the reactor proved very efficient in reducing the slip of gas. Another important issue related to mass transfer limitation is the phenomenon referred to as concentration polarization, which is due to mass transfer limitations in the gas phase close to a highly permeable membrane. This has recently been given more attention78,152 due to the development of thinner and highly permeable Pd-based membranes. Simulation and comparison with experimental data have confirmed that concentration polarization may readily occur for membranes in the range of 1–5 mm, but under certain conditions also for significantly thicker membranes. Generally, the tendency of concentration polarization increases with: decreasing membrane thickness, hydrogen molar fraction and Reynolds number, and increasing temperature and feed pressure.78,152,153 It should be noted that possible concentration polarization is not limited to the feed side of the membrane. Thick, porous supports with small pores and use of sweep gas will govern concentration polarization at the permeate membrane–support interface region. Since the tendency of concentration polarization depends on the mole fraction of hydrogen in the gas one may experience this effect in parts of the reactor. Since steam reforming of methane is carried out at fairly high temperature, the counteracting flux reduction by adsorption surface effects is less. Finally, transport limitations in the catalyst bed itself should be avoided by proper design of catalyst and reactor.

View Online

Palladium-based Membranes in Hydrogen Production

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

11.4.2

55

Water Gas Shift

WGS reactors allow for lower membrane temperatures and a broader application range that also include synthesis gas from gasification of coal and biomass, as well as the previously mentioned high pressure/temperature autothermal reformers. But as also mentioned, the gasification processes are more demanding because of impurities that can be detrimental to the membrane performance. Thus, current state-of-the-art Pd-based membrane technology appears as a more viable option for synthesis gas from pure feedstock or after purification upstream the WGS reactor. The WGS reaction is weakly exothermic and independent of pressure, and temperatures down to 180–200 1C are targeted in the last reaction step to approach close to complete conversion of CO. A WGS membrane reactor process, on the other hand, opens for somewhat higher WGS temperatures, possibly as a single step, while maintaining the CO conversion and hydrogen production. The two existing catalysts have, however, limited applicability in the whole temperature range, i.e. the iron– chromium-based catalyst has limited efficiency in the lower temperature range while the more active copper-based has limited stability in the upper temperature range.12,13 Adsorption effects are stronger in the WGS temperature range (200–500 1C), and this calls for attention towards, for example, CO competitive adsorption on the membrane surface. Both modeling and experimental WGS studies show that the CO conversion increases with pressure in the membrane reactor due to the extraction of hydrogen.42,154–160 The higher conversion may give rise to larger temperature gradients in the membrane reactor compared to the traditional reactor. Gosiewski and co-workers161 suggested that a catalyst with operation window as wide as 200–550 1C would be beneficial in conversion of synthesis gas at 2.5 MPa pressure. Temperature regulation may be obtained by combined control of the gas hourly space velocity in the reactor, steam addition, use of sweep as coolant as well as a sufficiently large heat transfer area.154,156 Table 11.3 shows experimentally reported results for WGS Pd-based membrane reactors. Concentration polarization under WGS conditions was experimentally verified by Peters and co-workers78 for B2 mm Pd-23wt%Ag membranes at 400 1C and 20 bar. For this condition the authors reported that 65% of the reduction in driving force for the separation was caused by surface adsorption of gas components, while only 20% was due to concentration polarization. These figures also clearly demonstrate that the resistance of the B2 mm membrane was relatively unimportant, and that thicker, possibly more stable membranes could be used. Pd-based membrane reactors are given increased attention as enabling technology in PCDC processes for power generation with CO2 capture. Integration of membrane reactors in integrated reforming or integrated gasification combined cycle (IRCC or IGCC) processes require hydrogen to be delivered at gas turbine inlet conditions (B250 1C, 20 bar), which implies that the membrane reactor should operate at high pressure to avoid or minimize hydrogen compression. From comparative studies between single operation methane

P (bar)

1 1.01 1 1 1 1 1.01

322 322 325 331 331 332 400

2.4 2.5 1 1 1 1 3

S/C (mol  mol1)

Pd/ceramic Pd Pd-Ag foil Pd Pd/Ag Pd23wt%Ag Pd/porous glass

Membrane 10 70 50 70 50 10 20

Thickness (mm) LT shift catalyst LT shift catalyst LT shift catalyst LT shift catalyst LT shift catalyst Haldor Topsoe Iron chromium oxide

Catalyst 98 100 98.2 94.18 100 100 98

CO conversion (%)

— — — — — 22.9 —

H2 flux (N m3  m2  h1)

Experimentally reported results for water gas shift reactions with Pd-based membrane reactors

T (1C)

Table 11.3

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

162 163 164 160 160 165 166

Ref.

View Online

56 Chapter 11

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

57

steam reforming and autothermal reforming (ATR) followed by a WGS membrane reactor, it is clear that the latter two-stage process scheme is economically more favorable with respect to cost of CO2 avoided/captured.167 A main reason for this lies in the limited partial pressure of hydrogen in the steam reforming process at high pressure, resulting in a low driving force for the separation of hydrogen and consequently a high membrane surface area. Estimations for a 400 MWe IRCC power plant either integrating the ATR þ WGS membrane reactor or the single stage steam reforming reactor give 4560 m2 and 28 000 m2 membrane surface area, respectively.168 Also compared to other promising technologies for PCDC processes with natural gas as fuel, a recent analysis shows that the two stage ATR þ WGS membrane reactor has potentially the lowest cost of CO2 avoided.139,167–169 Furthermore, the Pd-based membrane reactor concept enables co-production of electricity and pure hydrogen broadening the flexibility and applicability of the technology. Middleton and co-workers suggested for the WGS process that the reactor containing the catalyst could be separated from the membrane separation unit, in a three-stage sequential process of reaction and separation.4,170 They claimed several advantages with this concept: (i) each reactor and separation step can be optimized with respect to membrane area, catalyst type and volume, flow rate, sweep sizing, design, and to a certain degree, temperature; (ii) replacement of membrane and catalyst can be done independently; and (iii) problems related to chemical non-compatibility of catalyst and membrane can be eliminated. The down-side of the concept is an increase in catalyst volume (33%) and membrane area (29%) compared to a single stage WGS membrane reactor process designed for capture of 2 Mtonnes year1 CO2 in the production of low-pressure hydrogen for applications at the Grangemouth refinery in Scotland.

11.4.3

Reforming of Alcohols

Steam and autothermal reforming of ethanol to produce hydrogen have been studied in Pd-based membrane reactors.171–178 An extensive overview of possible reaction pathways and performance of various catalyst-support systems in ethanol steam reforming is given in references 16, 17, 171. The total reaction may be described by the generalized equation (eqn (11.6)). Experimental investigations at moderate operation pressures suggest that the hydrogen yield and conversion increase with membrane reactor pressure, as opposed to conventional reactor without hydrogen extraction from the reaction zone.123,179,180 This is also verified by simulation work by Gallucci and co-workers175 comparing traditional and membrane reactors in ethanol steam reforming. Their results demonstrate the importance of the steam content, which has different effects: (i) a diluting effect reducing the partial pressure of hydrogen and driving force of the separation and (ii) a thermodynamic effect increasing the conversion. Also considering the issue of carbon deposition on the catalyst,1 they suggested that the ideal steam/ethanol ratio should be as high as the one necessary for avoiding this problem. Other experimentally directed studies use even higher S/C ratios,171,178–181 also higher than required from thermodynamic

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

58

Chapter 11

Figure 11.5 Minimum hydrogen yields needed to achieve given process efficiencies, as a function of the steam-to-carbon molar ratio, after Papadias et al.171

considerations of carbon deposition,171 but over-stoichiometric amounts of steam influences the energy efficiency of the process. In Figure 11.5 the minimum hydrogen yield is shown as a function of steam/carbon ratio.171 The expression minimum hydrogen yield relates to the US Department of Energy efficiency target of 72% for distributed production of hydrogen from bio-derived renewable liquids for 2012. Any inefficiency due to heat loss or other power requirements are not included in the calculation. Only the energy required by the latent heat of vaporisation of inlet liquid water is considered since this heat is assumed not recovered.171 Referring to Figure 11.5 and eqn (11.10), it is clear that for an S/C ratio of 3, about 5.1 mol of hydrogen must be obtained for every mole of ethanol fed to the process: C2 H5 OHðlÞ þ aH2 OðlÞ $ bCO2 þ cCO þ dCH4 þ eH2 þ ða  3 þ c þ 2dÞH2 O ð11:10Þ Even though the hydrogen yield is improved in a membrane reactor, the authors concluded based on simulation that their employed 30 mm Pd-Ag membrane would need 20 times higher flux to generate hydrogen yields that match the DoE target. Table 11.4 shows experimentally reported results for membrane reactors employing ethanol as fuel. Co-feeding oxygen182–184 to provide heat, appears also to reduce the tendency of coking.183,184 The amount of oxygen co-fed with steam obviously has importance for the reaction path, with too little steam reforming will prevail,

P (bar)

2.5 1.3 3.04 2 2 2

T (1C)

400 450 450 450 450 450

Table 11.4

11 4.5 1 13 13 13

Pd-Ag Pd-Ag Pd-Ag/PSS Pd-Ag Pd-Ag Pd-Ag

S/C (mol  mol1) Membrane 50 50 20 50–60 50–60 50–60

Thickness (mm)

Ethanol conversion (%) 100 50 81 — — —

Catalyst Ru/Al2O3 Ru/Al2O3 Zn–Cu Ru/Al2O3 Pt/Al2O3 NiO/SiO2

8 — — 82.2 60 60

H2 recovery (%)

Experimentally reported results for membrane reactors employing ethanol as fuel

— — — — — —

H2 flux (N m3  m2  h1)

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

183 186 182 179 180 180

Ref.

View Online

Palladium-based Membranes in Hydrogen Production 59

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

60

Chapter 11

while too much may give higher selectivity to complete combustion products (CO2 þ H2O).182,185 Recent simulation studies of a fluidized bed membrane reactor show that autothermal reforming of ethanol may give close to 100% conversion and hydrogen recovery at high temperature and pressure.145 A comment about bio-ethanol as fuel is that since fermentation provides 12–15% ethanol in water, distillation could be avoided by employing this solution in steam reforming.176 The inherent flexibility of modulized membrane systems enables production of hydrogen on membrane area scales from 106 to 104 m2, i.e. from microreactors191,192 to large power plants with CO2 sequestration.4 The usefulness of the technology to convert liquid fuels like ethanol, methanol, LPG, etc.193–195 to high purity hydrogen up-front fuels cells opens up many applications at small and medium scale. Table 11.6 shows the hydrogen yield (g g1) and specific energy (Wh kg1) of some commercially available hydrogen sources.193 Of the liquid fuels, methanol conversion has been extensively studied in Pd-based membrane reactors, both in steam reforming188,189,196–199 and autothermal oxygen co-fed200,201 processes. Methanol has been regarded as a convenient on board fuel due to its availability, purity, fairly high energy density and low reaction temperature (200–500 1C). Too low temperature is not advantageous202 due to low permeance of the membrane and higher probability of flux limiting CO/H2O surface adsorption. However, methanol dissociates rapidly above 450 1C, thus, alternatively to steam or autothermal reforming at B200–280 1C, a more permeable membrane reactor with high temperature WGS catalyst is an option.193 Table 11.5 shows experimentally reported results for membrane reactors employing methanol as fuel. Current state-of-the-art PEM fuel cells can only tolerate a few ppm of CO in the fuel before poisoning of the anode Pt catalyst occurs. Integration of membrane reactors with fuel cells obviously becomes more demanding in dynamic than stable operation modus, and smart hydrogen storage buffer systems compensating for variations in fuel utilisation may be required.196 For stationary applications assuming 50% energy conversion efficiency of the fuel cell, about 17 kWh electricity can be produced from 1 kg of hydrogen. If, however, hydrogen is produced for storage, for example at a hydrogen filling-station for vehicles, considerable compression work is required. Assuming that hydrogen is produced at atmospheric pressure by the membrane reactor and thereafter compressed for storage to 425 bars using electricity from the grid (assuming transmission grid electricity efficiency of 32.5%), close to 40% of the fuel’s lower heating value (LHV) will be consumed.171 Thus, for electrically driven cars, on-board hydrogen production from renewable sources appears from this reason as an interesting solution.

11.4.4

Dehydrogenation and Coupled Endothermic and Exothermic Reactions

Dehydrogenation reactions with Pd-based membrane reactors have been reported in numerous publications.203–207 The commercial interest in typical

P (bar)

1.3

3

6.08

1.3 1.3

300

300

350

400 450

4.5 4.5

1.2

3

5

S/C (mol  mol1)

Pd-Ag Pd-Ag

Pd

Pd-Ag

Pd-Ag

Membrane

50 50

20

50

50

Thickness (mm) ZnOMgO/ CuO/Al2O3 Cu/Zn/Mgbased Cu/ZnO/ Al2O3 Ru/Al2O3 Ru/Al2O3

Catalyst

— —

50

95 100 100

93

8

H2 recovery (%)

—

100

Methanol conversion (%)

Experimentally reported results for membrane reactors employing methanol as fuel

T (1C)

Table 11.5

— —

3.70

—

2

H2 flux (Nm3  m2  h1)

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

186 190

189

188

187

Ref.

View Online

Palladium-based Membranes in Hydrogen Production 61

View Online

62

Chapter 11

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Table 11.6

Hydrogen yield and corresponding specific energy of fuel choices, after Damle193

Fuel

Hydrogen yield (g  g1)

Specific energy (Whe  kg1)

Compressed hydrogen Metal hydrides for storage NaBH4 hydrolysis Methanol Methanol (including water) Butane Butane (including water) Gasoline Gasoline (including water) JP-8 JP-8 (including water) Clearlites Clearlites (including water) Ammonia

0.01 0.013 0.108 0.188 0.120 0.448 0.129 0.444 0.125 0.435 0.123 0.430 0.120 0.176

150 200 1830 3190 2040 7620 2190 7500 2140 7400 2090 7350 2050 3000

dehydrogenation processes, however, is not the hydrogen by-product. Nevertheless, high purity hydrogen is being produced in addition to the higher conversion efficiency and yield often reported for such membrane reactor processes.204,205,207 The extraction of hydrogen may change the reaction paths compared to a traditional reactor thereby changing the weight distribution of products coming out of the reactor.204,205 Endothermic dehydrogenation reactions have been coupled with exothermic reactions in the same reactor unit, e.g. cyclohexane dehydrogenation coupled with methanol synthesis.208–210 These interesting reactors couple heat and mass transport very efficiently, but significant further work would be required realize the concept at a commercial level.

11.4.5

Decomposition of Ammonia

Ammonia is a potential hydrogen carrier (17.6 wt%) and is a liquid at room temperature around 6 atm pressure.211 As a fuel it will not generate CO, CO2 or C in combination with hydrogen, which is advantageous. Ammonia converts (>99%) to nitrogen and hydrogen in a weakly endothermic reaction via catalytic decomposition, as in eqn (11.11): 2NH3 $ N2 þ 3H2 ;

DH 0298 ¼ 46:22 kJ mol1

ð11:11Þ

Experimental evidence of 100% conversion at B320–400 1C and hydrogen production in Pd-membrane reactors are reported.211 Simulation confirms that ammonia conversion in the membrane reactor increases with increasing pressure in the lower pressure range, temperature, flow rate of sweep gas, and reducing membrane thickness.212–214

View Online

Palladium-based Membranes in Hydrogen Production

63

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

11.5 Demonstration of Up-scaled Hydrogen Production by Palladium-based Membrane Reactors Tokyo Gas has demonstrated the world’s largest membrane reformer (MRF) with a H2 production capacity of 40 N m3 h1 on natural gas. The MRF test unit has a multi-tube rectangular structure as shown in Figure 11.6 and the size including heat insulation is W1200  L750  T1350 mm. Unit reactor tubes are arrayed with uniform spacing and a combustion furnace with two burners is attached under the reformer. High-temperature combustion gas generated in the furnace flows upward and heats the reactor tubes to a selected temperature. The reformer has 112 reactor tubes, each of which has two planar-type membrane modules composed of stainless steel support and Pd-based films less than 20 mm thick. The system has demonstrated the potential advantages of the membrane reformer: simple system configuration by single-step production of high-purity hydrogen (99.999% level), compactness, and high-energy efficiency of 70–76%. In total, Tokyo Gas operated the membrane reformer during 3000 h in reforming of natural gas. In 2009 a new 40 N m3 h1 MRF-2, with improved durability and higher hydrogen production efficiency was installed.216 The hydrogen production efficiency was as high as 81.4%, which is the highest reported efficiency in terms of hydrogen production from natural gas. Since the CO2 concentration in the retentate is as high as 70–90%, CO2 can be liquefied and separated with a low energy loss. Over 90% of CO2 in the reactor off-gas has been captured by cryogenic separation, lowering the total hydrogen production efficiency to 78.6%, which is 5–10% higher than the conventional reforming technologies. A longer test with many start-and-stop cycles is also scheduled to prove improved durability. A reformer and membrane module system (RMM) (permeative-stage membrane reactor) with 20 N m3 h1 of hydrogen capacity has been designed

Figure 11.6

Structure of 40 N m3 h1 class membrane reformer, after Shirasaki et al.215

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

64

Chapter 11

and constructed to investigate, at an industrial scale, the performance of their architecture.217 The membrane reactor module consists of a reactor series where each reactor is made of two separate parts; a reformer and an external membrane separating module. The advantages of separating the membrane from the reforming part are that the Pd-Ag membrane is kept at a lower temperature than the reforming temperature, simpler geometry and more simplified maintenance and replacement of membranes and catalyst. The number of reforming modules has been optimized based on an evaluation of the minimum separation area needed to achieve a pre-set hydrogen production requirement. The gas turbine exhaust can be used to heat the convective RMM process. Different types of Pd-based membranes, two tubular and one planar shaped, have been installed in order to evaluate their thermochemical stability and hydrogen permeability. A noble metal catalyst supported on SiC foam is placed inside the reactor tubes in order to enhance heat conduction inside the bed. At a reaction temperature of 620 1C with S/C of 4.8 and a membrane temperature of 430 1C, an overall yield of 59% was achieved.218 By properly extending the design parameters within reasonable limits, it is possible to find conditions giving a methane conversion as high as 90%.

11.6 Examples of Up-scaled State-of-the-Art Palladium-based Membrane Technology For successful commercialization of Pd-based membranes, the membrane must have sufficient permeability, selectivity, robustness and durability in relevant environments, preferably in the presence of common contaminants such as H2S. Moreover, the production cost of membranes and modules must be low enough for integration in the process resulting in beneficial cost-effectiveness compared to alternative technologies and processes.

11.6.1

CRI-Criterion

During the last decade CRI-Criterion (a company owned by Shell) has made substantial efforts in developing and up-scaling Pd membrane technology. CRI has chosen to develop its technology around electroless plating on ceramiccoated tubular metal supports since it is an inexpensive method for deposition of palladium, silver and gold. CRI’s palladium membrane manufacturing effort is continuing to improve the quality and robustness and expands to alloy membranes including Pd-Ag and Pd-Au. The current membrane manufacturing process targets a 7 mm thick membrane. Tubular porous metal supports with 1 and 2 inch outside diameters (OD) were selected because they can be easily assembled with normal welding techniques, are not fragile like ceramic materials, and do not require special fittings to achieve leak-tight seals. CRI-Criterion currently produces membranes of ODs up to 2 inches and lengths up to 24 inches. Current focus is to extend the viable individual unit length to approximately 1 meter.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

65

Recycle procedures for spent membranes have been developed to improve membrane economics by recycled membranes to function in the same manner as the originals. The permeances of the membranes vary in a range of 30–70 N m3 m2 h1 bar0.5 and are affected by the support and process parameters. The CRI membranes operate in a pressure range of 30–45 bar dP and at temperatures of up to 550 1C. Typical examples of the performance of CRICriterion Pd/Stainless Steel H2 separation membrane are given in Table 11.7. A high temperature Pd/Stainless Steel H2 separation membrane manufactured by CRI-Criterion by deposition of Pd on sintered porous metal support with permeable dimensions of 1 inch OD and 6 inches length is shown in Figure 11.7. An SEM cross section of the top layers of the membranes showing the porous stainless support, the inter-metal diffusion barrier, and the selective Pd layer. Note that the picture lack pinholes in Figure 11.8. The membranes tested show long-term robustness and demonstrate stability during and after experiencing a number of uncontrolled shutdowns. The membranes consistently delivered hydrogen purity in a range of 98.16 to 100%. Figure 11.9 shows the performance of a membrane under various conditions up to a period of 7000 h.219,220 It was initially started with a mixture of steam nitrogen and hydrogen and then eventually switched to a feed containing H2, H2O, CO2, CO and CH4. During the test the membrane experienced a number of shut downs. Figure 11.10 shows the effect of varying CO partial pressure on the hydrogen flux. No effect of CO on the hydrogen flux is noted as opposed to some literature reports.

11.6.2

Pall Corporation

Pall Corporation221 develops Pd-alloy composite membrane technology for hydrogen separation and production. Central to Pall’s Pd-alloy composite membrane technology is the development of an appropriate substrate material consisting of a porous metal tube with a homogeneous fine pore size ceramic diffusion barrier coating that enables deposition of thin Pd-alloy films. Pall Corporation currently produces the base support porous metal tubes in lengths up to 8 feet. Welding non-porous metal tubes to the porous metal substrate tubes addresses the membrane sealing issues and allows fabrication of large-scale modules using conventional tube-sheet and shell-tube vessel manufacturing technology. Pall is currently producing the substrate tubes and membrane elements up to 12 inches in length, while composite membranes with a length of up to 1 m length are under way. Pall Corporation has different techniques available for preparing Pd-based composite membranes, and has the ability to vary alloy composition as well as membrane thickness. For a Pd-based composite membrane of a 3 mm nominal thickness, the typical pure hydrogen flux rate at 1.38 bar (20 psi) hydrogen differential pressure and at 400 1C membrane temperature is about 76 mL cm2 min1 (150 scfh ft2). The typical H2/Ar binary gas selectivity for the same membrane is about 10 000.

Test period (h)

7000

4000 2300 1600 900 800

CRI-221

CRI-248 CRI-223 CRI-269 CRI-253 CRI-251

99.3 99.75 100 in progress 99.94 in progress 100

98.16

H2 purity at end (%)

17.93–18.41 9.11–9.35 17.89–18.16 15.94–22.99 21.84–33.09

15.07–17.96

Permeate flux (N m3  m2  h1)

8.74–9.22 2.32–2.45 9.23–9.42 7.1–11.07 6.15–23.6

9.10–11.39

Effluent flux (N m3  m2  h1)

430–450 500–505 430–450 430–450 430–450

430–450

Temperature (1C)

29–40 28.6–29.4 28.85–29.15 29 29–45

30

Pressure (Barg)

(H2, N2, steam), (CO, CO2, CH4, H2, steam) H2, N2, steam Steam methane reforming H2, N2, steam CO, CO2, CH4, H2, steam CO, CO2, CH4, H2, steam

Conditions/temp gas composition

Typical examples of the performance of CRI-Criterion Pd/stainless steel H2 separation membrane

Membrane no.

Table 11.7

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

View Online

66 Chapter 11

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

67

Figure 11.7

High temperature Pd/stainless steel H2 separation membrane manufactured by CRI-Criterion by deposition of Pd on sintered porous metal support with permeable dimensions of 1 inch OD and 6 inch length.

Figure 11.8

SEM cross section of the top layers of the membranes showing the porous stainless support, the intermetal diffusion barrier, and the selective Pd layer.

Figure 11.11 shows a 12 inch length ceramic (zirconia) coated porous metal support tube and two Pd-based composite membrane elements of 4 inch and 12 inch length respectively. The current generation membrane elements have been successfully tested in typical synthesis gas environments for over 500 h and have been shown to be durable over 50 thermal cycles. The current membranes are designed for low sulfur natural gas and other low containing sulfur fuels such as methanol. Development of a Pd-alloy composition tolerant to up to 100 ppm H2S is currently ongoing. Pall Corporation is engaged in trials for small portable/ stationary hydrocarbon fuel reforming-based hydrogen and power generation applications.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

68

Chapter 11

Figure 11.9

Long-term performance of a Pd/stainless steel H2 separation membrane as manufactured by CRI-Criterion during 7000 h of operation.

Figure 11.10

Effect of CO partial pressure on the hydrogen flux obtained with a CRICriterion Pd membrane. During this period, the total gas flow (wet basis) was held constant at 794 SLPH. The CO/CO2 ratio was changed. This was done while holding all other flows, temperature (430 1C), pressure constant (30 barg).

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

Figure 11.11

11.6.3

69

Pd-based composite membranes and a substrate tube element.

Energy Centre of the Netherlands

Energy Centre of the Netherlands (ECN) is developing hydrogen separation membranes prepared by electroless plating of Pd-alloys with a thickness of B3– 9 mm on low-cost ceramic substrates. The membranes can be manufactured up to a length of 90 cm. The membranes are sealed at the tube ends by a patented leak tight sealing which can resist pressure differences of at least 30 bar with the highest pressure on the outside of the tube.222 Typical operating temperatures are up to 480 1C. Lifetimes of several thousands of hours have been shown under different conditions. The purities that can be reached range from 99.5 to 99.995% in a single step, depending on the initial composition. ECN has recently made their Pd-based membrane technology precommercially available through the Hyseps modules. These modules, currently with membrane area of 0.04 m2 0.1 and 0.5 m2, are only suitable for hydrogen separation. The nominal capacity of the largest membrane module equals 3.5–6 N m3 h1, based on the obtained hydrogen flux applying reformate with 33% H2, an inlet pressure of 25 bar and H2 outlet pressure of 4 bar.223 An example of the largest membrane Hyseps module and a SEM cross section of the Pd-based membrane can be seen in Figure 11.12.

Figure 11.12

(a) ECN hydrogen separation module type Hyseps 1308; (b) SEM cross section of the Pd-based membranes showing the porous ceramic support and the selective Pd layer. From Ref. 223.

View Online

70

Chapter 11

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

11.6.4

Membrane Reactor Technologies

One of the key activity areas for Membrane Reactor Technologies (MRT) has been the development of its own techniques for manufacturing hydrogenpermeable membranes and robust, functional membrane modules. Membranes are manufactured either as rolled foils or as deposited thin films. While Pd-based foils of common compositions are available commercially, MRT has developed its own alloy compositions for added performance and robustness. In addition, a patent-pending bonding technique has been developed to permanently attach membranes to support modules with a perfect, hydrogen-tight seal. MRT has for example manufactured two-side planar membrane panels consisting of 25 mm thick Pd-Ag foil mounted on a porous stainless steel base using proprietary protection techniques.224–226 For membranes thinner than 15 mm, MRT uses a proprietary coating technique. Prototype membranes as thin as 8 mm have been produced and show excellent performance and longevity. The state of this technology is that standalone purification systems have been developed for a range of pure and ultra-pure hydrogen end uses.

11.7 Concluding Remarks In this chapter we have shown that the main trend in Pd-based membrane development is towards thinner membranes, particularly composite membranes. The relatively demanding operation conditions in many important applications require further work in terms of membrane development in combination with optimisation of reactor design and operation conditions. The encouraging industrial involvement in the development of the membrane, module and reactor technology is a key factor for successful implementation. Current state-of-the-art shows capability of small scale industrial membrane production of the most common Pd-based composite membranes. Their stability has been verified for thousands of hours under realistic conditions, but in processes free of some of the most hazardous components, e.g. sulfur. A large number of reports demonstrates that compared to conventional reactors for hydrogen production, membrane reactors show some distinct advantages:  Direct production of high purity hydrogen eliminating downstream purification  Operating conditions (temperature, pressure) that reduce reactor material cost  Higher conversion and selectivity to desired products  More compact reactors with less heat and mass transfer limitation  High versatility in terms of applications and scale, i.e. from integration in micro-systems to large scale power plants with CO2 capture. Further research is required to establish more stable Pd-based membranes for different applications. New Pd alloys may lead to more demanding production,

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

71

which would subsequently require further development in membrane manufacturing. The membrane reactor design will also need further development to minimize operation and investment cost. With respect to sealing technology, little detailed open information is available, we believe that this issue should be addressed for both cost and reactor design reasons. Realization of compact membrane reactors will require cheap and efficient sealing solutions. Finally, further exploitation of the inherent advantages of the technology for hydrogen production and other applications should include more focus on technoeconomic assessments to ensure proper guidance of the development.

Acknowledgements The support from European Union, Statoil ASA through the Gas Technology Center NTNU-SINTEF, and the Research Council of Norway (RCN) through the following programs, RCN-RENERGI (Project No: 190779/S60), RCNKOSK (Contract No. 197709/431), and the EU-7FP CACHET-II project (Contract no.: 241342) (http://www.cachet2.eu/) is gratefully acknowledged.

References 1. A. Basile, Hydrogen production using Pd-based membrane reactors for fuel cells, Top. Catal., 2008, 51, 107–122. 2. R. Bredesen, K. Jordal and A. Bolland, High-temperature membranes in power generation with CO2 capture, Chem. Eng. Process., 2004, 43, 1129–1158. 3. G. Q. Lu, J. C. Diniz da Costa, M. Duke, S. Giessler, R. Socolow, R. H. Williams and T. Kreutz, Inorganic membranes for hydrogen production and purification: A critical review and perspective, J. Colloid Interface Sci., 2007, 314, 589–603. 4. R. Bredesen and T. A. Peters, in Membranes for Energy Conversion, ed. K. V. Peinemann and S. Pereira Nunes, Wiley-VCH Verlag, Weinheim, 2008, pp. 217–244. 5. A. Basile, F. Gallucci and S. Tosti, in Membrane Science and Technology Inorganic Membranes: Synthesis, Characterization and Applications, ed. R. Mallada and M. Menendez, Elsevier, 2008, pp. 255–323. 6. F. Gallucci, A. Basile, A. Iulianelli and J. A. M. Kuipers, A Review on Patents for Hydrogen Production Using Membrane Reactors, Recent Pat. Chem. Eng., 2009, 207–222. 7. B. N. Lukyanov, D. V. Andreev and V. N. Parmon, Catalytic reactors with hydrogen membrane separation, Chem. Eng. J., 2009, 154, 258–266. 8. C. A. Scholes, K. H. Smith, S. E. Kentish and G. W. Stevens, CO2 capture from pre-combustion processes - Strategies for membrane gas separation, Int. J. Greenhouse Gas Contr., 2010, 4, 739–755. 9. S. Tosti, Overview of Pd-based membranes for producing pure hydrogen and state of art at ENEA laboratories, Int. J. Hydrogen Energy, 2010, 35, 12650–12659.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

72

Chapter 11

10. J. D. Holladay, J. Hu, D. L. King and Y. Wang, An overview of hydrogen production technologies, Catal. Today, 2009, 139, 244–260. 11. J. N. Armor, The multiple roles for catalysis in the production of H2, Appl. Catal. A, 1999, 176, 159–176. 12. H. Hiller, R. Reimert, F. Marschner, H. J. Renner, W. Boll, E. Supp, M. Brejc, W. Liebner, G. Schaub, G. Hochgesand, C. Higman, P. Kalteier, W. D. Muller, M. Kriebel, H. Schlichting, H. Tanz, H. M. Stonner, H. Klein, W. Hilsebein, V. Gronemann, U. Zwiefelhofer, J. Albrecht, C. J. Cowper and H. E. Driesen, Gas Production, Ullmann’s Encyclopedia of Industrial Chemistry, 2006. 13. P. Haussinger, R. Lohmuller and A. M. Watson, Hydrogen, Ullmann’s Encyclopedia of Industrial Chemistry, 2000. 14. B. Silberova, H. J. Venvik, J. C. Walmsley and A. Holmen, Small-scale hydrogen production from propane, Catal. Today, 2005, 100, 457–462. 15. I. Aartun, T. Gjervan, H. Venvik, O. Gorke, P. Pfeifer, M. Fathi, A. Holmen and K. Schubert, Catalytic conversion of propane to hydrogen in microstructured reactors, Chem. Eng. J., 2004, 101, 93–99. 16. M. Ni, D. Y. C. Leung and M. K. H. Leung, A review on reforming bioethanol for hydrogen production, Int. J. Hydrogen Energy, 2007, 32, 3238–3247. 17. A. Haryanto, S. Fernando, N. Murali and S. Adhikari, Current status of hydrogen production techniques by steam reforming of ethanol: A review, Energy Fuels, 2005, 19, 2098–2106. 18. R. R. Davda, J. W. Shabaker, G. W. Huber, R. D. Cortright and J. A. Dumesic, A review of catalytic issues and process conditions for renewable hydrogen and alkanes by aqueous-phase reforming of oxygenated hydrocarbons over supported metal catalysts, Appl. Catal. B, 2005, 56, 171–186. 19. R. Peters, in Handbook of Heterogeneous Catalysis, ed. G. Ertl, 2008, pp. 3045–3080. 20. P. K. Cheekatamarla and C. M. Finnerty, Synthesis gas production via catalytic partial oxidation reforming of liquid fuels, Int. J. Hydrogen Energy, 2008, 33, 5012–5019. 21. G. J. Stiegel and M. Ramezan, Hydrogen from coal gasification: An economical pathway to a sustainable energy future, Int. J. Coal Geol., 2006, 65, 173–190. 22. Y. Kalinci, A. Hepbasli and I. Dincer, Biomass-based hydrogen production: A review and analysis, Int. J. Hydrogen Energy, 2009, 34, 8799–8817. 23. H. Balat and E. Kirtay, Hydrogen from biomass - Present scenario and future prospects, Int. J. Hydrogen Energy, 2010, 35, 7416–7426. 24. D. B. Levin and R. Chahine, Challenges for renewable hydrogen production from biomass, Int. J. Hydrogen Energy, 2010, 35, 4962–4969. 25. A. Tanksale, J. N. Beltramini and G. M. Lu, A review of catalytic hydrogen production processes from biomass, Renew. Sustain. Energy Rev., 2010, 14, 166–182.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

73

26. A. Abuadala, I. Dincer and G. F. Naterer, Exergy analysis of hydrogen production from biomass gasification, Int. J. Hydrogen Energy, 2010, 35, 4981–4990. 27. E. Rytter, B. Børresen, I. Aartun, B. Krogh and M. Rønnekleiv, Method and reactor for the production of hydrogen, Pat. Appl. 20100047158, 2010. 28. L. Deng, T. J. Kim and M. B. Hagg, PVA/PVAm blend FSC membrane for CO2 capture, Desalination, 2006, 199, 523–524. 29. S. J. Chung, J. H. Park, D. Li, J. I. Ida, I. Kumakiri and J. Y. S. Lin, Dual-phase metal-carbonate membrane for high-temperature carbon dioxide separation, Ind. Eng. Chem. Res., 2005, 44, 7999–8006. 30. T. Yamaguchi, T. Niitsuma, B. N. Nair and K. Nakagawa, Lithium silicate based membranes for high temperature CO2 separation, J. Membr. Sci., 2007, 294, 16–21. 31. H. Kawamura, T. Yamaguchi, B. N. Nair, K. Nakagawa and S. Nakao, Dual-ion conducting lithium zirconate-based membranes for high temperature CO2 separation, J. Chem. Eng. Jpn., 2005, 38, 322–328. 32. E. Ochoa-Fernandez, H. K. Rusten, H. A. Jakobsen, M. Rønning, A. Holmen and D. Chen, Sorption enhanced hydrogen production by steam methane reforming using Li2ZrO3 as sorbent: Sorption kinetics and reactor simulation, Catal. Today, 2005, 106, 41–46. 33. E. Ochoa-Fernandez, C. Lacalle-Vila, T. Zhao, M. Rønning and D. Chen, in Studies in Surface Science and Catalysis, Natural Gas Conversion VIII, ed. M. S. and E. F. S.-A. Fabio Bellot Noronha, Proceedings of the 8th Natural Gas Conversion Symposium, Elsevier, 2007, pp. 159–164. 34. A. D. Wright, V. White, J. R. Hufton, R. Quinn, P. D. Cobden and E. R. van Selow, CAESAR: Development of a SEWGS model for IGCC, Energy Proc., 2011, 4, 1147–1154. 35. K. Johnsen, H. J. Ryu, J. R. Grace and C. J. Lim, Sorption-enhanced steam reforming of methane in a fluidized bed reactor with dolomite as CO2-acceptor, Chem. Eng. Sci., 2006, 61, 1195–1202. 36. M.-L. Fontaine, T. Norby, Y. Larring, T. Grande and R. Bredesen, in Membrane Science and Technology Inorganic Membranes: Synthesis, Characterization and Applications, ed. R. Mallada and M. Menendez, Elsevier, 2008, 401–458. 37. Air Products Courtesy, in Air Products Courtesy (ed.), 2010. 38. G. Xomeritakis and Y. S. Lin, Fabrication of thin metallic membranes by MOCVD and sputtering, J. Membr. Sci., 1997, 133, 217–230. 39. G. Xomeritakis and Y. S. Lin, Fabrication of a thin palladium membrane supported in a porous ceramic substrate by chemical vapor deposition, J. Membr. Sci., 1996, 120, 261–272. 40. E. Kikuchi, Y. Nemoto, M. Kajiwara, S. Uemiya and T. Kojima, Steam reforming of methane in membrane reactors: comparison of electrolessplating and CVD membranes and catalyst packing modes, Catal. Today, 2000, 56, 75–81.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

74

Chapter 11

41. N. Itoh, T. Akiha and T. Sato, Preparation of thin palladium composite membrane tube by a CVD technique and its hydrogen permselectivity, Catal. Today, 2005, 104, 231–237. 42. A. Basile, E. Drioli, F. Santella, V. Violante, G. Capannelli and G. Vitulli, A study on catalytic membrane reactors for water gas shift reaction, Gas Sep. Purif., 1996, 10, 53–61. 43. V. Jayaraman and Y. S. Lin, Synthesis and hydrogen permeation properties of ultrathin palladium-silver alloy membranes, J. Membr. Sci., 1995, 104, 251–262. 44. V. Jayaraman, Y. S. Lin, M. Pakala and R. Y. Lin, Fabrication of ultrathin metallic membranes on ceramic supports by sputter-deposition, J. Membr. Sci., 1995, 99, 89–100. 45. B. McCool, G. Xomeritakis and Y. S. Lin, Composition control and hydrogen permeation characteristics of sputter deposited palladium-silver membranes, J. Membr. Sci., 1999, 161, 67–76. 46. J. O’Brien, R. Hughes and J. Hisek, Pd/Ag membranes on porous alumina substrates by unbalanced magnetron sputtering, Surf. Coat. Technol., 2001, 142, 253–259. 47. H. B. Zhao, G. X. Xiong, N. Stroh and H. Brunner, Preparation and characterization of Pd-Ag alloy composite membrane with magnetron sputtering, Sci. China, Ser. B: Chem., 1999, 42, 581–588. 48. H. B. Zhao, G. X. Xiong and G. V. Baron, Preparation and characterization of palladium-based composite membranes by electroless plating and magnetron sputtering, Catal. Today, 2000, 56, 89–96. 49. H. Y. Gao, J. Y. S. Lin, Y. D. Li and B. Q. Zhang, Electroless plating synthesis, characterization and permeation properties of Pd-Cu membranes supported on ZrO2 modified porous stainless steel, J. Membr. Sci., 2005, 265, 142–152. 50. K. Hou and R. Hughes, Preparation of thin and highly stable Pd/Ag composite membranes and simulative analysis of transfer resistance for hydrogen separation, J. Membr. Sci., 2003, 214, 43–55. 51. J. N. Keuler, L. Lorenzen and S. Miachon, Preparing and testing Pd films of thickness 1–2 micrometer with high selectivity and high hydrogen permeance, Sep. Sci. Technol., 2002, 37, 379–401. 52. H. B. Zhao, K. Pflanz, J. H. Gu, A. W. Li, N. Stroh, H. Brunner and G. X. Xiong, Preparation of palladium composite membranes by modified electroless plating procedure, J. Membr. Sci., 1998, 142, 147–157. 53. L. Shi, G. Zeng and H. Xu, Characterization and performance of highflux PdAu/ceramic composite membranes, Cuihua Xuebao/Chin. J. Catal., 2010, 31, 711–715. 54. L. C. Witjens, J. H. Bitter, A. J. van Dillen, W. M. Arnoldbik, F. H. P. M. Habraken and K. P. de Jong, Improving the control of the electroless plating synthesis of Pd/Ag membranes for hydrogen separation using Rutherford backscattering, J. Membr. Sci., 2005, 254, 241–248.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

75

55. N. Itoh, N. Tomura, T. Tsuji and M. Hongo, Deposition of palladium inside straight mesopores of anodic alumina tube and its hydrogen permeability, Micropor. Mesopor. Mater., 2000, 39, 103–111. 56. Y. Huang and R. Dittmeyer, Preparation and characterization of composite palladium membranes on sinter-metal supports with a ceramic barrier against intermetallic diffusion, J. Membr. Sci., 2006, 282, 296–310. 57. K. Zhang, H. Y. Gao, Z. B. Rui, P. Liu, Y. D. Li and Y. S. Lin, HighTemperature Stability of Palladium Membranes on Porous Metal Supports with Different Intermediate Layers, Ind. Eng. Chem. Res., 2009, 48, 1880–1886. 58. D. Yepes, L. M. Cornaglia, S. Irusta and E. A. Lombardo, Different oxides used as diffusion barriers in composite hydrogen permeable membranes, J. Membr. Sci., 2006, 274, 92–101. 59. C. L. Su, T. Jin, K. Kuraoka, Y. Matsumura and T. Yazawa, Thin palladium film supported on SiO2-modified porous stainless steel for a highhydrogen-flux membrane, Ind. Eng. Chem. Res., 2005, 44, 3053–3058. 60. J. H. Tong, Y. Matsumura, H. Suda and K. Haraya, Thin and dense Pd/ CeO2/MPSS composite membrane for hydrogen separation and steam reforming of methane, Sep. Purif. Technol., 2005, 46, 1–10. 61. S. E. Nam and K. H. Lee, Preparation and characterization of palladium alloy composite membranes with a diffusion barrier for hydrogen separation, Ind. Eng. Chem. Res., 2005, 44, 100–105. 62. D. Wang, H. H. Tong, H. Y. Xu and Y. Matsumura, Preparation of palladium membrane over porous stainless steel tube modified with zirconium oxide, Catal. Today, 2004, 93–95, 689–693. 63. J. H. Tong, H. Y. Xu, D. Wang and Y. Matsumura, Preparation of thin palladium membrane on porous stainless steel support modified with cerium hydroxide, J. Jpn. Pet. Inst., 2004, 47, 64–65. 64. Y. Guo, X. Zhang, H. Deng, X. Wang, Y. Wang, J. Qiu, J. Wang and K. L. Yeung, A novel approach for the preparation of highly stable Pd membrane on macroporous alpha-Al2O3 tube, J. Membr. Sci., 2010, 362, 241–248. 65. Y. H. Ma, B. C. Akis, M. E. Ayturk, F. Guazzone, E. E. Engwall and I. P. Mardilovich, Characterization of intermetallic diffusion barrier and alloy formation for Pd/Cu and Pd/Ag porous stainless steel composite membranes, Ind. Eng. Chem. Res., 2004, 43, 2936–2945. 66. I. P. Mardilovich, E. Engwall and Y. H. Ma, Presentation at the 9th International Conference on Inorganic Membranes, Lillehammer, Norway, 2006. 67. J. Okazaki, T. Ikeda, D. A. Pacheco Tanaka, T. M. Suzuki and F. Mizukami, In situ high-temperature X-ray diffraction study of thin palladium/alpha-alumina composite membranes and their hydrogen permeation properties, J. Membr. Sci., 2009, 335, 126–132. 68. I. P. Mardilovich, E. Engwall and Y. H. Ma, Dependence of hydrogen flux on the pore size and plating surface topology of asymmetric Pdporous stainless steel membranes, Desalination, 2002, 144, 85–89.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

76

Chapter 11

69. F. Roa, J. D. Way, R. L. McCormick and S. N. Paglieri, Preparation and characterization of Pd-Cu composite membranes for hydrogen separation, Chem. Eng. J., 2003, 93, 11–22. 70. G. B. Sun, K. Hidajat and S. Kawi, Ultra thin Pd membrane on alphaAl2O3 hollow fiber by electroless plating: High permeance and selectivity, J. Membr. Sci., 2006, 284, 110–119. 71. X. L. Pan, G. X. Xiong, S. S. Sheng, N. Stroh and H. Brunner, Thin dense Pd membranes supported on alpha-Al2O3 hollow fibers, Chem. Commun., 2001, 2536–2537. 72. S. F. Hou, K. Jiang, W. Z. Li, H. Y. Xu and L. X. Yuan, WO2005/ 065806, A metal palladium composite membrane or alloy palladium composite membrane and their preparation methods, 2005. 73. J. H. Tong and Y. Matsumura, Thin Pd membrane prepared on macroporous stainless steel tube filter by an in-situ multi-dimensional plating mechanism, Chem. Commun., 2004, 2460–2461. 74. G. Zeng, A. Goldbach and H. Xu, Defect sealing in Pd membranes via point plating, J. Membr. Sci., 2009, 328, 6–10. 75. A. W. Li, W. Q. Liang and R. Hughes, Repair of a Pd/alpha-Al2O3 composite membrane containing defects, Sep. Purif. Technol., 1999, 15, 113–119. 76. R. Bredesen and H. Klette, Method of manufacturing thin metal membranes U.S. Patent 6 086 729, 2000. 77. T. A. Peters, M. Stange and R. Bredesen, in Carbon Dioxide Capture for Storage in Deep Geological Formations, ed. L. I. Eide, CPL Press and BP, 2010, pp. 135–155. 78. T. A. Peters, M. Stange, H. Klette and R. Bredesen, High pressure performance of thin Pd-23%Ag/stainless steel composite membranes in water gas shift gas mixtures; influence of dilution, mass transfer and surface effects on the hydrogen flux, J. Membr. Sci., 2008, 316, 119–127. 79. R. J. Gouwen, C. J. G. M. de Kok and J. A. Z. Pieterse, Energy Centre Of The Netherlands, Noble Metal Membrane Preparation: Industrial Film Formation Techniques, ECN-E-10-050, 2010. 80. J. Okazaki, D. A. P. Tanaka, M. A. L. Tanco, Y. Wakui, F. Mizukami and T. M. Suzuki, Hydrogen permeability study of the thin Pd-Ag alloy membranes in the temperature range across the alpha-beta phase transition, J. Membr. Sci., 2006, 282, 370–374. 81. B. H. Howard, R. P. Killmeyer, K. S. Rothenberger, A. V. Cugini, B. D. Morreale, R. M. Enick and F. Bustamante, Hydrogen permeance of palladium-copper alloy membranes over a wide range of temperatures and pressures, J. Membr. Sci., 2004, 241, 207–218. 82. F. Roa and J. D. Way, The effect of air exposure on palladium-copper composite membranes, Appl. Surf. Sc., 2005, 240, 85–104. 83. F. Roa and J. D. Way, Influence of alloy composition and membrane fabrication on the pressure dependence of the hydrogen flux of palladiumcopper membranes, Ind. Eng. Chem. Res., 2003, 42, 5827–5835.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

77

84. F. Roa, M. J. Block and J. D. Way, The influence of alloy composition on the H2 flux of composite Pd-Cu membranes, Desalination, 2002, 147, 411–416. 85. P. M. Thoen, F. Roa and J. D. Way, High flux palladium-copper composite membranes for hydrogen separations, Desalination, 2006, 193, 224–229. 86. L. Peng, Y. Rao, L. Luo and C. Chen, The poisoning of Pd-Y alloy membranes by carbon monoxide, J. Alloys Compd., 2009, 486, 74–77. 87. D. T. Hughes and I. R. Harris, A comparative study of hydrogen permeabilities and solubilities in some palladium solid solution alloys, J. Less-Common Met., 1978, 61, 9–21. 88. D. Fort, J. P. G. Farr and I. R. Harris, A comparison of palladium-silver and palladium-yttrium alloys as hydrogen separation membranes, J. LesCommon Met., 1975, 39, 293–308. 89. S. K. Gade, O. Hatlevik, H. J. Park, A. E. Lewis and J. D. Way, Properties of palladium-gold alloys for hydrogen separation and purification, AIChE Annual Meeting, Conference Proceedings, 2009. 90. S. K. Gade, M. K. Keeling, A. P. Davidson, O. Hatlevik and J. D. Way, Palladium-ruthenium membranes for hydrogen separation fabricated by electroless co-deposition, Int. J. Hydrogen Energy, 2009, 34, 6484–6491. 91. L. Wang, R. Yoshiie and S. Uemiya, Fabrication of novel Pd-Ag-Ru/ Al2O3 ternary alloy composite membrane with remarkably enhanced H2 permeability, J. Membr. Sci., 2007, 306, 1–7. 92. D. Wang, T. B. Flanagan and K. L. Shanahan, Hydrogen permeation measurements of partially internally oxidized Pd-Al alloys in the presence and absence of CO, J. Membr. Sci., 2005, 253, 165–173. 93. D. R. Alfonso, A. V. Cugini and D. S. Sholl, Density functional theory studies of sulfur binding on Pd, Cu and Ag and their alloys, Surf. Sci., 2003, 546, 12–26. 94. A. Kulprathipanja, G. O. Alptekin, J. L. Falconer and J. D. Way, Pd and Pd-Cu membranes: inhibition of H2 permeation by H2S, J. Membr. Sci., 2005, 254, 49–62. 95. A. Kulprathipanja, G. O. Alptekin, J. L. Falconer and J. D. Way, Effects of water gas shift gases on Pd-Cu alloy membrane surface morphology and separation properties, Ind. Eng. Chem. Res., 2004, 43, 4188–4198. 96. C. H. Chen and Y. H. Ma, The effect of H2S on the performance of Pd and Pd/Au composite membrane, J. Membr. Sci., 2010, 362, 535–544. 97. C. P. O’Brien, B. H. Howard, J. B. Miller, B. D. Morreale and A. J. Gellman, Inhibition of hydrogen transport through Pd and Pd47Cu53 membranes by H2S at 350 1C, J. Membr. Sci., 2010, 349, 380–384. 98. A. L. Mejdell, D. Chen, T. A. Peters, R. Bredesen and H. J. Venvik, The effect of heat treatment in air on CO inhibition of a B3um Pd-Ag (23wt.%) membrane, J. Membr. Sci., 2010, 350, 371–377. 99. L. Yang, Z. X. Zhang, X. H. Gao, Y. J. Guo, B. F. Wang, O. Sakai, H. Sakai and T. Takahashi, Changes in hydrogen permeability and surface state of Pd-Ag/ceramic composite membranes after thermal treatment, J. Membr. Sci., 2005, 252, 145–154.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

78

Chapter 11

100. K. Yamakawa, M. Ege, B. Ludescher and M. Hirscher, Surface adsorbed atoms suppressing hydrogen permeation of Pd membranes, J. Alloy Compd., 2003, 352, 57–59. 101. A. Ramachandran, W. M. Tucho, A. L. Mejdell, M. Stange, H. J. Venvik, J. C. Walmsley, R. Holmestad, R. Bredesen and A. Borg, Surface characterization of Pd/Ag23 wt% membranes after different thermal treatments, Appl. Surf. Sci., 2010, 256, 6121–6132. 102. H. Li, A. Goldbach, W. Li and H. Xu, PdC formation in ultra-thin Pd membranes during separation of H2/CO mixtures, J. Membr. Sci., 2007, 299, 130–137. 103. H. Li, A. Goldbach, W. Li and H. Xu, CO2 Decomposition over Pd Membrane Surfaces, J. Phy. Chem. B, 2008, 112, 12182–12184. 104. F. L. Chen, Y. Kinari, F. Sakamoto, Y. Nakayama and Y. Sakamoto, Hydrogen permeation through palladium-based alloy membranes in mixtures of 10% methane and ethylene in the hydrogen, Int. J. Hydrogen Energy, 1996, 21, 555–561. 105. S. H. Jung, K. Kusakabe, S. Morooka and S. D. Kim, Effects of coexisting hydrocarbons on hydrogen permeation through a palladium membrane, J. Membr. Sci., 2000, 170, 53–60. 106. Y. Sakamoto, F. L. Chen, Y. Kinari and F. Sakamoto, Effect of carbon monoxide on hydrogen permeation in some palladium-based alloy membranes, Int. J. Hydrogen Energy, 2011, 21, 1017–1024. 107. L. Peng, Y. Rao, L. Luo and C. Chen, The poisoning of Pd-Y alloy membranes by carbon monoxide, J. Alloy Compd., 2009, 486, 74–77. 108. K. Hou and R. Hughes, The effect of external mass transfer, competitive adsorption and coking on hydrogen permeation through thin Pd/Ag membranes, J. Membr. Sci., 2002, 206, 119–130. 109. A. Unemoto, A. Kaimai, K. Sato, T. Otake, K. Yashiro, J. Mizusaki, T. Kawada, T. Tsuneki, Y. Shirasaki and I. Yasuda, Surface reaction of hydrogen on a palladium alloy membrane under co-existence of H2O, CO, CO2 or CH4, Int. J. Hydrogen Energy, 2007, 32, 4023–4029. 110. A. Unemoto, A. Kaimai, K. Sato, T. Otake, K. Yashiro, J. Mizusaki, T. Kawada, T. Tsuneki, Y. Shirasaki and I. Yasuda, The effect of coexisting gases from the process of steam reforming reaction on hydrogen permeability of palladium alloy membrane at high temperatures, Int. J. Hydrogen Energy, 2007, 32, 2881–2887. 111. Ø. Hatlevik, S. K. Gade, M. K. Keeling, P. M. Thoen, A. P. Davidson and J. D. Way, Palladium and palladium alloy membranes for hydrogen separation and production: History, fabrication strategies, and current performance, Sep. Purif. Technol., 2010, 73, 59–64. 112. S. K. Gade, M. K. Keeling, D. K. Steele, P. M. Thoen, J. D. Way, S. DeVoss and G. O. Alptekin, in Proceedings of the 9th International Conference on Inorganic Membranes, ed. R. Bredesen and H. Ræder, Lillehammer, Norway, 2006, pp. 112–117.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

79

113. L. Shi, A. Goldbach, G. Zeng and H. Xu, Preparation and performance of thin-layered PdAu/ceramic composite membranes, Int. J. Hydrogen Energy, 2010, 35, 4201–4208. 114. P. Kamakoti and D. S. Sholl, Towards first principles-based identification of ternary alloys for hydrogen purification membranes, J. Membr. Sci., 2006, 279, 94–99. 115. L. Semidey-Flecha, C. Ling and D. S. Sholl, Detailed first-principles models of hydrogen permeation through PdCu-based ternary alloys, J. Membr. Sci., 2010, 362, 384–392. 116. H. Klette, R. Bredesen and D. Slotfeldt-Ellingsen, Leak-proof membrane element and method of manufacturing such an element. International Patent Application, WO 2008/039080, 2008. 117. C. Y. Tsai, Presentation at the 11th International Conference on Inorganic Membranes, Washington, 2010. 118. S. Tosti, L. Bettinali, S. Castelli, F. Sarto, S. Scaglione and V. Violante, Sputtered, electroless, and rolled palladium-ceramic membranes, J. Membr. Sci., 2002, 196, 241–249. 119. F. Guazzone, E. A. Payzant, S. A. Speakman and Y. H. Ma, Microstrains and stresses analysis in electroless deposited thin Pd films, Ind. Eng. Chem. Res., 2006, 45, 8145–8153. 120. F. Guazzone and Y. H. Ma, Leak growth mechanism in composite Pd membranes prepared by the electroless deposition method, AIChE J., 2008, 54, 487–494. 121. T. A. Peters, W. M. Tucho, A. Ramachandran, M. Stange, J. C. Walmsley, R. Holmestad, A. Borg and R. Bredesen, Thin Pd-23%Ag/ stainless steel composite membranes: Long-term stability, life-time estimation and post-process characterisation, J. Membr. Sci., 2009, 326, 572–581. 122. R. Bredesen and J. Sogge, Presentation at the United Nation Seminar on the Ecological Applications of Innovative Membrane Technology in the Chemical Industry, Cetraro, Calabria, 1996. 123. A. Iulianelli, G. Manzolini, M. De Falco, S. Campanari, T. Longo, S. Liguori and A. Basile, H2 production by low pressure methane steam reforming in a Pd-Ag membrane reactor over a Ni-based catalyst: Experimental and modeling, Int. J. Hydrogen Energy, 2010, 35, 11514– 11524. 124. Y. M. Lin, S. L. Liu, C. H. Chuang and Y. T. Chu, Effect of incipient removal of hydrogen through palladium membrane on the conversion of methane steam reforming: Experimental and modeling, Catal. Today, 2003, 82, 127–139. 125. J. H. Tong, Y. Matsumura, H. Suda and K. Haraya, Experimental study of steam reforming of methane in a thin (6 mm) Pd-based membrane reactor, Ind. Eng. Chem. Res., 2005, 44, 1454–1465. 126. S. Uemiya, N. Sato, H. Ando, T. Matsuda and E. Kikuchi, Steam reforming of methane in a hydrogen-permeable membrane reactor, Appl. Catal., 1990, 67, 223–230.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

80

Chapter 11

127. Y. Chen, Y. Wang, H. Xu and G. Xiong, Hydrogen production capacity of membrane reformer for methane steam reforming near practical working conditions, J. Membr. Sci., 2008, 322, 453–459. 128. Y. Chen, Y. Wang, H. Xu and G. Xiong, Efficient production of hydrogen from natural gas steam reforming in palladium membrane reactor, Appl. Catal. B, 2008, 81, 283–294. 129. J. Tong and Y. Matsumura, Effect of catalytic activity on methane steam reforming in hydrogen-permeable membrane reactor, Appl. Catal. A, 2005, 286, 226–231. 130. P. Ferreira-Aparicio, M. Benito, K. Kouachi and S. Menad, Catalysis in membrane reformers: a high-performance catalytic system for hydrogen production from methane, J. Catal., 2005, 231, 331–343. 131. T. Ioannides and X. E. Verykios, Development of a novel heat-integrated wall reactor for the partial oxidation of methane to synthesis gas, Catal. Today, 1998, 46, 71–81. 132. H. F. Chang, W. J. Pai, Y. J. Chen and W. H. Lin, Autothermal reforming of methane for producing high-purity hydrogen in a Pd/Ag membrane reactor, Int. J. Hydrogen Energy, 2010, 35, 12986–12992. 133. B. Faroldi, C. Carrara, E. A. Lombardo and L. M. Cornaglia, Production of ultrapure hydrogen in a Pd-Ag membrane reactor using Ru/La2O3 catalysts, Appl. Catal. A, 2007, 319, 38–46. 134. L. Paturzo, F. Gallucci, A. Basile, G. Vitulli and P. Pertici, An Ru-based catalytic membrane reactor for dry reforming of methane–its catalytic performance compared with tubular packed bed reactors, Catal. Today, 2003, 82, 57–65. 135. E. Kikuchi and Y. Chen, in Studies in Surface Science and Catalysis, Natural Gas Conversion IV, ed. M. D. Pontes, Elsevier, 1997, pp. 547–553. 136. J. F. Mu´nera, C. Carrara, L. M. Cornaglia and E. A. Lombardo, Combined oxidation and reforming of methane to produce pure H2 in a membrane reactor, Chem. Eng. J., 2010, 161, 204–211. 137. C. S. Patil, PhD Thesis, University of Twente, Membrane Reactor Technology for Ultrapure Hydrogen Production, 2005. 138. A. Matzakos, S. L. Wellington, T. Mikus and J. M. Ward, Integrated flameless distributed combustion/steam reforming membrane reactor for hydrogen production and use thereof in zero emissions hybrid power system, U.S. Patent 6 821 501, 2004. 139. J. W. Dijkstra, G. Raju, G. Peppink and D. Jansen, Techno-economic evaluation of membrane technology for pre-combustion decarbonisation: water-gas shift versus reforming, Energy Proc., 2011, 4, 723–730. 140. A. Caravella, F. P. Di Maio and A. Di Renzo, Optimization of membrane area and catalyst distribution in a permeative-stage membrane reactor for methane steam reforming, J. Membr. Sci., 2008, 321, 209–221. 141. A. Li, C. J. Lim and J. R. Grace, Staged-separation membrane reactor for steam methane reforming, Chem. Eng. J., 2008, 138, 452–459. 142. J. R. Hufton, R. J. Allam, R. Chiang, P. Middleton, E. L. Weist and V. White, in Greenhouse Gas Control Technologies 7, ed. E. S. Rubin,

View Online

Palladium-based Membranes in Hydrogen Production

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

143.

144.

145.

146.

147.

148.

149.

150.

151.

152.

153.

154.

155.

81

D. W. Keith, C. F. Gilboy, M. Wilson, T. Morris, J. Gale and K. Thambimuthu, Elsevier Science Ltd, Oxford, 2005, pp. 253–261. G. Manzolini and F. Vigan, Co-production of hydrogen and electricity from autothermal reforming of natural gas by means of Pd-Ag membranes, Energy Proc., 2009, 1, 319–326. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, Theoretical comparison of packed bed and fluidized bed membrane reactors for methane reforming, Int. J. Hydrogen Energy, 2010, 35, 7142–7150. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, Pure hydrogen production via autothermal reforming of ethanol in a fluidized bed membrane reactor: A simulation study, Int. J. Hydrogen Energy, 2010, 35, 1659–1668. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, Autothermal reforming of methane with integrated CO2 capture in a novel fluidized bed membrane reactor. Part 2: Comparison of reactor configurations, Top. Catal., 2008, 51, 146–157. F. Gallucci, M. Van Sint Annaland and J. A. M. Kuipers, Autothermal reforming of methane with integrated CO2 capture in a novel fluidized bed membrane reactor. Part 1: Experimental demonstration, Top. Catal., 2008, 51, 133–145. C. S. Patil, M. Van Sint Annaland and J. A. M. Kuipers, Fluidised bed membrane reactor for ultrapure hydrogen production via methane steam reforming: Experimental demonstration and model validation, Chem. Eng. Sci., 2007, 62, 2989–3007. S. A. R. K. Deshmukh, S. Heinrich, L. Morl, M. Van Sint Annaland and J. A. M. Kuipers, Membrane assisted fluidized bed reactors: Potentials and hurdles, Chem. Eng. Sci., 2007, 62, 416–436. C. S. Patil, M. Van Sint Annaland and J. A. M. Kuipers, Experimental Study of a Membrane Assisted Fluidized Bed Reactor For H2 Production by Steam Reforming of CH4, Chem. Eng. Res. Des., 2006, 84, 399–404. S. A. R. K. Deshmukh, M. Van Sint Annaland and J. A. M. Kuipers, Effect of fluidization conditions on the membrane permeation rate in a membrane assisted fluidized bed, Chem. Eng. J., 2003, 96, 125–131. A. Caravella, G. Barbieri and E. Drioli, Concentration polarization analysis in self-supported Pd-based membranes, Sep. Purif. Technol., 2009, 66, 613–624. H. Li, J. W. Dijkstra, J. A. Z. Pieterse, J. Boon, R. W. van den Brink and D. Jansen, Towards full-scale demonstration of hydrogen-selective membranes for CO2 capture: Inhibition effect of WGS-components on the H2 permeation through three Pd membranes of 44 cm long, J. Membr. Sci., 2010, 363, 204–211. M. E. Adrover, E. Lopez, D. O. Borio and M. N. Pedernera, Simulation of a membrane reactor for the WGS reaction: Pressure and thermal effects, Chem. Eng. J., 2009, 154, 196–202. D. Mendes, V. Chibante, J. M. Zheng, S. Tosti, F. Borgognoni, A. Mendes and L. M. Madeira, Enhancing the production of hydrogen

View Online

82

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

156.

157.

158.

159.

160.

161.

162. 163.

164. 165.

166.

167. 168.

169.

Chapter 11

via water-gas shift reaction using Pd-based membrane reactors, Int. J. Hydrogen Energy, 2010, 35, 12596–12608. G. Chiappetta, G. Clarizia and E. Drioli, Theoretical analysis of the effect of catalyst mass distribution and operation parameters on the performance of a Pd-based membrane reactor for water-gas shift reaction, Chem. Eng. J., 2008, 136, 373–382. Y. Bi, H. Xu, W. Li and A. Goldbach, Water-gas shift reaction in a Pd membrane reactor over a Pt/Ce0. 6Zr0. 4O2 catalyst, Int. J. Hydrogen Energy, 2009, 34, 2965–2971. P. Pinacci, M. Broglia, C. Valli, G. Capannelli and A. Comite, Evaluation of the water gas shift reaction in a palladium membrane reactor, Catal. Today, 2010, 156, 165–172. S. S. Hla, G. J. Duffy, L. D. Morpeth, A. Cousins, D. G. Roberts and J. H. Edwards, Investigation of the effect of total pressure on performance of the catalytic water-gas shift reaction using simulated coal-derived syngases, Catal. Commun., 2009, 11, 272–275. A. Basile, G. Chiappetta, S. Tosti and V. Violante, Experimental and simulation of both Pd and Pd/Ag for a water gas shift membrane reactor, Sep. Purif. Technol., 2001, 25, 549–571. K. Gosiewski, K. Warmuzinski and M. Tanczyk, Mathematical simulation of WGS membrane reactor for gas from coal gasification, Catal. Today, 2010, 156, 229–236. A. Basile, V. Violante, F. Santella and E. Drioli, Membrane integrated system in the fusion reactor fuel cycle, Catal. Today, 1995, 25, 321–326. A. Criscuoli, A. Basile and E. Drioli, An analysis of the performance of membrane reactors for the water-gas shift reaction using gas feed mixtures, Catal. Today, 2000, 56, 53–64. S. Tosti, A. Basile, G. Chiappetta, C. Rizzello and V. Violante, Pd-Ag membrane reactors for water gas shift reaction, Chem. Eng. J., 2003, 93, 23–30. S. Tosti, V. Violante, A. Basile, G. Chiappetta, S. Castelli, M. De Francesco, S. Scaglione and F. Sarto, Catalytic membrane reactors for tritium recovery from tritiated water in the ITER fuel cycle, Fusion Eng. Des., 2000, 49, 953–958. S. Uemiya, N. Sato, H. Ando and E. Kikuchi, The water gas shift reaction assisted by a palladium membrane reactor, Ind. Eng. Chem. Res., 1991, 30, 585–589. R. Beavis, The EU FP6 CACHET Project - Final Results, Energy Proc., 2011, 4, 1074–1081. J. W. Dijkstra, J. A. Z. Pieterse, H. Li, J. Boon, Y. C. Van Delft, G. Raju, G. Peppink, R. W. van den Brink and D. Jansen, Development of membrane reactor technology for power production with pre-combustion CO2 capture, Energy Proc., 2011, 4, 715–722. D. Jansen, J. W. Dijkstra, R. W. van den Brink, T. A. Peters, M. Stange, R. Bredesen, A. Goldbach, H. Y. Xu, A. Gottschalk and A. Doukelis, Hydrogen membrane reactors for CO2 capture, Energy Proc., 2009, 1, 253–260.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

83

170. P. Middleton, H. Klette, R. Bredesen, Y. Larring, H. Ræder and C. Lowe, in 7th International Conference on Greenhouse Gas Control Technologies (GHGT-7), ed. E. S. Rubin, E. W. Keith and C. F. Gilboy, Vancouver, Canada. Elsevier Science, Oxford, 2004. 171. D. D. Papadias, S. H. D. Lee, M. Ferrandon and S. Ahmed, An analytical and experimental investigation of high-pressure catalytic steam reforming of ethanol in a hydrogen selective membrane reactor, Int. J. Hydrogen Energy, 2010, 35, 2004–2017. 172. S. Tosti, R. Borelli, A. Santucci and L. Scuppa, Pd-Ag membranes for auto-thermal ethanol reforming, Asia-Pacific J. Chem. Eng., 2010, 5, 207–212. 173. S. Tosti, A. Basile, R. Borelli, F. Borgognoni, S. Castelli, M. Fabbricino, F. Gallucci and C. Licusati, Ethanol steam reforming kinetics of a Pd-Ag membrane reactor, Int. J. Hydrogen Energy, 2009, 34, 4747–4754. 174. F. Gallucci, M. De Falco, S. Tosti, L. Marrelli and A. Basile, Co-current and counter-current configurations for ethanol steam reforming in a dense Pd-Ag membrane reactor, Int. J. Hydrogen Energy, 2008, 33, 6165–6171. 175. F. Gallucci, M. De Falco, S. Tosti, L. Marrelli and A. Basile, Ethanol steam reforming in a dense Pd-Ag membrane reactor: A modelling work. Comparison with the traditional system, Int. J. Hydrogen Energy, 2008, 33, 644–651. 176. K. Urasaki, K. Tokunaga, Y. Sekine, M. Matsukata and E. Kikuchi, Production of hydrogen by steam reforming of ethanol over cobalt and nickel catalysts supported on perovskite-type oxides, Catal. Commun., 2008, 9, 600–604. 177. C. Y. Yu, D. W. Lee, S. J. Park, K. Y. Lee and K. H. Lee, Study on a catalytic membrane reactor for hydrogen production from ethanol steam reforming, Int. J. Hydrogen Energy, 2009, 34, 2947–2954. 178. D. Mendes, S. Tosti, F. Borgognoni, A. Mendes and L. M. Madeira, Integrated analysis of a membrane-based process for hydrogen production from ethanol steam reforming, Catal. Today, 2010, 156, 107–117. 179. S. Tosti, A. Basile, F. Borgognoni, V. Capaldo, S. Cordiner, S. Di Cave, F. Gallucci, C. Rizzello, A. Santucci and E. Traversa, Low temperature ethanol steam reforming in a Pd-Ag membrane reactor: Part 1: Ru-based catalyst, J. Membr. Sci., 2008, 308, 250–257. 180. S. Tosti, A. Basile, F. Borgognoni, V. Capaldo, S. Cordiner, S. Di Cave, F. Gallucci, C. Rizzello, A. Santucci and E. Traversa, Low-temperature ethanol steam reforming in a Pd-Ag membrane reactor: Part 2. Pt-based and Ni-based catalysts and general comparison, J. Membr. Sci., 2008, 308, 258–263. 181. S. Tosti, F. Borgognoni and A. Santucci, Multi-tube Pd-Ag membrane reactor for pure hydrogen production, Int. J. Hydrogen Energy, 2010, 35, 11470–11477. 182. W. H. Lin, Y. C. Liu and H. F. Chang, Hydrogen production from oxidative steam reforming of ethanol in a palladium-silver alloy composite membrane reactor, J. Chin. Inst. Chem. Eng., 2008, 39, 435–440.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

84

Chapter 11

183. A. Iulianelli, T. Longo, S. Liguori, P. K. Seelam, R. L. Keiski and A. Basile, Oxidative steam reforming of ethanol over Ru-Al2O3 catalyst in a dense Pd-Ag membrane reactor to produce hydrogen for PEM fuel cells, Int. J. Hydrogen Energy, 2009, 34, 8558–8565. 184. A. Iulianelli, S. Liguori, V. Calabro, P. Pinacci and A. Basile, Partial oxidation of ethanol in a membrane reactor for high purity hydrogen production, Int. J. Hydrogen Energy, 2010, 35, 12626–12634. 185. W. H. Lin, Y. C. Liu and H. F. Chang, Autothermal reforming of ethanol in a Pd-Ag/Ni composite membrane reactor, Int. J. Hydrogen Energy, 2010, 35, 12961–12969. 186. F. Gallucci, A. Basile, S. Tosti, A. Iulianelli and E. Drioli, Methanol and ethanol steam reforming in membrane reactors: An experimental study, Int. J. Hydrogen Energy, 2007, 32, 1201–1210. 187. A. Basile, A. Parmaliana, S. Tosti, A. Iulianelli, F. Gallucci, C. Espro and J. Spooren, Hydrogen production by methanol steam reforming carried out in membrane reactor on Cu/Zn/Mg-based catalyst, Catal. Today, 2008, 137, 17–22. 188. A. Iulianelli, T. Longo and A. Basile, Methanol steam reforming in a dense Pd-Ag membrane reactor: The pressure and WHSV effects on COfree H2 production, J. Membr. Sci., 2008, 323, 235–240. 189. Y. M. Lin and M. H. Rei, Study on the hydrogen production from methanol steam reforming in supported palladium membrane reactor, Catal. Today, 2001, 67, 77–84. 190. A. Basile, S. Tosti, G. Capannelli, G. Vitulli, A. Iulianelli, F. Gallucci and E. Drioli, Co-current and counter-current modes for methanol steam reforming membrane reactor: Experimental study, Catal. Today, 2006, 118, 237–245. 191. M. K. Moharana, N. R. Peela, S. Khandekar and D. Kunzru, Distributed hydrogen production from ethanol in a microfuel processor: Issues and challenges, Renew. Sustain. Energy Rev., 2011, 15, 524–533. 192. A. L. Mejdell, T. A. Peters, M. Stange, H. J. Venvik and R. Bredesen, Performance and application of thin Pd-alloy hydrogen separation membranes in different configurations, J. Taiwan Inst. Chem. Eng., 2009, 40, 253–259. 193. A. S. Damle, Hydrogen production by reforming of liquid hydrocarbons in a membrane reactor for portable power generation–Experimental studies, J. Power Sources, 2009, 186, 167–177. 194. A. S. Damle, Hydrogen production by reforming of liquid hydrocarbons in a membrane reactor for portable power generation–Model simulations, J. Power Sources, 2008, 180, 516–529. 195. M. A. Rakib, J. R. Grace, C. J. Lim, S. S. E. H. Elnashaie and B. Ghiasi, Steam reforming of propane in a fluidized bed membrane reactor for hydrogen production, Int. J. Hydrogen Energy, 2010, 35, 6276–6290. 196. L. Capobianco, Z. Del Prete, P. Schiavetti and V. Violante, Theoretical analysis of a pure hydrogen production separation plant for fuel cells dynamical applications, Int. J. Hydrogen Energy, 2006, 31, 1079–1090.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

Palladium-based Membranes in Hydrogen Production

85

197. S. Sa, H. Silva, J. M. Sousa and A. Mendes, Hydrogen production by methanol steam reforming in a membrane reactor: Palladium vs carbon molecular sieve membranes, J. Membr. Sci., 2009, 339, 160–170. 198. A. Iulianelli, T. Longo and A. Basile, Methanol steam reforming reaction in a Pd-Ag membrane reactor for CO-free hydrogen production, Int. J. Hydrogen Energy, 2008, 33, 5583–5588. 199. I. Wieland, I. Melin and I. Lamm, Membrane reactors for hydrogen production, Chem. Eng. Sci., 2002, 57, 1571–1576. 200. H. Amandusson, L.-G. Ekedahl and H. Dannetun, Methanol-induced hydrogen permeation through a palladium membrane, Surf. Sci., 1999, 442, 199–205. 201. A. Basile, G. F. Tereschenko, N. V. Orekhova, M. M. Ermilova, F. Gallucci and A. Iulianelli, An experimental investigation on methanol steam reforming with oxygen addition in a flat Pd-Ag membrane reactor, Int. J. Hydrogen Energy, 2006, 31, 1615–1622. 202. H. Amandusson, L. G. Ekedahl and H. Dannetun, The effect of CO and O2 on hydrogen permeation through a palladium membrane, Appl. Surf. Sci., 2000, 153, 259–267. 203. R. Dittmeyer, V. Hollein and K. Daub, Membrane reactors for hydrogenation and dehydrogenation processes based on supported palladium, J. Mol. Catal. A: Chem., 2001, 173, 135–184. 204. W. Liang and R. Hughes, The catalytic dehydrogenation of isobutane to isobutene in a palladium/silver composite membrane reactor, Catal. Today, 2005, 104, 238–243. 205. H. Weyten, J. Luyten, K. Keizer, L. Willems and R. Leysen, Membrane performance: the key issues for dehydrogenation reactions in a catalytic membrane reactor, Catal. Today, 2000, 56, 3–11. 206. R. Dittmeyer, V. Hollein, P. Quicker, G. Emig, G. Hausinger and F. Schmidt, Factors controlling the performance of catalytic dehydrogenation of ethylbenzene in palladium composite membrane reactors, Chem. Eng. Sci., 1999, 54, 1431–1439. 207. J. N. Keuler and L. Lorenzen, The dehydrogenation of 2-butanol in a Pd-Ag membrane reactor, J. Membr. Sci., 2002, 202, 17–26. 208. M. H. Khademi, M. R. Rahimpour and A. Jahanmiri, Differential evolution (DE) strategy for optimization of hydrogen production, cyclohexane dehydrogenation and methanol synthesis in a hydrogenpermselective membrane thermally coupled reactor, Int. J. Hydrogen Energy, 2010, 35, 1936–1950. 209. M. H. Khademi, A. Jahanmiri and M. R. Rahimpour, A novel configuration for hydrogen production from coupling of methanol and benzene synthesis in a hydrogen-permselective membrane reactor, Int. J. Hydrogen Energy, 2009, 34, 5091–5107. 210. M. H. Khademi, P. Setoodeh, M. R. Rahimpour and A. Jahanmiri, Optimization of methanol synthesis and cyclohexane dehydrogenation in a thermally coupled reactor using differential evolution (DE) method, Int. J. Hydrogen Energy, 2009, 34, 6930–6944.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00040

86

Chapter 11

211. F. R. Garcia-Garcia, Y. H. Ma, I. Rodriguez-Ramos and A. GuerreroRuiz, High purity hydrogen production by low temperature catalytic ammonia decomposition in a multifunctional membrane reactor, Catal. Commun., 2008, 9, 482–486. 212. M. R. Rahimpour, H. R. Mottaghi and M. M. Barmaki, Hydrogen production from urea wastewater using a combination of urea thermal hydrolyser-desorber loop and a hydrogen-permselective membrane reactor, Fuel Process. Technol., 2010, 91, 600–612. 213. M. R. Rahimpour and A. Asgari, Production of hydrogen from purge gases of ammonia plants in a catalytic hydrogen-permselective membrane reactor, Int. J. Hydrogen Energy, 2009, 34, 5795–5802. 214. M. R. Rahimpour and A. Asgari, Modeling and simulation of ammonia removal from purge gases of ammonia plants using a catalytic Pd-Ag membrane reactor, J. Hazard. Mater., 2008, 153, 557–565. 215. Y. Shirasaki, T. Tsuneki, Y. Ota, I. Yasuda, S. Tachibana, H. Nakajima and K. Kobayashi, Development of membrane reformer system for highly efficient hydrogen production from natural gas, Int. J. Hydrogen Energy, 2009, 34, 4482–4487. 216. H. Kurokawa, Y. Shirasaki and I. Yasuda, Energy-Efficient Distributed Carbon Capture in Hydrogen Production from Natural Gas, Energy Proc., 2011, 4, 674–680. 217. D. Barba, F. Giacobbe, A. De Cesaris, A. Farace, G. Iaquaniello and A. Pipino, Membrane reforming in converting natural gas to hydrogen (part one), Int. J. Hydrogen Energy, 2008, 33, 3700–3709. 218. Tecnimont KT, Personal communication with E. Lollobattista, 2010. 219. A. Nijmeijer, E. Engwall, J. Saukaitis and A. Del Paggio, Presentation at the 11th International Conference on Inorganic Membranes, Washington, 2010. 220. E. Engwall, A. Nijmeijer, J. Saukaitis and A. Del Paggio, Presentation at the ACS 2010 Meeting, Boston, MA, 2010. 221. Pall Corporation, Personal communication with A. S. Damle, 2010. 222. F. T. Rusting, G. de Jong, P. P. A. C. Pex and J. A. J. Peters, EP 1128118, Sealing socket and method for arranging a sealing socket to a tube, 2001. 223. www.hysep.com, 2010. 224. Z. Chen, J. R. Grace, C. J. Lim and A. Li, Experimental studies of pure hydrogen production in a commercialized fluidized-bed membrane reactor with SMR and ATR catalysts, Int. J. Hydrogen Energy, 2007, 32, 2359–2366. 225. A. Li, J. R. Grace and C. J. Lim, Preparation of thin Pd-based composite membrane on planar metallic substrate: Part II. Preparation of membranes by electroless plating and characterization, J. Membr. Sci., 2007, 306, 159–165. 226. A. Li, J. Grace and C. J. Lim, Preparation of thin Pd-based composite membrane on planar metallic substrate: Part I: Pre-treatment of porous stainless steel substrate, J. Membr. Sci., 2007, 298, 175–181.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

CHAPTER 12

Membrane Reactors in Hydrogen Production A. BRUNETTI,a G. BARBIERI*a AND E. DRIOLIa,b a

Institute on Membrane Technology (ITM-CNR), National Research Council, c/o The University of Calabria, Cubo 17C, Via Pietro Bucci, 87036 Rende CS, Italy; b Department of Chemical Engineering and Materials, The University of Calabria, Cubo 44A, Via Pietro Bucci, 87036 Rende CS, Italy

12.1 Introduction In the last decade, energy demand has been growing by 1.2% a year and the fossil fuels still maintain a production share of ca. 75%. However, the ever stricter problems connected to a sustainable growth and to a lower environmental impact lead to the conclusion that the time of easy oil consumption is finished. Nowadays, the necessity to release energy production from oil and natural gas as primary energy sources is becoming more and more pressing. Indeed, more in general, diversifying such sources in order to assure supply makes the interest in membrane reactor (MR) technology more urgent. Moreover, the increasing effort dedicated to the reduction of environmental problems has recently led to the development of clean technologies, designed to enhance both the efficiency and environmental acceptability of energy production, storage and use, in particular for power generation.1 Among these technologies, the exploitation of light hydrocarbons is surely the main realistic energy source, since they allow both power generation and environmentally friendly fuel production. Specific reference should be made to hydrogen in this context. Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

87

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

88

Chapter 12

At present, global hydrogen production mainly relies on processes that extract hydrogen from fossil fuel feedstock. 96% of hydrogen is directly produced from fossil fuels and about 4% is produced indirectly by using electricity generated through them.2 The stream coming from a reformer or a coal gasification plant contains around 50% hydrogen (on a dry basis) that must be recovered and between 40 and 45% CO that, usually, is reduced in an upgrading stage, producing more hydrogen at the same time. In traditional applications (Figure 12.1), the upgrading of reformate streams is performed by using a multi-stage CO-shift process based on a series of catalytic reactors: the first, operating at high temperatures (about 350–400 1C) and taking advantage of the high reaction rate, converts a large portion of the carbon monoxide to give hydrogen and CO2; the other, operating at a low temperature (around 220–300 1C), refines the carbon monoxide conversion, thus allowing a lower final concentration of CO (less than 1% molar).3 This H2 rich stream coming out from the last reactor is fed to a pressure swing adsorption (PSA) unit for H2 separation from the other gases. It should be pointed out that the new utilization of H2 as feed in fuel cells for mobile power sources requires the anode inlet gas to have a CO concentration lower than 10–20 ppm4 in order to avoid catalyst poisoning with consequent drops in the fuel cell efficiency. Hence, the purification step of the H2 produced from hydrocarbon must be very efficient to fulfil the fuel cell requirements. Because of this in some cases, another reaction unit is added for oxidizing CO in CO2. One of the main challenges in the next few years will be the identification of new technologies able to provide a better exploitation of fossil fuels, e.g. hydrocarbons, in order to improve yield, energy saving and so on. The reduction of the reaction/separation/purification stages, which means lower footprint area occupied by the whole plant, less auxiliary devices required, reduction of the energetic load, and so forth, are fundamental aspects to be taken into account in the redesigning of hydrogen production processes. A promising approach for concretizing these technological aspects in the field of hydrogen production is the use of MRs, combining the reaction and H2

Steam Reforming

CO2 HT-WGS (350 – 400 °C)

LT-WGS (220 – 300 °C)

CO2 removal

CO clean up

Hydrocarbon

Hydrogen

Figure 12.1

Scheme of the traditional process for hydrogen production from light hydrocarbons.

View Online

89

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

separation by means of a selective membrane. Many studies are now focused on the analysis of MR performances where light hydrocarbon reforming or water gas shift (WGS) reaction are carried out. In these cases, for both these reactions the presence of the membrane allows the recovery of a hydrogen reach stream which does not require further separation. Moreover, the removal of the H2, reaction product, from the reaction volume shifts the reaction toward further conversion. This means the possibility of having an intensified process with a reduced plant size and higher yield. The traditional process can be thus redesigned as more compact and efficient (Figure 12.2) pursuing the logic of the process intensification strategy,1,2 which is an innovative methodology for process and plant design, proposing a new design philosophy to achieve significant reductions (by factors of 10 to 100 or more) in plant volume at the same production capacity or to improve overall efficiency. As Figure 12.2 shows, the integrated membrane system can be constituted of less reaction/separation units than the conventional one (Figure 12.1). A first MR can be used for carrying out the reforming of the light hydrocarbons and another one for the WGS reaction. The presence of the membrane in both the reactors allows the separation of a hydrogen rich stream from the two reaction volumes as well as the improvement in the conversion of the two stages. Obviously, the H2 purity level strictly depends on the membrane type used in both MRs. Actually, membranes can be distinguished by their selectivity which can be infinite or finite. The first ones, traditionally Pd-based, allow a pure hydrogen stream to be obtained, whereas the others provide a hydrogen-rich stream with a variable purity. If the recovered H2 stream does not have the purity required, the latter can be increased by adding another purification unit on the basis of the final use of the H2 stream. Selective CO oxidation is known as an interesting and economical approach for CO removal from H2-rich gas streams. Also in this field, new studies proposed in the literature have demonstrated how the use of an MR can improve the process by increasing CO conversion as well as the purity of the hydrogen stream.

CO2 concentrated stream

Natural gas

Water Gas Shift 300 °C

Steam reforming 500 °C Pure H2

Figure 12.2

Scheme of the integrated membrane process for hydrogen production from light hydrocarbons.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

90

Chapter 12

In this context, membrane engineering plays a fundamental role in the integration of these units in a single plant and, at the same time, in the definition of the knowledge necessary to drive the process by maximizing the gains both in terms of efficiency and plant size reduction. The synergic effects offered by MRs through combining reaction and separation in the same unit, their simplicity and the possibility of advanced levels of automation and control, offer an attractive opportunity to redesign industrial processes.5–7

12.2 Membranes and Membrane Reactors for Hydrogen Production MRs represent the most significant class of the so-called multifunctional reactors8 integrating reaction and separation in the same unit. The membrane, dividing the reactor into retentate (reaction side) and permeate volumes, can have three different functions (Figure 12.3):  Permselective separation  Catalytic  Distribution of a reactant in the reaction volume. A distinction might be made, in fact, when the MR is used to carry out a catalytic reaction, considering whether the membrane itself has a catalytic function or not. In the case of MRs for H2 production, most of the membranes used are permselective, which allows the selective removal of H2 from the reaction volume under the effect of a driving force. This is a function of the species partial pressures on both the membrane sides and can be created by means of an inert sweep gas in the permeate compartment (nitrogen, helium, water, etc.), or with the application of a pressure difference between the retentate and permeate sides. For Pd-alloy membranes, Sievert’s law (eqn (12.1)) is used worldwide for the mathematical description of H2 permeating flux in these types of membranes. The hydrogen permeating flux is a linear function of the permeability and driving force and reverse function of the membrane thickness. The permeation

Figure 12.3

A membrane reactor scheme.

View Online

Membrane Reactors in Hydrogen Production

91

driving force of Sievert’s law is the difference of the square root of the hydrogen partial pressure on the two membrane sides: qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Side Side  PPermeation PReactionm H2 H2 qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Permeability Side Side PReactionm ¼  PPermeation H2 H2 thickness

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Hydrogen permeating flux ¼ Permeance

ð12:1Þ The removal of a product such as hydrogen from the reaction volume implies a series of advantages:    

Conversion enhancement of equilibrium-limited reactions Depletion of undesired secondary reactions Recovery of a concentrated rich stream Coupling of two or more reactions, e.g. dehydrogenation (endothermic) with a hydrogenation (exothermic) on the two membrane sides  Better operating conditions (e.g. temperature).

The removal of a product from the reaction volume allows the thermodynamic equilibrium limit of a traditional reactor (TR) to be exceeded, obtaining higher conversion in analogous operating conditions. In other words, for endothermic reactions this allows the MR to achieve the same conversion of a TR, but at significantly lower temperatures. Another interesting aspect of MR use is the positive effect that the reaction pressure can have on the process, also for reactions taking place without mole number variation (e.g. WGS) or with a mole number increase (e.g. methane steam reforming, SMR), because of the favored removal of a product from the reaction volume. In hydrogen production dense or micro-porous membranes can be used depending on the role of the membrane, whether it be for H2 separation or purification. Most studies reported in the open literature show that the membranes can be distinguished in dense inorganic Pd-based membranes and ceramic membranes (silica, zeolite, etc.). The former show a permselective transport governed by a solution-diffusion mechanism. Micro-porous ceramic membranes can have both the permselective and non-permselective transport, depending on the size of the permeating molecules with respect to the membrane pore size as well as on the chemical nature of the permeating molecules and the membrane material.9 Table 12.1 reports a comparison between the main pros and cons of both the membrane types. The main advantage of the dense Pd-based membranes is their infinite selectivity versus H2. In the specific case, the important advantages of the Pd-alloy MR are the following:  Production of pure H2 permeate stream (when a pressure difference is used instead of a sweep gas to create the permeation driving force)  Retentate streams concentrate in the other species (e.g. CO2)

View Online

92

Chapter 12

Table 12.1

Advantages and drawbacks of dense and porous membranes

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Pd-based membranes

Microporous membranes

Advantages

Drawbacks

Advantages

Drawbacks

Infinite H2 selectivity High mechanical resistance High packing density

Low permeating fluxes High costs

High permeating fluxes High thermal resistance High chemical resistance No inhibition effect by gaseous species

Low selectivity Low packing density Fragility Instability due to H2O vapor

 Positive effect of reaction pressure on the equilibrium conversion also for reactions characterized by an increase in the number of moles. This also affects the performance of the catalyst in terms of activity and lifetime, also avoiding possible catalyst deactivation with significant advantages in terms of equipment costs. However, the high cost is actually the main problem that limits their diffusion at an industrial level. In order to overcome this drawback, several works are presenting new composite membranes consisting of a thin Pd-based layer deposited, with different techniques, on porous supports that can be ceramic or stainless steel. In this way, the Pd content is reduced as well as the related cost. Another typical problem of these dense membranes is the inhibiting or poisonous effects that some species like carbon monoxide10,11 or sulfur can exercise on the membrane, consequently reducing the permeance. On the contrary, the meso- and micro-porous membranes (alumina, silica, titania, zirconia, zeolites, etc.) are not affected by poisoning and are cheaper than the metallic membranes. However, they show finite selectivity versus all the chemical species, therefore it is not possible to obtain pure streams as permeate, as when using dense membranes.

12.3 Current and Potential Applications of Membrane Reactors for Hydrogen Production In the past, MR performances were studied by carrying out several gaseous phase reactions and using different membrane types, in particular for high temperature operations. Table 12.2 summarizes the main gaseous phase reactions for H2 production carried out in MRs, referring to some important works in the literature. It reports the membrane type used as well as the temperature range investigated in the experiments performed at laboratory scale. Most of the studies carried out so far on MRs have focused on equilibriumlimited reactions, where the permeation of a product enhances the conversion with respect to a TR. Other new applications propose the use of the membrane as a contactor for catalyst and reactants. However, even the MR studies also on

View Online

93

Membrane Reactors in Hydrogen Production

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Table 12.2

List of some membranes used in MR; the temperature range and the main gaseous phase reactions are also reported

Process

Membrane

Temperature range (1C)

Steam reforming of light hydrocarbons

Pd-Ag

400–600

Composite Pd-ceramic Silica supported Alumina Perovskite ceramics

800

Pd on silicalite Pd-ceramic Pd-Ag

700–750 250–500 200–400

Silica Ceramic Pd(60%)-Cu Catalytic zeolite Composite Pd-ceramic

180–250 350–500

13 14, 15 16, 17, 18, 19, 20, 21, 22, 23, 24, 25, 26 27 28 29 30, 31, 32 33, 34, 35

500 450–550 350–450 350–500 450–550 600

36 37 38 39 40, 41 42

600–700 800–900 450–600

43, 44 45, 46 47

Partial oxidation of light hydrocarbons Water gas shift

CO clean-up Dehydrogenation of light hydrocarbons

Composite Pd-metal Silica oxide ceramic Metal-modified alumina Pd-Ag Zeolite Oxidative dehydrogenation Pd-Ag of light hydrocarbons Metal–ceramic composite Perovskite Dehydrogenation of Composite Pd-ceramic ammonia

Reference 1, 2, 3, 4, 5, 7, 8, 9 10 11 11 12

the pilot plant gave promising results indicating wide-ranging potential for this technology; there are currently no large scale applications for their application. The next sections will focus on providing an overview on the MR technology application, considering only the main reaction stages constituting the traditional plant for hydrogen production from the reforming of light hydrocarbons (Figure 12.1).

12.3.1

Steam Reforming of Methane and other Light Hydrocarbons

Steam methane reforming (SMR) is the most common and cost-effective method for hydrogen production: 2H2 O þ CH4 ¼ CO2 þ 4H2

0 DH298 ¼ 191 kJ mol1

SMR is carried out commercially in externally heated fixed bed reactors on a very large scale. It is a reversible process usually operated at 700–900 1C using

View Online

94

Chapter 12

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

12

supported nickel catalysts. Owing to mass and heat transfer issues, the fixed-bed SMR suffers from significant disadvantages, such as low catalyst effectiveness factors and large temperature gradients.12 In addition, in most cases, to obtain hydrogen of the purity required, the reactor off-gas must go through a series of steps, such as high and low WGS reactions and PSA. In the last 20 years, membrane technology has been proposed as an alternative to improve the performance of the conventional process. Hydrogen can be produced in membrane steam reformer units. Owing to the selective and continuous removal of hydrogen in situ from the reaction zone, SMR reaction shifts towards the product side. Therefore also the feed pressure, which has a negative effect from a thermodynamic point of view, acts positively on the methane conversion, so that compared with the TRs, MRs can achieve higher conversions at the same temperature or the same conversion at lower temperatures.13 Moreover, the use of dense Pd-based membranes allows a pure hydrogen stream to be obtained so that there is no need for additional purification as in the conventional process. Up to now, many researchers have proposed applying Pd-based MRs for SMR reaction operating at milder operative conditions than the traditional reactors.5,12,14–22 The operating temperature of 500–550 1C used in most experiments is a compromise among several factors. Both membrane permeance and the SMR thermodynamics and kinetics are favored by higher temperatures but the membrane is more stable at lower temperatures.20 Figure 12.4 shows an example of the typical trend of CH4 conversion in MR as a function of the temperature for different feed conditions.

Figure 12.4

Methane conversion against temperature for membrane reactor. Comparison between experimental data (symbols) and model results (lines) for a 40 SCCM sweep flow rate. Reprinted from G. Barbieri, G. Marigliano, E. Drioli, Simulation of steam reforming process in a catalytic membrane reactor, Ind. Eng. Chem. Res., 36, 6, 2001, with permission of American Chemical Society.

View Online

95

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

The promising results obtained at laboratory scale encouraged application on a larger scale. MRT21 has developed a proven technology based on a patented fluidized-bed MR for high-purity hydrogen. The process combines hydrocarbon reforming, shift conversion and hydrogen purification in a single step. One unit supplied to Tokyo Gas gave improved performance in comparison with fixed bed membrane reformers. Another unit was also provided to BOC in a project sponsored by the S-DoE. Actually, the units are initially geared to capacities in the 15–50 N m3 h1 range, the technology can be readily scaled to higher and lower capacity as applications require. Shell Oil Company has recently patented22 a process and apparatus for the production of pure hydrogen by steam reforming. This process integrates the steam reforming and shift reaction to produce pure hydrogen with minimal production of CO and virtually no CO in the hydrogen stream, provides for CO2 capture by sequestration, uses a steam reforming MR and is powered by heat from a heater convection section. Hydrogen production by steam reforming of methanol, ethanol and other light hydrocarbons has become an attractive alternative to traditional operations. Especially attractive is their use in the decentralized production of clean electrical energy from fuel cells. The main differences as against the reforming of light hydrocarbons are the catalyst types used and the product distribution in the two reaction systems. Recent studies concern MR use in these reactions types. The main advantage is in performing both reaction and pure hydrogen recovery in the same device thus replacing the conventional system (reformer þ gas cleaning unit) with one MR.23 In both conventional and membrane systems, the main reaction products are hydrogen, carbon monoxide and carbon dioxide. Depending on the fuel used (e.g. ethanol, bioethanol) methane, acetaldehyde and ethylene could be also present. However, these reformers are still being studied only at laboratory scale.

12.3.2

Water Gas Shift

The water gas shift (WGS) reaction is a fundamental step for the upgrading of the streams exiting a reformer. The WGS role is in increasing H2 yield and decreasing the concentration of CO, which is, moreover, a poison for some catalysts used in downstream processing, such as for example ammonia synthesis as well as for the PEMFC electrodes. As already mentioned, WGS reaction is traditionally carried out in two fixed bed adiabatic reactors, operating at high (300–400 1C) and low (200–300 1C) temperature respectively, connected in series with a cooler between them. WGS reaction is exothermic and characterized by no variation in the number of moles. Thus, CO equilibrium conversion is favored by a low temperature and it does not depend, in a TR, on the reaction pressure: H2 O þ CO ¼ CO2 þ H2

0 DH298 ¼ 41 kJ mol1

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

96

Chapter 12

On the contrary, in an MR the reaction pressure has a positive effect on the equilibrium conversion since it facilitates permeation and therefore pushes the reaction towards further product formation.24,25 The use of an MR thus allows a higher conversion to be reached also at a higher temperature where the thermodynamic conversion is low, acting positively on the kinetics. As a consequence, the catalyst amount necessary for a given conversion can be significantly reduced.28 In the open literature, many studies were focused on MRs with palladium membranes. Seok and Hwang42 evaluated the performance of the WGS reaction by using Vycor glass coated with ruthenium(III) chloride trihydrate. The reaction was carried out under various operating temperatures, pressure and feed composition. The highest CO conversion obtained was 85% (equilibrium value 99.9%) at relatively low temperature (170 1C) and a high sweep factor equals 10, which means a sweep flow rate five times higher than the feed one. Complete conversion (100%) was obtained by Kikuchi et al.25 and Uemiya et al.26 at 400 K using a tube-in-tube MR, in which the inner tube was a thin palladium film, also by using a sweep factor equals 10. Tosti et al.27 added silver to palladium for decreasing membrane embrittlement and increasing the hydrogen permeability. They developed a WGS MR with a Pd-Ag film (50 mm thick) coated on the inside wall of a ceramic porous tube and achieved reaction conversions close to 100% (well above the equilibrium value of 80%) at 325–330 1C, owing to a high sweep gas flow rate. However, the majority of these studies proposed sweep gas use to promote H2 permeation, whilst only a few combine the use of a low feed pressure at the sweep gas for improving the permeation. Barbieri and co-workers studied the effect of feed pressure on MR performances, in different conditions of feed and MR configuration using in

Figure 12.5

CO conversion measured at 300 1C as a function of the feed pressure in a traditional reactor and in Pd-Ag membrane reactor. Reprinted from A. Brunetti, G. Barbieri, E. Drioli, Pd-based membrane reactor for syngas upgrading, Energy Fuels, 23, 10, 2009, with permission of American Chemical Society.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

97

addition to Pd-based membranes also silica supported on porous stainless steel ones.28,43 Figure 12.5 shows an example of the dependence of CO conversion on the feed pressure, in the upgrading of syngas stream using a Pd-alloy membrane reactor (MR) packed with a commercial catalyst CuO/CeO2/Al2O3. A Pdbased MR operated successfully exceeding the thermodynamic constraints of a traditional reactor and, specifically, the drawback introduced by the presence of hydrogen. In fact, a 90% CO conversion significantly exceeded (three times) the TR conversion and also the thermodynamics upper limit (o36%), owing to ca. 80% hydrogen permeated through the membrane. A fundamental aspect in the MR design is the effect of Damko¨hler’s number – the ratio between reaction rate at the reactor entrance and feed flow rate – on the MR performances28:

Da ¼

rCO jz¼0  VReaction Characteristic space time ¼ Feed Characteristic reaction time FCO

ð12:2Þ

Figure 12.6 shows CO conversion and H2 recovery of a tubular Pd-Ag MR in which the syngas stream upgrading by means of WGS reaction is performed. The simulations carried out with a 1D model analyse the effect of the both CO conversion and H2 recovery increase along the MR up to a plateau, which is reached faster for a higher Da. In particular, a high Damko¨hler (higher space time) favors CO conversion and thus the H2 produced, allowing higher H2 recovery to be obtained. On the contrary, a low Damko¨hler (reaction time much higher than space time) indicates that less H2 is produced by the reaction with consequent low recoveries.

Figure 12.6

CO conversion and H2 recovery profiles as a function of reactor length at different Damko¨hler’s numbers. Furnace temperature ¼ 280 1C. Feed pressure ¼ 1000 kPa. Permeate pressure ¼ 100 kPa.

View Online

98

Chapter 12

CO Conversion,-

1,500 kPa

PFeed 400 kPa

0.5

TREClocal TR

600

Temperature, °C

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

1

1,500 kPa

400

400 kPa TR P

200

0

Feed

0.5

1

Dimensionless Length, -

Figure 12.7

CO conversion and temperature profiles as a function of reactor at different feed pressures. Furnace temperature ¼ 280 1C. Damko¨hler ¼ 1.

By means of the model it was possible to analyse the temperature profiles correspondent to CO conversion and compare them with TR results, at set Damko¨hler, temperature and different feed pressures. As Figure 12.7 shows, the CO conversion profile increases along the reactor length. As the feed pressure increases, higher CO conversions are obtained at a shorter MR length, owing to the positive effect of the pressure on H2 permeation. Temperature profiles increase in the first part of the MR and then decrease because the heat exchange is higher than the heat generated by the reaction. Therefore, the maximum temperature and, thus, the slope of the temperature profile after the maximum depends on CO conversion establishing the heat production. The temperature reduces when CO conversion tends to a plateau, since when CO is almost completely converted the net heat flux leaves the system by conductive exchange with the external environment. As in the case of SMR the interesting results achieved at laboratory level arose in various, also industrial, patents.29–35 Among them, United Technologies Corporation29 patented the use of a WGS MR, comprising a WGS reaction region and a permeate volume, separated by an H2-separation membrane which allows H2 formed over a catalyst in the reaction region to be passed

View Online

99

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

selectively to the permeate region for delivery to a point of use (such as the fuel cell of a fuel cell power plant). Exxonmobil30 developed and patented a heat exchanged MR for electric power generation comprising an MR system that employs catalytic or thermal SMR and a WGS on one side of the membrane and hydrogen combustion on the other side for producing electricity. Also, General Electric Company31 patented a polygeneration system including (i) a syngas generator for producing a syngas, (ii) a syngas enrichment unit for removing undesired species from the syngas for producing an enriched syngas, and (iii) a syngas utilization system that utilizes the enriched syngas to produce useful products. In some embodiments, the polygeneration system includes MR, catalytic burner and power generation unit.

12.3.3

Carbon Monoxide Clean-up

Selective or preferential CO oxidation (Selox) is known as an interesting and economic approach for CO removal from H2-rich gas streams produced by reforming processes or available in petrochemical plants (e.g. in ethylene process), opening their use to some other operations, such as, ammonia synthesis, hydrogenations and fuel cell applications (proton-exchange membrane fuel cells). Traditionally, the reaction is carried out in conventional reactors; however, the presence of H2O and CO2 in the feed causes a significant decrease in the activity of the catalyst.36 In an integrated plant for hydrogen production, depending on the types of membranes (e.g. of silica) used in the SMR and WGS MRs the H2 stream recovered as permeate could require a further purification – particularly whether the final user is a PEMFC – to reduce the CO level to less than 10 ppm (Figure 12.8). The CO selox becomes thus a fundamental stage also in the integrated membrane systems for H2 production. Recently, in the literature it was demonstrated that the use of an MR for this reaction stage can improve the depletion of CO content. The membrane, constituted of a ceramic tube, most often zeolite, on which is deposited the catalyst, opportunely distributed in the pore, does not have the function of separating/purifying a stream, but to improve the reactant/catalytic phase contact, to reduce by-passing SMR

WGS CO2 rich stream

Methane CO selox Hydrogen CO free

Figure 12.8

Scheme of the integrated membrane process for hydrogen production from light hydrocarbons with a CO selox membrane unit for permeate purification.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

100

Chapter 12

and misdistribution generally shown in a packed bed and to provide a hot spot control that is beneficial since a local temperature increase can lead to promotion of reverse WGS reaction.37 Bernardo et al.38 successfully carried out the CO Selox in the presence of large quantities of hydrogen in continuous flow-through MRs using different Pt-loaded catalytic zeolite membranes. The catalytic MRs succeeded in reducing the outlet CO concentration from 10 000 ppm (1% molar) down to 10–50 ppm. These results confirmed the good potentiality of catalytic membranes for a deep purification of H2-rich streams, allowing the hydrogen final use, e.g. also in fuel cell applications.

12.4 New Indexes for the Comparison of Membrane and Traditional Reactors Many efforts have been made to measure progress towards sustainability, and others are still in progress to define indicators of industrial process and, in particular their impact on three specific areas.39–41 Membrane operations are well-known for their modularity, compactness and flexibility, therefore they can be considered as new operations developed in the logic of process intensification. Recently, new metrics for comparing membrane performances with those of conventional units have been introduced.39 With respect to the already existing indicators, these new metrics take into account the size, the weight, the flexibility, the yield and the modularity of the plants. They are useful for having an immediate indication of the eventual gain that a membrane operation can offer with respect to a conventional one. Particularly for MRs, some specific indexes were introduced, taking into account, among the several advantages connected to their use, the use related to the conversion improvement, which means better exploitation of raw materialand plant size reduction. This allows a better performance to be achieved than with a TR, with reduced reaction volumes and separation loads. These new metrics are:  Volume index (VI), defined as the ratio of the catalytic volume of an MR and that of a TR for reaching a set conversion of the reference reactant  Conversion index (CI), the ratio between the conversion of an MR and a TR, at a set reaction volume  Extraction index (EI), the ratio between H2 permeated through the membrane with respect to that totally fed to the reactor.

12.4.1

Case Study: Water Gas Shift Reaction in a Membrane Reactor

As widely described above (see Section 12.3.2), WGS is a fundamental stage in H2 production plants. Therefore, it was chosen as a case study because it can

View Online

101

well represent the possible advantages offered by MR use with respect to TRs. However, these indexes are useable in all reactions besides the WGS. The Volume index represents an important parameter in installing new plants. Future plants must be characterized by low sizes and high productivities: the VI is an indicator of the modularity of an MR and it compares the MR reaction volume with that of a TR, necessary to achieve a set conversion:  ðReaction VolumeÞMR  Volume Index ¼ VI ¼  ðReaction VolumeÞTR 

ð12:3Þ Set CO Conversion

VI ranges from 0 to 1. A low VI means that the reaction volume, required by an MR for reaching a set CO conversion, is much lower than that required by a TR. As a consequence, the catalyst weight necessary in MR is significantly reduced. In a Pd-Ag MR the hydrogen permeation through the membrane is promoted by the driving force, it being the difference of the square roots of the partial pressure on the two membrane sides. The increase of the feed pressure, keeping constant the atmospheric pressure on the permeate side, constitutes a possible solution for creating this driving force. Figure 12.9 shows the VI as a function of the feed pressure for a Pd-Ag MR where the WGS reaction is carried out. As can be seen, the higher the feed pressure the lower the VI, owing to the positive effect that it has on CO conversion in an MR. MR reaction volume is three quarters of that of TR at 600 kPa and goes down to one quarter at 1500 kPa, when an equimolecular mixture is fed and a final conversion of B80% is considered. VI further decreases when a stream coming out from a reformer is fed (50% H2, 10% CO2, 20% CO, 20% H2O) into the Pd-Ag MR, owing to the low value of the equilibrium conversion (35%) (Figure 12.10). As a consequence, the amount of catalyst necessary to reach a suitable conversion is drastically reduced with clear gain also in terms of plant size reduction.

50

Traditional Reactor

100

Volume Index, %

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

75%

40% 25%

0 Pd-based Membrane Reactor

200

600

1,000

1,500

Feed Pressure, kPa

Figure 12.9

Volume index as a function of feed pressure feeding an equimolecular mixture. Furnace temperature ¼ 280 1C. CO conversion sets at 90% of the TREC.

View Online

102

Chapter 12 100 Equimolecular mixture

Volume Index, %

50

0 0

Figure 12.10

400 800 Feed Pressure, kPa

3,000

Volume index as a function of feed pressure with two different feeds. Furnace temperature ¼ 280 1C. CO conversion sets at 90% of the TREC.

100 PFeed=1000 kPa

Volume Index, %

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Syngas mixture

50 Equimolecular mixture

Syngas mixture

0 200

Figure 12.11

300 250 Temperature, °C

350

Volume index as a function of feed pressure with two different feeds. Feed pressure ¼ 1000 kPa. CO conversion sets at 90% of the TREC.

Another operating parameter that strongly affects the reduction of the reaction volume is the temperature. As shown by Figure 12.11 feeding both an equimolecular mixture and a syngas stream, the higher the temperature the lower the VI. This must be attributed to the positive effect of the temperature on the hydrogen permeation, which implies higher CO conversion and, in other terms, less catalyst volume required for getting a set CO conversion. Of course,

View Online

103

this effect is much more emphasized in the case of a reformate feed stream containing also hydrogen. Another improvement in the MR configuration that leads to a further reduction of the VI was proposed by Barbieri et al.44 When the feed mixture enters the WGS stage it has a high CO content, such as in the case of the streams coming out from coal gasification, the traditional MR configuration does not allow the best exploitation of the whole membrane area because of the low H2 partial pressure at the inlet of the MR. For this reason the authors proposed the membrane only in the second part of the catalytic bed (Figure 12.12). In this way, a good exploitation of the whole membrane area for the permeation is assured. A VI of approx. 85% was calculated at 400 kPa and 280 1C for the ‘conventional’ MR; this innovative MR has shown a VI of about 65%, already at 300 kPa and, at 600 kPa, it has gone down at 40%, a lower value than the one shown by simple MR (65%) at the same operating conditions (Figures 12.13 and 12.14). This suggests that the innovative solution allows the problems related to the good exploitation of the membrane area to be overcome, consequently a further reaction volume reduction with respect to that achieved with the traditional MR is achieved. From the analysis reported above on a Pd-Ag MR it can be pointed out that, globally, the MR offers better performance than TR; similar behaviour can be Retentate CuO/CeO2 based Catalyst v

Feed

v

v

v

vv

TR section

Figure 12.12

Pd-Ag membrane

Permeate Pure H2

MR section

Innovative configuration of membrane reactor.

Figure 12.13

Traditional Reactor

PFeed= 600 kPa Feed: only CO and H2O

Volume Index, %

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

65% 40%

Comparison of the volume index of typical and innovative MR configuration. Elsevier, owner of the copyright of its original publication in Journal of Power Sources,44 is acknowledged for permission to use this figure.

View Online

~ 50 %

~ 30 % Pd-Ag MR

Silica based MR

Traditional Reactor

Chapter 12

Volume Index, %

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

104

Membrane reactor

Figure 12.14

Reaction volume reduction for different membrane types. Furnace Temperature ¼ 280 1C. Feed pressure ¼ 1200 kPa.

observed using hydrogen selective membranes different from Pd-Ag, such as silica. From the calculation, the reaction volume required by an MR with the silica membrane is half of the TR one but higher than that of a Pd-Ag MR, at the same operating conditions. This effect might be attributed to the fact that whereas Pd-Ag shows infinite hydrogen, a micro-porous membrane is permeable also by the other gases; hence more catalyst is required to reach the same CO conversion. The capability of reaching a conversion higher than a TR, exceeding the TR equilibrium limits is a typical property of an MR. The conversion index, defined as the ratio between the conversion achieved in an MR and that of a TR, for a set reaction volume, provides an evaluation of the gain in terms of conversion and its use is particularly indicated when the feed mixture also contains reaction products. A high CI implies a relevant gain in terms of conversion achieved in an MR with respect to that of the conventional reactor, with the same reaction volume, meaning better raw material exploitation and lower wastage:  ðConversionÞMR  Conversion Index ¼ CI ¼ ð12:4Þ  ð ConversionÞTR  Set Reaction Volume

MRs are pressure driven systems. Therefore, the higher the feed pressure the higher CI, as shown by Figure 12.15. In particular, when a reformate stream is fed into the Pd-Ag MR, the CI passes from 2 to 6, ranging the feed pressure from 200 kPa to 1500 kPa. However, already at 400 kPa, the CI is 5. When an equimolecular feed containing mainly reactants is fed to MR, the CI ranges from 1.5 to 1 since the TR conversion is itself high. However, a CI of 1.5 indicates around 95% CO conversion, implying not only a pure H2 stream in the permeate side but also a CO2 concentrated retentate stream and much easier CO2 recovery. Also considering the silica MR, fed with the syngas mixture, the

View Online

105

Membrane Reactors in Hydrogen Production 8

Conversion Index, -

4

Equimolecular mixture

0

0

1,000

2,000

3,000

Feed Pressure, kPa

Figure 12.15

Conversion index as a function of the feed pressure for different feeds. Furnace temperature ¼ 280 1C. Pd-Ag MR Silica MR

6 Conversion Index, -

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Syngas mixture

4

2

0

Figure 12.16

0

500 1,000 1,500 Feed Pressure, kPa

Conversion index as a function of the feed pressure for Pd-Ag and silica MR. Furnace temperature ¼ 280 1C.

CI ranges from 3 to 5 even if it is lower than the Pd-Ag MR owing to the permeation of the other gaseous species than H2 comprising the reactants, which causes a conversion depletion (Figure 12.16). In MR technology, the quantification of H2 recovered with respect to that totally extractable in the feed is an important issue. The extraction index (eqn (12.5)) defined as the ratio between H2 permeated through the membrane with respect to that totally fed to the reactor, gives an indication about the limitations of an MR in the achievement of a complete conversion. If the hydrogen is the permeating species, as in the case study considered, EI takes into account

View Online

106

Chapter 12

¼

a¼

8 >

:

per

FiFeed

FiPermeate Feed þ a  FLimiting

reactant

ð12:5Þ

mo1 fFor WGSg

1

per

m1

Feed The term a  FLimiting reactant of eqn (12.5) considers the maximum H2 extractable from the chemicals present in the system. The coefficient a takes into account the defecting reactant (CO or H2O) by means of the feed molar ratio H2O/CO. This coefficient equals the feed molar ratio if this latter is lower than 1 (CO defecting with respect to H2O), and equals 1 when CO exceeds the H2O. As defined, EI depends on the membrane properties, feed molar ratio and CO conversion achieved in the MR, at set operating conditions. In particular, EI increases with the temperature as well as the feed pressure, owing to the positive effect that they have on the permeation. Figure 12.17 shows some experimental results achieved by using a Pd-Ag MR fed with a syngas mixture. The highest EI measured is 75% at 600 kPa and 325 1C, which means that 75% of the hydrogen totally available in the feed, as molecule and also as reactant, was recovered as pure stream in the permeate side. This highlights the significant extractive capability of the MR that, thus, assures a good exploitation of the reactants.

1 600 kPa 300 kPa

H2 Extraction Index, -

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

the hydrogen fed as H2 molecules and the one contained in the feed stream in other chemicals (e.g. H2O): F Permeate Extraction Indexi ¼ EIH2 ¼ Available in the feed Fi

0.5

0

Figure 12.17

275

300 Temperature, °C

325

Extraction index at 2600 h1 for different temperatures and feed pressures.

View Online

107

Membrane Reactors in Hydrogen Production

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

12.5 Concluding Remarks Nowadays, membrane reactors are a promising innovative technology in the field of hydrogen production from light hydrocarbons. Their use allows better performance to be achieved than conventional reactors in terms of high recovery of pure hydrogen streams, higher conversion and reduced catalyst amount. The traditional process can thus be redesigned as more compact and efficient thereby obtaining an intensified process with a reduced plant size and higher yield. In this context, membrane engineering plays a fundamental role in the integration of these units in a single plant as well as in the definition of the knowledge necessary to drive the process by maximizing the gains both in terms of efficiency and plant size reduction.

12.6 List of Symbols, Abbreviations and Dimensionless Numbers Symbols A DPSievertH2 I J k keq L m n P — — — RI T VI WCatalyst

Surface area H2 permeation Sievert’s driving force Sweep factor Permeating flux Kinetic constant Equilibrium constant Length Reactant feed molar ratio Number of moles Pressure Permeability Permeance Permeating flux Recovery index Temperature Volume index Catalyst weight

m2 Pa0.5 — mol (m2 s)–1 see related equation — m — — Pa mol (m s Pa0.5)–1 mol (m2 s Pa0.5)–1 mol (m2 s)–1 % 1C or K g

Abbreviations GHSV MR MREC SMR

STP TR TREC WGS

Gas hourly space velocity (STP) Membrane reactor Membrane reactor equilibrium conversion Methane steam reforming Standard temperature (25 1C) and pressure (100 kPa) Traditional reactor Traditional reactor equilibrium conversion Water gas shift

s1

Dimensionless numbers Da

Damkho¨ler number

—

View Online

108

Chapter 12

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Acknowledgements The joint bilateral project ‘Functionalized ZnO nano tubular membrane assembly for CO2 capturing: design, fabrication and performance assessment studies’ between ITM-CNR and Sol-gel Nano ceramics section of Materials and Minerals Division of the National Institute for Interdisciplinary Science & Technology (NIIST) of the CSIR (India), is gratefully acknowledged for co-funding this research.

References 1. S. Wadhwani, A. K. Wadhwani and R. B. Agarwal, First International Conference on Clean Coal Technologies for Our Future, Chia Laguna, Sardinia, 21–23 October 2002. 2. R. Kothari, D. Buddhi and R. L. Sawhney, Renew. Sust. Energy Rev., 2008, 12, 553. 3. G. Raggio, A. Pettinau, A. Orsini, M. Fadda, D. Cocco, P. Deiana, M. L. Pelizza and M. Marenco, Second International Conference on Clean Coal Technologies for Our Future, Castiadas, Sardinia, 10–12 May 2005. 4. F. Barbir, Sol. Energy, 2005, 78, 661. 5. M. E. E. Abashar, K. I. Alhumaizi and A. M. Adris, Chem. Eng. Res. Des., 2003, 81, 251. 6. T. T. Tsotsis, A. M. Champagnie, S. P. Vasileiadis, Z. D. Ziaka and R. G. Minet, Chem. Eng. Sci., 1992, 47, 2903. 7. A. M. Adris, C. J. Lim and J. R. Grace, Chem. Eng. Sci., 1997, 52, 1609. 8. A. Stankiewicz, Chem. Eng. Proc., 2003, 42, 137. 9. R. Dittmeyer, V. Ho¨llein and K. Daubb, J. Mol. Catal. A: Chem., 2001, 173, 135. 10. A. Caravella, F. Scura, G. Barbieri and E. Drioli, J. Phys. Chem. B, 2010, 114, 12264–12276. 11. A. Caravella, G. Barbieri and E. Drioli, Sep. Purif. Technol., 2009, 66, 613. 12. A. Li, C. J. Lim and J. R. Grace, Chem. Eng. J., 2008, 138, 452. 13. J. Shu, B. P. A. Grandjean and S. Kaliaguine, Appl. Catal. A: Gen., 1994, 119, 305. 14. J. Shu, B. P. A. Grandjean and S. Kaliaguine, Catal. Today, 1995, 25, 327. 15. J. Tong and Y. Matsumura, Appl. Catal. A: Gen., 2005, 286, 226. 16. E. Kikuchi, Catal. Today, 2000, 56, 97. 17. Y. M. Lin, S. L. Liu, C. H. Chuang and Y. T. Chu, Catal. Today, 2003, 82, 127. 18. S. Jorgensen, P. E. H. Nielsen and P. Lehrmann, Catal. Today, 1995, 25, 303. 19. Y. Chen, Y. Wang, H. Xu and G. Xiong, Appl. Catal. B: Environ., 2008, 80, 283. 20. J. R. Grace, S. Elnashaie and C. J. Lim, Int. J. Chem. React. Eng., 2005, 3, 128. 21. http://www.membranereactor.com

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00087

Membrane Reactors in Hydrogen Production

109

22. J. Cui, M. L. Joshi and S. L. Wellington, System and process for making hydrogen from a hydrocarbon stream, U.S. Patent 2009/0180949 Al, 2009. 23. J. Han, I. S. Kim and K. S. Choi, J. Power Sources, 2000, 86, 223. 24. G. Barbieri, G. Marigliano, G. Perri and E. Drioli, Ind. Eng. Chem. Res., 2001, 40, 2017. 25. E. Kikuchi, S. Uemiya, N. Sato, H. Inoue, H. Ando and T. Matsuda, Chem. Lett., 1989, 489. 26. S. Uemiya, N. Sato, H. Inoue, H. Ando and E. Kikuchi, Ind. Eng. Chem. Res., 1991, 30, 585. 27. S. Tosti, A. Basile, G. Chiappetta, C. Rizzello and V. Violante, Chem. Eng. J., 2003, 93. 28. A. Brunetti, A. Caravella, G. Barbieri and E. Drioli, J. Membr. Sci., 2007, 306, 329. 29. M. Gummalla, T. H. Vanderspurt, Y. She, Z. Dardas and B. Olsommer, U.S. Patent 20100104903, 2010. 30. H. W. Deckman, J. W. Fulton, J. M. Grenda and F. Hershkowitz, European Patent 1294637, 2005. 31. W. Wei, Polygeneration Systems, U.S. Patent 20090084035, 2009. 32. D. Tsay, S. E. Weiss and T. C. Tsay, U.S. Patent 20070157517, 2007. 33. S. Tosti, A. Basile, D. Lecci and C. Rizzello, European Patent 1 829 821 A1, 2007. 34. A. Lamm, T. Poschmann and J. Schaefer, U.S. Patent 20050039401, 2005. 35. R. S. Willms and S. A. Birdsell, U.S. Patent 6165438, 2000. 36. G. Avgouropoulos, T. Ioannides and H. Matralis, Appl. Catal. B, 2005, 56, 87. 37. X. Ouyang, L. Bednarova, R. S. Besser and P. Ho, AIChE J., 2005, 51, 1758. 38. P. Bernardo, C. Algieri, G. Barbieri and E. Drioli, Sep. Purif. Technol., 2008, 62, 629. 39. A. Criscuoli and E. Drioli, Ind. Chem. Eng. Res., 2007, 46, 2268. 40. S. K. Sikdar, AIChE J., 2003, 49, 1928. 41. IChemE. Sustainable Development Progress Metrics: Recommended for use in the Process Industries. [Available: http://www.icheme.org/sustainability/ metrics.pdf.] Institution of Chemical Engineers, Rugby, 2006, pp. 1–28. 42. D. R. Seok, S. T. Hwang, in Future Opportunities in Catalytic and Separation Technology, Y. Morooka, S. Kimura, (Eds.), Elsevier, Amsterdam, 1990, 248–267. 43. A. Brunetti, G. Barbieri, E. Drioli, K.-H. Lee, B. Sea and D. W. Lee, Chem. Eng. Proc., 2007, 46, 119–126. 44. G. Barbieri, A. Brunetti, G. Tricoli, E. Drioli, J. Power Sources, 2008, 182, 160–167.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

CHAPTER 13

Palladium-based Selective Membranes for Hydrogen Production G. IAQUANIELLO,*a M. DE FALCOb AND A. SALLADINI c a

Tecnimont-KT S.p.A., Viale Castello della Magliana 75, 00148 Roma, Italy; b Faculty of Engineering, University Campus Bio-Medico of Rome, via Alvaro del Portillo 21, 00128 Rome, Italy; c Processi Innovativi, Corso Federico II, 67100 L’Aquila, Italy

13.1 Basic Features of Membrane Reactors The original idea of coupling catalyst and membranes dates back to the 1960s. Among the first, Michaels1 suggested that through the use of a semi-permeable membrane a considerable increase in the conversion of thermodynamically limited reactions could be achieved. In a membrane reactor one or more chemical reactions, generally catalytically promoted, are carried out in the presence of a membrane selectively permeated by one of the reaction products. As result of a lower reaction temperature, another major advantage emerges, i.e. the possibility of a better heat integration, as the use of gas exhausts from a gas turbine or solar heated molten salts.2 In view of the significant potential advantages, attention hereafter is paid mostly to membrane reactor engineering focusing on the most interesting applications. Criticism of membrane integration has to be carefully faced. If the selective membrane is directly integrated in the reaction environment, coupling catalyst and membrane operating conditions leads to the necessity to define a compromise Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

110

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

Palladium-based Selective Membranes for Hydrogen Production

111

optimization in order to promote both the kinetics and permeability, without damaging the membrane always requesting stringent thermal threshold. On the other hand, the membrane can be integrated externally, by an architecture which foresees reaction and separation steps in series. In this way, catalyst and membrane operating conditions are independent and their optimal operating conditions can be defined separately, but the membrane integration benefits are reduced. It is a worth assessment that the development of such innovative reactors requires ad hoc design criteria definition, which is one of the main scope of the present work.

13.1.1

Selective Membranes

Figure 13.1 shows a general classification scheme for membranes.3 It has to be noticed that membranes can be divided into two macro-categories: inorganic membranes and organic membranes. When a selective membrane is applied in a chemical process for the separation of high temperature gases, inorganic membranes represent the only suitable solution. In the viewpoint of the morphology and membrane structure categorization, the inorganic membranes can be divided into ceramic and metallic. In particular, ceramic membranes differ according to their pore diameter in microporous (dpo2 nm), meso-porous (2 nmodpo50 nm) and macro-porous (dp450 nm), while metallic membranes can be categorized into supported and unsupported. Generally, inorganic membranes are stable between 200 and 8001C and in some cases they can operate at elevated temperatures (ceramic membranes over 1000 1C).4 Depending on their geometry, the membranes can be subdivided in tubular, hollow fiber, spiral wound and flat sheet:5 tubular membranes are the most common solution, even if they require relatively high volume per membrane area unit and present high costs.

Figure 13.1

A general classification scheme of the membranes.3

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

112

Figure 13.2

Chapter 13

Hydrogen permeability through different metals.6

Focusing on hydrogen selective membranes, which are one of the most interesting and promising applications, dense metallic membranes are the dominant technology as a result of the high selectivity and good permeability. Figure 13.2 illustrates H2 permeability through different metals:6 although niobium (Nb), vanadium (V) and tantalum (Ta) offer higher hydrogen permeability than palladium in a temperature range between 0 and 700 1C, these metals give a stronger surface resistance to hydrogen transport than palladium (Pd). For this reason, dense palladium membranes are preferentially used. The molecular transport of hydrogen in palladium membranes occurs through a solution/diffusion mechanism, which follows six steps:     

Dissociation of molecular hydrogen at the gas/metal interface Adsorption of the atomic hydrogen on membrane surface Dissolution of atomic hydrogen into the palladium matrix Diffusion of atomic hydrogen through the membrane Re-combination of atomic hydrogen to form hydrogen molecules at the gas/metal interface  Desorption of hydrogen molecules. Depending on temperature, pressure, gas mixture composition and thickness of the membrane, each one of these steps may control hydrogen permeation through the dense film. As a result, the hydrogen permeating flux can be expressed by means of the following equation: JH2 ¼ PeH2 =d  Pn H2;ret  Pn H2;perm




ð13:1Þ

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

Palladium-based Selective Membranes for Hydrogen Production

113

where n (variable in the range 0.5–1) is the dependence factor of the hydrogen flux to the hydrogen partial pressure, JH2 the hydrogen flux permeating through the membrane, Pe the hydrogen permeability, d the membrane thickness, pH2-ret and pH2-perm the hydrogen partial pressures in the retentate (high pressure side) and permeate sides (low pressure side), respectively. When the pressure is relatively low and Pd layer is relatively thick (410 mm), the diffusion is assumed to be the rate-limiting step and the factor n is equal to 0.5. In this case, eqn (13.1) becomes the Sieverts–Fick law:
JH2;SievertsFick ¼ PeH2 =d  P0:5 H2;ret  P0:5 H2;perm ð13:2Þ On the contrary, at high pressures or with much thin selective layer, the hydrogen–hydrogen interactions in the palladium bulk are not negligible and factor n becomes equal to 1:
JH2 ¼ PeH2 =d  PH2;ret  PH2;perm ð13:3Þ The relationship between hydrogen permeability and temperature follows an Arrhenius behavior while n generally does not depend on the temperature: PeH2 ¼ Pe0 H2 expðEa=RTÞ

ð13:4Þ

where Pe0 is the pre-exponential factor, Ea the apparent activation energy, R the universal gas constant and T the absolute temperature. The most important problem associated with the use of pure palladium membranes is the ‘hydrogen embrittlement’ phenomenon. When the temperature is below 300 1C and the pressure below 2.0 MPa, the b-hydride phase may nucleate from the a-phase, resulting in severe lattice strains (see Figure 13.3), so that a pure palladium membrane becomes brittle after a few cycles of a$b transitions.7–9 Such a problem can be overcome by using Pd-alloy containing another metal, such as silver. The palladium alloys have a reduced critical temperature for the a–b phase transition. Pd-Ag membranes can operate in hydrogen atmosphere at temperatures below 300 1C without observing

Figure 13.3

Equilibrium solubility isotherms of PdHn for bulk Pd at different temperatures.13

View Online

114

Chapter 13

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

10

embrittlement rather than pure palladium membranes. Moreover, in some cases, the hydrogen permeability of palladium alloys is higher than pure palladium. The hydrogen flux through the Pd-Ag membranes reaches the maximum value at 350 1C and 2.2 MPa with a 23% Ag content.11 In details, the permeability is 1.7 times higher than one of a pure Pd membrane. The Pd-Cu alloy even shows a maximum value of hydrogen flux with 40% Cu content, although these membranes suffer a permeation decrease when exposed at 900 1C for a long time.12 Generally, Pd-based membranes are supported, in which the Pd-based selective layer is deposited on a support able to assure the mechanical strength required to operate within chemical processes temperature/pressure ranges. The main supports applied are: ceramics as SiO2, Al2O3 and B2O3, or metallic as porous stainless steel.

13.1.2

Membrane Fabrication Methods

Various methods can be applied to produce Pd-based membranes, depending on some factors such as the nature of the selective layer metal, the manufacturing facilities, the required thickness, surface area, shape, purity, etc. Nevertheless, no one method can produce a membrane, which combines advantageously all these factors. Therefore, the choice of the production method is a compromise among these factors.13 Table 13.1 summarizes the main fabrication techniques, with benefits and drawbacks. To lower the production costs, it is important to develop a proper membrane manufacturing strategy which involves two main aspects:  The manufacturing process itself, which will give the business a distinct advantage in the market-place through unique technology, for example  Manufacturing associated activities in terms of infrastructure design, such as controls, procedures, selection of subcontractors etc. that are involved in the main process aspects of manufacturing. To simplify the manufacturing process, one way would be to separate the preparation of the Pd-based selective layer from the its integration on the support, as demonstrated by Bredesen and Klette.14 Moving ahead with such a concept, the manufacturing process should consist of a batch process of three independent steps. The choice of a batch process is a logical one because it provides similar items on a repeat basis, usually in larger volume. A batch procedure divides the manufacturing task into a series of appropriate operations, which together will make the product involved. It is then not so difficult to define the main steps in such a process:  Preparation of the Pd-alloy selective layer  Preparation of the support  Assembling and testing of the membrane module.

Based on the controlled auto-catalyzed decomposition or reduction of meta-stable metallic salt complexes on target surfaces

Not simple to control

Not simple to control

Relatively low H2/other gases permselectivity

Adherence problem between selective layer and support

Very versatile method

Very thin layer deposition

Very simple

Low cost

The thickness of deposited films can be mastered by controlling electroplating time and current density Uniformity of deposits on complex shapes and hardness

Difficult thickness control, costly losses of palladium in the bath, non-guaranteed purity of the deposit

Large domains of alloy composition are not easy to control

Expensive method Able to coat a complexshaped component with a uniform thickness layer

Drawbacks

Benefits

Palladium-based Selective Membranes for Hydrogen Production

Electroless plating deposition

Physical vapor The solid material to be deposited is evaporated in a deposition (PVD) vacuum system through physical techniques, followed by condensation and deposition as a thin film on a cooler substrate A sputtering system consists of a vacuum chamber Sputtering and containing a target (a plate of the material to be deposited) magnetron and the substrate (i.e. the membrane), in which a sputtering sputtering gas (an inert gas such as argon) is introduced to provide the medium in which a glow discharge, or plasma, may be initiated and maintained. Afterwards, positive ions strike the target and remove target atoms and ions by momentum exchange. The condensation of these species over the support produces a thin film Spray pyrolysis A metal salt solution is sprayed into a heated gas stream and, then, pyrolyzed  Melting the raw materials with chosen composition at Cold rolling very high temperature  Ingot casting  High temperature homogenization  Hot and cold forging or pressing, followed by repeated sequences of alternate cold rolling and anneals, down to the required thickness Chemical vapor A chemical reaction involving a metal complex in the gas deposition (CVD) phase is performed at a controlled temperature and the produced metal deposits as a thin film by nucleation and growth on the substrate Electroplating A substrate, used as a cathode, is coated with a metal or an alloy in a plating bath

Procedure

Pd-based selective membrane fabrication methods

Fabrication method

Table 13.1

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

View Online

115

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

116

Chapter 13

One way to approach the membrane and support fabrication is to consider two distinct lines, each one delivering the distinctive product, which at the end of the process is tested for quality control before moving to the integration.3 In the last few years, the main aim of researchers and industries has been to reduce the thickness of the Pd-based films, with crucial benefits in terms of hydrogen permeation flux and reduced cost. Nowadays, Pd-based membranes of thickness 2–5 mm, with a good layer uniformity and adherence on the support, have been produced.

13.1.3

Palladium-based Membranes Available on the Market

Only a few companies are able to supply Pd-based membranes ‘modules’ since the market is still limited at laboratory scale membranes or modules for small pilot units. In the following, the main membranes providers are listed.

13.1.3.1

ECN Hydrogen Separation Modules

The Energy research Centre of the Netherlands (ECN) produces hydrogen separation modules (Hysep) on a pre-commercial basis for evaluation purposes. An essential element of the Hyseps technology is the use of thin-film palladium composite membranes to enable low cost and reliable hydrogen separation. The supported palladium layer in the Hyseps module has a thickness as low as 3–9 mm, a substantial improvement over current commercial available palladium membranes, which are based on self supporting metal foils with a thickness of 20–100 mm.

13.1.3.2

MRT Hydrogen Separation Modules

MRT is a Vancouver-based private company interested in hydrogen purifiers to provide high-purity hydrogen and to recover hydrogen from mixed gas streams. MRT produced membranes either as rolled foils or as deposited thin films (8–15 mm). In addition, patented bonding techniques have been developed to permanently attach membranes to support modules with a perfect, hydrogentight seal. For membranes thinner than 15 mm, MRT uses a proprietary coating technique. Prototype membranes as thin as 8 mm, tested by MRT, have been produced and show excellent performance and longevity.

13.1.3.3

Hydrogen Selective Membranes Produced in Japan

An important Japanese company (JC) is developing a gas separation membrane, which efficiently recovers hydrogen, by forming a film on a porous ceramic substrate using palladium alloy known for its feature of selective permeation of hydrogen. The key is to simultaneously achieve cost effectiveness and high hydrogen selectivity, by making expensive palladium membranes thinner. The membranes are produced in a three step procedure: at first Pd is

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

Palladium-based Selective Membranes for Hydrogen Production

117

deposited onto the Al2O3 support by electroless plating technique, and then Ag is layered on by electroplating using the Pd layer as electrode. The layered Pd-Ag membrane is finally heat treated to obtain the Pd-Ag alloy membrane. The resulting membranes are tubular with an external diameter of about 1.0 cm, an effective length of about 9.0 cm and a Pd-Ag coating deposited on the external surface with a selective layer of about 28.3 cm2.

13.1.3.4

SINTEF Hydrogen Selective Membranes

SINTEF has developed a technique for the manufacturing of Pd-based hydrogen separation membranes based on a two step process allowing a reduction of membrane thickness. First, a defect free Pd-alloy thin film is prepared by magnetron sputtering onto a silicon wafer. In a second step the film is removed from the wafer. These films may subsequently either be used self-supported or integrated with various supports of different pore size, geometry and size. This allows, for example, the preparation of very thin (approximately 2–3 mm) high-flux membranes supported on macro-porous substrates, which can operate at high pressures.

13.1.4

Membrane Cost Analysis

In order to forecast a production cost for thin Pd-based membranes, it is important to introduce the concept of ‘economics of learning’ in understanding the behaviour of all added costs of membranes as cumulative production volume increased. Such economics of learning or law of the experience may be expressed as follow: cn ¼ c1 na

ð13:5Þ

where c1 is the cost of the unit production (square meter of membrane for instance), cn is the cost of the nth unit of production, n is the cumulative volume of production and a is the elasticity of cost with regard to output. Then, an experience curve can be constructed by using the data available, which are limited, for the Pd-based or ceramic membrane, to minimal surface (less than 1 m2). Another issue associated with drawing an experience curve is that cost and production data must be related to a ‘standard product’, which is not the case due to the fact that in the membrane technology no standard has yet emerged. It is, however, a fact that costs decline systematically with increases in cumulative output. The assumptions made are: c1 ¼ 50 000 euros and a ¼ 0.25, where c1 value derived by Tecnimont-KT recent experience in building pilot units, meanwhile the ‘a’ factor is assumed as average value typically between 20 and 30%. Using such data foreseeing the cost for m2 of membrane module as a function of the cumulative value of production, expressed in terms of m2, is possible as reported in Figure 13.4.3

View Online

118

Chapter 13 10000

€ per m2

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

8000

6000

4000

2000

0 0

2000000

4000000

6000000

8000000

10000000

12000000

2

Cumulative Production m

Figure 13.4

Cost for square meter of membrane module versus cumulated production.3

Obviously, the experience curve is characterized by a progressively declining gradient. The size of the experience effect is measured by the proportion by which costs are reduced with subsequent doublings of aggregate production. For Pd-based membranes, it is foreseen a cost of 8900 euros per m2 for a cumulative production of 1000 m2, while the cost would be reduced to 900 euros per m2 if the cumulative production is risen to 10 000 000m2 worldwide. This final value fit well with the cost reported by ECN.15 Now, the main question is how to reach cumulative production of millions of square meters of Pd-based membrane in a few years. The R&D efforts on membrane cost reduction, on the reliability of thin film fabrication methods and on the benefits demonstration of selective membrane application in chemical processes would push the introduction of such a new technology in the market.

13.2 Membrane Reactor Architectures Selective membrane application in chemical processes represents one of the most interesting scientific and technological topic of the last years, in the context of industrial process intensification tendency aimed to improve process efficiency. Generally, integrating a selective membrane in a reaction environment allows reducing operating temperature to obtain a specific reactant conversion, improving global process efficiency and, consequently, reducing the fuel to be burned to supply process heat duty. Refer to Section 13.3 for the quantification of membrane integration benefits in the main chemical processes devoted to hydrogen production.

View Online

Palladium-based Selective Membranes for Hydrogen Production

119

In this paragraph, two membrane reactor configuration are proposed and evaluated:

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

 The integrated membrane reactor (IMR) configuration  The staged membrane reactor (SMR) configuration.

13.2.1

Configuration Layouts

13.2.1.1

Integrated Membrane Reactor

An integrated membrane reactor (IMR) is a compact device in which a selective membrane is directly assembled inside the reaction environment.16–21 The simplest configuration is composed of two concentric tubes, where catalyst pellets are packed in annular zone while the inner tube is the membrane itself, as shown in Figure 13.5. Through the inner tube a sweeping gas is fed, co-currently or counter-currently, in order to carry out the hydrogen permeated. Obviously, the membrane integration can be made also assembling many smaller tubes, so thus increasing the specific membrane surface on reactor volume and consequently the global permeated hydrogen flow. Two different zones can be recognized: the reaction zone, which is the annular section where catalyst is packed, and the permeation zone, where the sweeping gas is fed. For industrial applications, the IMR is a tubes-and-shell shaped reactor where a bundle of membrane reactor tubes are assembled in a shell through which a heating fluid is sent to supply heat duty required by the reactions (Figure 13.6).

13.2.1.2

Staged Membrane Reactor

Another method for integrating a selective membranes in a chemical process is to place them outside the reactor, in proper units located downstream. Figure 13.7 shows a staged membrane reactor (SMR) composed of two-step reactionseparation units.

Figure 13.5

Integrated membrane reactor draft.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

120

Chapter 13

Figure 13.6

Tubes-and-shell shaped IMR.

Figure 13.7

SMR process layout.

The feedstock is sent to a first reactor where it is partially converted into the products; then one of the products (as an example, hydrogen) is recovered through a selective membrane separation module, while the retentate is sent to the next step or recycled to the first module. By means of a heat recovery system, the operating temperature can be reduced before the membrane unit, in order to assure a thermal level suitable for a proper operation of the membrane

View Online

Palladium-based Selective Membranes for Hydrogen Production

121

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

unit, and again increased before the second reactor at values suitable for supporting the reactions. It is possible to replicate the reaction-separation steps until the desired natural gas conversion is achieved.

13.2.2

Benefits and Drawbacks

Generally, both the configurations assure crucial benefits in respect to the traditional technology. Taken as an example the natural gas steam reforming reaction, which is the most used process for the massive production of hydrogen and which is a strongly endothermic reaction supported only at very high temperature (850–950 1C), a Pd-based integration would lead to the following main benefits.

13.2.2.1

A Strong Reduction of the Reaction Temperature

Steam reforming reactions are very fast on Ni-based catalyst and quickly reached equilibrium conditions, leading to the necessity of increasing temperature to obtain high conversion. By integrating a selective membrane able to separate a reaction product as H2, the equilibrium conditions are not reached and this allows the promotion of the reactions at lower temperatures (450–650 1C). Consequently, a different heat duty supplying strategy can be implemented, i.e. a heat exchanger with a heating fluid could be used in the place of the furnace needed if very high temperatures are required, with the following advantages:  Higher efficiency of the heat transfer from the external source to the reactor  Lower energy of the heating fluid in comparison with the high temperature combustion gas used in the furnace, which means lower heating costs  The possibility to use different heating fluids, depending on their availability  The easy scalability (scale up or scale down) of the system and therefore its applicability in many fields (small, medium or large scale)  Use of cheaper alloy steel.

13.2.2.2

Increase in Process Efficiency

The lower temperature results in an increase of overall process efficiency, since the heat supplied is better exploited. It is foreseen that the global process efficiency should increase from the 65–80% of today’s technology up to 85% and more for all the plant sizes.

13.2.2.3

Large Saving of Methane

Reduction of reaction temperature means that the heat duty requirement is lower than for the traditional process. The heat flux from the external source to the catalytic bed should be about 30–40 kW m2 instead of 80 kW m2 and more of the traditional process.22 Consequently, a smaller amount of methane

View Online

122

Chapter 13

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

has to be burned. The lower thermal level could allow the coupling of the reformer with a different and clean energy source, i.e. solar energy, nullifying the fraction of natural gas to be burned for process heat requirements.23,24

13.2.2.4

Reduction of CO2 Emissions

The methane saving leads to a reduction of the greenhouse gas emissions, since less or no carbon dioxide is produced by the methane combustion. In a traditional process, the ratio (CO2 released)/(H2 produced) is 8–12 kgCO2 kgH21, depending on the process efficiency. An increase of the efficiency could lead to a reduction of GHG emissions within the range 20–55%, up to 5.5 kgCO2 kgH21 if a renewable energy source is used for process heat duty.

13.2.2.5

Easier CO2 Purification

Membrane integration into the reaction environment ensures a first substantial hydrogen separation step (up to 90% of the hydrogen produced can be removed); as for CO2 separation, because of the higher carbon dioxide partial pressure in the reformer outlet stream, due to the hydrogen removal, physical separation methods could be used to separate CO2 rather than the chemical adsorption in mono-diethanol ammine (MDEA).

13.2.2.6

Reduced Dependence on the Cost of Natural Gas

Increasing the reaction efficiency and reducing the amount of methane to be burned to supply the process heat duty requirements lead to a reduction of the total amount of methane required for producing a mass unit of hydrogen. Therefore, although a higher plant cost has to be supported because of the increasing reactor complexity, the hydrogen price would be less dependent on the natural gas.

13.2.2.7

The Two Configurations

Concerning with the two configurations proposed, Table 13.2 summarizes benefits and drawbacks of both the architectures. Globally:25  At the same operating conditions, IMR leads to better performance, since an integrated membrane reactor is equivalent to an infinite series of reactor þ separator modules  On the other hand, if a dense supported membrane (as Pd-based membranes for hydrogen separation) is assembled, a stringent temperature threshold (T4 500 1C) has to be respect to guarantee a proper selective layer–support adherence. For IMR, this leads to a limit for reaction operating conditions, as well. In SMR, reaction and separation operating conditions can be defined, and optimized, separately, leading to crucial performance benefits.

SMR configuration

 Ease of scalability: scale-up and scale-down of the membrane reactor are easy through an increase or a decrease of the number of parallel tubular reactors or of the length of the single membrane reactor.  Less useless catalyst: for strongly endothermic reactions, as for natural gas steam reforming, catalyst pellets placed in the central zone usually work bad since the temperature is too low for promoting the reactions due to the large radial temperature gradient. In membrane reactors, the central zone does not contain catalyst but the membrane tube devoted to separate the hydrogen produced.  Decoupling of reaction and separation operating conditions: the reactor and separation units temperatures can be optimized independently, both increasing reactant conversion for each reaction step and membrane stability and lifetime.  Simplification of the mechanical design of membrane tubes relative to those embedded in catalyst tubes.  Simplification of membrane modules maintenance and of catalyst replacement.

 High cost, mainly for selective membranes surface required.

 High membrane surfaces required.

 Compactness of the process: SMR configuration is composed by reactors, separation modules, heat exchangers.

 No easy reactor maintenance.

Drawbacks  Technological problem in designing the reactor: a sensible component as the selective membrane has to be inserted in a critical environment.  Coupling between reaction and separation operating conditions.

Benefits

 Compactness of the process: the reactor is able to support simultaneously the reactions and the product separation.

IMR configuration

Benefits and drawbacks for IMR and SMR architectures

Architecture

Table 13.2

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

View Online

Palladium-based Selective Membranes for Hydrogen Production 123

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

124

Chapter 13

From these considerations it can be concluded that, at the actual membrane state-of-the-art, SMR is the leading architecture for membrane safety, for the possibility of increasing reaction temperature up to value equal to 650 1C for natural gas steam reforming and for its maintenance easiness. Surely, a future improvement of membrane performance, mainly for stability, would promote the applications of IMRs. In the next paragraph some applications of both configurations are reported and evaluated for the main hydrogen production processes.

13.3 Case Studies 13.3.1

Natural Gas Steam Reforming

Steam reforming of natural gas is a well-established technology for hydrogen production in refining and fertilizer industries. The process is controlled by chemical equilibrium and significant hydrogen yields are achieved only at high temperatures (850–900 1C) according to the following reactions: CH4 þ H2 O $ CO þ 3H2

Methane steam reforming

DH 0 298K ¼ 206 kJ mol1

CO þ H2 O $ CO2 þ H2

Water gas shift

DH 0 298K ¼ 41 kJ mol1

CH4 þ 2H2 O $ CO2 þ 4H2

Overall reaction

DH 0 298K ¼ 165 kJ mol1

In order to sustain the global endothermic reaction, a part of methane feedstock has to be burned in furnaces, reducing the process global efficiency, increasing the greenhouse gas (GHG) emissions and strengthening the dependence of hydrogen cost on the natural gas cost. The integration of hydrogen selective membranes allows to shift the chemical equilibrium toward the righthand side of the reaction enhancing hydrogen yield at lower temperatures. Several theoretical18,26–32 and experimental works33–40 investigated membrane integration with natural gas steam reforming reactors. Among the typologies of hydrogen selective membranes, thin Pd-based supported membranes seem to be the most promising thanks to the high selectivity and high permeation flow. They can be integrated in steam reforming process by means of two potential configurations:  In configuration 1, the reformer and membrane module (RMM) where hydrogen selective membrane is assembled in separation modules applied downstream to reaction units so that the process scheme is composed by a series of reaction-separation units (staged membrane reactor architecture)  In configuration 2, the membrane reactor (MR) where the selective membrane is assembled directly inside the reaction environment, so that

View Online

Palladium-based Selective Membranes for Hydrogen Production

125

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

the hydrogen produced is immediately removed from reaction zone (integrated membrane reactor architecture). With respect to a conventional steam reforming process, several differences can be distinguished when employing palladium membranes: (i) the reforming reaction is performed at temperature lower than 650 1C, thus reducing the energy consumption and enabling the use of less expensive materials for the reforming tubes; (ii) the H2 removal through the membrane enables for the achievement of a CH4 conversion up to 90%; (iii) the retentate gas mixture has a high calorific value and is available under pressure slightly lower than the reforming inlet pressure; and (iv) the water gas shift reaction can be performed in the reforming itself due to the low temperature. In this way the overall duty of steam reforming is reduced and water gas shift reactor is no longer required. The milder condition makes it possible to locate membrane assisted reformer downstream of a gas turbine with a consequent reduction in energy saving. An interesting application of membrane steam reforming reactor to co-generative systems was reported by Iaquaniello et al.41,42 The process layout is shown in Figure 13.8.43 In this application, the high pressure retentate is directly routed to a gas turbine and its exhausts can be used to supply the reforming duty, preheat other process services and raise steam with minimum post-combustion firing.

Figure 13.8

Schematic representation of membrane reforming reactor integrated with a gas turbine: (1) Combustion chamber; (2) hydro-desulfurization unit; (3) convective furnace; (4) and (5) membrane modules for H2 separation; (6) H2 compressor.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

126

Chapter 13

Two schemes with two reforming modules, with and without CO2 removal, has been analyzed, taking into account a natural gas feed and a H2 production capacity of 20 000 N m3 h1. Compressed natural gas at 20 barg is first mixed with hydrogen recycled from PSA section, and preheated before entering the desulfurization unit, where sulfur compounds are removed. The desulfurized feed is then mixed with steam in order to achieve a steam carbon ratio ranging from 2.5 to 4 depending on the type and the quantity of the C21 hydrocarbons present in the feed. First stage reforming takes place at a temperature of 630 1C. The reformed gas product is cooled down at 450 1C: such a temperature is suitable for membrane separation. A retentate, recycled to the second reformer stage, and a mixture of H2 and sweeping steam are produced by membrane separation. The second stage reformer effluent is cooled down from 580 to 450 1C and sent to the second membrane separator. Permeate stream is mixed with that coming from the first stage before cooling. The retentate from the second stage (purge gas) is cooled above dew point and part of it is recycled to the first reformer stage. The residual is further cooled to condense the steam and separate the water. The purge gas is fired in the gas turbine, together with external fuel, to co-generate electric power for export. The exhausts from the gas turbine, after re-heating, are sent as heating medium to the first and second reforming stages to supply the reaction heat, and the effluent from the reforming stage is sent to the heat exchange network to provide heat for generate export, dilution and sweeping steam and provide the heat required to close the heat balance of the plant. Hydrogen streams are cooled down to condensate steam that is separated and recycled back to the degasifier. The hydrogen stream is then compressed and further cooled before final purification with a dedicated PSA. Calculations made on the proposed scheme showed that: (i) there is a substantial amount of export power associated with hydrogen production equal to 1.1 MW per 1000 N m3 of H2; (ii) overall estimated energy consumption is about 10% less than that of the conventional steam reforming process; and (iii) there is a reduction of reforming duty from 84.07 M kJ h1 of conventional scheme to 74.94 of hybrid scheme without CO2 removal. Providing the steam reforming duty by sources as solar heated molten salts or a fluid heated in a nuclear reactor may further increase the overall energy efficiency of the system and pave the way for producing large amount of hydrogen with minimum environmental impact. De Falco et al.2 studied a three-step reforming and membrane modules plant powered by a nuclear reactor or by solar heated molten salts where heat is supplied via an indirect exchanger. Results showed that integration makes the production of hydrogen a 10% less expensive than in the case of the conventional steam reforming and reduces the overall CO2 production of 28%. A two step reformer and membrane modules (RMM) test plant having the capacity of 20 N m3 h1 has been designed and constructed to investigate at an industrial scale level the potentialities of this kind of architecture.3,44 At a reaction temperature of 620 1C, steam to carbon ratio of 4.8 and a membrane temperature of 430 1C, an overall yield of 59% is achieved. By properly

View Online

127

Palladium-based Selective Membranes for Hydrogen Production 100% 95%

Methane conversion, %

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

90% 85% 80% 75% 2R STAGE Tr=650°C; Tm=470°C

70% 3R STAGE Tr=650°C; Tm=470°C

65% 4R STAGE Tr=650°C; Tm=470°C

60% Exp values 2R STAGE Tr=620°C;Tm=430°C

55% 0

0.5

1

1.5

2

2.5

3

3.5

Total membrane area, m2

Figure 13.9

Operating conditions extension based on experimental results for a twostep reformer and membrane modules (RMM) test plant.

extending the design parameters within reasonable limits a yield as high as 90% can be calculated as reported in Figure 13.9.

13.3.2

Water Gas Shift Reactor

The water gas shift (WGS) reaction is an important step in many chemical processes for the hydrogen production such as catalytic steam reforming of hydrocarbons, coal gasification, production of ammonia as well as for fuel cell technology. An interesting application of WGS reaction is related to membrane assisted integrated gasification combined cycle in a pre-combustion configuration (IGCC) for clean energy or hydrogen production with zero CO2 emissions.45 Most of the theoretical and experimental work is focused on a Pd-based WGS reactor for the highest flux and H2 selectivity46–48 but recent studies also support the use of porous inorganic membranes such those based on silica49,50 although the low hydrothermal stability. A possible scheme for a membrane assisted IGCC is shown in Figure 13.10. It is achieved with hydrogen separation modules located upstream of each conventional WGSR reactor. The membrane separators allow an increase in CO conversion to selectively remove H2 from the syngas feed. A H2-rich permeate stream and a medium pressure CO2-rich stream are obtained as end products. The hydrogen removal before the WGSR reduces the volume of the syngas stream and consequently the steam requirement to avoid sintering of the catalyst. Overcoming thermodynamics constraints, WGSR may be performed at a higher inlet temperature with a consequent more favorable kinetics.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

128

Chapter 13

Figure 13.10

Schematic of membrane assisted WGSR reactor.3

To meet the requirements of the catalyst, additional steam is introduced before the second reactor lowering the stream temperature. The coupling of membrane separation modules and the conventional WGSR reactors through this kind of architecture results in a better overall efficiencies (97.5% as compared to 91% for the reference case). By this configuration, the fuel stream is enriched in H2 by the membrane reactor and requires only polishing by PSA.

13.3.3

Propane Dehydrogenation

Nowadays, olefins are widespread as the raw materials for a variety of chemical industries. This is the reason for interest towards intensification methods of olefins production within the last half a century. Catalytic dehydrogenation is one of the most important among the currently available methods. Dehydrogenation of light alkanes is an endothermic process limited by the thermodynamic equilibrium. To overcome the thermodynamic constraints high temperature and low pressure conditions are employed in the commercial dehydrogenation of hydrocarbons. Such conditions result in side reactions lowering the yields and in coke formation with following catalyst deactivation. The lower temperature conversion can be increased by the selective removal of H2 through selective membranes. Several works report on catalytic dehydrogenation of propane to propylene using membrane reactors.51–59 A possible scheme also enabling the reduction in CO2 emissions is reported in Figure 13.11.43 In this application, the separated hydrogen is fed to a gas turbine and to a post-combustion chamber using its exhausts to supply the dehydrogenation reactor thermal duty. The plant is designed for a propylene capacity of

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

Palladium-based Selective Membranes for Hydrogen Production

Figure 13.11

129

Two-step propane dehydrogenation to propylene with minimum CO2 emissions: (1) Combustion chamber; (2) post-combustion chamber; (3) convective furnace; (4) and (5) membrane modules for H2 separation; (6) boiler; (7) propane/propylene separator; (8) H2 compressor.

1000 kmol h1 (42 ton h1). It is supposed that the gas turbine exhausts are at a temperature of 500 1C and that they are heated up to 600 1C in a post combustion chamber burning a portion of the produced hydrogen. Further, the exhausts are cooled up to 450 1C supplying the thermal duty in the dehydrogenation. Hydrogen is separated by two membrane modules that are external to reactor. Thermal duty at the outlet of the combustion chamber, required to ensure a power production of 60 MWel, may be produced by a combined firing of CH4 and H2 while thermal duty necessary to increase the exhausts temperature from 500 1C to 600 1C is provided by the firing of the remaining portion of the produced H2. In this way the total power production is equal to 72 MWel, with a total CO2 emissions equal to 0.6 kg of CO2 per kg of propylene.

13.3.4

Catalytic Partial Oxidation

The catalytic partial oxidation (CPO) achieved by short contact time is emerging as one of the more promising technology to compete against the autothermal reforming for syngas production in the gas to liquids (GTL) and steam reforming for hydrogen production. Coupling the CPO reaction with H2 selective membranes will allow an increase in the process competitiveness. Successful experiences dealing with membrane applications are reported in literature, such as Galuszka et al.60 who observed considerable enhancement of

View Online

130

Chapter 13

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

61

CH4 conversion and CO and H2 yield, and Basile et al. who achieved a maximum methane conversion of 96% at 550 1C, both using a palladium membrane reactor. A possible scheme of CPO membrane assisted is reported in Figure 13.12. Compressed natural gas is mixed with hydrogen recycle and preheated before entering the HDS reactor. The desulfurized feed is mixed with steam in a ratio lower from 0.2 to 0.5 and oxygen in a ratio of 0.1–0.15 and preheated before entering the first CPO reactor. The steam to carbon ratio (S/C) and oxygen to carbon ratio (O/C) are slightly different from conventional CPO application due to the hybrid configuration (syngas þ hydrogen production) and presence of the membrane separators. The first stage CPO takes place at 750 1C. The syngas product enters the first module at 450/500 1C and splits in retentate, recycled to the second CPO reactor, and H2 þ sweeping steam. The retentate from the second stage is added with CO2 and enters the third CPO stage where the reactions take place at around 920 1C to reach the target feed þ CO2 conversion. Effluent is sent to a CO2 removal section where CO2 is partially recycled to the third CPO stage. Hydrogen streams are cooled down to condense the steam, then compressed and partially routed to the syngas to readjust the H2/CO ratio to 2:1. Capturing carbon dioxide and recycling it back to process to transform it into CO by dry reforming not only will reduce the greenhouse effect, but will also enhance the production of the syngas and reduce the costs.

13.3.5

Catalytic Decomposition of Hydrogen Sulfide

Hydrogen sulfide has potentially high economic value if both sulfur and hydrogen can be recovered. Due to thermodynamic constraints, thermal catalytic decomposition is a good candidate for membrane reactor application. Silica membranes appear more suitable than noble metal membranes which are affected by chemical attack by hydrogen sulfide.62,63 A possible scheme of membrane assisted H2S decomposition is based on Claus process to generate heat for the thermal decomposition and on a lower temperature decomposition step equipped with H2 permeable membrane. In this way unconverted hydrogen sulfide, recycled to the Claus reactor, produces the required reaction heat according to the following reaction: H2 S þ heat $ H2 þ 12 S2 10H2 S þ 5O2 $ 2H2 S þ SO2 þ 72 S2 þ 8H2 O þ heat A three-step catalytic reactors membrane separation (CRMS) configuration based on this concept is shown in Figure 13.13. A H2S reach stream would be compressed to 8 bar, preheated to 550 1C downstream of the third module, to be fed to the first reaction step consisting of fixed bed catalytic tubes. These catalytic tubes are immersed in the Claus reaction chamber, where Claus gases provide thermal duty required to carry out the H2S decomposition reaction.

Figure 13.12

Process scheme for a membrane assisted CPO for syngas and hydrogen coproduction.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

View Online

Palladium-based Selective Membranes for Hydrogen Production 131

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

132

Figure 13.13

Chapter 13

Three-step CRMS configuration for the novel H2S decomposition process.3

The product stream, containing H2S, H2, and S2, leaves through the top of the catalytic reactor and enter the first membrane module. Hydrogen is removed in the separation module and the retentate is cooled to the dew point temperature to separate sulfur. The decomposition gas leaving the third separator is recycled to the Claus reactor to treat the unconverted H2S and produce the required reaction heat through the Claus process. Hydrogen streams are cooled, compressed and further cooled before a final purification with a dedicated PSA. Since no natural gas is used the proposed scheme allows hydrogen production without CO2 emissions.

13.4 Concluding Remarks This chapter has reported the basic features of a membrane reactor: the properties of selective membranes, fabrication methods, actual markets and a cost analysis are described and assessed. Two membrane reactor configurations, the staged membrane reactor (SMR) and the integrated membrane reactor (IMR), are presented and compared, with the following main outcomes:  Under the same operating conditions, an IMR leads to better performance, since it is equivalent to an infinite series of reactor þ separator modules  In an SMR, the reaction and separation operating conditions can be defined, and optimized, separately, leading to crucial performance benefits.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

Palladium-based Selective Membranes for Hydrogen Production

133

From these considerations it can be concluded that, at the actual membrane state-of-the-art, the SMR is the leading architecture for membrane safety, for the possibility of increasing reaction temperature and for its maintenance easiness. Future improvement in membrane performance, mainly for stability, would promote the applications of IMRs. The case studies analysis, which focused on hydrogen production processes, demonstrates the potential of membrane reactor technology in terms of improving performance and reducing operating temperature, even if some crucial obstacles have yet to be overcome, as developing a reliable fabrication method and reducing membrane costs. With regard to the costs of Pd-based membranes, it is foreseen that there would be a cost of 8900 euros per m2 for a cumulative production of 1000 m2, while the cost would be reduced to 900 euros per m2 if the cumulative production is increased to 10 000 000 m2 worldwide. The main question is how to reach cumulative production of millions of square meters of Pd-based membranes in a few years. The R&D efforts directed towards reducing membrane costs, ensuring the reliability of thin film fabrication methods and demonstrating the benefits of selective membranes in chemical processes would push the introduction of such a new technology into the marketplace.

References 1. A. S. Michaels, Chem. Eng. Prog., 1968, 64, 31. 2. M. De Falco, D. Barba, S. Cosenza, G. Iaquaniello and L. Marrelli, Int. J. Hydrogen Energy, 2008, 33, 5326. 3. M. De Falco, L. Marrelli and G. Iaquaniello, Membrane Reactors for Hydrogen Production Processes, Springer, 2010, ISBN 978-0-85729-150-9. 4. H. M. Van Veen, M. Bracht, E. Hamoen and P. T. Alderliesten, in Feasibility of the Application of Porous Inorganic Gas Separation Membranes in some Large-Scale Chemical Processes. Fundamentals of Inorganic Membrane Science and Technology, ed. A. J. Burggraaf, L. Cot, Elsevier, 1996, 14, pp. 641. 5. J. Mallevialle, P. E. Odendaal and M. R. Wiesner, Water Treatment Membrane Processes, ed. McGraw-Hill Publishers, New York, 1998. 6. F. Gallucci, M. De Falco, S. Tosti, L. Marrelli and A. Basile, Int. J. Hydrogen Energy, 2007, 32, 4052. 7. G. J. Grashoff, C. E. Pilkington and C. W. Corti, Plat. Met. Rev., 1983, 27, 157. 8. F. A. Lewis, K. Kandasamy and B. Baranowski, Plat. Met. Rev., 1988, 32, 22. 9. H. P. Hsieh, AIChE Symp. Ser., 1989, 85, 53. 10. D. J. Edlund and W. A. Pledger, J. Membr. Sci., 1993, 77, 255. 11. S. T. Hwang and K. Kammermeyer, Techniques in Chemistry: Membranes in Separation, Wiley Interscience, New York, 1975.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

134

Chapter 13

12. B. H. Howard, R. P. Killmeyer, K. S. Rothenberger, A. V. Cugini, B. D. Morreale, R. M. Enick and F. Bustamante, J. Membr. Sci., 2004, 241, 207. 13. J. Shu, B. P. A. Grandjean, A. Van Neste and S. Kaliaguine, Can. J. Chem. Eng., 1991, 69, 1036. 14. R. Bredesen and H. Klette, U.S. Patent 6.086.729, 2000. 15. Y. C. Van Delft and L. A. Correia, Palladium Membrane Reactors for Large Scale Production of Hydrogen, 8th International Conference of Catalysis in Membrane Reactors, 18–21 December 2007, Kolkata, India. 16. J. Shu, B. P. A. Grandjean and S. Kaliaguine, Appl. Catal. A: Gen., 1994, 119, 305. 17. Y. Lin, S. Liu, C. Chuang and Y. Chu, Catal. Today, 2003, 82, 127. 18. F. Gallucci, L. Paturzo and A. Basile, Int. J. Hydrogen Energy, 2004, 29, 611. 19. J. Oklany, K. Hou and R. Hughes, Appl. Catal. A: Gen., 1998, 170, 13. 20. G. Madia, G. Barbieri and E. Drioli, Can. J. Chem. Eng., 1999, 77, 698. 21. (a) W. Yu, T. Ohmori, T. Yamamoto, E. Endo, T. Nakaiwa, T. Hayakawa and N. Itoh, Int. J. Hydrogen Energy, 2005, 30, 1071; (b) M. Chai, M. Machida, K. Eguchi and H. Arai, Appl. Catal. A: Gen., 1994, 110, 239. 22. I. Dybkjaer, Fuel Process. Technol., 1995, 42, 85. 23. M. De Falco, D. Barba, S. Cosenza, G. Iaquaniello, A. Farace and F. G. Giacobbe, Membrane Reactors, Special Issue Asia-Pacific J. Chem. Eng., 2009, doi:10.1002/apj.241 24. M. De Falco, A. Basile and F. Gallucci, Membrane Reactors, Special Issue Asia-Pacific J. Chem. Eng., 2010, 5, 179. 25. M. De Falco, G. Iaquaniello, B. Cucchiella and L. Marrelli, Reformer and membrane mod-ules plant to optimize natural gas conversion to hydrogen, in Syngas: Production Methods, Post Treatment and Economics, Nova Science Publishers Inc., 2009, ISBN 978-1-60741-841-2. 26. M. De Falco, Int. J. Hydrogen Energy, 2008, 33, 3036. 27. F. A. N. Fernandes and A. B. Soares, Fuel, 2006, 85, 569. 28. M. De Falco, L. Di Paola and L. Marrelli, Int. J. Hydrogen Energy, 2007, 32, 2902. 29. F. Gallucci, A. Comite, G. Capannelli and A. Basile, Ind. Eng. Chem. Res., 2006, 45, 2994. 30. A. Li, C. J. Lima and J. R. Grace, Chem. Eng. J., 2008, 138, 452. 31. S. Hara, K. Haraya, G. Barbieri and E. Drioli, Desalination, 2008, 233, 359. 32. A. Caravella, F. P. Di Maio and A. Di Renzo, J. Membr. Sci., 2008, 321, 209. 33. P. Bernardo, G. Barbieri and E. Drioli, Chem. Eng. Sci., 2010, 654, 1159. 34. Y. Chen, Y. Wang, H. Xu and G. Xiong, Appl. Catal. B: Environ., 2008, 80, 283. 35. Y. Chen, Y. Wang, H. Xu and G. Xiong, J. Membr. Sci., 2008, 322, 453. 36. F. Gallucci, L. Paturzo, A. Fama’ and A. Basile, Ind. Eng. Chem. Res., 2004, 43, 928.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

Palladium-based Selective Membranes for Hydrogen Production

135

37. A. Mahecha-Boteroa, T. Boyd, A. Gulamhusein, N. Comyn, C. J. Lima, J. R. Gracea, Y. Shirasakic and I. Yasudac, Chem. Eng. Sci., 2008, 63, 2752. 38. Y. Matsumura and J. Tong, Top. Catal., 2008, 51, 123. 39. J. Tong, Y. Matsumura, H. Suda and K. Haraya, Ind. Eng. Chem. Res., 2005, 44, 1454. 40. J. Tong and Y. Matsumura, Catal. Today, 2006, 111, 147. 41. D. Barba, F. Giacobbe, A. D. Cesaris, A. Farace, G. Iaquaniello and A. Pipino, Int. J. Hydrogen Energy, 2008, 33, 3700. 42. G. Iaquaniello, F. Giacobbe, B. Morico, S. Cosenza and A. Farace, Int. J. Hydrogen Energy, 2008, 33, 6595. 43. G. Iaquaniello, P. Ciambelli, V. Palma and E. Palo, Application of membrane reactors to distributed cogenerative systems, in Cogeneration: Types, Technologies and Costs, Nova Science Publishers, Inc. 44. M. De Falco, G. Iaquaniello and A. Salladini, Experimental tests on steam reforming of natural gas in a reformer and membrane modules (RMM) plant, J. Membr. Sci., 2011, 368, 264. 45. P. Chiesa, T. G. Kreutz and G. Lozza, J. Eng. Gas Turbines Power, 2007, 129, 123. 46. S. Uemiya, N. Sato, H. Ando and E. Kikuchi, Ind. Eng. Chem. Res., 1991, 30, 585. 47. A. Basile, A. Criscuoli, F. Santella and E. Drioli, Gas. Sep. Purif., 1996, 10, 243. 48. A. Basile, G. Chiappetta, S. Tosti and V. Violante, Sep. Purif. Technol., 2001, 25, 549. 49. G. Q. Lu, J. C. Diniz da Costa, M. Duke, S. Giessler, R. Socolow, R. H. Williams and T. Kreutz, J. Colloid Interface Sci., 2007, 314, 589. 50. A. Colin Scholes, K. H. Smith, S. E. Kentish and G. W. Stevens, Int. J. Greenhouse Gas Contr., 2010, 4, 739. 51. V. S. Bobrov, N. G. Digurov and V. V. Skudin, J. Membr. Sci., 2005, 253, 233. 52. J. S. Chang, H. S. Roh, M. S. Park and S. E. Park, Bull. Korean Chem. Soc., 2002, 23(5). 53. J. P. Collins, R. W. Schwartz, R. Sehgal, T. L. Ward, C. J. Brinker, G. P. Hagen and C. A. Udovich, Ind. Eng. Chem. Res., 1996, 35, 4398. 54. R. Dittmeyer, V. Ho¨llein and K. Daub, J. Mol. Catal. A: Chem., 2001, 173, 135. 55. P. Quicker, V. Ho¨llein and R. Dittmeyer, Catal. Today, 2000, 56, 21. 56. R. Scha¨fer, M. Noack, P. Ko¨lsch, M. Sto¨hr and J. Caro, Catal. Today, 2003, 82, 15. 57. M. Sheintuch, D. Moshe and R. M., Chem. Eng. Sci., 1996, 51, 535. 58. H. Weyten, J. Luyten, K. Keizer, L. Willems and R. Leysen, Catal. Today, 2000, 56, 3. 59. R. Schafer, M. Noack, P. Kolsch, S. Thomas, A. Seidel-Morgenstern and J. Caro, Sep. Purif. Technol., 2001, 25, 3.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00110

136

Chapter 13

60. J. Galuszka, R. N. Pandey and S. Ahmed, Catal. Today, 1998, 46, 83. 61. A. Basile, L. Paturzo and F. Lagana`, Catal. Today, 2001, 67, 65. 62. J. Galuszka and T. Giddings, Silica membranes-preparation by chemical vapour deposition and characteristics, in Membranes for Membrane Reactors: Preparation, Optimization and Selection, ed. A. Basile, Wiley, Chapter 12. Wiley, 2011, ISBN 978-0-470-74652-3. 63. K. Akamatsu, M. Nakane, T. Sugawara, T. Hattori and S. Nakao, J. Membr. Sci., 2008, 325, 16.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

CHAPTER 14

Polarization and Inhibition by Carbon Monoxide in Palladiumbased Membranes GIUSEPPE BARBIERI,*a ALESSIO CARAVELLAb AND ENRICO DRIOLIc a

National Research Council, Institute on Membrane Technology, ITM-CNR, c/o University of Calabria, Cubo 17C, 87036, Rende (Cosenza), Italy; b National Institute of Advanced Industrial Science and Technology, National Institute for Innovation in Sustainable Chemistry, AIST-ISC, Central 5, 1-1-1 Higashi, 305-8565, Tsukuba (Ibaraki), Japan; c University of Calabria, Department of Chemical and Materials Engineering, Via P. Bucci, Cubo 44A, 87036, Rende (Cosenza), Italy

14.1 Palladium-based Membranes: Overview and Potential for Hydrogen Purification Hydrogen production is becoming progressively more important because of the strategic role that hydrogen could play as an energy carrier.1,2 For example, high-purity hydrogen can be conveniently used to feed proton exchange membrane fuel cells (PEM-FCs) for civil and/or military transport3 more efficiently than traditional internal combustion engines and with a much lower impact on the environment. Hydrogen is produced, essentially, from oxidation and steam reforming of hydrocarbons and water gas shift, which need some stages of upgrading to Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

137

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

138

Chapter 14

produce hydrogen with the required purity. Among the hydrogen purification techniques, membrane-based processes provide significant advantages with respect to more traditional ones, because they could significantly improve the performances of hydrogen production and purification steps, and at the same time reduce the cost related to these processes. Furthermore, the volume plant can be drastically reduced by integrating the membrane technology into the production steps, this being one of the objectives of the process intensification strategy,4 which aims at a better exploitation of the raw materials and resources. For this reason, several membrane types have been studied continuously and tested to investigate their performances in terms of permeating flux, stability and selectivity. In particular, a number of polymeric, metallic and/or ceramic materials have been identified to prepare good membranes for hydrogen separation, but, among these, palladium and its alloys are very important thanks to their characteristic of being permeable to hydrogen only, i.e. infinitely selective with all the other gases.5 However, the cost of these membranes is relatively high, which is the problem that could severely affect their massive development and commercial/industrial applications. Therefore, research is very active in its attempts to make high performance membranes, which could make up for their high fixed costs, and reduce the palladium content, with a consequent decrease of the material cost. The use of Pd-alloys not only allows a decrease of the palladium content in the membrane, but also provides some advantages in performances and/or mechanical resistances. For example, other metals present in the Pd-alloy (e.g. Ag) strongly limit its embrittlement caused by hydrogen permeation in the metallic lattice, and allow a good resistance towards the poisoning species (the Pd-Cu membranes are quite resistant to the contamination of S-, Cl- and Hg-based compounds, especially of H2S).6 Moreover, another way to improve the performances of the Pd-based membranes consists in decreasing its thickness as much as possible, as this reduction allows an increase of permeating flux. However, making good thin membranes, i.e. o3 mm approximately, is generally difficult, because the probability of finding surface defects is higher as the membrane thickness is reduced. In fact, the thin metal layer necessarily has to be deposited on an appropriate porous support in order to acquire the required mechanical resistance. The importance of considering thin membranes is increasing, since the techniques to make very thin selective Pd-based layers have been improving. The traditional deposition processes consist in depositing the metal film directly on the support (e.g. by electroless plating, magnetron sputtering, chemical vapor deposition, physical vapor deposition, etc.). However, in this way, the irregularities of the support surface can strongly limit the very thin membranes required. Another more recent technique consists of (i) depositing the metal film on a silicon wafer, (ii) removing the wafer by an appropriate solvent, and

View Online

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

139

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

7–10

(iii) putting the thin metal layer on the support. This technique is known as micro-fabrication and presents the advantage of producing membranes with a uniform surface. Improving the fabrication techniques is crucial for Pd-based membrane development, because eventual surface defects (in terms of holes) can drastically reduce the membrane selectivity, heavily affecting the purity of the permeation stream.11

14.2 Objectives Mass transfer from bulk to membrane surface is affected by external resistance much more in thin membranes than in thicker ones and, moreover, in the presence of inhibiting species for the membrane, the effective membrane area becomes smaller, thereby causing an additional reduction of the permeating flux. All these phenomena, negative for membrane performances, can cause the validity of Sieverts law (eqn (14.1)) to be compromised because the bulk properties (permeance and permeation driving force) are generally different from those evaluated immediately close to the membrane surfaces: qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi H2 Flux ¼ Permeance   PDownstream PUpstream H2 H2

ð14:1Þ

However, Sieverts law is very simple and, therefore, it would be very useful to be able to continue using it even when it is not strictly valid. In this sense, the aim of this chapter is to show a systematic method to take into account the presence of both inhibition and polarization by keeping the form of Sieverts law, but modifying its original structure by means of an appropriate overall ‘permeation reduction coefficient’, PRC (eqn (14.2)):12 H2 Flux¼ ð 1PRC ÞPermeanceMembrane qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi  PUpstream  PDownstream H2 H2

ð14:2Þ

Bulk

In this equation, whose development and implications will be analyzed in detail throughout this chapter, the only quantity necessary to calculate the flux is justly PRC, since intrinsic Sieverts permeance – i.e. the one not affected by inhibition and/or polarization – and bulk driving force can be simply evaluated from experimental data.

14.3 Gas–surface Interactions for Palladium-based Membranes The hydrogen streams to be purified usually contain several impurities, which can interact with the membrane surface in different ways. Understanding the nature of these interactions is fundamental to analyze the membrane behavior and correctly model the permeation process, which leads to a better design of

View Online

140

Chapter 14

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

the separation/purification modules. With respect to Pd-based membranes, but, also, more generally, to other types of metal membranes, all the nonpermeating gas species (all gases except hydrogen) can be basically divided into the following two groups:  Species adsorbing on metal surface so weakly that the intrinsic membrane permeability is not affected by them. Therefore, these species are commonly regarded as inert species (e.g. N2 or CH4 on palladium).  Species that interact more strongly with the metal surface. This type of interaction, which is competitive with that of hydrogen, causes a decrease of the effective membrane area, because of the consequent reduction of the adsorption sites. The species behaving in this way are usually regarded as inhibiting or poisoning species (e.g. CO, CO2 or H2S on palladium13). However, the analysis of these interactions is made more complicated by the fact that all the non-permeating species, i.e. inert, inhibiting/poisoning ones, are also responsible of a reduction of driving force along the permeation direction. This phenomenon, well known in membrane technology, is termed ‘concentration polarization’ and is generally caused by an accumulation of the lesspermeating species to the surface due to the motion of the more-permeating ones towards the membrane surface. The overall negative effect on membrane performances is that the permeation driving force between the membrane surfaces of upstream and downstream is lower than that between the respective bulks, causing a consequent reduction of flux. A significant presence of concentration polarization can make the advantages of preparing very thin membranes partially useless, because the mass transfer resistance tends to be located in the gas phase external to the membrane rather than in the membrane itself.

14.4 Concentration Polarization in Gas Separation Concentration polarization is present in all membrane processes, even though it has been neglected for a long time in the gas separation field. As mentioned above, this phenomenon occurs because the permselectivity of the membrane itself consists in a decrease of driving force of the most permeable species owing to the presence of the least permeable ones. Its negative effect on membrane performances can be relevant when the selective layer is very thin (o5 mm approx.), as already shown in several works in the literature (see, for example, references 14–23). The first field for which concentration polarization was deeply investigated (since the 1960s17) was liquid separation by membrane processes, such as ultrafiltration and reverse osmosis. On the contrary, for a long time it had been generally accepted that concentration polarization had only a negligible effect on membrane performance in gas separation. This was justified by the fact that membranes were quite thick and permeating flux very low, and, moreover, that

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

141

gas diffusivities are four to five orders of magnitude higher that those of the liquids.17 However, the progress obtained in the last few decades in membrane preparation for gas separation has allowed the transmembrane flux to be significantly improved. High permeating flux causes the external mass transfer to be capable of seriously affecting membrane performances, because greater driving force tends to be concentrated out of the membrane and not in the selective layer. In order to understand this phenomenon and correctly quantifying its effects on permeation under different operating conditions several studies have been conducted. Lu¨dtke et al.18 considered the effect of the concentration polarization on the separation of n-butane/nitrogen by means of a composite (three layers) polymeric membrane. Under their operating conditions, they found that the relative resistance in the boundary layer was so significant as to exceed that in the membrane at a sufficiently high total pressure of feed. He et al.17 used a binary mixture-based film model to perform a theoretical analysis on the concentration polarization in a generic membrane. They defined a concentration polarization coefficient for both the two species involved in the separation as the ratio of the actual flux to the ideal one (without polarization), quantifying the polarization effect by means of the ratio of the actual fluxes of the components. Although this is a simplified approach that cannot be generalized to multi-component systems, nevertheless, under some operating conditions, the authors predicted a significant influence of the external mass transfer on the process. Other researchers used CFD calculation with complex 3D models to study reactive and separating systems with Pd-based membranes20,24 to take into account the concentration polarization phenomenon. However, their attention was strongly directed towards a specific process and, thus, it is very difficult to achieve and extrapolate general considerations about it. Concerning the hydrogen purification by Pd-based membranes, the effect of concentration polarization was specifically investigated experimentally by Peters et al.19 and later by Caravella et al.14 from a modeling point of view. In the first of these works, where the influence of CO and CO2 content in feed was also analyzed, the concentration polarization level was estimated as the ratio between the hydrogen pressure drop in the external gas phases and the whole hydrogen pressure drop through the membrane. A different approach was adopted by Caravella et al., who used a model that divides the hydrogen permeation process into several different steps,25 by considering the presence of a multi-component mixture (six species) on the feed side.14 Once numerically solved, this model provided the hydrogen partial pressure profile through the membrane, as well as the pressure profiles of all the species in the gas phases of the feed and permeate side. From these pressure profiles, the concentration polarization level was evaluated by means of a concentration polarization coefficient (CPC), whose definition was made according to Sieverts permeation driving force.14 Since this is one of the main topics of this chapter, it will be discussed in detail later.

View Online

142

Chapter 14

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

14.5 Inhibition by Carbon Monoxide in Palladium-based Membranes It is well known that CO inhibits the ability of hydrogen to permeate through Pd membranes. The CO-Pd bond involves a physical adsorption phenomenon26–29 that is favored at low temperatures. Several studies on the performance decrease of Pd-based membranes and Pd-based membrane reactors due to inhibition have been carried out,19,30–40 highlighting the under-estimation of the membrane area caused by the non-consideration of the inhibition by CO in the design of a reactor and/or a separation equipment. In fact, the membrane inhibition phenomenon can be relevant not only in gas separation, but also in reactive processes like the water gas shift carried out in Pd-based membrane reactors,41 where CO is involved as a principal reactant at a relatively low temperature (280–350 1C). Some authors observed the effects on superficial modifications induced on Pdbased surfaces by interfacial reactions,42–44 which also modify the intrinsic membrane performances by causing a change in the number of superficial active sites. The importance of understanding CO adsorption on Pd surfaces and its alloys is shown by the number of theoretical and experimental works performed and published in the last four decades. Basically, the main aim of these papers is an investigation of the mutual interaction of CO with several Pd and Pd-based surfaces – Pd(111), Pd(100), etc. – characterizing the adsorption kinetics in terms of activation energy and pre-exponential factor of the kinetic constant as functions of surface coverage and temperature by means of several different techniques.45–68 If correctly used, this information can be very useful to describe correctly the hydrogen permeation reduction due to the competitive co-adsorption of CO on the different types of superficial sites characterizing the membrane surface (hollow sites, bridge–bridge sites and on-top sites51,52,54,61,64–67). However, to perform a quantitative investigation on the hydrogen permeation reduction due to the presence of CO, these data should be input in an appropriate permeation model able to take into account not only the competitive adsorption of hydrogen and CO, but also the concentration polarization in the gas phase caused by the presence of inert species and CO itself.

14.6 Coupled Effect of Concentration Polarization and Inhibition by Carbon Monoxide In order to discuss on the separate effect of concentration polarization and inhibition by CO, the system shown in Figure 14.1 is considered.12 Referring to this sketch, three different zones can be identified:  Upstream side, where a multi-component gas mixtures (three species: H2, CO, N2) is present  Pd-based membrane, which is the actual selective layer responsible for separation  Downstream side, where only pure hydrogen is present.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

Figure 14.1

143

Sketch of the system considered for discussion. The solid and dashed lines indicate the forward and backward flux, respectively.

Both forward and backward permeation are considered in order to analyze both cases where the membrane behaves as hydrogen purifier (forward permeation) or hydrogen supplier (backward permeation). The permeation process is considered to be divided into several elementary steps, each of which is characterized by its own model equations. The details of such a model can be found elsewhere12,25 where the model introduced in14,25 was modified in some steps to take into account the presence of an inhibitor species (CO). The same approach was followed by Catalano et al., who analyzed the inhibiting effect of CO without considering the concentration polarization.69 In the following sub-section, concentration polarization coefficient (CPC) and inhibition coefficient (IC) will be introduced separately in order to show their different contribution on the overall permeance reduction.

14.6.1 14.6.1.1

Concentration Polarization Coefficient Overview on the Definition of Concentration Polarization Coefficient

In order to provide a quantitative indication of the polarization level, several different concentration polarization coefficients can be found in the literature. The most used definitions are summarized in Table 14.1. The first thing to

View Online

144

Table 14.1

Chapter 14

The most common concentration polarization coefficients in literature

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Reference Wijmann et al.70 Yeom et al.71,72 Nagy15,16 Zhao et al.73

Wang et al.74 Zhang et al.20

Haraya et al.21 Takaba et al.22

Expression of the concentration polarization coefficient Upstream Ci;Film   Ci;Bulk  Upstream and  xi;Film Downstream xi;Bulk  Upstream xi;Film  1  xi;Bulk  N1 defined for one component: 0 ; N1 N1 defined for a binary mixture: N2   xi;Film Upstream  xi;Bulk Downstream xi;Bulk jUpstream  xi;Bulk jDownstream

Range of values (0, 1)

(0, 1)

(0, 1) (0, 1) (0, N) (0, 1)

C ¼ molar concentration, x ¼ molar fraction, N ¼ molar flux, N0 ¼ molar flux without polarization.

notice in this table is that all coefficients refer to a key permeating species through the membrane. In particular, considering that there is no significant variation of the total pressure along the flux direction, the definitions by Wijmann et al.,70 Yeom et al.,71,72 Nagy15 and Nagy and Kulcsa`r,16 who use the molar concentration, in fact coincide with that by Zhao et al.,73 who use the molar fractions. All these authors adopt a single coefficient for each membrane side. Hence, the effect of mass transfer resistance in the upstream and downstream side is measured by two different coefficients. With this definition, as polarization becomes more important, the coefficient tends to be progressively closer to zero, going towards the unitary value in the opposite situation (no polarization). Wang et al.74 uses a similar approach to define a concentration polarization coefficient, with the only difference that in their definition the coefficient is close to zero when the polarization is negligible. In order to link concentration polarization and membrane selectivity, Zhang et al.20 used the permeating fluxes in the definition of the coefficient. In the case where only one component is able to permeate through the membrane, this definition is very useful, because it does not depend on the particular permeation mechanism and, thus, is general. However, this approach is valid only for a binary mixture and cannot simply be generalized for a multi-component mixture. Additionally, the overall coefficient varies between zero and infinity, therefore not allowing an immediate perception of the polarization level. A useful definition is provided by Haraya et al.21 and Takabaand Nakao,22 who used a unique coefficient expressed in terms of molar fractions to measure the polarization on both the membrane sides. However, the expressions

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

145

presented in Table 14.1 are empirical and, thus, the concentration polarization coefficient is not explicitly linked to any quantity related to the permeation process, like driving forces (usually the difference between partial pressures) and/or the permeances of the species to be separated. In fact, the polarization coefficient involving molar fractions are equivalent to those expressed as a function of the partial pressure only if the total pressures of feed and permeate are equal, something that does not generally occur.

14.6.1.2

Defining Concentration Polarization Coefificient According to Hydrogen Permeation Driving Force

All of the considerations made in the previous section lead to the necessity of defining a concentration polarization coefficient that could be directly connected to specific permeation quantities.12,14 Here, this approach will be direct to hydrogen permeation through Pd-based membranes. Generally speaking, the transmembrane flux can be expressed by the product of a permeance and the corresponding driving force (eqn (14.3)): H2 Flux ¼ Permeance  Driving Force

ð14:3Þ

This expression is not a way to calculate the permeating flux (known from experiments), but is the simple definition of permeance, which can be calculated by the ratio between flux and driving force (eqn (14.4)): Permeance ¼

H2 Flux Driving Force

ð14:4Þ

Since eqns (14.3) and (14.4) are general they can be written in different locations, i.e. between the membrane surfaces and/or between the bulks of the gas phases on both membrane sides (eqns (14.5) and (14.6)): PermeanceMembrane ¼ PermeanceBulk ¼

H2 Flux Driving ForceMembrane

ð14:5Þ

H2 Flux Driving ForceBulk

ð14:6Þ

It is important to remark that the bulk permeance reported in eqn (14.6) is not a membrane property, but is a coefficient taking into account the ignorance of the transport phenomena occurring in the film adjacent to the membrane surface. In fact, in the presence of concentration polarization, which causes the bulk conditions to be different from those on the membrane surface, the only known quantities are the membrane permeance and the bulk driving force. In fact, the former can be obtained from pure hydrogen tests, whilst the latter is chosen by the user by setting the external pressure conditions. Therefore, it would be useful to have these two quantities in the definition of the concentration

View Online

146

Chapter 14

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Table 14.2

Limit values of CPC according to its definition

Coefficient

Extreme values

Physical meaning

CPC

0 -1

No polarization Total polarization

polarization coefficient, in order for it to assume a more concrete physical meaning. According to these considerations, the concentration polarization coefficient (CPC) can be conveniently defined as reported in eqn (14.7):14 H2 Flux  ð1  CPC Þ  PermeanceMembrane  Driving ForceBulk

ð14:7Þ

By comparing the expressions in eqn (14.7), this definition can be read in two ways: when concentration polarization is negligible, i.e. CPC E 0, membrane permeance is, in fact, coincident with that of bulk, or, dually, the permeation driving force between the bulks is the same as that between the membrane surfaces. It should also be noticed that the condition of maximum polarization, i.e. CPC ¼ 1, is an asymptotic conditions where permeating flux tends to zero because the hydrogen partial pressure at membrane surface approaches to zero. This situation is summarized in Table 14.2. Hence, according to its definition (eqn (14.7)), CPC can be evaluated from eqns (14.8) to (14.10) (all of them equivalent to each other), where it is underlined that the product between membrane permeance and bulk driving force corresponds to the hydrogen flux evaluated in presence of pure hydrogen on both membrane sides: CPC ¼ 1 

H2 Flux PermeanceMembrane  Driving ForceBulk

ð14:8Þ

H2 Flux H2 FluxjPure H2

ð14:9Þ

CPC ¼ 1 

CPC ¼ 1 

Driving ForceMembrane PermeanceBulk ¼1 Driving ForceBulk PermeanceMembrane

ð14:10Þ

This fact is very important, since it shows that the CPC value is independent of the choice of permeance and driving force. In fact, both the fluxes (general and pure hydrogen one) can be simply evaluated experimentally or from the numerical solution of an appropriate permeation model, once the external conditions of temperature and pressures have been set.14,25,75 However, to relate CPC to the membrane permeance according to eqn (14.1), it is necessary to choose an appropriate form of the bulk driving force. Concerning this issue, it should be considered that, when the rate-determining step of the overall permeation process is the diffusion in the Pd-based layer18,23,25,75,76 and, also, the atomic hydrogen concentration inside membrane

View Online

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

147

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

77

is low enough to consider the Pd-H system as infinitely diluted, Sieverts law (eqn (14.1)) describes very well the behavior of the permeating flux. Since Sieverts law is valid or the deviation from it is small in most of the operating conditions under which the Pd-based membranes work, the square root difference of hydrogen partial pressure (Sieverts driving force) seems a good choice for the characteristic driving force to define the concentration polarization coefficient. This particular choice of driving force is not a limitation and does not affect the generality of this approach. In fact, CPC can be calculated for each operating condition, even for those where Sieverts law is not strictly valid. In practice, CPC represent an ‘ignorance’ correction factor by means of which it is possible to continue to use Sieverts law even when diffusion in the selective metal layer is not the rate-determining step anymore. According to the CPC definition reported in eqn (14.8), CPC was evaluated as a function of several operating conditions, like temperature, feed and permeate pressure, feed composition and fluid-dynamic conditions by means of an elementary steps permeation model (eqn (14.11)):14    H2 Flux   ð14:11Þ CPCModel ¼ 1   H2 Flux Pure H2 From solution of the model

The simulation results obtained were reported in terms of polarization maps, which were developed in order for CPC to be read directly and used under the considered operating conditions to estimate the transmembrane flux once the intrinsic membrane permeance and Sieverts driving force of bulk are known (eqn (14.12)). The results from the model solution have been confirmed by a comparison with the experimental ones in order to validate the analysis: H2 FluxjPredicted ¼ ð1 CPCModel Þ  PermeanceMembrane qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi  PDownstream  PUpstream H2 H2

ð14:12Þ

Bulk

14.6.1.3

Definition of Concentration Polarization Coefficient in the Presence of Inhibition

The definition of concentration polarization coefficient adopted in the previous section (eqns (14.8) to (14.10)) is general and, thus, can be conveniently applied to different situations. In particular, CPC can be also used in a situation in which permeation is affected by inhibition, as reported in eqn (14.13) in terms of both driving forces and permeances: CPC ¼ 1  or

Driving ForceInhibited Membrane Driving ForceInhibited Bulk

CPC ¼ 1 

PermeanceInhibited Bulk PermeanceInhibited Membrane

(14.13)

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

148

Chapter 14

These definitions, formally different with respect to those reported in eqn (14.10), approach each other in the case of negligible inhibition. This approach is very useful because, with this choice, CPC is directly related to the flux reduction due only to external mass transfer resistance and not to the inhibition phenomenon. Moreover, the polarization and inhibition effect are able to be separately identified and split into their own different contributions, which can thus be analyzed to provide a better understanding of the coupled influence of these two phenomena. This is done by defining another coefficient, i.e. the inhibition coefficient, IC, which is on the other hand a quantitative indicator of the inhibition phenomenon only (see next section).

14.6.2

Inhibition Coefficient

As anticipated in the previous section, an inhibition coefficient is necessary in order to evaluate conveniently the inhibition effect only without the polarization one. Analogously to what was done for CPC, an inhibition coefficient, IC, was defined according to eqn (14.14), where the superscript ‘clean’ and on permeance indicates that membrane is not affected by inhibition12: IC ¼ 1 

PermeanceInhibited Membrane PermeanceClean Membrane

ð14:14Þ

This definition, applied in the specific case to the inhibition by CO, is general and reflects the deep difference existing between the effect of polarization and inhibition on permeation. In fact, polarization acts outside the membrane, whereas inhibition affects its intrinsic permeance by competitively occupying the adsorption site and decreasing the effective nominal membrane area. However, the two phenomena are strongly linked to each other, since polarization favors the inhibiting species towards the membrane surface and, on the other hand, inhibition tends to reduce the permeating flux decreasing to a certain extent the polarization itself. This is the reason why coefficients by which to decouple polarization and inhibition are useful to represent easily this complex membrane behavior. The effect of inhibition by CO on hydrogen permeation has recently become the object of several investigations, since CO is typically present in significant amounts in the hydrogen streams to be separated. In particular, a macroscopic model equation was developed to describe the influence of CO on permeation. This equation modifies Sieverts law to take into account the permeance reduction due to CO as reported in eqn (14.15), where the quantities a and yCO represent a coefficient accounting for other additional effects of CO on adsorbed hydrogen and the membrane surface coverage by CO, respectively:36 H2 Flux ¼ ½1 aðTÞ  yCO ðT; PCO Þ  Permeance Membrane qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi Upstream  PH2  PDownstream H2

ð14:15Þ

View Online

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

149

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

While the parameter a is substantially an empirical factor, analytical forms are provided for the surface coverage, yCO,36 as shown in eqn (14.16), where the hydrogen adsorption is considered undisturbed by the presence of CO: yCO ¼

KCO ðTÞPCO 1 þ KCO ðTÞPCO

ð14:16Þ

Rigorously speaking, the CO partial pressure present in eqn (14.16) is that immediately adjacent to the membrane surface. Therefore, in order to estimate correctly the parameters that appear in this expression, experimental systems that minimize the concentration polarization are necessary. This can be done by using a thick membrane36 or choosing the fluid-dynamic conditions in such a way for the external mass transfer resistance to be negligible. However, if the first method is quite simple to perform, the second one requires much more effort, because turbulence promoters or very high flow rate are necessary. The expression reported in eqn (14.16) allows the inhibition coefficient to be expressed explicitly. In fact, comparing eqns (14.14) to (14.16), it is possible to note that IC assumes the form reported in eqn (14.17), where the subscript ‘Membrane’ has been added to the hydrogen and CO partial pressures to highlight that all the variables should be evaluated in correspondence of the membrane surface and not of the fluid bulk: IC ¼ aðTÞ

KCO ðTÞPCO;Membrane 1 þ KCO ðTÞPCO;Membrane

ð14:17Þ

Hence, eqn (14.15) can be re-written (eqn (14.18)) to show that, analogously to CPC, IC also allows Sieverts law to be used even when the validity of Sieverts hypothesis is questionable: H2 FluxInhibited ¼ ð1  ICÞ  PermeanceClean Membrane  or

qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi qffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi PUpstream PDownstream H2;Membrane H2;Membrane 

H2 FluxInhibited ¼ ð1  ICÞ  H2 FluxClean Sieverts

(14.18)

The practical importance of CPC and IC is relevant for membrane engineering, because the performance of a Pd-based membrane (in terms of permeating flux and/or permeance) can be estimated easily once they are known. The values of CPC and IC were calculated under several working conditions12 by reporting them in the form of maps from which to read their values directly (see Section 14.6.4.2).

14.6.3

Overall Permeation Reduction Coefficient

According to their definitions, CPC and IC provide quantitative measures of the flux decrease due to concentration polarization and inhibition, respectively. However, it is sometimes useful to have an indication of the overall flux

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

150

Chapter 14

reduction due to both phenomena, in order to make the membrane module design more easily. For this aim, an overall ‘permeance reduction coefficient’, PRC,12 was introduced, to take into account the permeation reduction owing to both concentration polarization and inhibition at the same time. Its definition, reported in eqn (14.19), arises from the simple consideration that the lowest permeance is evaluated at bulk conditions when both polarization and inhibition affect the membrane system, whereas, on the other hand, the ideal membrane behavior and the maximum permeance can be found under pure hydrogen conditions:

PRC ¼ 1 

PermeanceInhibited Bulk PermeanceClean Membrane

ð14:19Þ

Therefore, it represents the distance of the actual system status from the best possible conditions, providing an immediate quantitative measure of the membrane performance loss. Since, due to its nature, this coefficient takes into account polarization and inhibition, it can be expressed in terms of the other two coefficients, CPC and IC, by combining eqns (14.13), (14.14) and (14.19) to obtain eqn (14.20): PRC ¼ 1  ð1  CPCÞð1  ICÞ or PRC ¼ CPC þ IC  CPC  IC

(14.20)

From these equations, it is clearly shown that the effects of polarization and inhibition are not simply additive, but are coupled to each other in a more complex relation. In particular, analyzing the second expression in eqn (14.20), it is possible to notice that the second-order term CPC  IC is subtractive. This remarkable fact is not only a mere mathematical aspect, but represents a very important physical aspect. In fact, as mentioned above, when polarization and inhibition occur at the same time, the presence of inhibition causes the flux to be lower that that measured with polarization only. Since the effect of polarization itself is larger at higher permeating flux, the overall reduction is less than the sum of the two Table 14.3

Summarization of the reduction coefficients considered in this chapter

Coefficient Meaning CPC IC PRC

Definition

Concentration polar- PermeanceInhibited ¼ ð1  CPCÞ  PermeanceInhibited Bulk Membrance ization coefficient Clean Inhibition coefficient PermeanceInhibited Membrance ¼ ð1  ICÞ  PermeanceMembrance Clean Permeation reduction PermeanceInhibited ¼ ð1  PRCÞ  Permeance Bulk Membrance coefficient

View Online

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

151

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

single contributions of inhibition and polarization. A summary of all the reduction coefficients introduced here is reported in Table 14.3. Similarly to what was done for the coefficients CPC and IC, also for PRC some permeation reduction maps were calculated and developed. These maps are reported and discussed in Section 14.6.4.2.

14.6.4

Main Results of Analysis

In this section, some of the main results obtained on the separate and combined effect of inhibition and polarization are presented and discussed. Further details concerning the validation of the analysis can be found elsewhere.12

14.6.4.1

Permeance and Flux Analysis

In this section, the analysis is focussed on the direct effect that inhibition by CO and concentration polarization have on hydrogen permeance and transmembrane flux. To do that, the operating conditions reported in Table 14.4 are considered.12 As shown in the table, two mixtures are considered to separate the CO contribution to concentration polarization and inhibition. Figure 14.2 shows the hydrogen permeance as a function of the hydrogen upstream composition for the two mixtures, whose behavior is qualitatively similar. In this figure, several situations are analyzed:    

Permeation Permeation Permeation Permeation

without polarization and inhibition by CO with polarization only with inhibition only by CO with both polarization and inhibition at the same time.

As regards the first case (horizontal dashed lines), it is possible to notice that the membrane permeance is constant. This is because it represents an intrinsic property of the material, which, thus, cannot be influenced by any external fluid-dynamic factor. When concentration polarization is taken into account, the permeance evaluated according to the bulk driving force (dashed-line named ‘Bulk’) is significantly different with respect to the membrane one. This is due to the external mass transfer influence, whose resistance becomes higher as the composition of the other species progressively increases. Table 14.4

Operating conditions considered by Caravella et al.12 Molar fraction, Mixture 1

Mixture 2

Zone

H2

CO

N2

H2

CO

Total pressure, kPa

Fluid dynamic conditions

Upstream Downstream

0. . .1 1

0. . .0.5 —

0. . .0.5 —

0. . .1 1

0. . .1 —

400. . .1000 200

Re E 1200 Not influent

Temperature ¼ 374 1C; membrane thickness ¼ 60 mm.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

152

Figure 14.2

Chapter 14

Bulk and membrane hydrogen permeance as a function of the hydrogen upstream molar fraction. PUpstream ¼ 1000 kPa and PDownstream ¼ 200 kPa. Adapted from Caravella et al.12

In the third case, the presence of inhibition only by CO (indicated in the figure by the upper continuous line) is considered. It should be remarked that, rigorously speaking, this represents an ideal case, because, in a more or less significant way, the inhibiting species also generate concentration polarization. Differently from the latter, inhibition affects directly the adsorption properties of the membrane surface, causing the occupation of the active sites on which the hydrogen molecules dissociates. Also in this case, the gap between membrane permeance with and without inhibition represents a direct quantification of this phenomenon. The presence of a plateau is because inhibition by CO is an equilibrium-limited phenomenon. In the latter case, both polarization and inhibition are considered. As stated before, the presence of these two phenomena at the same time makes the system analysis more complex, since they are mutually dependent on each other. Analyzing the profiles in the plot, for high hydrogen molar fraction (approximately higher than 0.9) no appreciable difference between inhibition alone and inhibition affected by polarization can be observed. This occurs because of the small content of the other species (CO and N2 for mixture 1 and CO for mixture 2), which do not provide significant polarization effects. However, the CO presence strongly affects the membrane performance in terms of inhibition, which is significant even at high hydrogen content. As this

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

Figure 14.3

153

Hydrogen flux as a function of hydrogen upstream molar fraction. PUpstream ¼ 1000 kPa and PDownstream ¼ 200 kPa. Adapted from Caravella et al.12

becomes lower, the effect of polarization becomes progressively more important, whereas inhibition approaches its equilibrium value. This situation can be analyzed also in terms of hydrogen flux (Figure 14.3) under the same conditions as previously considered. Examining the case of forward flux, it is possible to notice that the flux with the sole polarization is higher than the one with the inhibition alone, both of them being higher than the flux evaluated in presence of both contributions. In the case of back flux, a different situation occurs, which shows the inhibition and polarization fluxes to be significantly different with respect to each other only up to a hydrogen molar fraction of about 0.03, both approaching to the same value towards the zero-flux line. This occurs because, along its path, the back flux encounters first the resistance of the membrane (larger) and, then, that provided by the mixture on the other side (smaller), whereas the opposite situation is found for the forward-flux. The net result of these resistances in series favors the membrane used as hydrogen distributor more than the membrane used as selective means for purification/separation. When both inhibition and polarization are considered, another effect favors the back flux with respect to the forward one. In fact, when the driving force is high enough in conditions of back-permeation, the CO partial pressure is moved away by the effect of the flux from the membrane surface to the bulk: a fact that reduces the negative effect of CO.

14.6.4.2

Permeation Reduction Maps

In this section, the reduction coefficients introduced before (CPC, IC and PRC) are presented in the form of the so-called ‘permeation reduction maps’, which

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

154

Chapter 14

represent the most useful way to analyze the membrane behavior in the presence of polarization and/or inhibition, as well as to read the coefficients values and use them directly to estimate the actual permeating flux. For this goal, their values have been evaluated by Caravella et al.12 under the operating conditions reported in Table 14.4, even though other conditions can be considered, depending on the particular membrane systems specifications. According to its definition, the coefficient PRC is able to deal with the possible situations in which a membrane system affected by concentration polarization and inhibition can operate (Table 14.5). In Figure 14.4, all the three reduction coefficients are reported as functions of upstream CO partial pressure. The dashed line indicates the concentration polarization presence without considering the inhibition by CO. Considering mixture 1 (plot on the left hand side), at high hydrogen molar fraction, the polarization caused by CO and N2 does not provide an appreciable contribution. Therefore, the dashed and solid lines are practically coincident and, consequently, the system is more sensitive to the CO presence, evidenced by Table 14.5

Different situations dealt with by PRC

Permeation reduction coefficient PRC ¼

Figure 14.4

Equivalent coefficient

Considered phenomena

0 IC CPC PRC

Neither inhibition nor polarization Only inhibition Only polarization Inhibition and polarization at the same time

PRC, IC and CPC as functions of CO partial pressure. The dashed lines related to CPC are evaluated in absence of inhibition. PUpstream ¼ 1000 kPa and PDownstream ¼ 200 kPa. Adapted from Caravella et al.12

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

155

an important inhibition effect, whose level increases significantly up to about 100 kPa, from which its profile approaches a plateau value. For these reasons, the system is controlled by inhibition and, thus, the values of the overall permeation reduction coefficient PRC are practically the same as those of IC. On the contrary, from a CO partial pressure of 100 kPa on, polarization influence starts to be relevant and PRC curves begin to be separated from the IC one (split point) tending in this way towards the behavior of CPC coefficient: a fact that can be seen from the slope change in the curve. Therefore, the ‘split point’ shown by IC and PRC represents the lower limit of a range of conditions within which the system is driven by both polarization and inhibition. However, considering that the plateau value of IC indicates the maximum level after which the system cannot be affected by inhibition anymore, a method to quantify separately the effect of inhibition and polarization can be developed, even when only CO and H2 are present (mixture 2). In fact, the case represented by the second mixture clearly shows that it is theoretically impossible to reach the plateau value by performing only permeation tests, since after a certain level of CO the polarization begins to be appreciable. However, the polarization effect is progressively lower as higher membrane thicknesses are considered. Therefore, the equilibrium value can be estimated with good approximation and, thus, the difference between thin and thick membrane behavior can be attributed to concentration polarization. It is also interesting to notice that the polarization coefficient CPC is higher without an inhibition effect. This aspect can be discussed also considering Figure 14.5, where CPC is reported as a function of the hydrogen molar fraction for different total upstream pressures.

Figure 14.5

CPC reported as a function of the hydrogen upstream molar fraction. PUpstream ¼ 1000 kPa and PDownstream ¼ 200 kPa. Adapted from Caravella et al.12

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

156

Chapter 14

From these plots it is possible to see that the CPC coefficient is a decreasing function of the upstream hydrogen molar fraction. Furthermore, high upstream total pressures favor concentration polarization because of a higher permeating flux. A novel result shown in this and in the previous figure is represented by the difference between systems with and without inhibition (i.e. with CO considered as inert). This last (virtual) case, shown at 1000 kPa only to evidence that inhibition and concentration polarization affect each other, clearly demonstrates that in the presence of inhibiting species the polarization decreases because of a lower permeating flux. Analyzing the behavior of the inhibition coefficient IC (Figure 14.6), it is shown that it continuously decreases for both mixtures with hydrogen molar fraction (i.e. increasing with CO partial pressure), approaching the plateau value according to the Sievert–Langmuir equation (first expression in eqn (14.16)).36 Concerning the dependence on upstream total pressure, the inhibition coefficient increases with it because of a higher CO partial pressure. However, the total pressure effect on IC is progressively lower, this being a direct consequence of the interaction of the nature of the inhibition phenomenon, which is equilibrium limited. Finally, in order to observe quantitatively both the influence of concentration polarization and inhibition by CO, the behavior of the overall permeation reduction coefficient is shown in Figure 14.7, which represents a map analogous to the ones presented for CPC and IC. As seen before, PRC has a nonlinear dependence on CPC and IC, which, therefore, cannot simply be added to each other to measure the overall permeation reduction. It has also been shown that PRC assumes the characteristic

Figure 14.6

IC reported as a function of the hydrogen upstream molar fraction. PUpstream ¼ 1000 kPa and PDownstream ¼ 200 kPa. Adapted from Caravella et al.12

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

Figure 14.7

157

PRC reported as a function of the hydrogen upstream molar fraction. PUpstream ¼ 1000 kPa and PDownstream ¼ 200 kPa. Adapted from Caravella et al.12

shape of CPC or IC in dependence on the mixture content in terms of hydrogen molar fraction and CO partial pressure. That is the reason for the typical sigmoidal shape of the PRC curves and is an explanation for why the effect of the total pressure is progressively less appreciable also for this coefficient, becoming negligible from about 1000 kPa on. From this fact, considering that the permeating flux progressively increases with the total pressure because of a higher driving force, it is possible to conclude that the most convenient operating conditions to reduce the effect of the permeation reduction in a membrane purification system are established at high total pressure. In fact, in these conditions, the incremental decrease of membrane permeance owing both to polarization and inhibition is minimized. Moreover, it would certainly be convenient to adopt module configurations capable of favoring the mass transfer between bulk and membrane surface, because in this way not only can the concentration polarization influence be minimized, but also the inhibition effect, which on the other hand, cannot be avoided at all.

14.7 Concluding Remarks In this chapter, an innovative approach dealing with the combined effect of concentration polarization and inhibition by CO in Pd-based membranes is discussed and analyzed by means of appropriate coefficients measuring the permeation reduction due to these phenomena. According to their definition, these coefficients – namely, concentration polarization coefficient (CPC), inhibition coefficient (IC) and the overall

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

158

Chapter 14

permeation reduction coefficient (PRC) – allow Sieverts law to be extended to other conditions where the hypotheses on which its validity is based are not strictly fulfilled (i.e. high polarization and/or inhibition level). A remarkable result obtained from such an investigation consisted in the quantification of the mutual interactions between concentration polarization and inhibition, for which the non-additivity of the effects was shown in terms of their respective coefficients. In fact, it was mathematically demonstrated that the physical phenomenon according to which the concentration polarization in the presence of inhibition is less than in its absence due to a lower permeating flux. All reduction coefficients dealt with in this chapter were calculated from the numerical solution of a complex hydrogen permeation model, opportunely validated by a number of experimental data,12,14,25 and reported in form of ‘permeation reduction maps’ for several operating conditions. From these maps, the values of CPC, IC and PRC can be read directly and used simply in the modified forms of Sieverts law shown above to evaluate the hydrogen permeating flux, providing in this way a possible novel strategy of membrane module design. The approach to the permeation reduction by polarization and inhibition introduced in this chapter for Pd-based membranes is a general concept that. in principle. can be applied to all the other types of membranes considering several reduction factors in different situations. This step can be helpful in developing a standard procedure of membrane module design.

14.8 List of Symbols and Abbreviations Symbols KCO P a

Equilibrium constant in eqn (14.14) Pressure Parameter in eqn (14.14)

Pa–1 Pa —

Concentration polarization coefficient Inhibition coefficient Permeation reduction coefficient

— — —

Abbreviations CPC IC PRC

Acknowledgement The project FIRB-CAMERE (RBNE03JCR5) ‘Repubblica Italiana, Ministero dell’Universita` e della Ricerca (MUR)’ is gratefully acknowledged for co-funding this research.

References 1. B. D. Morreale, M. V. Ciocco, R. M. Enick, B. I. Morsi, B. H. Howard, A. V. Cugini and K. S. Rothenberger, J. Membr. Sci., 2003, 212, 87–97.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

159

2. M. V. Ciocco, B. D. Morreale, K. S. Rothenberger, B. H. Howard, A. V. Cugini, R. P. Killmeyer and R. M. Enick, in Proceedings 19th Pittsburgh Coal Conference, 2002, Paper 49–3. 3. Fuel Cell Handbook 7th edn., EG&G Technical Services, Inc., 2004. 4. J. K. Keil, Modeling of Process Intensification, Weinheim: Wiley-VCH Verlag GmbH KGaA, 2007. 5. R. Dittmeyer, V. Ho¨llein and K. Daub, J. Mol. Catal. A: Chem., 2001, 173, 135–184. 6. B. H. Howard, R. P. Killmeyer, K. S. Rothenberger, A. V. Cugini, B. D. Morreale, R. M. Enick and F. Bustamante, J. Membr. Sci., 2004, 241, 207–218. 7. R. Bredesen and H. Klette, Method of Manufacturing Thin Metal Membrane. U.S. Patent 6,086,729, 2000. 8. H. Klette, E. Raeder, Y. Larring and R. Bredesen in GRACE: Development of Supported Palladium Alloy Membranes, (ISBN 008044570), ed. D. C. Thomas and S. M. Benson, Elsevier, 2005, 21, 377–384. 9. H. D. Tong, F. C. Gielens, J. W. Berenschot, M. J. de Boer, J. G. E. Gardeniers, W. Nijdam, C. J. M. van Rijn and M. C. Elwenspoek, J. Microelectromech. Syst., 2003, 12, 622–629. 10. H. D. Tong, F. C. Gielens, J. G. E. Gardeniers, H. V. Jansen, J. W. Berenschot, M. J. de Boer, J. H. de Boer, C. J. M. van Rijn and M. C. Elwenspoek, J. Microelectromech. Syst., 2005, 14, 113–123. 11. A. Caravella, F. P. Di Maio and A. Di Renzo, Asia-Pacific J. Chem. Eng., 2010, 5, 213–225. 12. A. Caravella, F. Scura, G. Barbieri and E. Drioli, J. Phys. Chem. B, 2010, 114, 12264–12276. 13. C. P. O’Brien, B. H. Howard, J. B. Miller, B. D. Morreale and A. J. Gellman, J. Membr. Sci., 2010, 349, 380–384. 14. A. Caravella, G. Barbieri and E. Drioli, Sep. Purif. Technol., 2009, 66, 613–624. 15. E. Nagy, Sep. Purif. Technol., 2010, 73, 194–201. 16. E. Nagy and E. Kulcsa`r, Desalin. Water Treat., 2010, 14, 220–226. 17. G. He, Y. Mi, P. L. Yue and G. Chen, J. Membr. Sci., 1999, 153, 243–258. 18. O. Lu¨dtke, R. D. Behling and K. Ohlrogge, J. Membr. Sci., 1998, 146, 145–157. 19. T. A. Peters, M. Stange, H. Klette and R. Bredesen, J. Membr. Sci., 2008, 316, 119–127. 20. J. Zhang, D. Liu, M. He, H. Xu and W. Li, J. Membr. Sci., 2006, 274, 83–91. 21. K. Haraya, T. Hakuta and H. Yoshitome, Sep. Sci. Technol., 1987, 22, 1425–1438. 22. H. Takaba and S. Nakao, J. Membr. Sci., 2005, 249, 83–88. 23. F. C. Gielens, H. D. Tong, M. A. G. Vorstman and J. T. F. Keurentjes, J. Membr. Sci., 2007, 289, 15–25. 24. T. P. Tiemersma, C. S. Patil, M. van Sint Annaland and J. A. M. Kuipers, Chem. Eng. Sci., 2006, 61, 1602–1616.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

160

Chapter 14

25. A. Caravella, G. Barbieri and E. Drioli, Chem. Eng. Sci., 2008, 63, 2149–2160. 26. G. Ertl and J. Koch in Adsorption-Desorption Phenomena: Proceedings of the II International Conference, held in Florence in April 1971, F. Ricca (Ed), Academic Press, London, 1972, pp. 345–357. 27. P. W. Davies and R. M. Lambert, Surf. Sci. Lett., 1981, 111, L671–L674. 28. D. Prigge, W. Schlenk and E. Bauer, Surf. Sci. Lett., 1982, 123, L698–L702. 29. A. Ortega, F. M. Huffman and A. M. Bradshaw, Surf. Sci., 1982, 119, 79–94. 30. S. Hara, K. Sakaki and N. Itoh, Ind. Eng. Chem. Res., 1999, 38, 4913–4918. 31. D. A. King and M. G. Wells, Proc. R. Soc. London, Ser. A, 1974, 339, 245–269. 32. B. E. Koel and R. J. Smith, Surf. Sci., 1990, 231, 325–332. 33. A. L. Mejdell, H. Klette, A. Ramachandran, A. Borg and R. Bredesen, J. Membr. Sci., 2008, 307, 96–104. 34. J. A. Rabo, A. P. Risch and M. L. Poutsma, J. Catal., 1978, 53, 295–311. 35. F. Roa and J. D. Way, Appl. Surf. Sci., 2005, 240, 85–104. 36. Barbieri, F. Scura, F. Lentini, G. De Luca and E. Drioli, Sep. Purif. Technol., 2008, 61, 217–224. 37. A. L. Mejdell, M. Jondahl, T. A. Peters, R. Bredesen and H. J. Venvik, Sep. Purif. Technol., 2009, 68, 178–184. 38. A. L. Mejdell, D. Chen, T. A. Peters, R. Bredesen and H. J. Venvik, J. Membr. Sci., 2010, 350, 371–377. 39. F. C. Gielens, R. J. J. Knibbeler, P. F. J. Duysinx, H. D. Tong, M. A. G. Vorstman and J. T. F. Keurentjes, J. Membr. Sci., 2006, 279, 176–185. 40. F. Scura, G. Barbieri, G. De Luca and E. Drioli, Int. J. Hydrogen Energy, 2008, 33, 4183–4192. 41. G. Barbieri, A. Brunetti, G. Tricoli and E. Drioli, J. Power Sources, 2008, 182, 160–167. 42. H. Uchikawa, T. Okazaki and K. Sato, Jpn. J. Appl. Phys., 1993, 32, 5095–5096. 43. F. Guazzone, E. E. Engwall and Y. H. Ma, Catal. Today, 2006, 118, 24–31. 44. D. Wang, J. D. Clewley, T. B. Flanagan, B. Balasubramaniam and K. L. Shanahan, J. Alloys Compd., 2000, 298, 261–273. 45. F. Abild-Pedersen and M. P. Andersson, Surf. Sci., 2007, 601, 1747–1753. 46. S. E. Mason, I. Grinberg and A. M. Rappe, Phys. Rev. B, 2004, 69, 161401(R)-1-4. 47. E. Ozensoy and D. W. Goodman, Phys. Chem. Chem. Phys., 2004, 6, 3765–3778. 48. V. V. Kaichev, I. P. Prosvirin, V. I. Bukhtiyarov, H. Unterhalt, G. Rupprechter and H. J. Freund, J. Phys. Chem. B, 2003, 107, 3522–3527. 49. H. Cordatos, T. Bunluesin and R. J. Gorte, Surf. Sci., 1995, 323, 219–227. 50. V. V. Gorodetskii, A. V. Matveev, P. D. Cobden and B. E. Nieuwenhuys, J. Mol. Catal. A: Chem., 2000, 158, 155–160. 51. I. V. Yudanov, R. Sahnoun, K. M. Neyman, N. Rosch, J. Hoffmann, S. Schauermann, V. Johanek, H. Unterhalt, G. Rupprechter, J. Libuda and H. J. Freund, J. Phys. Chem. B, 2003, 107, 255–264.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00137

Polarization and Inhibition by Carbon Monoxide in Palladium-based Membranes

161

52. K. Honkala, P. Pirila¨ and K. Laarsonen, Surf. Sci., 2001, 489, 72–82. 53. N. Lopez and J. K. Nørskov, Surf. Sci., 2001, 477, 59–75. 54. S. Surnev, M. Sock, M. G. Ramsey, F. P. Netzer, M. Wiklund, M. Borg and J. N. Andersen, Surf. Sci., 2000, 470, 171–185. 55. B. Gleich, M. Ruff and R. J. Behm, Surf. Sci., 1997, 386, 48–55. 56. K. Fukui, H. Miyauchi and Y. Iwasawa, J. Phys. Chem., 1996, 100, 18795–18801. 57. B. Hammer, Y. Morikawa and J. K. Nørskov, Phys. Rev. Lett., 1996, 76, 2141–2144. 58. D. R. Rainer, M. C. Wu, D. I. Mahon and D. W. Goodman, J. Vac. Sci. Technol. A, 1996, 14, 1184–1188. 59. L. Pedocchi, M. R. Ji, S. Lizzit, G. Comelli and G. Rovida, J. Electron Spectrosc. Relat. Phenom., 1995, 76, 383–387. 60. R. D. Ramsier, K. W. Lee and J. T. Yates Jr., Surf. Sci., 1995, 322, 243–255. 61. X. Xu, P. Chen and D. W. Goodman, J. Phys. Chem., 1994, 98, 9242–9246. 62. J. Szanyi, W. K. Kuhn and D. W. Goodman, J. Vac. Sci. Technol. A, 1993, 11, 1969–1974. 63. X. Guo and J. T. Yates Jr., J. Chem. Phys., 1989, 90, 6761–6766. 64. M. A. Chesters, G. S. McDougall, M. E. Pemble and N. Sheppard, Surf. Sci., 1985, 164, 425–436. 65. M. P. Kishinova and G. M. Bliznakov, Surf. Sci., 1982, 123, 61–76. 66. R. J. Behm, K. Christmann and G. Ertl, Surf. Sci., 1980, 99, 320–340. 67. R. J. Behm, K. Christmann, G. Ertl and M. A. Van Hove, J. Chem. Phys., 1980, 73, 2984–2995. 68. H. Conrad, G. Ertl, J. Koch and E. E. Latta, Surf. Sci., 1974, 43, 462–480. 69. J. Catalano, M. G. Baschetti and G. C. Sarti, J. Membr. Sci., 2010, 362, 221–233. 70. J. G. Wijmann, A. L. Athayde, R. Daniels, J. H. Ly, H. D. Kamaruddin and I. Pinnau, J. Membr. Sci., 1996, 109, 135–146. 71. C. K. Yeom, S. H. Lee, J. M. Lee and H. Y. Song, J. Membr. Sci., 2002, 204, 303–322. 72. C. K. Yeom, S. H. Lee, J. M. Lee and H. Y. Song, J. Membr. Sci., 2002, 205, 155–174. 73. S. Zhao, Z. Li, Y. Liu and L. Wang, Desalination, 2008, 233, 310–318. 74. R. Wang, S. L. Liu, T. T. Lin and T. S. Chung, Chem. Eng. Sci., 2002, 57, 967–976. 75. T. Ward and T. Dao, J. Membr. Sci., 1999, 153, 211–231. 76. R. A. Oriani, In Proceedings of 4th. International Conference on Cold Fusion. Lahaina, Maui: Electric Power Research Institute, Palo Alto, CA, 1993. 77. S. Hara, M. Ishitsuka, H. Suda, M. Mukaida and K. Haraya, J. Phys. Chem. B., 2009, 117, 9795–9801.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

CHAPTER 15

Carbon Molecular Sieve Membranes for Gas Separation MAY-BRITT HA¨GG* AND XUEZHONG HE Department of Chemical Engineering, Norwegian University of Science and Technology, N-7491 Trondheim, Norway

15.1 Introduction Carbon molecular sieve (CMS) membranes have been studied for more than 20 years as a promising candidate for energy-efficient gas separation technology. Much interest has been shown in the preparation of carbon membranes for separation of gas mixtures such as CO2–N2, O2–N2 and CO2–CH4. The first carbon membranes were prepared from the carbonization of cellulose hollow fibers by Koresh and Soffer.1 After that, many different polymer precursors were used to prepare the CMS membranes, including polyimide,2–4 polyacrylonitrile (PAN),5 poly(phthalazinone ether sulfone ketone),6 poly(phenylene oxide)7,8 and cellulose derivatives.9,10 The carbon molecular sieve membranes are more expensive than polymeric membranes due to the increased need for man-hours in the production processes; however, theyhave the advantages of better permeability and selectivity as well as higher thermal and chemical stability.6,11–17 These key advantages have encouraged many researchers since the 1980s to investigate and develop carbon molecular sieve membranes for gas separation. The attention has focused on the carbon membranes that exhibit molecular sieving properties, which can exceed the Robeson upper boundary of permeability versus selectivity trade-off relationship,18 as shown in Figure 15.1.7,15,19–21 where also the region for industrial applicability is suggested by Hillock et al.22 Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

162

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

Figure 15.1

163

Comparing the CO2/CH4 Robeson upper bound for dense and thermally rearranged (TR) polymer membranes18 to the carbon membranes,7,15,19–21 and the region for industrial applicability was suggested by Hillock et al.22 (Data for CMS membranes and industrial applicability region added to the original Robeson plot.)

Carbon molecular sieve membranes can be divided into two categories: unsupported and supported carbon membranes.23 Unsupported membranes have three different configurations: flat film, hollow fiber and capillary tubes, while the supported carbon membranes involve two configurations: flat and tube. Detailed descriptions of these two categories can be found in the review of Ismail and David.24 The supported carbon membranes have better mechanical stability than the unsupported carbon membranes, but the preparation process is much more complex. The supported carbon membranes are typically prepared by coating the supports with a thin, uniform polymer layer. Although many different techniques can be used such as ultrasonic,25 dip coating,26 vapor deposition,27 spin coating28 and spray coating,29 there are still some challenges to successfully prepare the supported membrane: (i) control the amount of material being deposited on the support; (ii) produce a uniform layer; and (iii) produce the defect-free layer. The unsupported carbon membranes, i.e. hollow fibers, are prepared from the unsupported polymeric precursors. The spinning conditions are crucial for making the precursor fibers, and should be well controlled in order to form a good fiber for carbonization. The choice of supported or unsupported carbon membranes will mainly depend on the application. Normally, the hollow

View Online

164

Chapter 15

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

fiber modules are chosen for most of gas separation processes due to the high packing capacity compared to flat-sheet membranes.5 Hence, also the unsupported hollow fiber carbon membranes are promising for future application.

15.2 Production of Carbon Molecular Sieve Membranes Preparing a carbon membrane from a precursor is easy, but producing a high performance carbon membrane is a quite difficult task, since it includes many steps that must be well controlled and optimized. The fabrication process for CMS membranes normally consists of six important steps, i.e. material selection, material functionalization, precursor preparation, pre-treatment, carbonization and post-treatment, as illustrated in Figure 15.2. Each step includes many parameters which need to be optimized in order to obtain a high performance membrane. Among these steps, the carbonization process is the most important and can be regarded as the heart of the CMS membrane fabrication process.30 Methods to control the carbonization conditions for making an optimized carbon membrane are described in Section 15.2.5.

15.2.1

Material Selection

The chemical structure and physical properties of the polymer should be primarily considered for the choice of polymer materials. However, there are only a few literature reports on the influences of chemical structure of polymers on the properties of the derived carbon membranes, e.g. by Park et al.31 and Xiao et al.32 The latter reported the structure and properties relationship for polymers based on the experiment and simulation approaches, which provided considerable information for the choice of suitable polymers for carbon

Figure 15.2

Schematic diagram of fabrication process for CMS membranes.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

165

membrane preparation in various targeted applications. Hence it would be an efficient way to choose the polymer by investigating the factors of chemical structure and physical properties for determining the carbon membrane performance based on experiments and molecular simulation method. A suitable polymer material for preparation of carbon membranes should not cause pore holes or any defects after the carbonization. Up to now, various precursor materials such as polyimide,2,3 polyacrylonitrile (PAN),5 poly(phthalazinone ether sulfone ketone)6 and poly(phenylene oxide)7,8 have been used for the fabrication of carbon molecular sieve membranes. Likewise, aromatic polyimide and its derivatives have been extensively used as precursor for carbon membranes due to their rigid structure and high carbon yields. The membrane morphology of polyimide could be well maintained during the high temperature carbonization process. A commercially available and cheap polymeric material is cellulose acetate (CA, MW 100 000, DS ¼ 2.45); this was also used as the precursor material for preparation of carbon membranes by He et al.19 They reported that cellulose acetate can be easily dissolved in many solvents to form the dope solution for spinning the hollow fibers, and the hollow fiber carbon membranes prepared showed good separation performances.

15.2.2

Material Functionalization

In order to enhance the selectivity and permeability of the carbon molecular sieve (CMS) membrane, the addition of other components to the carbon matrix is considered. For CMS membranes, two types of additives have been reported. The first type of additive increases the micropore volume of the carbon membrane by degrading during the carbonization process (temperature range 500–1000 1C), and leave behind specific spacing within the mass of the carbon. Such additives are often referred to as porogens and can serve as templates in the formation of microporosity in carbon, e.g. polyvinylpyrrolidone (PVP).33,34 The second kind of additive is the nano-functional additives and can be thought of as thermally stable compounds incorporated into the carbon membrane precursor, either before or after casting or spinning. These components may enhance the gas separation process if the interactions between the additive and the penetrants can be exploited for an enhanced transport rate through the carbon membranes. Obviously, the variety of nano-functional additives which can be incorporated into the CMS membrane is limited due to the high temperatures applied during carbonization which will result in the degradation of all organic components. Possible additives include metals (added as metal salts) which show a high affinity to one of the permeating gas species, silica nano-particles and carbon nanotubes; these are all listed in Table 15.1.6,35–46 The addition of metal nitrates have the additional porogen effect as the nitrates degrade releasing gases during the carbonization procedure.20 Metal oxides are thought to be unsuitable due to their low solubility in organic solvents such as NMP and their affinity for water which may result in blocking of the metallic sites. The use of nickel for the adsorption of CO2 has previously

View Online

166

Chapter 15

Table 15.1

Representative examples of functionalization additives Additives

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Metal nitrates and metal oxides Alkali metals Transition metals Silver Nickel Other metal additives Silica nano-particle Zeolite Carbon nanotube

Function

Fe(NO3)3, Cu(NO3)2, Increase the polarity and/or Ag NO3, MgO, form interlayer spaces and/ or have an affinity towards CaO, SiO2, Fe2O3 target gases Na1 and Mg21 Cause steric hindrance in the carbon matrix Cu21, Ni21 or Zn21 Increase macropore volume Ag nanoclusters behaving as AgNO3, AgC2H3O2 spacers within the carbon matrix Nickel nano-particles A strong chemisorption of the H2 on the nickel particles Palladium and Palladium acting as a gas platinum permeation barrier to other gases Show molecular sieving SiO2 properties Provide transport pathways for specifics gasses Single-walled or Changing of charge patterns, multi-walled used as a compacting agent nanotube for the polymer blends

References 36–38

38 38 35 6 44, 45 43 40, 42 39, 41, 46

been demonstrated and this may help to enhance CO2 transport through the membrane. Also, the alkali metals have been found to increase gas selectivity and would be a useful additive if the decrease in permeability could be overcome. Another promising additive appears to be the carbon nanotubes which can be tailored to various sizes and are highly temperature resistant. Example of enhanced separation properties is shown in Figure 15.3.3,20,47

15.2.3

Precursor Preparation

The general process for preparation of the precursors consists of four steps, i.e. dope formation, casting/spinning, dehydration and post-treatment. There are many parameters that will affect the precursor properties during the preparation process. An example for the optimization of spinning condition was reported by He et al., who reported that the optimal conditions for spinning cellulose acetate hollow fiber membranes was found to be as follows: bore fluid, water þ NMP (85%); air gap; 25 mm; bore flow rate, 40% of dope flow rate (2.2 mL min1); and temperature of quench bath, 50 1C.48

15.2.4

Pretreatment

The precursor membranes are often pre-treated prior to the carbonization/ carbonization process. This step can be helpful to ensure the stability of the

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

Figure 15.3

167

Separation performance of carbon and metal loaded carbon for the CO2/ CH4 gas pair at 30 1C. Dots are literature values for carbonized polyimides,3,47 the other marks are from a cellullosic precursor.

precursor and retain the chemical structure during the carbonization. In some degree, the performance of CMS membranes can be adjusted by specific pretreatment for a given precursor membrane. Saufi and Ismail30 reviewed the pre-treatment methods for precursor membranes published in the open literature before 2003. The pre-treatment approaches can be divided into physical and chemical methods.

15.2.4.1

Physical Pre-treatment

The physical pre-treatment methods for hollow fiber membranes mainly consist of stretching or drawing. This technique used in CMS membranes is sometimes referred to as a post-spinning treatment, which can remove the surface defects and enhance the retention of molecular orientation prior to the carbonization so as to obtain a good balance of stiffness and strength. The draw can take place during the spinning process or after it, and the draw ratios can become very high if the fiber is not ruptured. Yoneyama and Nishihara49 reported that the drawing can be carried out under conditions which give three times or higher total draw by multi-stage drawing method.

15.2.4.2

Chemical Pre-treatment

Chemical pre-treatment includes air oxidation and use of chemical reagents. The oxidation pre-treatment is considered very important and can have a

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

168

Chapter 15

substantial effect on the resulting membrane performance, and the aim is to contribute to the stabilization of the asymmetric structure of the precursor and provide sufficient dimensional stability to undergo the high carbonization temperature. Some researchers reported that different oxidation conditions can be applied at various ranges of thermal soak times, depending largely on the precursor choices. The chemical pre-treatment can enhance the uniformity of the pore system formed in the carbonization process. Hydrazine, DMF, hydrogen chloride (HCl) and ammonium chloride (NH4Cl) can be used for chemical pre-treatment. Schinedler and Maier50 reported that an aqueous solution of hydrazine was used to pre-treat the acrylic precursor, which can improve the dimensional stability of membrane during the subsequent process. Tin et al.51 pointed out that the as-spun hollow fiber membrane was immersed in p-xylenediamine/methanol solution in order to form the cross-sectional morphology. The deacetylation of cellulose acetate precursor was carried out by He et al.19 to obtain the optimal precursor before the carbonization as reported.

15.2.5

Carbonization

The CMS membranes are prepared by carbonizing (under pyrolysing conditions) the precursor membranes in a high temperature tube furnace, as shown in Figure 15.4. The step-by-step method (several dwells) most commonly used as the protocol for the carbonization process is described elsewhere.20,51,52 Many researchers report different carbonization conditions in their research works illustrating very well that each precursor will need different protocols in order to be pore tailored for specific applications.9,13,53,54 The carbonization process is the most important step for fabrication of CMS membranes and is used to tailor the pore size and structure of the carbon membranes. Therefore, how to control the carbonization conditions is crucial for the resulting CMS membrane performance. Su and Lua reported that the statistical 24–1 factorial

Figure 15.4

A schematic overview of the furnace set-up.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

169

experiment design was used to evaluate the influence of carbonization conditions on the membrane transport properties. The parameters of purge gas, carbonization temperature, heating rate and thermal soak time were employed, and the influence of importance for each parameter and interaction between them were found.55 Wang et al. reported the influence of different carbonization degrees on the carbon membrane performance. They pointed out that the CO2 diffusion coefficient in the CMS membranes depends on the surface heterogeneity of the membrane sample and increases with the degree of carbonization.56 Geiszler et al. investigated the effect of the polyimide carbonization conditions such as purge gas, purge flow rate and temperature on the carbon membrane performance.57 They concluded that the vacuum carbonization could prepare more selective but less productive CMS membranes than the inert gas carbonized membranes, and the high purge flow rates could result in a much higher permeability, but lower selectivity membranes. Moreover, by increasing the final temperature, the membranes become more selective but less productive. In order to systematically investigate the influences of carbonization parameters on the membrane properties, the orthogonal experimental design (OED) and conjoint analysis was executed by He et al.19 They reported that an orthogonal experimental design for four factors with three levels was conducted to optimize the carbonization process. The conjoint analysis in Statistical Package for the Social Sciences (SPSS) software method was employed for the statistical analysis of OED results. The influence importance for these factors was sorted out in the following order: purge gas 4 final temperature 4 heating rate 4 final soak time. The optimal carbonization condition was obtained with CO2 pyrolyzing atmosphere, final temperature 823 K, heating rate 4 K min1 and 2 h final soak time. The high performance hollow fiber CMS membranes were then prepared under these optimal conditions.

15.2.6

Post-treatment

After the carbonization process, the precursor membranes are transformed into the CMS membranes, which have different porosity, structure and separation performance that depend on the carbonization conditions. The CMS membrane performances can be partly adjusted by the application of various posttreatment methods, i.e. post-oxidation, chemical vapor deposition (CVD),58 coating, post-carbonization. The post-oxidation is the most used method to change the carbon membrane pore structure. Kusakabe et al.26 studied the post-oxidation of a polyimide-derived carbon membrane in the air. They found that the CO2 permeability increased without any significant change in selectivity. The chemical vapor deposition can be used to introduce the organic species into the carbon matrix and can give three different results: homogeneous deposition, adlayer deposition and in-layer deposition. The coating technique is mainly used to repair the defects in the carbon matrix in order to compose the high selectivity. However, the coating will typically result to the decrease of permeability as reported by Liang et al.59 Some literature reported

View Online

170

Chapter 15

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

the use of different post-treatment methods for altering the membrane structure in order to improve the membrane performance.60–63

15.3 Characterization for Carbon Molecular Sieve Membranes 15.3.1

General Characterization Techniques

15.3.1.1

Scanning Electron Microscopy

Scanning electron microscopy with a resolution of about 0.5 mm is often used as a standard technique to examine the membrane morphology. The cross section views the thickness and diameter of the membranes, as shown in Figure 15.5.19 The carbon membrane forms a symmetric structure and the dimensions were significantly smaller than the precursors a shrinkage of 30% could be detected.

15.3.1.2

Fourier Transform Infrared Spectroscopy

Fourier transform infrared spectroscopy (FTIR) can be used to determine the chemical functional group in the carbon membranes. The FTIR spectra of precursor and carbon membranes with different final carbonization temperatures using CO2 as purge gas as well as in vacuum environment is illustrated in

Figure 15.5

SEM for deacetylated CA precursor and carbon membrane.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

Figure 15.6

171

FTIR spectra of precursor and carbon membranes obtained at different conditions.

Figure 15.6. Most peaks disappeared for carbon membranes when the temperature was higher than 550, and the new characteristic absorption peaks were found at 2350 cm1 and 670 cm1, which contribute to the CO2 adsorbed in carbon matrix or C¼O bond formed in the membrane surface and the aromatic ¼C-H out of plane deformation.15 In vacuum condition, the characteristic absorption peak of CO2 also appears in the FTIR spectrum which indicates the CO2 comes out during the decomposition of deacetylated cellulose acetate and adsorbs strongly in the carbon matrix.

15.3.1.3

X-ray Photoelectron Spectroscopy

X-ray photoelectron spectroscopy (XPS) is a powerful tool to study the surface elemental compositions of materials, and can be used to determine the trend of carbon content followed by the change of carbonization temperature. Figure 15.7 shows the XPS spectra of the original PPESK membrane and the carbon membranes obtained at different carbonization temperatures, revealing that for all membranes, carbon, oxygen, nitrogen and sulfur are the main elements on the membrane surface as reported by Zhang et al.6 They reported that in this carbonization step, rearrangement reactions between the poly- and heterocyclic aromatic nitrogen-containing compounds in carbon structure take place.

15.3.1.4

X-ray Diffraction

X-ray diffraction (XRD) is a useful tool for studying the arrangement of carbon atoms at molecular level. The inter-planar distance and its variation can be

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

172

Chapter 15

Figure 15.7

XPS spectra of the original PPESK membrane and as-prepared carbon membranes.

Figure 15.8

XRD pattern for carbon membranes (a) C(PPO), (b) C(TMS80).7

monitored by XRD, so it has been well established that the d-spacing can serve as indicative of the graphitization degree of the examined carbon membranes since the d-spacing of graphite is 0.335 nm. The XRD patterns for the carbon membrane C(PPO) and C(TMS80) are shown in Figure 15.8.7 The average d-spacing (d002) values were calculated from the Bragg equation, providing the inter-layer distance of the carbon membranes as 0.41 nm and 0.40 nm, respectively. The inter-layer distance can be considered as a diffusional path for gas molecules through the carbon membranes, which is helpful to evaluate the microstructure of the carbon membranes. The intensity of d002 perk of C(TMS80) was somewhat lower than that of C(PPO), indicating that the microstructure of C(TMS80) is arranged less orderly and tightly.

View Online

Carbon Molecular Sieve Membranes for Gas Separation

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

15.3.2

173

Gas Sorption

The gas sorption can be well used to determine the carbon membrane structural parameters such as micropore volume, pore size and pore size distribution. The gas adsorption isotherm data for the carbon membranes can be obtained by a Robutherm magnetic suspension balance (MSB) having a 0.01 mg resolution and 0.02 mg reproducibility. The MSB overcomes the other conventional gravimetric sorption instruments by separating the microbalance from the sample and adsorbed gases.64 The sample is placed in a suspended basket by a permanent magnet through an electromagnet in a closed system. The MSB instrument can perform the sorption measurements within a wide pressure range up to 35 bar and 150 bar for CO2 and N2, respectively. Moreover, the temperature can be well controlled within the range 298–423 K using a Julabo thermostatic circulator. The system can automatically measure the weight change of the samples over time at a certain temperature and pressure according to the measurement procedure described elsewhere.65 The pore size distribution for the carbon membranes can be determined by the method proposed by Nguyen and co-workers.66,67 CO2 isotherm adsorption at 301 K up to 5 bar was executed by Lagorsse et al.68 The adsorption equilibrium isotherms and Dubinin–Radushkevich regression are shown in Figure 15.9. The structural parameters were estimated by the above method and a narrow pore size distribution was found for the tested carbon membranes.

Figure 15.9

Adsorption equilibrium of CO2 on sample MS1 (m), MS1-T600 (.) and MS2 (K and J) at 301 K. The solid lines correspond to Dubinin– Radushkevich fitting equations.

View Online

174

15.3.3

Chapter 15

Gas Permeation

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

The hollow fiber CMS membranes prepared were mounted into a module, and the schematic drawing of the carbon membrane module is shown in Figure 15.10.30

15.3.3.1

Single Gas Tests

Single gas tests are quite important for CMS membranes as they will also give an indication of the membranes pore size. The tests are conducted at varying temperature and feed pressure (permeate side evacuated) in a standard pressure-rise set-up (MKS Baratron pressure transducer, 0–100 mbar range) with LabView data logging, and the schematics of the gas permeation set-up was shown in Figure 15.11. The order of testing was always H2, N2, CH4 and finally CO2 in order to prevent the strongly adsorbing gases from disturbing the performance of the more ideal or non-interacting gases in carbon membranes.16 The tests can be run from several minutes to several hours, to ensure that the transient phase of diffusion is passed and the diffusion steady state had been obtained (dp/ dt is constant). The gas permeability, Pe (in Barrer, where 1 Barrer ¼ 1010 cm3

Figure 15.10

A typical carbon module structure.

Figure 15.11

A schematic diagram for gas permeation test setup.

View Online

Carbon Molecular Sieve Membranes for Gas Separation 2

175

1

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

(STP) cm (cm s cmHg) ) can be calculated using standard procedures reported elsewhere.15 The ideal selectivity is defined as the ratio of the single gas permeabilities and can be evaluated as follows: aA= ¼ B

PeA PeB

ð15:1Þ

Figure 15.12 illustrates the gas permeability values of H2, CO2, O2, N2 and CH4 versus the gas molecule kinetic diameters for the carbon membranes of C(PPO) and C(TMS80) carbonized at 923 K.7 The gas permeability values of the selected gases are in this order: H2 (2.89 A˚) 4 CO2 (3.3 A˚) 4 O2 (3.46 A˚) 4 N2 (3.64 A˚) 4 CH4 (3.8 A˚) at 298 K, which clearly indicated that the molecular sieving transport mechanism was dominated for the gas penetrates through the carbon membranes. In order to compare the performance for polymeric and carbon membranes, Figure 15.13 shows a CO2/CH4 trade-off line for P84 and Matrimid precursors and their carbon membranes as reported by Tin et al.15 It is clear that carbon membranes possess excellent permeation properties, where both of the permeability and ideal selectivity access the Robeson upper-bound curve. Moreover, some researchers have also investigated the influence of temperature on the gas permeability.7,69 They concluded that the gas permeability values increased with the increase of temperature due to the activated process for the CMS membranes. They also found that the apparent activation energies for CO2 calculated from the Arrhenius equation (Pe ¼ Pe0 exp(–Ea/RT)) was much smaller than the other gas species of O2, N2 and CH4, thereby indicating that CO2 has much higher permeability.

Figure 15.12

Single gas permeabilities of C(PPO) (K) and C(TMS80) (’) as function of the gas molecules kinetic diameter at 298 K.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

176

Chapter 15

Figure 15.13

15.3.3.2

CO2/CH4 trade-off line for P84 and matrimid precursors and their carbon membranes.

Mixed Gas Measurements

A permeation cell and a gas chromatograph (GC) were combined in order to allow straightforward determination of gas permeability. The permeability of component i in the gas mixture under steady state of diffusion can be calculated according to the following equations: Pei ¼

273  107 yi V l ðdp=dtÞ ðxi PH  yi PL ÞATexp xi ¼

xF;i  xR;i xF;i ln xR;i

ð15:2Þ ð15:3Þ

where Pei is the permeability of component i. PH and PL are the upstream and downstream pressure (bar), and xF,i is the feed composition of component i, and xR,i and yi are the molar fraction of component i in retentate and permeate stream, respectively, which were measured by GC. The process selectivity (Si/j) for the gas mixture was calculated using the equation: Si=j ¼

Pei Pej

ð15:4Þ

The pure gas tests are normally used to indicate the ideal separation performance for carbon membranes. However, the separation properties will be affected by the presence of other penetrants in a gas mixture.15 Since the transport for gas mixture will be much different from that in pure gas, especially in the presence of strong adsorbable gas like CO2, the adsorption of gas molecules in carbon membranes matrix will significantly affect the penetration of other less or

View Online

Carbon Molecular Sieve Membranes for Gas Separation

177

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

non-absorbable gas molecules. Tin et al. reported the CO2/CH4 separation properties for carbon membrane (CMP84-800).15 They found that the selectivity of CO2/CH4 in the binary mixture is about 10% higher than the ideal selectivity. This is due to the hindrance effect on the CH4 permeation brought upon by the CO2 molecules.

15.3.4

Aging and Regeneration

Although the carbon membrane has high thermal and chemical resistance, they may present significant problems related to performance stability which appears to be more vulnerable to oxidation, humidity and blockage of the pores. A small change of the pore size will dramatically affect the permeability. Therefore, the carbon membrane aging should always be investigated, and the corresponding regeneration methods should be conducted to recover the membrane performance periodically.

15.3.4.1

Humidity Effect

Aging is the change of membrane performance over time or in different environments. The most relevant aging effects include the physisorption (e.g. N2, CO2 and water) and chemisorption (propylene and O2). Jones and Koros studied the influences of water vapor on the carbon membrane performance.70 They found that the performance loss increased with the humidity. The vulnerability of CMS membranes to humidity is a complex phenomenon considering the weak character of the water–carbon dispersion forces and the tendency of water molecules to form hydrogen bonds within the bulk phase.71 Water will initially adsorb onto hydrophilic sites and further chemisorb the penetrants. The hydrophilic sites are much more reactive than the atoms in the interior of the carbon matrix. Once the first water molecule is adsorbed onto the carbon matrix, the adsorbate–adsorbate interactions will promote the adsorption of further molecules through hydrogen bonds.68 The water vapor adsorption and the gas permeance exposure to the different relative humidity have been investigated by Lagorsse et al.71 They concluded that the humidity effect must not be considered as a pore blocking mechanism associated with a slowly diffusing strongly adsorbed species, but as a competitive multi-component diffusion process.

15.3.4.2

Chemisorption of Oxygen

Lagorsse et al.72 further reported the long-term exposure to different dry environments, and they concluded that the membrane performance losses were mainly caused by the chemisorption of oxygen. The reaction of oxygen with carbon matrix is believed to involve dissociative adsorption of molecular oxygen to form oxygen surface groups and subsequent desorption of the surface oxides to the carbon monoxide and dioxide.

View Online

178

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Table 15.2

Chapter 15

Summary of different regeneration method

16

Method

Energy demand

Operation

Complexity

Thermal Chemical (e.g. C3H6) Electrothermal Ultrasonic Microwaves

High Medium Low Low Medium

Offline Offline Online Online Offline

Medium Medium Low Low Low

15.3.4.3

Regeneration Techniques

Most carbon membranes do not have a long-time stable permeability, especially not the high flux membranes. The carbon membrane performance will be gradually reduced due to the pore blockage or aging effects on the carbon matrix. This may be compared to the fouling of membranes in liquid separation. Therefore, regular regeneration techniques such as thermal, chemical, electrothermal, ultrasonic or microwave regeneration are needed. Menendez and Fuertes reported regeneration in vacuum at 600 1C for 1 h to improve the N2 permeance,73 while Jones and Koros investigated the chemical regeneration with propene.74 They supposed that the propene may interact with the carbon matrix in two ways. First, it may act as a solvent, dissolving the penetrants that are adsorbed in the carbon matrix, and. second, the propene may swell the carbon matrix, and the arrangement of the carbon skeleton may release some of the adsorbed gas molecules. Lie and Ha¨gg reported that in-line electrothermal regeneration method can be used to efficiently desorb the adsorbed CO2 with a direct current.9 A review about the detail regeneration technique can be found elsewhere.16 Choosing a suitable regeneration method will mainly depend on the energy demand, operation type and complexity as given in Table 15.2.

15.4 Theory on Transport Mechanisms for Carbon Molecular Sieve Membranes The ability of a micro-porous carbon fiber to separate gases depends on the pore size of the membrane, the physicochemical properties of the gases and surface properties of the membrane pore. The pore size of a carbon fiber for gas separation is usually within the range of 3.5–10 A˚ depending on the conditions for preparation of the membrane during the carbonization or treatment afterwards (post-oxidation or chemical vapor deposition). The transport mechanism for CMS membranes are basically taking place according to one of the three mechanisms listed below, and as described by Ha¨gg et al.75 in:  Knudsen diffusion, hence the square root of the ratio of the molecular weights will give separation factor  Selective surface diffusion governed by a selective adsorption of the larger non-ideal components on the pore surface, hence retaining the smaller components from permeation

View Online

Carbon Molecular Sieve Membranes for Gas Separation

179

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

 Molecular sieving, hence the smallest molecules will permeate and the larger be retained.

15.4.1

Knudsen Diffusion

For Knudsen diffusion to take place, the lower limit for pore diameter has usually been set to dpore 4 20 A˚.76 Gilron and Soffer77 have, however, discussed thoroughly how Knudsen diffusion may contribute to transport in even smaller pores, and from a model considering pore structure, shown that contributions to transport may both come from activated transport and Knudsen through one specific fiber. It may thus be difficult to know exactly when transport due to Knudsen diffusion is taking place. One way to approach this problem is to calculate the Knudsen number, NKnudsen, for the system, which is l/dpore, where l is the mean free path. If NKnudsen Z 10, then the separation can be assumed to take place according to Knudsen diffusion.78 Therefore, if the preparation of the carbon membranes has been unsuccessful, one may get Knudsen diffusion.

15.4.2

Selective Surface Flow

The driving force for separation according to a surface selective flow is basically the difference in the concentration of the adsorbed phase of the diffusing components. This means that a large driving force can be attained even with a small partial pressure difference for the permeating component. The larger molecules (more condensable, e.g. hydrocarbon) in a gas mixture will be selectively adsorbed, hence the smaller molecules will be retained due to reduced pore size. The pore size region where selective surface flow is expected to take place is about 5 A˚ o dpore o 10 A˚ or up to 3  diameter of the molecule.76 The movement of gas molecules in a carbon membrane can be well described by Fick’s first law which gives for the uni-dimensional flux Ji of component i through the membrane: dci ð15:5Þ Ji ¼ Di dxi Here Di is the diffusion coefficient for component i and dci/dxi is the driving force. The activated diffusion can be described by an Arrhenius type of equation: Da ¼ D0  expðEd =RT Þ

ð15:6Þ

where Ed is the activation energy for diffusion. Now if Henry’s law is assumed to apply, the integrated flux equation may be written as in eqn (15.7):     Dp ðEa;S  Eads Þ Dp DES Ja ¼ D0  exp D0  exp ¼ RT  l RT RT  l RT

ð15:7Þ

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

180

Chapter 15

DES, the difference in transport activation energy and adsorption energy may be positive or negative. When DES o 0, transport due to selective surface flow will increase with decreasing temperature; with DES 4 0 it will decrease. Plainly stated, adsorption and hence selectivity increase with decreasing temperature. This is the opposite of temperature influence for molecular sieving separation (see below).

15.4.3

Molecular Sieving

Molecular sieving is the dominating transport mechanism where carbon membranes are applied; this has also given the name to these membranes, CMS. The pore size is usually within the range of a few angstrom (3–5). The dimensions of a molecule are usually described either with the Lennard-Jones radii or the Van der Waal radii. The sorption selectivity has little influence on the separation when molecular sieving is considered. Equation (15.6) is still valid for the activated transport, but now  attention should be drawn to the pre-exponential term,79 D0 ðD0 ¼ el2 kT h expðSa;d =RÞ. So the flux for single component (eqn (15.5)) can be expressed as follows if Henry’s law is introduced: Ja ¼

  Dp Ea;MS D0 exp RTl RT

ð15:8Þ

Here Ea,MS is the activation energy for diffusion in the molecular sieving process for CMS membranes. Nguyen et al. reported that the CMS membrane presents reasonable sieving effect for gas molecules with different kinetic diameters, which suggests that the CMS membrane is predominantly micro-porous with no major contribution from Knudsen diffusion or viscous flow in its overall mass transfer.67

15.5 Module Construction The choice of module design for CMS membranes will typically be the hollow fiber module with counter-current flow. Membrane module construction is, however, seldom referred in open literature as details on this will typically be confidential information for a company producing membrane modules. To date, only tubular and hollow fiber laboratory scale modules have been reported for carbon membranes.13,58,80 The potential industrial use of these membranes has been reported by two companies: Carbon Membranes Ltd (Israel) in the late 1990s, and, later, Blue Membranes GmbH (Germany). Carbon Membranes Ltd produced hollow fibers on a pilot scale and demonstrated successful separation for various applications, while Blue Membranes developed a new concept based on the honeycomb membrane module configuration (HM)81 for their carbon membranes. Neither of these two companies succeeded in taking their CMS membranes all the way to the market. There are, however, new companies which will take advantage of the superior separation

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

181

properties the CMS membranes have, and will be able to develop them for various applications in the market, as suggested in Section 15.6. Saufi and Ismail30 reported that all system designs for module must consider the factors of production cost, maintenance, efficiency. A typical laboratory scale module published in the open literature is shown in Figure 15.10. Although the hollow fiber configuration for module is most commonly used in commercial applications,62 their assembly in high packing density modules has also proved to be difficult for the CMS membranes due to the challenge of the relatively brittle fibers. The mechanical strength of the fibers is therefore a main focus for the commercial development of carbon membranes this can be improved both by choice of a good precursor, and also by developing an optimized carbonization protocol.

15.6 Potential Industrial Applications for Carbon Molecular Sieve Membranes There are several potential industrial applications for the CMS membranes some which are close to market, others which may be more future applications. This may be a function of both the volume of the gas streams, and/or challenging process conditions. Closest to market is the upgrading of biogas to vehicle fuel and separation of air by the use of carbon membranes.

15.6.1

Biogas

Biogas is the gas mixture produced by microbial digestion of organic waste (from households, agriculture, fish industry, waste water treatment etc.) without the presence of oxygen, also called anaerobic decomposition of organic matter. The biogas consists mainly of CH4 (50–75%) usually referred to as biomethane, and CO2, but will also contain some NH3 and H2S. If produced in a more open landfill, there will also be some N2 present due to leakage into the system. However, in the European Union, it is no longer allowed (by 2009) to dispose of organic waste in landfills; handling of organic waste is strictly regulated. Under controlled conditions the amount of biomethane produced from organic waste can be optimized by using a micro-aerated digester as described by Bakke et al.82 The biomethane is a valuable energy carrier, and the use of this gas gives no net contribution of CO2 to the atmosphere when burnt. Since CH4 is a very potent greenhouse gas (around 24 times stronger than CO2), an actual reduction in greenhouse gas emissions is achieved when biomethane is burnt. Biogas is already being utilized in various ways as energy carrier, such as electricity production, being burnt for local heating or combined heat and power generation. By upgrading the biogas to the quality of vehicle fuel and purified natural gas, it may be injected into a natural gas grid or used in the transport sector for cars, buses and trucks. If used for vehicle fuel (ignition engines), the CH4 content must be minimum 96%, while it is sufficient with

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

182

Chapter 15

Figure 15.14

A typical biogas upgrading process.84

80–90% CH4 for so-called dual-fuel engines (tractors, machinery).83 The upgrading process to high quality biomethane is illustrated in Figure 15.14.84As illustrated in the flow sheet, H2S and water vapor must be removed before it is led to the membrane for separating CO2 from CH4. The compression of the gas may vary depending on whether it goes to the gas grid or will be used for vehicle fuel. Biomethane for vehicle fuel must be compressed to around 200 bar, while the pressure will be less if injected into the gas grid (o80 bar). The dew point of the final gas should beo–80 1C. There are several technologies available for upgrading of biogas, such as pressure swing adsorption, physical/chemical absorption and cryogenic separation. These technologies have high energy demands and waste issues, and are not economical for gas streamso200 N m3 h1. The upgrading of biogas using CMS membranes has, however, been found to be especially favorable for gas streams for these smaller to medium gas streams, and the company MemfoACT85 is currently starting production of carbon membrane modules for production of high purity biomethane for a steadily increasing market within the transport sector.

15.6.2

Natural Gas

The purification of natural gas by removal of CO2 (natural gas sweetening) is, in principle, the same separation process as for upgrading of biogas, although it

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

183

is at a much higher feed pressure which is also very favorable for the process. Depending on where in the world the gas production field is found, the pressure, temperature and composition will vary to a large extent. The content of CO2 is typically very low in the North Sea gas fields (o5%), while in other places, such as the Far East and the Mexican Gulf, it can be very high (>40%). As the driving force for the separation of gases when using a membrane typically depends on a high partial pressure difference over the membrane, it is usually favorable with a high CO2 content combined with the high feed pressure of the natural gas (often up to 100 bar). The selectivity for CO2/CH4 measured with hollow fiber carbon membranes prepared from cellulosic precursors, was found to be 4100, and with a CO2 permeability around 100 Barrer.19 For natural gas, water needs to be removed to avoid formation of hydrates during pipeline transport, and since the CMS membranes may be sensitive to high contents of water in the gas stream, it is also favorable for the membranes that the gas is dried. The main challenge for using CMS membranes in this application will most likely be the price for the membranes since large gas volumes are usually involved. However, with the very good separation performance these membranes have, a membrane process would potentially be very compact and have a small footprint.

15.6.3

Flue Gas

In a fossil fuel power plant, the chemical energy stored in coal, fuel oil, natural gas or oil shale is converted successively into thermal energy, mechanical energy and, finally, electrical energy for continuous use and distribution. The complete combustion of fossil fuel using air as the oxygen source is summarized in the following chemical reaction:    y y y y Cx Hy þ x þ O2 þ3:76 x þ N2 ! xCO2 þ H2 O þ 3:76 x þ N2 ð15:9Þ 4 4 2 4 The combustion of the hydrocarbon fossil fuels will generate water vapor, carbon dioxide and the non-reactive N2 when burned. Some by-products for combustion are sulfur dioxide (predominantly in coal) and oxides of nitrogen. If the combustion is not complete, the residual O2 will also be present in the flue gases. Different approaches such as physical absorption (Selexol) and chemical absorption (MEA, DMEA, ammonia) and membrane technology can potentially be used to capture CO2 from flue gas in post-combustion process. The MEA technology has been widely used in natural gas for over 60 years and produce relatively high purity CO2 stream. However, if used for CO2 capture in flue gas, it will be very costly and the challenges will be different from those related to natural gas application. The National Energy Technology Laboratory (NETL)86 estimates that this method will increase the cost of electricity production by 70%. Some literature reported that an alternative way to use the membrane technology for CO2 capture in power plant.10,87–89 He et al. investigated the application of the hollow fiber carbon membrane for CO2

View Online

184

Chapter 15

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

10

capture from flue gas. They reported that the capital cost using the carbon membranes was 197 $ per tonne CO2 avoided, which is still higher than the traditional chemical method like MEA (59 $ per tonne CO2 avoided reported by Rao and Rubin90), but the referred carbon membranes had a clear potential of further optimization. In any case, the environmentally friendly technique with further improved membrane performance could promote the hollow fiber carbon membranes as a promising candidate for CO2 capture in future.

15.6.4 15.6.4.1

Air Separation Nitrogen Production

The feed air is normally compressed to 8–10 bar with a low cost screw compressor and then passed through a bore-side hollow-fiber module. In a membrane nitrogen-from-air plant, approximately two-thirds of the total plant cost is associated with the air compressors; 20% or less is associated with the membrane modules, which indicates that reducing the size of the feed gas compressor will significantly decrease the cost for nitrogen production. Baker91 suggested that the compressor size can be reduced by 20% if the membrane performance for O2/N2 selectivity is improved from 8 to 12. This might cut nitrogen production costs by 10–15%. The Robeson upper bound (Figure 15.15) indicates that the O2/N2 selectivity for most of polymeric membranes is below 8 with relative high permeability (higher than 1 Barrer), which is the commercially interesting area. An alternative way for O2/N2 separation by carbon molecular sieve membranes have been investigated, and Figure 15.15 gives some representative results.18,92

15.6.4.2

Oxygen Production

Unlike the production of nitrogen, the production of oxygen is more difficult since a certain amount of nitrogen will always permeate together with the oxygen, resulting in oxygen-enriched air rather than pure oxygen. This can be easily understood become of the relatively low content of O2 in air (21%) which will in any process result in a relatively low driving force over the membrane. The pressure differential across the membrane can be evaluated either by pulling a vacuum on the permeation side or using a compressor on the feed side. For the second option, all of the feed air must be compressed, while only a small portion permeates through the membrane as oxygen-enriched product. The energy consumption for a vacuum pump on the permeate side is about onehalf that of a feed compressor, because the only gas that needs to pass through the pump is the oxygen-enriched product. However, it should be noted that vacuum operation requires a larger membrane area to produce the same flow of product gas. To make this operating mode economical, high-flux membranes and low-cost membrane modules are required.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

Figure 15.15

15.6.5

185

Comparing the O2/N2 Robeson upper bound for dense polymer membranes18 to the carbon membranes (C(PPO);7 C(PPESK);6 C(Kapton);92 C(Cellulose);20 C(Co-polyimide)35 and C(Cellulose acetate)),19 and the region for industrial applicability was suggested by Zhang et al.6 (Data for CMS membranes and industrial applicability region added to the original Robeson plot.)

Petrochemical Industry

At petrochemical plants there are numerous gas streams that contain valuable components which need to be recovered and reused. These are typically non-reacted monomers, by-products from reactors, inerts, solvents and carrier gas. There is a nice potential for using CMS membranes for many of these applications, and thereby also save money if complicated systems with columns, refrigeration and compressors can be avoided. A study on separation of alkanes–alkenes was performed by Hagg et al.75 Their systems were the separation of propane–propene and propan–ethene. As the alkanes–alkenes are chemically and physically quite similar compounds with almost identical critical properties, they must be separated on the basis of their molecular size. The Lennard-Jones diameter is 4.7 A˚ and 5.1 A˚ for propene and propane, respectively; hence, a carefully tailored CMS membrane would be able to separate these two components according to the molecular sieving mechanism. A selectivity of 23 for this gas pair was documented at 30 1C, and even much higher selectivity at 50 1C; this is believed to be a result of a transition of separation mechanisms for propane: at lower temperature propane will

View Online

186

Chapter 15

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

permeate faster according to the SSF mechanism, while at higher temperature, it will no longer sorb on the wall, and permeance goes down, selectivity goes up. As the hydrocarbons will more easily clog the membrane at low temperature, it is favorable to run this separation at higher temperature, and possibly regenerate the membrane on-line (see Section 15.3.4).

15.6.6

High-temperature Applications

Although the dense metal membranes (Pd, Ag, including their alloys) or solid electrolytes can be used for high temperature application such as hydrogen/ hydrocarbon/CO2 separation, they are still found to be too expensive for commercial applications, although they will show high selectivity but low permeability.93 In a search for a highly selective and relatively inexpensive membrane, the carbon membrane may be a candidate for use in membrane reactors for the hydrogen separation together with equilibrium-limited reaction. The potential interesting applications include hydrocarbon dehydrogenation and steam methane reforming (SMR) for H2 production. A schematic diagram of the carbon membrane reactor is shown in Figure 15.16. The reactant is fed into the system from the shell side of the carbon membranes. The driving force for the transport through the carbon membranes can be achieved by compression of the feed stream or using sweep gas or vacuum on permeate side. The membrane reactor can be heated to high working temperature for the application, while there is a cooler in each module end to protect the overheating of the carbon membrane sealing. The sealing of the module at high temperatures is typically not yet solved, so the temperature should be o150 1C at the ends.

15.6.6.1

Dehydrogenation

The application of carbon membrane reactors for the dehydrogenation of cyclohexane into benzene was investigated by Itoh and Haraya.94 They found a higher conversion for the carbon membrane reactor comparing to the normal reactor, which was caused by the chemical reaction shifti ng to the product side due to the preferential permeation of H2. Sznejer and Sheintuch studied the dehydrogenation of isobutane to isobutene in a membrane reactor equipped

Figure 15.16

Schematic diagram of hollow fiber carbon membrane reactor.

View Online

Carbon Molecular Sieve Membranes for Gas Separation

187

with the carbon membranes. The conversion achieved in the counter-current flow operation met hod was achieved a maximum of 85% at 500 1C, which is much higher than in the corresponding PFR.95

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

15.6.6.2

Steam Methane Reforming

The steam methane reforming (SMR) technology is the major route to industry’s production of merchant H2 on a worldwide scale. This is a very endothermic reaction (eqn (15.10)) that operates at 800 1C and at 20 bar pressure in order to achieve near equilibrium conversions and to meet the customers need for high pressure H2. CH4 þ H2 O ! CO þ 3H2

ð15:10Þ

By using a membrane reactor, it can shift the reaction to produce more H2 at lower operating temperatures. Some literature reported to use the Pd and ceramic membrane reactor for steam methane reforming to H2 production.96–98 The results showed that both the overall CH4 conversion and the conversion to CO, indicative of the extent of the water gas shift reaction, exceed the thermodynamic equilibrium values. It appear to be greater opportunity for application of membrane reactors with regard to SMR, but some challenges still remain due to the higher cost for membranes based on the process economic analysis.99 Although the application of the carbon membrane reactor in the steam methane reforming process has not been investigated, a carbon membrane reactor used for methanol steam reforming reaction to generate a product with high-purity H2 was reported by Zhang et al.100 Their results showed that the carbon membrane reactor provided a higher methanol conversion than the fixed bed reactor at all investigated operating conditions, while the overall yields of hydrogen in the carbon membrane reactor and the fixed bed reactor are identical. However, a CO-free hydrogen stream can be produced with the carbon membrane reactor, which could be directly used in a proton-exchange membrane fuel cell. Therefore, the carbon membranes reactor can also become a promising candidate for the application in high temperature methane steam reforming reaction.

15.7 Concluding Remarks There is a growing interest in the development of CMS membranes for gas separation based on the selected precursor that provide better selectivity, thermal and chemical stabilities than polymeric membranes. Although the cost of production and module units for CMS membranes are more expensive than polymeric membranes, the CMS membranes have nice potential to fill niches in the membrane market due to their high performance and ability to separate efficiently gas mixtures containing molecules of almost similar size.24 In order to compensate for their higher cost, a superior performance must be achieved. Due to the many factors that may affect the CMS membrane performance, an

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

188

Chapter 15

effective way to resolve this problem is based on combining the experimental research with computer simulation. As for the experiments, the multi-factorial design can be used to optimize the conditions for spinning the precursor membranes and fabricating CMS membranes (carbonization conditions) due to the interaction between these parameters. With respect to computer simulation, the following aspects should be in focus for further research:  Investigate the gas separation process for precursor membranes and CMS membranes using computational fluid dynamics  Give an insight into the mechanism of carbonization process and formation of carbon membranes that are unavailable by experimental techniques based on the molecular simulation  Process simulation for membrane gas separation by process system engineering method. Moreover, to decrease the material cost for CMS membrane will also be an important factor in future research. A dual-layer CMS membrane, which uses the polyimide and cheaper materials (i.e. PSf, CA) as the inner and outer layers respectively, could be an opportunity to decrease the membrane cost while ensuring high membrane performance.

Acknowledgement The authors gratefully acknowledge the financial support of the NanoGloWa project.

References 1. 2. 3. 4. 5. 6. 7. 8. 9. 10. 11.

J. E. Koresh and A. Soffer, Sep. Sci. Technol., 1983, 18, 723–734. H. Suda and K. Haraya, J. Phys. Chem. B, 1997, 101, 3988–3994. K. M. Steel and W. J. Koros, Carbon, 2003, 41, 253–266. J. N. Barsema, N. F. A. van der Vegt, G. H. Koops and M. Wessling, Adv. Funct. Mater., 2005, 15, 69–75. L. I. B. David and A. F. Ismail, J. Membr. Sci., 2003, 213, 285–291. B. Zhang, T. Wang, S. Zhang, J. Qiu and X. Jian, Carbon, 2006, 44, 2764–2769. M. Yoshimune, I. Fujiwara and K. Haraya, Carbon, 2007, 45, 553–560. H.-J. Lee, H. Suda, K. Haraya and S.-H. Moon, J. Membr. Sci., 2007, 296, 139–146. J. A. Lie and M.-B. Hagg, J. Membr. Sci., 2006, 284, 79–86. X. He, J. A. Lie, E. Sheridan and M.-B. Ha¨gg, Energy Procedia, 2009, 1, 261–268. J. N. Barsema, N. F. A. van der Vegt, G. H. Koops and M. Wessling, J. Membr. Sci., 2002, 205, 239–246.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

189

12. E. P. Favvas, G. C. Kapantaidakis, J. W. Nolan, A. C. Mitropoulos and N. K. Kanellopoulos, J. Mater. Process. Technol., 2007, 186, 102–110. 13. D. Q. Vu, W. J. Koros and S. J. Miller, Ind. Eng. Chem. Res., 2002, 41, 367–380. 14. S. M. Saufi and A. F. Ismail, Songklanakarin J. Sci. Technol., 2002, 24, 843–854. 15. P. S. Tin, T.-S. Chung, Y. Liu and R. Wang, Carbon, 2004, 42, 3123–3131. 16. J. A. Lie, PhD thesis, Norwegian University of Science and Technology, 2005. 17. J. N. Barsema, PhD thesis, University of Twente, 2007. 18. L. M. Robeson, J. Membr. Sci., 2008, 320, 390–400. 19. X. He, J. A. Lie, E. Sheridan and M.-B. Ha¨gg, ICOM 2008, Honolulu, Hawaii, 2008. 20. J. A. Lie and M.-B. Hagg, Carbon, 2005, 43, 2600–2607. 21. H. B. Park and Y. M. Lee, Adv. Mater., 2005, 17, 477–483. 22. A. M. W. Hillock, S. J. Miller and W. J. Koros, J. Membr. Sci., 2008, 314, 193–199. 23. A. B. Fuertes and T. A. Centeno, J. Membr. Sci., 1998, 144, 105–111. 24. A. F. Ismail and L. I. B. David, J. Membr. Sci., 2001, 193, 1–18. 25. M. B. Shiflett and H. C. Foley, J. Membr. Sci., 2000, 179, 275–282. 26. J.-I. Hayashi, H. Mizuta, M. Yamamoto, K. Kusakabe and S. Morooka, J. Membr. Sci., 1997, 124, 243–251. 27. H. Wang, L. Zhang and G. R. Gavalas, J. Membr. Sci., 2000, 177, 25–31. 28. A. B. Fuertes and T. A. Centeno, Micropor. Mesopor. Mater., 1998, 26, 23–26. 29. M. Acharya and H. C. Foley, J. Membr. Sci., 1999, 161, 1–5. 30. S. M. Saufi and A. F. Ismail, Carbon, 2004, 42, 241–259. 31. H. B. Park, Y. K. Kim, J. M. Lee, S. Y. Lee and Y. M. Lee, J. Membr. Sci., 2004, 229, 117–127. 32. Y. Xiao, T. S. Chung, M. L. Chng, S. Tamai and A. Yamaguchi, J. Phys. Chem. B, 2005, 109, 18741–18748. 33. Y. K. Kim, H. B. Park and Y. M. Lee, J. Membr. Sci., 2004, 243, 9–17. 34. H.-J. Lee, H. Suda and K. Haraya, Sep. Sci. Technol., 2007, 42, 59–71. 35. J. N. Barsema, J. Balster, V. Jordan, N. F. A. van der Vegt and M. Wessling, J. Membr. Sci., 2003, 219, 47–57. 36. Y. K. Kim, H. B. Park and Y. M. Lee, J. Membr. Sci., 2003, 226, 145–158. 37. M. S. Silverstein, Y. Najary, Y. Lumelsky, I. von Lampe, G. S. Grader and G. E. Shter, Polymer, 2004, 45, 937–947. 38. M. Yoshimune, I. Fujiwara, H. Suda and K. Haraya, Desalination, 2006, 193, 66–72. 39. S. Banerjee, S. Murad and I. K. Puri, Chem. Phys. Lett., 2007, 434, 292–296. 40. L. Y. Jiang, T. S. Chung, C. Cao, Z. Huang and S. Kulprathipanja, J. Membr. Sci., 2005, 252, 89–100. 41. S. Kim, T. W. Pechar and E. Marand, Desalination, 2006, 192, 330–339. 42. Q. Liu, T. Wang, C. Liang, B. Zhang, S. Liu, Y. Cao and J. Qiu, Chem. Mater., 2006, 18, 6283–6288.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

190

Chapter 15

43. H. B. Park, I. Y. Suh and Y. M. Lee, Chem. Mater., 2002, 14, 3034–3046. 44. H. Suda, S. Yoda, A. Hasegawa, T. Tsuji, K. Otake and K. Haraya, Desalination, 2006, 193, 211–214. 45. S. Yoda, A. Hasegawa, H. Suda, Y. Uchimaru, K. Haraya, T. Tsuji and K. Otake, Chem. Mater., 2004, 16, 2363–2368. 46. C. Zhou, T. Liu, T. Wang and S. Kumar, Polymer, 2006, 47, 5831–5837. 47. L. Shao, T.-S. Chung, G. Wensley, S. H. Goh and K. P. Pramoda, J. Membr. Sci., 2004, 244, 77–87. 48. X. He, M.-B. Ha¨gg, Using Orthogonal Experimental Design for Spinning Defect-Free Hollow Fiber Membranes, in Advanced Membrane Technology IV, ECI Conference, Trondheim, 2009. 49. H. Yoneyama and Y. Nishihara, E.P. Pat. 0394449, 1990. 50. E. Schindler and F. Maier, U.S. Pat. 4919860, 1990. 51. P. S. Tin, T.-S. Chung, S. Kawi and M. D. Guiver, Micropor. Mesopor. Mater., 2004, 73, 151–160. 52. D. W. Wallace, C. Staudt-Bickel and W. J. Koros, J. Membr. Sci., 2006, 278, 92–104. 53. V. M. Linkov, R. D. Sanderson and E. P. Jacobs, J. Membr. Sci., 1994, 95, 93–99. 54. C. W. Jones and W. J. Koros, Carbon, 1994, 32, 1419–1425. 55. J. Su and A. C. Lua, J. Membr. Sci., 2006, 278, 335–343. 56. K. Wang, L. S. Loo and K. Haraya, Ind. Eng. Chem. Res., 2007, 46, 1402–1407. 57. V. C. Geiszler and W. J. Koros, Ind. Eng. Chem. Res., 1996, 35, 2999–3003. 58. S. Lagorsse, F. D. Magalhaes and A. Mendes, J. Membr. Sci., 2004, 241, 275–287. 59. C. Liang, G. Sha and S. Guo, Carbon, 1999, 37, 1391–1397. 60. T. Yoshinaga, H. Shimazaki, Y. Kusuki and Y. Sumiyama, E.P. Pat. 0459623, 1991. 61. A. Soffer, A. Azariah, A. Amar, H. Cohem, D. Golub and S. Saguee, U.S. Pat. 5695818, 1997. 62. J. Petersen, M. Matsuda and K. Haraya, J. Membr. Sci., 1997, 131, 85–94. 63. W. J. Koros and C. W. Jones, U.S. Pat. 5288304, 1994. 64. F. Dreisbach, R. Seif and H. W. Lo¨sch, J. Therm. Anal. Calorim., 2003, 71, 73–82. 65. Y. Belmabkhout, R. Serna-Guerrero and A. Sayari, Chem. Eng. Sci., 2009, 64, 3721–3728. 66. C. Nguyen and D. D. Do, J. Phys. Chem. B, 1999, 103, 6900–6908. 67. C. Nguyen, D. D. Do, K. Haraya and K. Wang, J. Membr. Sci., 2003, 220, 177–182. 68. S. Lagorsse, M. C. Campo, F. D. Magalhaes and A. Mendes, Carbon, 2005, 43, 2769–2779. 69. X. He and M.-B. Ha¨gg, Euromembrane 2009, Montpellier, France 2009. 70. C. W. Jones and W. J. Koros, Ind. Eng. Chem. Res., 1995, 34, 158–163.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00162

Carbon Molecular Sieve Membranes for Gas Separation

191

71. S. Lagorsse, F. D. Magalha˜es and A. Mendes, J. Membr. Sci., 2008, 310, 494–502. 72. M. R. M. Abu-Zahra, J. P. M. Niederer, P. H. M. Feron and G. F. Versteeg, Int. J. Greenhouse Gas contr., 2007, 1, 135–142. 73. I. Menendez and A. B. Fuertes, Carbon, 2001, 39, 733–740. 74. C. W. Jones and W. J. Koros, Carbon, 1994, 32, 1427–1432. 75. M.-B. Hagg, J. A. Lie and A. Lindbrathen, Ann. N. Y. Acad. Sci., 2003, 984, 329–345. 76. M. B. Rao and S. Sircar, J. Membr. Sci., 1996, 110, 109–118. 77. J. Gilron and A. Soffer, J. Membr. Sci., 2002, 209, 339–352. 78. C. J. Geankoplis, Transport Processes and Unit Operations. 3rd edn., Prentice-Hall, Englewood Cliffs, 1993. 79. S. Glasstone, K. J. Laidler and H. Eyring, The Theory of Rate Processes. 1st edn., McGraw-Hill, New York, 1941. 80. Y. Kusuki, H. Shimazaki, N. Tanihara, S. Nakanishi and T. Yoshinaga, J. Membr. Sci., 1997, 134, 245–253. 81. S. Lagorsse, A. Leite, F. D. Magalhaes, N. Bischofberger, J. Rathenow and A. Mendes, Carbon, 2005, 43, 809–819. 82. P. G. Rathasiri, D. Botheju and R. Bakke, in Effects of Biofilm Formation on Oxygen Transfer Membranes in Membrane Micro-Aerated Anaerobic Digesters, March/April 2009. 83. http://www.lupower.at/index.php/cng-power, Accessed in September 2010. 84. M.-B. Ha¨gg, J. A. Lie, in Upgrading of Biogas by Membranes for Vehicle Fuel or Injection into the European Gas Network, 2006. 85. MemfoAct, http://www.memfoact.no/, Accessed September 2010. 86. L. C. Elwell and W. S. Grant, Accessed in September 2010, 2006. 87. R. Bredesen, K. Jordal and O. Bolland, Chem. Eng. Process., 2004, 43, 1129–1158. 88. M. B. Hagg and A. Lindbrathen, Ind. Eng. Chem. Res., 2005, 44, 7668–7675. 89. A. Hussain and M.-B. Ha¨gg, J. Membr. Sci., 2010, 359, 140–148. 90. A. B. Rao and E. S. Rubin, Environ. Sci. Technol., 2002, 36, 4467/75. 91. R. W. Baker, Ind. Eng. Chem. Res., 2002, 41, 1393–1411. 92. A. C. Lua and J. Su, Carbon, 2006, 44, 2964–2972. 93. G. A. Sznejer, I. Efremenko and M. Sheintuch, AIChE J., 2004, 50, 596–610. 94. N. Itoh and K. Haraya, Catal. Today, 2000, 56, 103–111. 95. G. Sznejer and M. Sheintuch, Chem. Eng. Sci., 2004, 59, 2013–2021. 96. E. Kikuchi, S. Uemiya and T. Matsuda, in Stud. Surf. Sci. Catal., ed. K. J. J. A. Holmen and S. Kolboe, Elsevier, 1991, vol. 61, pp. 509–515. 97. E. Kikuchi, Catal. Today, 1995, 25, 333–337. 98. Z. Chen, J. R. Grace, C. Jim Lim and A. Li, Int. J. Hydrogen Energy, 2007, 32, 2359–2366. 99. J. N. Armor, J. Membr. Sci., 1998, 147, 217–233. 100. X. Zhang, H. Hu, Y. Zhu and S. Zhu, Ind. Eng. Chem. Res., 2006, 45, 7997–8001.

CHAPTER 16

Perovskite Membranes for High Temperature Oxygen Separation F. LIANG AND J. CARO* Leibniz University Hannover, Institute of Physical Chemistry and Electrochemistry, Callinstr. 3A, D-30167 Hannover, Germany

16.1 Introduction If the oxygen partial pressure on the two sides of an oxygen transporting membrane (OTM) is different, an oxygen flux is started to compensate this difference. The rate of oxygen transport through an OTM is described by the Wagner equation,1,2 (eqn (16.3)) assuming as rate determining step the diffusion of oxygen ions/electrons, or in the Kroeger–Vink notation, the diffusion of oxygen lattice vacancies/electron holes. However, if the walls of the OTM become sufficiently thin, the rate of surface processes (exchange of oxygen with the bulk OTM) gain influence. The molecular oxygen flux through a perovskite OTM is related to the chemical potential gradient rm(O2): JO 2 ¼ 

1 4ðz0 FÞ

2



sO  ðsh þ se Þ  rmðO2 Þ sO þ ðsh þ se Þ

ð16:1Þ

where F is Faraday constant, sO, se and sh denote the partial conductivities of oxygen ions, electrons and electron holes, respectively. The chemical potential Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

192

Perovskite Membranes for High Temperature Oxygen Separation

193

gradient can be expressed by the derivative of the chemical activity of molecular oxygen: @ lnaðO2 Þ ð16:2Þ rmðO2 Þ ¼ RT  @x where R and T denote the gas constant and temperature, respectively. Assuming that both surfaces are in chemical equilibrium with the adjacent gas phases, the surface oxygen activities can be expressed by the oxygen partial pressures. For the oxygen flux through a membrane of thickness L with the permeate oxygen partial pressures Pfeed one obtains: O2 and PO2 permeate0

JO 2 ¼ 

RT  4 F2 L 2

lnPO2

Z

0 ln Pfeed O2

sO  sel dln PO2 sO þ sel

ð16:3Þ

with summarizing the partial conductivities sh and se of the electrons and electron holes as electronic partial conductivity sel. Assuming that the oxygen ion migration governs the bulk transport (selcso) and the ionic conductivity sO depends in an exponential way – as is often observed experimentally
n – on the oxygen partial pressure following the power law so ¼ s0o  PO2 , the integration of eqn (16.3) gives: JO 2

  RT s0O permeate n feed n ¼ 2 2 Þ  ðPO2 Þ ðPO2 4F L n

ð16:4Þ

For details see Schroeder.3 For oxygen transport through perovskite membranes two consecutive steps have to be considered: (i) the surface oxygen exchange, i.e. the incorporation and release of molecular oxygen, and (ii) the solid-state diffusion of oxygen vacancies/oxygen ions. A characteristic membrane thickness Lc can be introduced for which the rate of the surface exchange is comparable to the rate of the diffusion-controlled oxygen ion transport through the membrane.4,5 Depending on the perovskite material under study and on the temperature (note that surface exchange and bulk diffusion have different activation energies) Lc can be of the order of 0.01–10 mm. That is to say, for membranes thinner than Lc there is no linear relationship between the oxygen flux JO2 and the reciprocal membrane thickness L–1 predicted by the Wagner equation in its usual form for bulk diffusion-controlled oxygen transport through an OTM of thickness L.5 Consequently, extremely thin supported perovskite layers often do not show the expected high oxygen flux since the surface exchange reaction becomes rate limiting rather than the oxygen bulk diffusion. However, for mixed oxygen ion–electron conducting materials the surface processes can become rate limiting for oxygen transport through the membrane rather than bulk diffusion. Under surface reaction we understand the

194

Chapter 16

incorporation of oxygen into the lattice (adsorption, dissociation, ionization, incorporation on a lattice vacancy) and vice versa on the permeate side. For our perovskite hollow fiber membrane with 170 mm wall thickness we found that the main transport mechanism is bulk diffusion with a contribution of the surface reaction.6 From the value of n of eqn (16.4) the rate-limiting step of the oxygen permeation can be identified.7,8 For negative values of n, the oxygen permeation is dominated by the bulk diffusion, while for n Z 0.5 the exchange processes at the membrane surfaces can be assumed as rate limiting. In the case of the surface exchange as rate-limiting step, a catalytic coating of the membrane can accelerate the oxygen flux. In the surface exchange current model,9 0:5 JO2 gives a linear relationship with ðPpermeate =PO2 Þ0:5  ðPfeed if oxygen O2 O2 =PO2 Þ transport is controlled by the surface exchange reaction. In the case of bulk =Pfeed diffusion control, JO2 is a linear function of lnðPpermeate O2 O2 Þ. However, a reliable way to determine the relative contributions of the surface processes and the oxygen bulk diffusion to the oxygen transport is the preparation of discs of different thickness. However, this way is time consuming it is and experimentally extremely difficult to prepare membranes thinner than 0.5 mm but the interesting critical thickness, where the change between bulk diffusion and surface exchange takes place, is often found to be around a few hundreds of micrometers.

16.2 Materials Aspects of Oxygen Transporting Membranes Despite much R&D effort, 20 years after the pioneering papers from Teraoka and co-workers,10,11 who first reported oxygen permeation through a La1–xAxCo1–yFeyO3–d perovskite membrane, there is still no industrial application of perovskite type membranes. This lack of applications is mainly due to the long-term stability problems of perovskites especially at low oxygen partial pressure at the usual operation temperature near 850 1C. Often high oxygen fluxes reduce the stability and vice versa. As a rough summary of the current R&D status of OTM one can state that either the stability or the oxygen flux are insufficient: Examples for OTM materials with high oxygen flux but low stability are La1–x(Ca,Sr,Ba)xCo1–yFeyO3–d10,11 and Ba(Sr)Co1–xFexO3–d;12 typical materials with high stability but low oxygen flux are Sr(Ba)Ti(Zr)1–x–yCoyFexO3–d13 and La1–xSrxGa1–yFeyO3–d.14 The aim of the material synthesis is to substitute, e.g. the lattice positions A and/or B of the basic ABO3 perovskite structure, by cations of similar size but lower charge to create an oxygen transporting material with both ionic and electronic conductivity and sufficient chemical and mechanical stability. Whereas most OTM materials presently studied are of perovskite type structure (ABO3), also fluorite (AO2), Brownmillerite (A2B2O5) and pyrochlore (A2B2O7) structures are evaluated since the latter two show oxygen transport already in the undoped formulation.15

195

Perovskite Membranes for High Temperature Oxygen Separation

To improve the stability of the OTM materials, the reducible B-site ions (Co41/Co31, Fe41/Fe31) can be partially substituted by cations with constant valence (Ta51, Nb51, Zr41, Al31). Examples are BaCo0.7Fe0.2Ta0.1O3–d,16,17 Ba1.0Co0.7Fe0.2Nb0.1O3–d,18 BaCo0.4Fe0.6–xZrxO3–d,19 BaCo0.4Fe0.4ZnxZr0.2–x O3–d,20 La0.3Sr0.7(Fe,Ga)O3–d,21 Sr(Fe, Al) O3–d.22 Cobalt-based perovskite oxides like (LaxSr1–x)(CoyFe1–y)O3–d or (BaxSr1–x)(CoyFe1–y)O3–d are often thought to be the most promising materials because they show very good oxygen permeability.23 However, Co-based membranes usually have a stability problem under reducing conditions. Therefore, there are activities in the development of cobalt-free membranes with high oxygen permeabilities. So far, some cobalt-free membranes, such as La1–xSrxFeO3–d,24 Ba0.3Sr0.7FeO3–d,25 BaFe1–yCeyO3–d,26 Ba0.5Sr0.5Fe0.8 Zn0.2O3–d,27 Ba0.5Sr0.5Fe1–xAlxO3–d,28 (Ba0.5Sr0.5)(Fe0.8Cu0.2)O3–d,29 BaFe1–y ZryO3–d30 and Ba0.95La0.05FeO3–d31 have been developed (Table 16.1). An alternative to the OTM can be dual phase membranes. Here a nanoscale ion conductor is in intimate contact with an electron conductor. The first dual phase membranes were made of oxide conductors and noble metal: (Bi2O3)0.74SrO0.26-Ag,40 Bi1.5Y0.3Sm0.2O3-Ag,41,42 Bi1.5Er0.503-Ag,43 Bi1.6Y0.4 O3-Ag,44 YSZ-Pd.45 However, the application of these materials is limited due to high materials costs, a mismatch of the coefficient of thermal expansion (CTE) between ceramic and metallic phase, and poor oxygen permeabilities. In another concept, perovskite- or fluorite-type oxides are used instead of noble metals as electron conductors. Kharton and co-workers investigated the oxygen permeability of Ce0.8Gd0.2O1.9-La0.7Sr0.3MnO3–d (CGO-LSM), Ce0.8Gd0.2O1.9La0.8Sr0.2Fe0.8Co0.2O3–d (CGO-LSCF), Ce0.8Gd0.2O1.9-La0.8Sr0.2Fe0.8Cr0.2O3–d (CGO-LSFC), La0.8Sr0.2Fe0.8Co0.2O3-(La0.9Sr0.1)0.98Ga0.8Mg0.2O3–d (LSCFLSGM).40,46,47 Zhu and Yang developed Ce0.8Gd0.2O1.9-Gd0.2Sr0.8FeO3–d (CGO-GSF),48 Ce0.85Sm0.15O1.95-Sm0.6Sr0.4FeO3–d (CSO-SSF),49 Ce0.85Sm0.15O1.95Sm0.6Sr0.4Fe0.7Al0.3O3–d (CSO-SSFA),50 La0.15Sr0.85Ga0.3Fe0.7O3–d-Ba0.5Sr0.5Fe0.2

Table 16.1

Examples for newly developed perovskite materials, their oxygen flux and stability in the partial oxidation of methane to synthesis gas (POM)39

Material

Temperature (1C)

Oxygen permeation flux (mL min1cm2)

Stability in the POM (h) Reference

Initial La0.2Sr0.8Co0.4Fe0.6O3–d Improved La0.2Sr0.8Co0.4Fe0.6O3–d La0.8Sr0.2Co0.1Fe0.6Cr0.1O3–d La0.3Sr1.7Ga0.6Fe1.4O6–d SrCo0.5Fe0.5O3–d SrCo0.8Fe0.2O3–d Ba0.5Sr0.5Co0.8Fe0.2O3–d BaZr0.2Co0.4Fe0.4O3–d AxA#2–xByB#2–yO5–d

850 850 900 900 850 87 875 850 900

— 3–4 14.5 2–4 2–4 — 11.5 5.6 1012

o1 500 340 8000 1000 o1 500 2200 48700

32 33 34 35 34 36 37 13 38

196

Chapter 16 51

Co0.8O3–d. Chen and co-workers reported Ce0.8Sm0.2O2–d-La0.8Sr0.2 CrO3–d (CSO-LSC),52 Ce0.8Sm0.2O2–d-La0.8Sr0.2MnO3–d (CSO-LSM),53 Y0.2Zr0.8 O0.9-La0.8Sr0.2CrO3–d (YSZ-LSC),54 Y0.16Zr0.84O0.9-La0.8Sr0.2MnO3–d (YSZ-LSM).55 However, these dual phase membranes have a perovskite phase (ABO3, A ¼ alkaline earth metals or lanthanide element; B ¼ transition metal) as electronic and oxygen ionic mixed conductor, in which the A site is usually occupied by alkaline earth metals, which easily generate impermeable carbonates if CO2 is present. The oxygen permeation of the different geometry MIEC membrane is usually tested in devices as shown in Figure 16.1. Tubular membranes were developed to reduce the engineering difficulties, especially the problem of the high temperature sealing. However, their small membrane area per unit volume and relatively thick wall limit them in practical application. A hollow fiber membrane with a thin wall can solve the above problem. By keeping the two sealed ends outside the high temperature zone, polymer O-rings can be used. Compared to the conventional disk and tubular membranes, hollow fiber membranes give very high values of the membrane area per unit volume. Furthermore, due to the thin wall, the materials costs of the hollow fiber membrane are reduced. During the last 10 years, remarkable progress has been achieved in the development of spinning techniques for hollow fiber production. As an example, Figure 16.2 shows the perovskite Ba(Cox,Fey,Zr1–x–y)O3–d (BCFZ) hollow fiber membrane which was recently developed by a spinning process.56 The BCFZ hollow fibers were used in a catalytic membrane reactor (Figure 16.3), which was on-line coupled to a gas chromatograph. The two ends of the hollow fiber can be coated with a gold paste (C5754, Heraeus) and calcined at 950 1C for 5 h, this procedure was repeated three times. Such Au-coated BCFZ hollow fiber can be sealed by silicon rubber rings outside the oven whereas a 3 cm long non-coated part of the hollow fiber (between 0.35 and 0.86 cm2 area) can be kept in the middle of the oven thus ensuring isothermal conditions.

16.3 Oxygen Separation by Oxygen Transporting Membranes Oxygen can be separated from air as the cheapest source of oxygen, if the oxygen partial pressure on the feed side of the OTM is higher than that on the permeate side. This difference in the oxygen partial pressures can be achieved by (i) using sweep gases like steam, which can be easily separated from the permeated oxygen by condensation, (ii) using vacuum pumps to drain off the permeated oxygen, (iii) pressurized air on the feed side of the OTM having pure oxygen on the permeate side at atmospheric pressure, or (iv) combining of the pumping on the permeated side and having pressurized air on the feed side.

16.3.1

Using Sweep Gases

If the oxygen partial pressure on the two sides of an OTM is different, according to eqns (16.3) and (16.4), an oxygen flux starts to compensate this

Perovskite Membranes for High Temperature Oxygen Separation

197

Figure 16.1 Different types of MIEC laboratory scale membrane reactors: (a) short hollow fiber with Au sealing in the hot zone,56 (b) short tubular membrane,57 and (c) disk membrane.58

198

Chapter 16

Figure 16.2

SEM micrograph of the dense BCFZ hollow fiber after sintering at T>1200 1C.

Figure 16.3

One-hollow-fiber membrane reactor/permeator for oxygen separation and catalysis.

difference. The most simple concept to establish oxygen partial pressure differences is the use of a sweep gas in order to reduce the oxygen partial pressure on the permeate side and having air as oxygen source on the feed side. Various sweep gases like He, Ar or N2 are used. However, the separation of these sweep gases from the permeated oxygen turns out as a problem. It has been proposed, therefore, to use steam as a sweep gas which can be easily separated from the permeated oxygen by condensation. A corresponding experiment of the oxygen separation from air using He and steam as sweep gases, respectively, is shown in Figure 16.4 using a hollow fiber BCFZ membrane reactor as shown in Figure 16.3. Increasing sweep gas flows increase the oxygen partial pressure gradient across the membrane as driving force of oxygen transport (see eqns (16.3) and (16.4)). On the other hand, there is no remarkable influence of the kind of sweep gas on the oxygen transport (Figure 16.4). Surprisingly, the BCFZ membrane showed a very good long-time stability at 850 1C in the presence of steam (Figure 16.5).59

Perovskite Membranes for High Temperature Oxygen Separation

199

Figure 16.4

Oxygen separation from air using inert sweep gases: Oxygen flux through a BCFZ hollow fiber membrane as a function of temperature for steam (J) and He (B) as sweep gases (after Tablet et al.6). Experimental details: Air flow rate on the core side ¼ 150 mL min–1, steam or He flow rate on the shell side ¼ 10 mL min–1, 0.35 cm2 effective membrane area. For comparison, different sweep gas flows of 5 (&) and 20 (n) mL min–1 He.

Figure 16.5

Oxygen separation from air with steam as sweep gas: Oxygen flux through a BCFZ hollow fiber membrane as a function of time.59 Experimental details: Air flow rate on the core side ¼ 150 mL min–1, steam flow rate on the shell side ¼ 10 mL min–1, 0.35 cm2 effective membrane area, 875 1C.

Whereas most perovskite material have operation temperatures above 800 1C, our BCFZ perovskite hollow fiber membrane could be successfully used for oxygen separation from air between 400 and 500 1C (Figure 16.6).60 As Figure 16.6 shows, the oxygen permeation flux increases from 0.03 to 0.45 mL cm–2 min–1 as the temperature rises from 400 to 500 1C. However, in

200

Chapter 16

Figure 16.6

Low temperature oxygen separation from air: Oxygen flux from air through a BCFZ hollow fiber membrane with He as sweep gas for different low temperatures.60 Experimental details: Air flow rate on the shell side ¼ 150 mL min–1, He flow rate on the core side ¼ 30 mL min–1, 0.43 cm2 effective membrane area.

Figure 16.7

Low temperature oxygen separation from air: Oxygen permeation flux through a BCFZ hollow fiber membrane as a function of time at 500 1C. At this temperature a continuous decrease of the oxygen flux is observed. For regeneration, the fiber was heated to 925 1C, kept at this temperature for 1 h in air, then cooled down to 500 1C for the next run.60 Experimental details: Air flow rate on the shell side ¼ 150 mL min–1, He flow rate on the core side ¼ 30 mL min–1, 0.43 cm2 effective membrane area.

this low temperature region, the oxygen flux is not stable and decreases with time (Figure 16.7). The origin of this reversible oxygen decrease is still unclear. Because of the use of synthetic air, the reversible formation of a carbonate layer, which would reduce the oxygen flux61 can be excluded in the present case. The following explanation is more realistic. At temperatures below 900 1C, cobalt ions show a temperature dependent change of their spin and/or oxidation state.62 The oxidation of cobalt ions combined with the spin-state

Perovskite Membranes for High Temperature Oxygen Separation

201

transition with decreasing temperatures leads to a significant diminution of their ionic radius and result in the formation of hexagonal perovskite polytypes.63 The hexagonal phases are disadvantageous for oxygen transport due to their lower crystal symmetry and less amount of mobile disordered oxygen vacancies as compared to cubic perovskites.

16.3.2

With Evacuation on the Permeate Side

Instead of a sweep gas, in a second approach for having a driving force for oxygen permeation according to eqns (16.3) and (16.4), the oxygen partial pressure on the permeate side can be decreased by applying vacuum. In this case, dead end perovskite hollow fibers or capillaries/tubes can be used. A typical ‘one dead-end membrane permeator’ with the BCFZ hollow fiber used for the production of pure oxygen is shown in Figure 16.8.64 The dead-end hollow fiber membrane is obtained by sealing one end of the BCFZ hollow fiber with a gold plug using Au paste. The dead-end BCFZ hollow fiber was coated by dense gold film except the 3 cm part close to the end with gold plug. The uncoated end of the hollow fiber was put in the middle of the oven thus ensuring isothermal conditions. In our previous study, we have already reported some results on the production of oxygen-enriched air using these BCFZ hollow fiber membranes.65 Oxygen-enriched air can be used in many industrial processes, such as methane combustion at high temperatures, the synthesis of ammonia, the Claus process, steel plants and waste burning.66 In order to obtain a high oxygen permeation flux and, thus, a high space–time yield, a large oxygen partial pressure gradient across the membrane is required, which can be achieved by using oxygenenriched air as feed. Figure 16.9 shows the influence of the temperature on the production of pure oxygen from oxygen-enriched air of different oxygen content in a dead-end permeator with a BCFZ hollow fiber membrane. The pressure on the feed/shell side was 1 bar, and the oxygen pressure on the permeate/core side was kept at 0.05 bar by using a vacuum pump. The oxygen production rate increases with increasing temperature and with increasing

Figure 16.8

Permeator in dead-end geometry for the production of pure O2.64

202

Chapter 16

Figure 16.9

Influence of the temperature on the production of pure oxygen from O2enriched air with different oxygen concentration. Experimental details: Air flow rate on feed/shell side ¼ 150 mL(STP) min–1 at 1 bar; pure O2 at about 0.05 bar on the permeate/core side.64

oxygen content of the oxygen-enriched air which can be used instead of normal air with about 21% oxygen content. For an oxygen concentration 21% as it is found in normal air, the oxygen permeation rate increases from 1.8 to 3.7 mL(STP) cm–2 min–1 when increasing the temperature from 750 to 950 1C. However, at 950 1C the oxygen permeation rate can be almost doubled from 3.7 to 6.8 mL(STP) cm–2 min–1 when increasing the oxygen concentration in the fed air from 20 to 50 vol.%.

16.3.3

Applying Elevated Pressure on the Permeate Side

Oxygen is transported across the OTM if there are different oxygen partial pressures on both membrane sides. If the permate side of the membrane is at atmospheric pressure, oxygen will be transported across the OTM, if the product of the gas pressure on the feed side and its oxygen content is larger than 1 bar. These oxygen partial pressures >1 bar can be established by compressing normal air above 6 bar, the oxygen partial pressure of the feed side is then about 1.2 bar. For oxygen-enriched air lower compression pressures are necessary, to get oxygen partial pressure higher than atmospheric pressure. Figure 16.10 shows the influence of the oxygen partial pressure difference on the production of pure O2 in a dead-end permeator with one BCFZ hollow fiber membrane at different temperatures. When the pressure on the permeate side is fixed at 1 bar, the oxygen partial pressure gradient across the membrane can be established by elevating the pressure of oxygen-enriched air used as feed. The oxygen permeation rates increase with increasing oxygen partial pressure difference and with increasing temperature. At an oxygen partial pressure difference of 1 bar, an oxygen flux of 3.0 mL(STP) cm–2 min–1 is obtained at 950 1C when using air with 50 vol.% O2 as feed.

Perovskite Membranes for High Temperature Oxygen Separation

203

Figure 16.10

Influence of the oxygen partial pressure difference on the production of pure oxygen at different temperatures. Experimental details: O2enriched air flow rate on the feed/shell side ¼ 150 mL(STP) min–1 at different pressure; pure O2 at 1 bar on the permeate/core side.64

Figure 16.11

Influence of the oxygen partial pressure difference on the production of pure oxygen at different temperatures. Experimental details: Air flow rate on feed/shell side ¼ 150 mL(STP) min–1 with 20 vol.% O2 at different pressure; Pure O2 at about 0.05 bar on the permeate/core side.64

16.3.4

Combining Evacuation of the Permeate Side and Elevated Pressure on the Feed Air Side

According to the Wagner theory, the oxygen permeation flux can be enhanced by increasing the oxygen partial pressure gradient across membrane. To obtain a higher oxygen permeation flux, the pressure on the permeate side was reduced to 0.05 bar by using a vacuum pump, and the pressure on the feed side was simultaneously elevated. Figure 16.11 shows the influence of the air pressure on the feed side on the production of pure oxygen in a dead-end permeator with BCFZ hollow fiber membrane at different temperatures. It can be seen that the

204

Chapter 16

oxygen permeation flux increased with raising the pressure on the feed side while keeping the pressure at 0.05 bar on the permeate side. For the same oxygen partial pressure difference, the oxygen permeation flux increases with rising temperature. At 950 1C and 5 bar of total pressure difference, the oxygen permeation rate reaches 8.5 mL(STP) cm–2 min–1. Ito et al.67 reported an oxygen flux of 9 mL(STP) cm–2 min–1 of an asymmetrical tubular perovskite membrane with a B50 mm dense layer at 900 1C and 10 bar air pressure. Zhu et al.68 reported an oxygen flux of 9.5 mL(STP) cm–2 min–1 in a dead-end tube perovskite membrane permeator at 925 1C and vacuum of B100 Pa on the permeate side and 7 bar air on the feed side.

16.4 Oxygen Separation from Air with its Immediate Consumption in a Partial Oxidation The oxygen separated from air can be continuously consumed in situ in a hydrocarbon partial oxidation. On the one hand side, this concept is a clever combination of oxygen separation and its use for the hydrocarbon catalytic partial oxidation (HCPO) in one reactor following the concept of process intensification, provided that both oxygen separation through the OTM and the HCPO require the same temperature window near 800 1C. On the other hand, the hydrocarbon side of the OTM is extended to a strongly reducing atmosphere (CH4, H2, CO), which can cause a reduction of the oxidic perovskite (cf. Ellingham diagram for the minimum O2 partial pressure necessary for the oxide stability). Protective layers are expected to prevent this reductive damage of the perovskite by maintaining a non-zero oxygen partial pressure direct at the perovskite surface.

16.4.1

Partial Oxidation of Methane to Syngas

As an ambitious goal, the separation of oxygen from air and the catalytic partial oxidation of methane (CPOM) to synthesis gas has been realized in one apparatus. On the one hand side, by tuning the oxygen supply via the membrane and the processing of this oxygen by tuning the methane flux, kinetic compatibility between oxygen supply and its consumption can be reached, and the effluent product flux consists of about one-third CO and two-thirds H2.69 Figure 16.12 shows that for the reaction parameters applied, in a temperature window 850–900 1C high-quality synthesis gas can be produced.70 On the other hand, the non-reacted feed methane and the products hydrogen and carbon monoxide as reducing gas can damage the perovskite OTM by reducing the metal oxides of the perovskite to its pure metals (cf. Ellingham diagram). This damage process can be suppressed, by having an inert porous coating direct on the OTM. In this case, the released of oxygen from the perovskite has a longer residence time in the porous layer and avoids thus the direct contact of the reducing gases like CH4, H2, or CO with the perovskite surface.71

Perovskite Membranes for High Temperature Oxygen Separation

Figure 16.12

205

In situ separation of oxygen from ambient air for the CPOM: Performance of a BCFZ hollow fiber membrane reactor in the CPOM as a function of temperature with CH4 conversion (J), CO selectivity (&), CO2 selectivity (B) and H2/CO ratio (n).70 Experimental details: Flow rate on the core side of the BCFZ hollow fiber membrane ¼ 150 mL min– air, flow rate on the shell side ¼ 20 mL min– (10 mL min– CH4 þ 10 mL min– He). 0.88 cm2 effective membrane area, 0.8 g Ni/Al2O3 steam reforming catalyst as packed bed on the shell side.

The problem of the hot sealing of tubular OTMs is avoided in a reactor design by Air Liquid.72 In this design (Figure 16.13), the OTM is used in a dead end configuration. The OTM was La0.8Sr0.2Fe0.7Ga0.3O3–d with a porous catalyst layer of La0.8Sr0.2Fe0.7Ni0.3O3–d for CPOM brought about by dip coating. The membrane reactor was operated at a pressure of 3 bar with a steam:methane ratio 1:1 for at least 142 h. During this operation, severe microstructural and chemical evolutions on both catalyst and OTM were observed. However, these degradations seem to have little impact on the reactor performance.

16.4.2

Oxidative Coupling of Methane

The oxidative coupling of methane (OCM) to ethane and ethene is an attractive alternative to indirect routes involving the sequential steps CH4 reforming and Fischer–Tropsch synthesis.73 Intense R&D activities on the highly exothermic catalytic OCM reaction started in the early 1980s with the pioneering works of Keller and Bhasin.74 In the past two decades, the C2 yield of OCM on the packed-bed reactor with novel catalysts was limited to about 25%. High concentrations of oxygen are disadvantageous to high C2 selectivities. Low oxygen concentrations are, however, unfavorable for high degrees of methane conversion and high C2 yields. An alternative to staged oxygen delivery is the use of OTM membrane reactors. A perovskite membrane can distribute the oxygen

206

Figure 16.13

Chapter 16

Schema of the CPOM to synthesis gas in an OTM reactor with deadend-tubes for elevated pressure (up to 5 bar) with air on the core side of the OTM, and methane as feed on the shell side.72

homogeneously over the whole. Note that the OTM membrane material itself can show an intrinisic catalytic activity for OCM.75 There are two types of the OTM reactor for OCM that with or without other catalyst. Olivier et al.76 and Wang et al.57 investigated the OCM in a BSCF disc and a tubular perovskite membrane reactors with a LaSr/CaO catalyst. In the BSCF disc membrane at 950 1C, a C2 yield of more than 18% with a C2 selectivity of more than 65% was achieved. In the tubular membrane reactor at 825 1C, a C2 yield of 15% with a C2 selectivity of 70% was achieved. In Bi1.5Y0.3Sm0.2O3 (BYS) tubular membranes reactor, a C2 yield of 35% with a C2 selectivity of 54% achieved for OCM at 900 1C.77 Recently, Bhatia et al.78 investigated the OCM in three reactors configurations: catalytic Ba0.5Ce0.4Gd0.1Co0.8Fe0.2O3–d (BCGCF) tubular membrane reactor, catalyst packed bed reactor, and catalyst packed bed BCGCF tubular membrane reactor. The catalytic membrane reactor performed best among three reactors with a C2 yield of 34.7%, a C2 selectivity of 67.4%, and a methane a conversion of 51.6% at 850 1C. This C2 yield are higher than the 28% under conventional, packed-bed, continuous-feed operation and 30% threshold for commercial feasibility. So far, there are only a few reports of perovskite hollow fiber membrane reactors in the oxidative coupling of methane. Tan and Li79 studied the catalytic perovskite La0.6Sr0.4Co0.2Fe0.8O3 (LSCF) hollow fiber membrane as

Perovskite Membranes for High Temperature Oxygen Separation

Figure 16.14

207

Schematic drawing of the reactor set-up and an incorporated BCFZ hollow fiber for OCM.80 At both ends the 30 long fiber was coated with gold to obtain a 3 cm long gold-free isothermal oxygen permeation zone. The active surface area for the oxygen permeating BCFZ hollow fiber is 0.78 cm2. The OCM catalyst was dipersed between the outer dense alumina tube and the fiber, which was inserted into a porous alumina tube.

separator and reactor’s catalyst without or with additional OCM SrTi0.9Li0.1O3 catalyst for the OCM. A higher C2 selectivity up to 71.9% was obtained in the blank LSCF hollow fiber membrane reactor without catalyst and a maximum C2 yield close to 21% can be achieved in the packed hollow fiber membrane reactor with catalyst.79 BCFZ hollow fiber membranes were used for investigation of the OCM reaction with a 2 wt% Mn/5 wt% Na2WO4 on SiO2 catalyst.80 Figure 16.14 shows schematic drawing of the reactor set-up and an incorporated BCFZ hollow fiber for OCM. In this BCFZ hollow fiber membrane reactor at 800 1C, a high C2 selectivity of approximately 75% was observed at methane conversion of 6%. C2 yield of 17% was obtained at 50% C2 selectivity. Figure 16.15 illustrates the effect of diluting the air on core side of the hollow fiber membrane while feeding diluted methane.80 Compared to discs and tubular membranes, such hollow-fiber membranes possess much larger membrane area per unit volume values for oxygen permeation.

16.4.3

Oxi-dehydrogenation of Alkanes to the Corresponding Olefins

A new concept in olefin production is the multi-step thermal–catalytic dehydrogenation of short chain alkanes with selective hydrogen combustion using oxygen which is in situ separated from air.81 To overcome the thermodynamic restriction of alkan conversion, repeatedly the alkane reaction system is brought to the thermodynamic equilibrium by catalytic dehydrogenation due to C3 H8 ! C3 H6 þ H2 by using the dehydrogenation catalyst Pt/Sn on Al2O3.

208

Figure 16.15

Chapter 16

Improved C2 selectivities and C2 yields by feeding diluted oxygen on core side while feeding 10% methane on shell side of the hollow fiber membrane at a total flow rate of 25 mL min–1. Methane conversion, product selectivities and C2 yield as a function of the oxygen concentration on the core side of the membrane at 1 bar and T ¼ 800 1C. Experimental details: Flow rate on the shell side ¼ 5 mL min–1 methane þ 20 mL min–1 helium; on the core side Ftotal ¼ 50 mL min–1 air diluted with helium; 0.78 cm2 effective membrane surface; 0.25 g of catalyst; WHSV ¼ 0.86 h–1).80

Periodically, oxygen is separated from air by the hollow fiber BCFZ membrane and sent to the shell side of the hollow fiber where it selectively combusts the hydrogen to water as schematically shown in Figure 16.16. It follows from Figure 16.17 that for an oxygen concentration of 20% which corresponds to air, propylene yield of 34% can be obtained.81 In the regime of kinetic compatibility between H2 formed in the catalytic dehydrogenation and O2 transported through the perovskite membrane, TAP shows that hydrogen reacts exclusively with lattice oxygen of the perovskite after Mars and van Krevelen.82 If hydrogen, alkan and olefin compete for lattice oxygen of the BCFZ perovskite membrane, hydrogen is the only source of water formation because of the high reaction rate of hydrogen with lattice oxygen of the perovskite surface since the alkane and olefin are less reactive than hydrogen. The principle of catalytic dehydrogenation coupled with the selective hydrogen combustion has been also demonstrated for ethylene production.83

Perovskite Membranes for High Temperature Oxygen Separation

209

Figure 16.16

Schema of the multi-step oxidative dehydrogenation of propane. Parts of a tubular OTM are blocked for O2 transport, e.g. by coating with gold. Here the catalytic dehydrogenation takes place and leads to equilibrium. Then oxygen enters through a non-passivated part and this O2 selectively combusts the H2 in the presence of the hydrocarbon. Now, again, a catalytic dehydrogenation starts and so on (see Czuprat et al.81,83).

Figure 16.17

Impact of oxygen partial pressure on the core side on propane conversion, product selectivities and propene yield. Experimental details: 5 mL min–1 propane þ 10 mL min–1 steam þ 35 mL min–1 helium on the shell side; Ftotal ¼ 50 mL min–1 oxygen diluted with helium on the core side; 0.44 cm2 effective membrane surface; 0.25 g of Pt/Sn catalyst; WHSV ¼ 2.4 h; T ¼ 675 1C.81

16.5 Oxygen Separation from Oxygen-containing Gases and its in situ Consumption in a Partial Oxidation Instead of air, other oxygen-containing molecules like H2O, N2O or NO can also be used as oxygen source for an oxygen consuming reaction on the other

210

Chapter 16

side of the membrane. On one side of the OTM, these gases undergo a selfdissociation or decomposition and produce oxygen due to H2O # 12O2 þ H2 or N2O-12O2 þ N2. This oxygen is transported through an OTM to the other side of the membrane where it is consumed by the CPOM or an oxi-dehydrogenation. Additionally to the production of synthesis gas, other aims like the hydrogen production or the abatement of NxOy are followed in this concept.

16.5.1 16.5.1.1

Water as an Oxygen Source for Hydrogen Production Coupled with Synthesis Gas or Ethylene Production Production of Hydrogen and Synthesis Gas by Combining Thermal Water Spaltimg with Partial Oxidation of Methane

In this case study, water undergoes a thermal self-dissociation into hydrogen and oxygen. The latter is continuously extracted which results in a hydrogen production on the feed side of the OTM. At high temperatures, water undergoes a thermal self dissociation due to H2O # H2 þ 12O2. According to the low equilibrium constant of KP E 2 10–8 at 950 1C,84 only low equilibrium concentration of PO2 E 4.6  10–6 bar and PH2 E 9.2  10–6 bar are present. However, if we have on one side of an OTM the thermal water splitting, there will be a continuous oxygen flux through the membrane, if oxygen is removed on the other side of the OTM either by a sweep gas or consumed by a CPOM to synthesis gas.85–88 Since the thermal water splitting is kinetically fast, the rate of synthesis gas formation is determined by the rate of the oxygen flux through the OTM. Equation (16.3) shows that the latter can be high, if the permeated oxygen is consumed quickly by a combustion process, which is in our case the CPOM. In summary, the reaction seems to looks at first sight like steam reforming (SR) according to CH4 þ H2O-3H2 þ CO. However, different to SR, in our case one mol hydrogen is gained from the retentate after water condensation and the synthesis gas produced on the permeate side has the right ratio H2:CO ¼ 2:1 as it is necessary for methanol or Fischer–Tropsch syntheses. That is to say, in the case of the membrane reactor the water gas shift steps are not necessary. Figure 16.18 shows that industrially interesting hydrogen fluxes of >1 m3 (STP) m–2 h–1 (1.67 mL cm–2 min–1) can be obtained on the retentate side. From Figure 16.19 it follows that synthesis gas production is in principle possible but needs optimization to get higher CO selectivities and CH4 conversions.87

16.5.1.2

Production of Hydrogen and Ethylene by Combining Thermal Water Splitting with Oxidative Dehydrogenation

Water splitting and oxidative dehydrogenation of ethane have been studied in the BCFZ oxygen-permeable membrane reactor at moderate temperatures (700–800 1C). Water dissociates into hydrogen and oxygen on the core side of

Perovskite Membranes for High Temperature Oxygen Separation

211

Figure 16.18

CPOM with thermal water splitting as oxygen source: Amount of H2 obtained on the retentate side after H2O condensation (J) and amount of O2 transported through the OTM (B) to the methane side where the CPOM takes place as a function of temperature.87 Experimental details: Flow rate on the cores side of the BCFZ hollow fiber membrane ¼ 40 mL min–1 (30 mL min–1 H2O þ 10 mL min–1 He). Flow rate on the shell side ¼ 50 mL min–1 (3 mL min–1 Ne þ 47 mL min–1 CH4). 0.88 cm2 effective membrane area, 0.8 g Ni/Al2O3 SR catalyst as packed bed on the shell side.

Figure 16.19

CPOM with thermal water splitting as oxygen source: CH4 conversion (J) and CO selectivity (J).65,87 Experimental details: Flow rate on the cores side of the BCFZ hollow fiber membrane ¼ 40 mL min–1 (30 mL min–1 H2O þ 10 mL min–1 He). Flow rate on the shell side ¼ 50 mL min–1 (3 mL min–1 Ne þ 47 mL min–1 CH4). 0.86 cm2 effective membrane area, 0.8 g Ni/AlO3 SR catalyst as packed bed on the shell side, 950 1C.

212

Chapter 16

Figure 16.20

H2 production on the core side and the simultaneous ethylene production on the shell side at 800 1C.89 Experimental details: FH2O ¼ 30 mL min–1 and FHe ¼ 10 mL min–1 on the core side; 40 mL min–1 (FC2H6 ¼ 3 mL min–1 þ FHe ¼ 36 mL min–1 þ FNe ¼ 1 mL min–1) on the shell side.

the perovskite BCFZ membrane. The produced oxygen is then removed as oxygen ion (O2–) to the ethane side of the membrane, where it is consumed to convert ethane to ethylene according to C2H6 þ O2–-C2H4 þ H2O þ 2e. Local charge neutrality is maintained by the counter diffusion of electrons. Thus, the produced oxygen can be continuously removed via the BCFZ membrane, and more hydrogen from water splitting can be produced even under equilibium-controlled conditions. Simultaneously, ethylene can be obtained after steam condensation on the shell side of the membrane. The advantage of this coupling strategy is that the produced hydrogen from water splitting and the ethylene from ethane oxi-dehydrogenation are inherently separated in the membrane reactor. Figure 16.20 presents the production of hydrogen and ethylene in the BCFZ hollow fiber membrane reactor as a function of time. During a period of 100 h, about 60% ethane conversion and 90% ethylene selectivity were obtained. Simultaneously, the hydrogen production rate on the steam side reaches89 0.8 mL (STP) min–1 cm–2.

16.5.2 16.5.2.1

Decomposition of N2O and NO into Nitrogen and using the Abstracted Oxygen for Synthesis Gas Production N2O as Oxygen Source: N2O Abatement and Syngas Production

Since N2O has a greenhouse impact, which is 300 times higher than that of CO2, the N2O release from anthropological sources has to be stopped. Perovskites

Perovskite Membranes for High Temperature Oxygen Separation

Figure 16.21

213

Synthesis gas formation by CPOM with N2O as oxygen source: CH4 conversion (J) and CO selectivity (&).90 Experimental details: Flow rate on the core side ¼ 30 mL min–1 (6 mL min–1 N2O þ 1.5 mL min–1 O2 þ 22.5 mL min–1 He), Flow rate on the shell side ¼ 40 mL min–1 (23 mL min–1 CH4 þ 17 mL min–1 H2O), 0.86 cm2 membrane area, 1.2 g of the Ni-SR catalyst as packed bed on the shell side.

can be used as catalysts for the decomposition following N2O-N2 þ 12O2. From thermodynamic considerations, this reaction should proceed. However, as in the case of the decomposition of NO, the oxygen as product is strongly adsorbed on the perovskite surface thus blocking surface sites for N2O adsorption and following decomposition. The kinetic blockade of the reaction can be released if this surface oxygen is in situ removed via an OTM (Figure 16.21). To have a high driving force for oxygen permeation, the oxygen partial pressure on the permeate side of the OTM can be decreased by continuously consuming oxygen in the CPOM.90

16.5.2.2

NO as Oxygen Source: NO Abatement and Syngas Production

Figure 16.22 shows that also NO can be used for the CPOM to synthesis gas. On one side of the OTM, oxygen is liberated by catalytic decomposition at the perovskite surface due to 2NO-N2 þ O2. Because of the strong oxygen adsorption to the perovskite surface, this reaction is kinetically blocked. However, these surface oxygen species can be continuously removed if there exists a low oxygen partial pressure on the other side of the OTM. The surface oxygen species is incorporated into the lattice of the perovskite OTM and thus transported to the methane side where CPOM takes place. Both the CH4 conversion and the CO selectivity gradually increased with increasing temperature. At 900 1C, a CH4 conversion of 93% with a CO selectivity of 80% can be obtained (Figure 16.22).

214

Figure 16.22

Chapter 16

Use of NO as oxygen source in synthesis gas production: CH4 conversion (J) and CO selectivity (B) in the CPOM on the shell side of the BCFZ perovskite hollow fiber membrane as a function of temperature. Experimental details : Flow rate on the core side ¼ 30 mL min–1 (3 mL min–1 NO þ 27 mL min–1 He), Flow rate on the shell side ¼ 50 mL min–1 (8 mL min–1 CH4 þ 5 mL min–1 H2O þ 1 mL min–1 Ne þ 36 mL min–1 He), 0.86 cm2 membrane area, 0.5 g of the Ni-SR catalyst as packed bed on the shell side.88

16.6 Engineering and Scale-up Aspects So far, mainly relative thick disk-shaped membranes with a limited membrane area were studied because disks can be easily fabricated by a conventional pressing method. Although a multiple planar stack can be adopted to enlarge the membrane area to an industry-relevant scale, many problems, such as the high-temperature sealing and the pressure resistance, have to be faced.91 Tubular membranes with diameters in the centimeter range with thick walls were developed to reduce the engineering difficulties, especially the problems associated with the high-temperature sealing (Figure 16.23).92 However, their small surface area to volume ratio and their relatively thick walls lead to a low oxygen flux and make them unfavorable in practice. A membrane with a thin wall in hollow fiber geometry can solve the problems mentioned above. Compared to the disk and tubular membranes, hollow fiber membranes possess much larger membrane area per unit volume for oxygen permeation.93,94 Furthermore, the resistance of the hollow-fiber membrane as a full material (i.e. non-supported) to oxygen permeation is very much reduced due to the thin wall as it is the case for supported thin perovskite films. Other module solutions are multichannel monoliths, tube-and-plate assemblies, or single-hole tube prepared by extrusion.95 The so-called tube-and-plate concept of Air Products consists of 10 cm  10 cm plates connected with a central support tube for the pure oxygen.96 Hydro Oil and Energy developed a multichannel monolith for oxygen separation.97

Perovskite Membranes for High Temperature Oxygen Separation

Figure 16.23

215

Ba0.5Sr0.5Co0.8Fe0.2O3– tubes (1 cm) of hitk (now Fraunhofer IKTS Dresden) in the pilote module for O2 separation from air.

In recent years, many perovskite hollow fiber membranes have been successfully prepared by a phase-inversion/sintering process.56,94,98,99 Compared to other configurations such as planar or tubular membranes, the hollow fiber membranes can provide a much larger area per unit volume. Thus it is possible to reduce the membrane system size remarkably. Furthermore, the hollow fiber OTM can solve the problem of the high-temperature sealing in fabricating membrane modules.79 Recently, Tan et al.100 investigated the pilotscale production of oxygen from air using perovskite La0.6Sr0.4Co0.2Fe0.8O3–d (LSCF) hollow fiber membranes. The separation performances, stability, scaling up effect and the energy consumption of an OTM system prepared 889 one dead-end hollow fiber have been investigated both theoretically and experimentally. Oxygen (3.1 L (STP) min–1) with the purity of 99.9% was obtained in the OTM system at 1070 1C. Figure 16.24 shows the LSCF hollow fiber membrane module.

16.7 Comparing Cryogenic Air Distillation, Pressure Swing and Permeation with Organic and Inorganic Membranes: Economic Evaluation Oxygen is a key requirement for many advanced industrial processes with the third largest volume of chemicals produced worldwide. Depending on the amount and the purity of oxygen needed, several techniques are used to separate oxygen from air. The cryogenic fractionation technology uses the boiling point differences to separate oxygen from nitrogen and other constituents of air. The cryogenic fractionation technology after Linde can produce pure oxygen with an oxygen concentration >99 vol.%. The first patent on

216

Figure 16.24

Chapter 16

Modul with 889 LSCF hollow fiber membranes.100

cryogenic air separation plant for the production of oxygen was issued in 1902, now it is a proven and reliable process through continuous technology research and development. Due to the overall thermodynamic efficiency of the modernday cryogenic air separation technology, this technique is likely reaching its theoretical limits, the cost of producing oxygen by the cryogenic air separation technology are unlikely to get reduced remarkably. The energy input of competing technologies has to be lower 0.28 kWh kg–1 oxygen of 95 vol.% purity. Pressure swing adsorption (PSA) is a common technology. Adsorption processes are based on the adsorption ability differences of gas species. Zeolites are typical adsorbent materials in PSA processes for oxygen production. In PSA, oxygen is produced by passing air through a vessel containing adsorbent materials. PSA can give oxygen with a purity of up to 95–97 vol.%. Bed size is the controlling factor in the capital costs due to the cyclic nature of the adsorption process. Since production rate depends on the bed volume, capital costs increase linearly as a function of production rate compared to cryogenic plants. Efficiency limitations inherent in adsorption technology restrict its application to relatively small plants. Using organic polymeric hollow fiber membranes, the separation processes are based on the difference in diffusion rates of oxygen and nitrogen through a polymeric material. Due to the smaller size of the oxygen molecule, most membrane materials are more permeable to oxygen than to nitrogen. The O2 concentration in the permeate is typically of the order of 30–50 vol.% under a pressure difference of about 10 bar. Carbon dioxide and water usually appear in

Perovskite Membranes for High Temperature Oxygen Separation

217

the permeate, since they are more permeable than oxygen for most polymeric membranes. Polymeric membranes allow the production of oxygen enriched air and do not provide the separation factor and flux required for economical large-scale production of oxygen. OTM produces oxygen by the passage of oxygen ions through the ceramic crystal structure. High purity oxygen (theoretically up to 100%) can be obtained using OTM. Oxygen can be separated from air using OTM at high temperatures and pressures with high flux and purity in a single-stage operation. The OTM oxygen technology provide a radically different approach to producing low-cost, high-purity tonnage oxygen at temperatures synergistic with power production and many other oxygen intensive applications. If OTM oxygen systems combine cycle coal gasification plants and other advanced power generation systems, the cost of large-scale oxygen production from air can be reduced by approximately one-third compared to conventional cryogenic air separation technologies.101

16.8 Concluding Remarks Oxygen is a basic chemical which is produced by established separation technologies like cryo-distillation, pressure swing adsorption or polymeric membrane permeation from air. Especially the purity and the amount of the oxygen to be produced decide about the economics of these separation technologyes. Dense mixed oxygen ions and electrons conducting materials are mixed oxides in the crystallographic perovskite, brownmillerite or chlorite structures. The advantage of these materials is their infinite oxygen selectivity which is linked to the transport mechanism of oxygen ions via vacancies in the oxygen framework. Moderate oxygen fluxes are found which can provide in highsurface permeators, e.g. when using hollow fibers, oxygen production rates of industrial relevance. The disadvantages of these perovskite-type membranes are the high temperatures above 800 1C to get sufficient oxygen ion mobility as a basis for high oxygen fluxes. This high-temperature operation causes several materials aspects not only of the membrane itself but also for the gas tight potting of the membranes into the housing. Dead-end and flow-through permeators operating with slightly pressurized air as feed and vacuum pumps on the permeate side for the recovery of the permeated oxygen are used. The main energy consumption of the permeator is due to the power for the furnace to heat the air feed to permeation temperature. The energy consumption for pressurizing the air or the vacuum pump is much less. Therefore, heat exchangers have to be integrated to recover the heat energy in exhaust gas and oxygen product. Whereas a catalytic membrane reactor on an industrial scale is still an aim that is difficult to realize, the implementation of a perovskite-based plant for the production of oxygen-enriched air or of air separation seems to be more feasible. By Vente et al.102 the membrane geometries of single-hole tubes, multichannel monoliths, and hollow-fibers for air separation are evaluated.

218

Chapter 16

Acknowledgements The EU is thanked for financing in the 7th Framework Program the IP Innovative Catalytic Technologies & Materials for the Next Gas to Liquid Processes (NEXT-GTL). G. Centi, Messina, and G. Iaquaniello, Rome, are thanked for stimulating discussions. J.C. thanks H.H. Wang (Guangzhou), A. Feldhoff (Hannover), K. Efimov (Hannover) for co-operation during the last few years.

References 1. L. Basini, Catal. Today, 2006, 117, 384. 2. M. Schroeder, Phys. Chem. Chem. Phys., 2005, 7, 166. 3. M. Schroeder, On the Behaviour of Mixed-Conducting Complex Metal Oxides in Oxygen Potential Gradients, Habil Thesis, RWTH Aachen, 2006. 4. C. G. Guizard and A. C. Julbe, in Membrane Science and Technology Series, ed. N. K. Kanellopoulos, Elsevier, Amsterdam, 2000, vol. 6, p. 435. 5. H. J. M. Bouwmeester and A. J. Burggraaf, in Membrane Science and Technology Series, ed. A. J. Burggraaf and L. Cot, Elsevier, Amsterdam, 1996, vol. 4, p. 435. 6. C. Tablet, G. Grubert, H. H. Wang, T. Schiestel, M. Schroeder, B. Langanke and J. Caro, Catal. Today, 2005, 104, 126. 7. H. Wang, T. Schiestel, C. Tablet, M. Schroeder and J. Caro, Solid State Ionics, 2006, 177, 2255. 8. K. Q. Huang, M. Schroeder and J. B. Goodenough, in Proceedings of the International Symposium, Solid State Ionic Devices, ed. E. D. Wachsman, J. R. Akridge, M. Liu and N. Yamazoe, The Electrochemical Society Proceedings Series, 1999, PV 99-13, p. 95. 9. S. Kim, Y. L. Yang, A. J. Jacobson and B. Abeles, Solid State Ionics, 1999, 121, 31. 10. Y. Teraoka, H. M. Zhang, S. Furukawa and N. Yamazoe, Chem. Lett., 1985, 1743. 11. Y. Teraoka, T. Nobunaga and N. Yamazoe, Chem. Lett., 1988, 503. 12. Z. P. Shao, W. S. Yang, Y. Cong, H. Dong, J. H. Tong and G. X. Xiong, J. Membr. Sci., 2000, 172, 177. 13. J. H. Tong, W. S. Yang, R. Cai, B. C. Zhu and L. W. Lin, Catal. Lett., 2002, 78, 129. 14. www.eltronresearch.com. 15. R. Bredesena, K. Jordal and A. Bolland, Chem. Eng. Process., 2004, 43, 1129. 16. H. X. Luo, B. B. Tian, Y. Y. Wei, H. H. Wang, H. Q. Jiang and J. Caro, AIChE J., 2010, 56, 604. 17. H. X. Luo, Y. Y. Wei, H. Q. Jiang, W. H. Yuan, Y. X. Lv, J. Caro and H. H. Wang, J. Membr. Sci., 2010, 350, 154.

Perovskite Membranes for High Temperature Oxygen Separation

219

18. H. Makoto, D. Kazunari, H. Michhikazu and T. Takashi, Chem Lett., 2006, 35, 1326. 19. J. H. Tong, W. S. Yang, B. C. Zhu and R. Cai, J. Membr. Sci., 2002, 203, 175. 20. X. Z. Chen, H. F. Liu, Y. Y. Wei, J. Caro and H. H. Wang, J. Alloys Compd., 2009, 484, 386. 21. V. V. Kharton, A. A. Yaremchenko, M. V. Patrakeev, E. N. Naumovich and F. M. B. Marques, J. Eur. Ceram. Soc., 2003, 23, 1416. 22. V. V. Kharton, A. L. Shaula, F. M. M. Snijkers, J. F. C. Cooymans, J. J. Luyten, A. A. Yaremchenko, A. A. Valente, E. V. Tsipis, J. R. Frade, F. M. B. Marques and J. Rocha, J. Membr. Sci., 2005, 252, 215. 23. Y. Y. Liu, X. Y. Tan and K. Li, Catal. Rev. Sci. Eng., 2006, 48, 145. 24. J. E. ten Elshof, H. J. M. Bouwmeester and H. Verweij, Solid State Ionics, 1995, 81, 97. 25. Y. Teraoka, H. Shimokawa, C. Y. Kang, H. Kusaba and K. Sasaki, Solid State Ionics, 2006, 177, 2245. 26. X. F. Zhu, H. H. Wang and W. S. Yang, Chem. Commun. (Cambridge, U. K.), 2004, 1130. 27. H. H. Wang, C. Tablet, A. Feldhoff and J. Caro, Adv. Mater. (Weinheim, Ger.), 2005, 17, 1785. 28. J. Martynczuk, F. Y. Liang, M. Arnold, V. Sepelak and A. Feldhoff, Chem. Mater., 2009, 21, 1586. 29. K. Efimov, T. Halfer, A. Kuhn, P. Heitjans, J. Caro and A. Feldhoff, Chem. Mater., 2010, 22, 1540. 30. K. Watanabe, D. Takauchi, M. Yuasa, T. Kida, K. Shimanoe, Y. Teraoka and N. Yamazoe, J. Electrochem. Soc., 2009, 156, E81. 31. T. Kida, D. Takauchi, K. Watanabe, M. Yuasa, K. Shimanoe, Y. Teraoka and N. Yamazoe, J. Electrochem. Soc., 2009, 156, E187. 32. U. Balachandran, J. T. Dusek, R. L. Mieville, R. B. Poeppel, M. S. Kleefisch, S. Pei, T. P. Kobylinski, C. A. Udovich and A. C. Bose, Appl. Catal. A: Gen., 1995, 133, 19. 33. U. Balachandran, J. T. Dusek, S. M. Sweeney, R. B. Poeppel, R. L. Mieville, P. S. Maiya, M. S. Kleefisch, S. Y. Pei, T. P. Kobylinski, C. A. Udovich and A. C. Bose, Am. Ceram. Soc. Bull., 1995, 74, 71. 34. H. J. M. Bouwmeester, Catal. Today, 2003, 82, 141. 35. M. Schwartz, J. H. White and A. F. Sammels, (Eltron Research, Inc., Boulder, CO.), U.S. Pat. 6,214,757, 2001. 36. S. Pei, M. S. Kleefisch, T. P. Kobylinski, J. Faber, C. A. Udovich, V. Zhangmccoy, B. Dabrowski, U. Balachandran, R. L. Mieville and R. B. Poeppel, Catal. Lett., 1995, 30, 201. 37. H. Dong, Z. P. Shao, G. X. Xiong, J. H. Tong, S. S. Sheng and W. S. Yang, Catal. Today, 2001, 67, 3. 38. A. F. Sammells, M. Schwartz, R. A. Mackay, T. F. Barton and D. R. Peterson, Catal. Today, 2000, 56, 325. 39. J. Caro, Chem.-Ing.-Tech., 2006, 78, 899.

220

Chapter 16

40. V. V. Kharton, A. V. Kovalevsky, A. P. Viskup, A. L. Shaula, F. M. Figueiredo, E. N. Naumovich and F. M. B. Marques, Solid State Ionics, 2003, 160, 247. 41. F. T. Akin and J. Y. S. Lin, J. Membr. Sci., 2004, 231, 133. 42. J. Kim and Y. S. Lin, J. Membr. Sci., 2000, 167, 123. 43. J. E. ten Elshof, N. Q. Nguyen, M. W. den Otter and H. J. M. Bouwmeester, J. Electrochem. Soc., 1997, 144, 4361. 44. K. Kobayashi and T. Tsunoda, Solid State Ionics, 2004, 175, 405. 45. C. S. Chen, Fine grained zirconia-metal dual phase composites, Ph.D. thesis, University of Twente, 1994. 46. V. V. Kharton, A. V. Kovalevsky, A. P. Viskup, F. M. Figueiredo, A. A. Yaremchenko, E. N. Naumovich and F. M. B. Marques, J. Electrochem. Soc., 2000, 147, 2814. 47. A. L. Shaula, V. V. Kharton and F. M. B. Marques, J. Eur. Ceram. Soc., 2004, 24, 1309. 48. X. F. Zhu and W. S. Yang, AIChE J., 2008, 54, 665. 49. Q. M. Li, X. F. Zhu and W. S. Yang, J. Membr. Sci., 2008, 325, 11. 50. X. F. Zhu, Q. M. Li, Y. Cong and W. S. Yang, Catal. Commun., 2008, 10, 309. 51. H. H. Wang, W. S. Yang, Y. Cong, X. F. Zhu and Y. S. Lin, J. Membr. Sci., 2003, 224, 107. 52. J. X. Yi, Y. B. Zuo, W. Liu, L. Winnubst and C. S. Chen, J. Membr. Sci., 2006, 280, 849. 53. W. Li, T. F. Tian, F. Y. Shi, Y. S. Wang and C. S. Chen, Ind. Eng. Chem. Res., 2009, 48, 5789. 54. B. Wang, M. C. Zhan, D. C. Zhu, W. Liu and C. S. Chen, J. Solid State Electrochem., 2006, 10, 625. 55. W. Li, J. J. Liu and C. S. Chen, J. Membr. Sci., 2009, 340, 266. 56. T. Schiestel, M. Kilgus, S. Peter, K. J. Caspary, H. Wang and J. Caro, J. Membr. Sci., 2005, 258, 1. 57. H. H. Wang, Y. Cong and W. S. Yang, Catal. Today, 2005, 104, 160. 58. H. H. Wang, Y. Cong and W. S. Yang, Catal. Lett., 2002, 84, 101. 59. H. H. Wang, P. Kolsch, T. Schiestel, C. Tablet, S. Werth and J. Caro, J. Membr. Sci., 2006, 284, 5. 60. H. H. Wang, C. Tablet and J. Caro, J. Membr. Sci., 2008, 322, 214. 61. M. Arnold, H. H. Wang and A. Feldhoff, J. Membr. Sci., 2007, 293, 44. 62. M. Arnold, Q. Xu, F. D. Tichelaar and A. Feldhoff, Chem. Mater., 2009, 21, 635. 63. M. Arnold, T. M. Gesing, J. Martynczuk and A. Feldhoff, Chem. Mater., 2008, 20, 5851. 64. F. Y. Liang, H. Q. Jiang, T. Schiestel and J. Caro, Ind. Eng. Chem. Res., 2010, 10.1021/ie101311g. 65. H. H. Wang, S. Werth, T. Schiestel and A. Caro, Angew. Chem., Int. Ed., 2005, 44, 6906.

Perovskite Membranes for High Temperature Oxygen Separation

221

66. R. Baker, in Membrane Technology in the Chemical Industry: Further Directions, ed. S. P. Nunes and K. V. Peinemann, Wiley-VCH, Weinheim, 2001, p. 280. 67. W. Ito, T. Nagai and T. Sakon, Solid State Ionics, 2007, 178, 809. 68. X. F. Zhu, S. M. Sun, Y. Cong and W. S. Yang, J. Membr. Sci., 2009, 345, 47. 69. J. Caro, K. J. Caspary, C. Hamel, B. Hoting, P. Kolsch, B. Langanke, K. Nassauer, T. Schiestel, A. Schmidt, R. Schomacker, A. SeidelMorgenstern, E. Tsotsas, I. Voigt, H. H. Wang, R. Warsitz, S. Werth and A. Wolf, Ind. Eng. Chem. Res., 2007, 46, 2286. 70. H. H. Wang, C. Tablet, T. Schiestel, S. Werth and J. Caro, Catal. Commun., 2006, 7, 907. 71. H. Q. Jiang, F. Y. Liang, O. Czuprat, K. Efimov, A. Feldhoff, S. Schirrmeister, T. Schiestel, H. H. Wang and J. Caro, Chem.–Eur. J., 2010, 16, 7898. 72. C. Delbos, G. Lebain, N. Richet and C. Bertail, presented in part at the Proceedings, The 9th International Conference on Catalysis in Membrane Reaction., Lyon, 28 June to 2 July, 2009. 73. J. A. Lercher, J. H. Bitter, A. G. Steghuis, J. G. van Ommen and K. Seshan, Environmental Catalysis, Imperial College Press, London, 1999. 74. G. E. Keller and M. M. Bhasin, J. Catal., 1982, 73, 9. 75. S. M. Liu, X. Y. Tan, K. Li and R. Hughes, Catal. Rev. Sci. Eng., 2001, 43, 147. 76. L. Olivier, S. Haag, C. Mirodatos and A. C. van Veen, Catal. Today, 2009, 142, 34. 77. F. T. Akin and Y. S. Lin, AIChE J., 2002, 48, 2298. 78. S. Bhatia, C. Y. Thien and A. R. Mohamed, Chem. Eng. J. (Lausanne), 2009, 148, 525. 79. X. Y. Tan and K. Li, AIChE J., 2007, 53, 838. 80. O. Czuprat, T. Schiestel, H. Voss and J. Caro, Ind. Eng. Chem. Res., 2010, 10.1021/ie100282g. 81. O. Czuprat, S. Werth, T. Schiestel and J. Caro, AIChE J., 2010, 56, 2390. 82. O. Czuprat, J. Caro, V. A. Kondratenko and E. V. Kondratenko, Catal. Commun., 2010, 11, 1211. 83. O. Czuprat, S. Werth, S. Schirrmeister, T. Schiestel and J. Caro, ChemCatChem, 2009, 1, 401. 84. S. Ihara, Bull. Electrotechnol. Lab., 1977, 41, 259. 85. U. Balachandran, T. H. Lee, S. Wang and S. E. Dorris, Int. J. Hydrogen Energy, 2004, 29, 291. 86. U. Balachandran, T. H. Lee and S. E. Dorris, Int. J. Hydrogen Energy, 2007, 32, 451. 87. H. Q. Jiang, H. H. Wang, S. Werth, T. Schiestel and J. Caro, Angew. Chem., Int. Ed., 2008, 47, 9341. 88. H. Q. Jiang, L. Xing, O. Czuprat, H. H. Wang, S. Schirrmeister, T. Schiestel and J. Caro, Chem. Commun., 2009, 6738.

222

Chapter 16

89. H. Jiang, Z. Cao, S. Schirrmeister, T. Schiestel and J. Caro, Angew. Chem., Int. Ed., 2010, 49, 5656. 90. H. Q. Jiang, H. H. Wang, F. Y. Liang, S. Werth, T. Schiestel and J. Caro, Angew. Chem., Int. Ed., 2009, 48, 2983. 91. P. N. Dyer, R. E. Richards, S. L. Russek and D. M. Taylor, Solid State Ionics, 2000, 134, 21. 92. H. H. Wang, Y. Cong and W. S. Yang, J. Membr. Sci., 2002, 210, 259. 93. X. Y. Tan, Y. T. Liu and K. Li, AIChE J., 2005, 51, 1991. 94. S. M. Liu and G. R. Gavalas, J. Membr. Sci., 2005, 246, 103. 95. R. Van Hassel Prasad, R. Chen and J. Lane, U.S. Pat. 6,565,632, 2001. 96. P. A. Armstrong, D. L. Bennet, E. P. T. Forster and E. E. van Stein, ITM Oxygen for Gasification, in Proceedings of the Gasification Technology Conference, Washington DC, 3–6 October 2004. 97. T. Bruun, Design Issues for High Temperature Ceramic Membrane Reactors, in Proceedings of the 6th International Catalysis in Membrane Reactors, Lahnstein, Germany, 2004. 98. X. Y. Tan, Y. T. Liu and K. Li, Ind. Eng. Chem. Res., 2005, 44, 61. 99. Z. G. Wang, N. T. Yang, B. Meng, X. Y. Tan and K. Li, Ind. Eng. Chem. Res., 2009, 48, 510. 100. X. Y. Tan, Z. G. Wang, B. Meng, X. X. Meng and K. Li, J. Membr. Sci., 2010, 352, 189. 101. P. A. Armstrong, D. L. Bennet, E. P. T. Forster and E. E. van Stein, ITM Oxygen: The New Oxygen Supply for the New IGCC Market, in Proceedings of the Gasification Technology Conference, San Francisco, California, 9–12 October 2005. 102. J. F. Vente, W. G. Haije, R. Ijpelaan and F. T. Rusting, J. Membr. Sci., 2006, 278, 66.

CHAPTER 17

Zeolite Membranes for Gas Separations C. ALGIERI,*a G. BARBIERIa AND E. DRIOLIa,b a

National Research Council – Institute for Membrane Technology (ITM– CNR), c/o The University of Calabria, cubo 17C, Via Pietro BUCCI, 87036 Rende CS, Italy; b The University of Calabria – Department of Chemical Engineering and Materials, cubo 44A, Via Pietro BUCCI, 87036 Rende CS, Italy

17.1 Introduction to Zeolite Membranes Zeolites are alumino-silicate micro-porous materials composed of AlO45– and SiO44– tetrahedra which build a network of channels and cavities. The size of the channels is typically 3–10 A˚ in the range of molecular dimensions. Their adsorption properties can be changed by varying the Si/Al ratio set for the synthesis. Figure 17.1 shows the topology of the MFI and FAU zeolites. Zeolite crystals are useful for a large number of industrial applications such as petrochemical cracking, cation exchange and for separation and removal of gases and solvents.1 In the last two decades, zeolite membranes have attracted considerable interest in both industrial and academic sectors because they can be used at high temperatures and with organic solvents when polymeric membranes cannot operate. They might separate in continuous gas and liquid mixtures on the basis of differences in the molecular size and shape (e.g. isomers,2 azeotropic mixtures3), on the basis of different adsorption properties4 owing to their crystalline structure and also on the basis of different diffusivity.

Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

223

224

Chapter 17

(a)

(b)

Figure 17.1

Topology of (a) MFI and (b) FAU zeolites.

Table 17.1

Pervaporation (PV) and vapor permeation (VP) properties of zeolite membranes in water–organic mixtures7

Membrane NaA membrane Methanol Ethanol

i-Propanol Acetone T membrane Methanol Ethanol

i-Propanol

Separation T (1C) Feed water (wt%) J (kg m2h1) Separation factor PV PV VP VP VP VP PV PV

50 75 135

75 50

10 10 10 5 0.5 0.1 10 10

0.305 1.69 13.7 8.34 0.758 0.133 2.28 0.832

2100 10 000 15 000 30 000 3 500 1 800 10 000 5 600

PV PV VP VP VP PV VP

50 75 135 135 135 75 110

10 10 10 5 0.5 10 10

0.275 1.33 10.9 6.46 0.567 2.01 7.16

27 2 100 3 500 5 200 1 500 10 000 20 000

Zeolite membrane application at industrial level is strongly limited by costs (specifically support cost5) and reproducibility problems in the preparation stage.6 Today, the industrial applications are only represented by LTA and T zeolite membranes for organic solvent dehydration by means of pervaporation and vapor permeation processes.7 The performance of these commercial membranes is summarized in Table 17.1.7 Both zeolite membranes were highly water-permeable for these organic mixtures and exhibited extremely high permeation fluxes and separation factors. The first large scale pervaporation plant was built in Japan by Mitsui and Mitsui Engineering & Shipbuilding Co.8 Today, Mitsui has also installed a pilot plant in Brazil and Daurala (India) for ethanol dewatering by LTA membranes.5 In Europe, Inocermic produces LTA

Zeolite Membranes for Gas Separations

Figure 17.2

225

Inocermic semi-technical pervaporation plant (UK). Reprinted from J. Caro, M. Noack and P. Kolsch, Zeolite membranes: from the laboratory scale to technical applications, Adsorption 11, 215–227, 2005, with permission of SpringerLink.

supported zeolite membranes for pervaporation processes.9 Figure 17.2 shows the plant of Inocermic (UK). The current commercial zeolite membranes, developed for pervaporation, are not yet useful in gas separations (H2/CO2 selectivity for NaA membranes of Mitsui10 and Inocermic11 are 6 and 5.6, respectively) because of the presence of large inter-crystalline defects. They, furthermore, participate in the separation process. During pervaporation the water fills the intra-crystalline and intercrystalline pathways.12 However, much effort is in progress to produce defect free zeolitic membrane also for gas separations. In this chapter the application of zeolite membranes in gas separations is reported and deeply discussed. The main strategic methods used for the membrane preparation and mass transport through zeolite membranes are also dealt with.

17.2 Preparation of Zeolite Membranes Thin, self-standing zeolite layers larger than a few square centimeters are difficult to synthesize and the resulting structures are brittle.13 Therefore, the zeolitic membranes are synthesized on porous supports. They provide mechanical resistance without introducing additional mass transfer resistance. The most used support material is alumina since the availability of high quality micro-, ultra- and nano-filtration ceramic membranes with smooth top surfaces14 and here used ‘just’ a support. This last property is significant for the preparation of zeolite layers without defects. Stainless steel supports with rougher surfaces and larger pore sizes (>100 nm) than those in alumina are not often used.14 Another drawback of this material is its higher thermal expansion coefficient with respect to alumina. Therefore, stainless steel supported zeolite membranes are more susceptible to cracks during thermal treatments and excursions. The cost of these support types makes the membranes expensive. For instance, considering the membranes produced by Inocermic GmbH the

226

Chapter 17

ceramic support is responsible for at least 70% of the zeolite membrane cost.5 Therefore, it is also important to study the possibility of using cheaper support materials. Mullite is a good candidate owing to its low cost and ease of processing, giving a support with a regular enough structure.15 The one step and the secondary growth are the two methods used in the preparation of zeolitic membranes. The traditional ‘one step’ method consists in immerging the support surface in a suitable hydrothermal chemical system and repeating the deposition if the membrane has defects.16 This method produces very thick membranes with a low concentration of defects. The secondary growth method, decoupling zeolite nucleation from crystal growth, allows the optimization the operating conditions for each step independently of each-other, suppressing the secondary nucleation. The first step consists in the deposition of the crystal seeds on the surface of a support, followed by a crystal growth by means of hydrothermal treatment. The secondary growth method has different advantages with respect to the one step method such as easier operation, higher controllability of the crystal orientation and the film thickness giving a much better reproducibility. Seeding is a very crucial step because it influences the membrane quality. Many seeding procedures have been considered for covering the support surface with seed crystals and the main ones are listed in Table 17.2. However, more controllable seeding procedures described in the literature involve the filtration (Table 17.3) of water slurry from the zeolitic crystals through a porous support. Dead-end filtration causes excessive crystal accumulation whereas the crossflow mode allows a more uniform and compact zeolite layer, because the suspension flows tangentially along the support surface. In cross-flow filtration separation systems, a high velocity is typically used for increasing the sweeping action through the membrane. This is not desired for seeding purposes which

Table 17.2

Seeding procedures reported in different papers and patents

Seeding procedure

References

Rubbing Dip coating Spin coating Cationic polymeric use

17, 18 19, 20 21 22

Table 17.3

Seeding procedures by filtration reported in different papers and patents

Seeding procedure

References

Cross-flow filtration Dead-end filtration (vacuum seeding)

23, 24 25

Zeolite Membranes for Gas Separations

227

require low linear velocity, to enable uniform support coverage thereby avoiding shearing forces on the deposited layer. Coverage uniformity problems can happen when tubular supports are used horizontally and without any rotation. In this case, zeolite seeds will be deposited preferentially in the bottom of the support owing to the effect of the gravitational force. To overcome this problem a new seeding procedure for tubular membranes has been designed15 combining for the first time:  Cross-flow filtration of a water suspension of zeolite seeds through a porous support  Support tilting with respect to the horizontal plane  Support rotation along its longitudinal axis. Suited tilting allows air bubbles, eventually formed inside the support during filtration and responsible for local defects, to go up. The rotation is required to achieve the deposition over the whole membrane area. These three different procedure aspects guarantee a uniform and sufficient coverage of the support with zeolitic seeds. Very recently, Tsapatsis and co-workers26 discovered the possibility of improving the separation performance of the zeolite membranes by eliminating grain boundary defects by means of heat treatment, named ‘rapid thermal processing’, once the membrane top-layer film was formed by hydrothermal treatment. This methodology enables silicalite-1 membranes with high separation performance to be obtained for aromatic and linear hydrocarbons versus their branched isomers. The researchers hypothesized a condensation of Si-OH groups between adjacent crystal grains during the rapid thermal processing. Figure 17.3 illustrates the positive effect of the rapid thermal processing on membrane performance in terms of separation factors. If beneficial effects of the rapid thermal processing on the membrane performance can be demonstrated for other zeolites, compositions, and microstructures, this thermal treatment could contribute, in combination with fast one-step deposition methods, to the realization of large-scale production of zeolite membranes.26 In the last 20 years, great progress on the zeolite membrane has been made, but only 20 structures are used for membrane preparation (Table 17.4) even if 170 zeolitic structures are indicated by the International Zeolite Association. Different techniques can be used in order to evaluate the morphological quality of the zeolite membranes. Scanning electron microscopy is used to study the surface morphology and the cross section. The top view shows the shape and size of the crystals and also the eventual presence of large defects and a cross-sectional view shows the thickness of the zeolite layer. Figure 17.4 shows cross section and top-view of a supported FAU zeolite membrane. The penetration of siliceous species in the porous support is investigated by means of energy dispersive X-ray. Phase identification and determination of any preferred orientation is performed by using X-ray diffraction.

228

Chapter 17

Fluorescence confocal optical microscopy gives the possibility of observing the polycrystalline network of the zeolite membrane and so the internal defects not observable with SEM analysis.27

Figure 17.3

Permeances and separation factors of c-oriented MFI membranes for xylene (A and B), butane (C and D) and hexane (E and F). (A, C, E) transport properties of conventionally calcined membranes and (B, D, F) transport properties of membranes treated by rapid thermal processing. Reprinted from J. Choi, H.-K. Jeong, A. Snyder, J. A. Stoeger, R. I. Masel and M. Tsapatsis, Grain boundary defect elimination in a zeolite membrane by rapid thermal processing, Science, 325, 590–593, 2009, with permission of AAAS.

229

Zeolite Membranes for Gas Separations

Table 17.4

Zeolite membrane topologies present in the literature

Zeolite type

Max pore size (A˚)

FAU NaX NaY LTA MFI Silicalite ZSM-5 CHA SAPO-34 DD3R BEA T

7.4

PV, VP, GS, MR

4.2 5.6

PV GS, PV, VP, MR

3.8

GS

3.6  4.4 7.6 3.6  5.1

Applications

GS PV; GS; MR PV, GS

PV, pervaporation; VP, vapor permeation; GS, gas separation; MR, membrane reactor.

(a)

Figure 17.4

(b)

SEM micrographs of (a) cross section and (b) top view of a FAU supported zeolite membrane.

17.3 Mass Transport in Zeolite Membranes Transport properties (e.g. flux and selectivity) require single and mixture separation tests also for evaluating the presence of defects in the membrane structure. The permeance of a single gas through a zeolite membrane depends not only on the characteristics of the gas molecules and membrane, but also on the temperature and pressure. Figure 17.5 shows the typical curve of the permeance as a function of the temperature of a defect-free zeolitic membrane. This trend is a combination of two mechanisms: the surface diffusion in the inner surface of the pores and the activated gaseous diffusion.28 In the part A–B, the permeance increases because the mobility of the gas molecules increases. However, at a higher temperature the coverage degree of the inner surface of the membrane pores by gas molecules decreases and consequently

230

Chapter 17

Permeance B D

A

C

Temperature

Figure 17.5

Single gas permeance trend as a function of the temperature.

the permeance decreases (B–C). At high temperatures the adsorption mechanism is negligible and the permeance is controlled by an activated mechanism (C–D). The permeation follows the Knudsen diffusion and viscous flow in the mesoporous and macroporous defects, eventually present, respectively. The Knudsen mechanism is present when the pore size is smaller than the mean free path of the diffusing molecules and the collisions among molecules of gaseous species are less frequent than their collisions with the pore wall. Equation 17.1 describes the permeance when the Knudsen diffusion takes place: rffiffiffiffiffiffiffiffiffi 1 g Knudsen 1 g RT D dpore Permeancei ¼ ¼ ð17:1Þ RT d i RT d 3Mi where g is a geometric factor, d is the membrane thickness, dpore is the pore diameter, T is the temperature, R the constant of the ideal gas law, and Mi and DKnudsen are is the molecular weight and the Knudsen diffusion coefficient of the i i-th gas, respectively. The viscous flow occurs when gas molecules collide exclusively with eachanother. In this case no separation can be obtained. Equation 17.2 describes the gas permeance in presence of the viscous flow: Permeancei ¼

2 1 g dpore    Pi RT d 2Zi

ð17:2Þ

where Zi is the viscosity of the i-th gas and Pi is the average pressure. The permeating flux (Ji) is the product of permeance and driving force (DPi) acting on the two membrane surface for each single chemical species: Ji ¼ Permeancei DPi

ð17:3Þ

231

Zeolite Membranes for Gas Separations

The ratio of the permeances of two pure gases (i and j) measured at the same temperature is the ideal selectivity: ideal selectivity i;j ¼

Permeancei Permeancej

ð17:4Þ

In the case of a mixture is fed to the membrane, the selectivity is defined as: selectivity i;j ¼

xPermeate =xPermeate i j xFeed =xFeed j i

ð17:5Þ

Two types of gas permeation measurements are generally performed through zeolite membranes to identify their mass transport properties: the pressure drop and the concentration gradient methods. By the pressure drop method an absolute pressure is applied through the membrane and the single gas fed is forced to permeate the membrane. The concentration gradient method (CGM) is applied when a gas mixture, eventually humidified, is fed to the membrane module. The separation can be performed by pressure difference and/or using a sweep gas such as helium or argon. The sweep gas diffuses during the permeation into the feed side. Van der Graaf et al.29 comparing the two methods found that the concentration gradient one presents disadvantages with respect to the former. First of all in it the sweep gas diffuses from the permeate side to the feed side, thus affecting the permeation of the considered species. Moreover, the pressure drop method reveals any contribution of viscous flow of large membrane defects, if any. The experimental results are sensitive to the technique used, so a comparison of the membrane performance based on measurements by using different techniques might be misleading.

17.4 Zeolite Membranes and Gas Separations Zeolite membranes give the possibility of separating different gaseous and liquid mixtures in a continuous way owing to their uniform pore size, uniform distribution and good adsorption properties. Moreover, owing to their good thermal and chemical stability they are also good candidates for application in membrane reactors.30 In this paragraph, the separations of light gases by using zeolite membranes of different topologies will be described in detail.

17.4.1

Carbon Dioxide Separation

Carbon dioxide is produced by different industrial processes reported below:  Combustion of fossil fuel (energy production)  Hydrogen production  Cement production

232

Chapter 17

 Iron and steel production  Aluminium production. Purification of natural gas is another process requiring carbon dioxide separation. Carbon capture and storage seems to be a technological option for reducing CO2 emissions of existing power plants.31 Conventional technology for gas separation such as cryogenic distillation, absorption and adsorption are energy intensive due to high heating and cooling load and also the installation of the units usually requires a very high capital cost. Over the past half century, there has been a great interest in alternative technology for gas separation in the industries and this has led to the development of membranebased separation. Polymeric membranes are the most widely used due to their economical manufacture. However, the main disadvantages of these membranes are their low stability at high temperature and in the presence of highly sorbent species.32 Moreover, at a high partial pressure of CO2 there is the polymer plasticization with consequently higher permeability and lower selectivity. These disadvantages disappear when zeolite membranes are considered. Zeolite membranes with different topology were studied for CO2 separation from light gases. Table 17.5 reports some of the different results present in the open literature about CO2/N2 and CO2/CH4 separations. Flue gas from fossil fuel power generation is the single greatest contribution to CO2 emissions. Therefore, the separation and capture of CO2 from flue gas is one of the most important measures to control greenhouse gas emission. Generally, the flue gas from power plants has a large volume and a relatively

Table 17.5

Comparison of CO2/N2 and CO2/CH4 separations with zeolite membranes CO2 permeance (nmol m2 s1 Pa1) CO2/N2

Zeolite type–support

T (1C)

NaY–alumina tube KY–alumina tube RbY–alumina tube CsY–alumina tube NaY–alumina disk NaX–alumina disk ZSM-5–alumina tube T–mullite tube T–mullite tube DD3R DD3R–alumina tube SAPO-34–stainless steel tube SAPO-34–stainless steel tube SAPO-34–stainless steel tube

30 35 35 35 30 23 25 35 35 27 28 25

40–300 — — — 21 0.54 36 38 46 60

20–100 39 40 34 20 0.12 54.3 104 — —

160

25 22

CO2/CH4 Reference — — — — — — —

—

400 500 220 67

35 37 37 37 38 42 49 53 53 56 57 58

36

21

—

60

1200

21–32

—

62

Zeolite Membranes for Gas Separations

233

low concentration of carbon dioxide and a high concentration of nitrogen (typically 10–18% CO2 and 60–80% N2, on a molar basis).33 For CO2 separation from flue gas a CO2/N2 selectivity higher than 70 and a permeance higher than 330 nmol m2 s1 Pa1 are required for economical operation.34 Membranes with FAU topology are extensively studied for the separation of carbon dioxide from nitrogen. Very high separation factors were obtained by Kusakabe et al.35 using high quality FAU zeolite membranes prepared by a secondary growth method. For this high quality membrane the transport of the gas species takes place only through the zeolite pores so the contribution of the limited defects is relatively small. CO2 has a quadrupole moment36 (14.31040 C m2) whilst the N2 is a non-polar species and so the carbon dioxide preferentially adsorbed in the zeolite pores blocks significantly reduces the N2 permeation. The authors found that CO2/N2 selectivity decreased with the temperature because of a lower CO2 adsorption and an increase of the N2 permeation as shown in Figure 17.6.35 The effect of different counter-ions (K, Rb and Cs) in the supported NaY zeolite membranes on CO2/N2 separations has also been studied.37 For an equimolar mixture the selectivity is increased in the order Rb>K>Cs>Na due to an increase of CO2/N2 sorption selectivity. In Table 17.6 the sorption, diffusion and overall permeation selectivity are reported for ion exchanged membranes.37 The sorption selectivity always increased for single and binary systems, however, the increase is very pronounced with the binary system and a zeolite membrane loaded with Rb ions shows the highest selectivity. Gu et al.38 studied the effects of the seeding procedures and the conditions of hydrothermal treatment on the FAU membranes quality. For the seeding stage, they used rubbing and dip-coating. They found that dip-coating combined with

Figure 17.6

Temperature effect on the selectivity for FAU membrane A (D m) and FAU membrane B (& ’). Reprinted from K. Kusakabe, T. Kuroda, A. Murata and S. Morooka, Formation of a Y-type zeolite membrane on a porous a-alumina tube for gas separation, Industrial and Engineering Chemistry Research, 36, 649–655, 1997, with permission from ACS.

234

Table 17.6

Chapter 17

Overall permeation, sorption and diffusion selectivity for cation exchanged membranes37

Membrane

system

Overall permeation selectivity

Sorption selectivity

Diffusion selectivity

Na-Y

Single gas Binary mixture Single gas Binary mixture Single gas Binary mixture Single gas Binary mixture

5.4 19 5.6 39 3.2 40 2.2 34

13 125 16 207 28 362 24 291

0.41 0.15 0.34 0.19 0.12 0.11 0.19 0.11

K-Y Rb-Y Cs-Y

a repeated short-duration hydrothermal treatment gave defect free membranes. They asserted that the dip-coating procedure gives a good control of the seed coverage and also a good reproducibility. The short duration hydrothermal treatment is necessary because of the FAU crystallization is much faster than its competing phases NaP and ANA at 100 1C.39,40 The presence of NaP in the zeolite layer can decrease the permeance value and reduce the thermal stability of the membranes.41 The authors also studied the effect of water on CO2/N2 separation. For the system without water, the separation factor declines with the temperature reaching a value of 1.9 at 200 1C. For the humidified system, the CO2/N2 selectivity enhanced in the temperature range 110–200 1C but significantly decreased below 80 1C (Figure 17.7). It is possible to explain the latter result because water adsorption in the low temperature range (below 80 1C) is very high with consequently very low CO2 and N2 permeances. Above 110 1C there is large water desorption with an increase of N2 and especially of CO2 permeances. The positive effect of water at high temperature shows the potential industrial applications of these membranes in the separation of water gas shift product streams and carbon dioxide capture from the power plant. By using supported NaX zeolite membranes Caro and co-workers42 found that considering equimolar mixtures at 23 1C, CO2 is the less permeable gas with respect to nitrogen. The lower permeance of the carbon dioxide can be explained by the electrostatic interaction of the quadrupolar molecules with the accessible cations of the NaX zeolite. NaX and NaY zeolites have the same topology but different Si/Al ratios (NaX from 1 to 1.5; NaY higher than 1.5) and as a result the X structure has a number of Na1 cations higher than NaY.43 MFI has been extensively studied in zeolite membranes preparation due to its pore size suitable for several industrially important separations.44–47 Using MFI supported membranes it was demonstrated that the CO2/N2 separation factor increases with CO2 feed composition because of the higher CO2 adsorption on the zeolite wall, which consequently limits the N2 transport in zeolitic channels.48 The selectivity of this gas species reaches the value of 20 at 180 1C when the carbon dioxide composition is higher than 60% in the feed. Other researchers using membranes with the same topology found the same effect of the CO2 feed concentration on its separation from nitrogen.49

Zeolite Membranes for Gas Separations

Figure 17.7

235

Gas permeance and CO2/N2 separation factor as a function of the temperature for equimolar CO2–N2 mixture in dry and moist conditions. Reprinted from X. Gu, J. Dong and T. M. Nenoff, Synthesis of defect free FAU-type zeolite membranes and separation for dry and moist CO2/ N2 mixtures, Industrial and Engineering Chemistry Research, 44, 937–944, 2005, with permission from ACS.

236

Chapter 17

MFI membranes are also studied for CO2 separation from CH4 (the mainly component of natural gas) because the removal of the acidic gases such as CO2 and H2S from natural gas is another very important operation to avoid pipe-line corrosion and also to improve the gas energy density.50 Zhu et al.51 investigated the adsorption role of CO2 and CH4 feeding a binary mixture as a function of the feed pressure and temperature through MFI supported membranes. The flux of both gases increases with the pressure and the selectivity remaining almost constant at 30 1C. CO2 flux monotonically decreases with the temperature (at 101.3 kPa) whereas the CH4 flux increases up to 65 1C and then decreases; consequently, the selectivity also decreases. The results obtained show that the selectivity is favorable for the CO2 owing to the higher adsorption affinity of the zeolite for it. Figure 17.8 shows the trend of the CO2/CH4 separation factor as a function of the pressure and temperature. Other zeolite membranes such as zeolite T, DD3R and SAPO-34 have been studied for these gaseous separations, since they give the possibility of separating small species on the basis of molecular sieving mechanism due to their small pore size. Zeolite T membranes are hydrophilic and also acid resistant and they would be preferred for the pervaporation dehydration in the presence of organic acid, such as the esterification hybrid process.52 However, there are fewer reports about the zeolite T membranes preparation. Cui et al.53,54 reported the hydrothermal synthesis of these zeolite membranes on pre-seeded mullite tubes. Figure 17.9 shows the feed composition effect on the permeance and CO2/N2 and CO2/CH4 separation factors at 35 1C and 100 kPa.53 By increasing the molar fraction of carbon dioxide the nitrogen permeance decreases, whereas for the other mixture, that of methane slightly decreases. The separation factor for both mixtures increases with the CO2 concentration. The same result was found

Figure 17.8

CO2/CH4 separation factor (a) as function of the total feed pressure at 303 K; and (b) as a function of the temperature at a total feed pressure of 101.3 kPa. (Open symbols are the ideal selectivity). Reprinted from W. Zhu, P. Hrabanek, L. Gora, F. Kapteijn and J. Moulijn, Role of adsorption in the permeation of CH4 and CO2 through a silicalite-1 membrane, Industrial and Engineering Chemistry Research, 45, 937–944, 2006, with permission from ACS.

Zeolite Membranes for Gas Separations

Figure 17.9

237

Feed composition dependence on permeance of each component and CO2/N2 (closed symbol) and CO2/CH4 (open symbol) separation factors at 35 1C and a feed pressure of 100 kPa. Reprinted from Y. Cui, H. Kita and K. Okamoto, Preparation and gas separation properties of zeolite T membranes, Chemical Communications, 2154–2155, 2003, with permission from RSC.

by Lovallo and co-workers48 using MFI zeolite membranes. However, with T zeolite membranes very high selectivity values were found for the binary systems CO2/N2 and CO2/CH4 (Table 17.5) owing to the combination of competitive adsorption, molecular sieving and diffusion through the zeolite pores.54 In the CO2/N2 system the competitive adsorption is more pronounced, whilst the molecular sieving effect has little influence. For the CO2/CH4 system the competitive adsorption and above all the molecular sieving mechanism are effective. The sieving mechanism for the latter system is marked because the methane could permeate, but with some difficulties, through T zeolite membrane since its kinetic diameter equals 3.8 A˚. Defect free DD3R supported membranes have been newly prepared by NGK insulators (Japan).55 The great advantage of this zeolite type with respect to the SAPO-34 and T zeolites, should be the chemical and thermal stability because of its all silica structure. Kapteijn and co-workers56 studied the permeation of various gases (carbon dioxide, nitrous oxide, methane, nitrogen, oxygen,

238

Figure 17.10

Chapter 17

Separation factor of equimolar mixtures through the DD3R membrane as a function of the temperature at 101 kPa (A) and total feed pressure at 30 1C (B). CO2/CH4 (’); N2/CH4 (J); CO2/air (E); N2O/air (}); Air/Kr (,) and N2O/CO2 (m). Reprinted from J. van der Bergh, W. Zhu, J. Gascon, J. A. Moulijn and F. Kapteijn, Separation and permeation characteristics of a DD3R zeolite membrane, Journal of Membrane Science, 316, 35–45, 2008, with permission from Elsevier.

argon, krypton, neon) through DD3R membranes in a temperature range from –197 1C to 100 1C. They found that the selectivity decreases with the temperature, whereas it remains constant or decreases slightly with the pressure. The decreasing trend was observed when one of the components is strongly adsorbed on the zeolite pores as illustrated in Figure 17.10. The performances of these membranes are a result of a molecular sieving effect owing to the opening of the pore size and the competitive adsorption effects when the mixture is

239

Zeolite Membranes for Gas Separations 56

considered. These membranes exhibited very high separation factors for CO2/CH4 mixtures, in fact at room temperature it equals 500, and in the range from –230 1C to 100 1C the values are in the range 1000 to 100. A membrane with the same topology formed on a porous alumina substrate also showed high separation factor (220) at 28 1C and the permeation properties were slightly influenced by water addition in the feed.57 The water adsorbed in the zeolite pores blocks the CO2 and CH4 permeation. The DD3R presents an all silica structure, furthermore, it is expected to be affected less by water adsorption. However, the membrane used in this work was not perfectly hydrophobic (Si/Al ¼ 980), probably owing to the support dissolution in the synthesis gel. SAPO 34 zeolite has pores similar to the kinetic diameter of the CH4 (3.8 A˚) but larger than that of CO2 (3.3 A˚) and for this reason exhibited also a good CO2/CH4 separation factor.58,59 SAPO-34 membranes, prepared by the one step method, present a maximum in CO2 permeance at 2 1C for both single and binary systems. Moreover, the CO2/CH4 selectivity decreases with the temperature in the range from 2 to 200 1C. The highest selectivity value (67) was found at 25 1C.58 Noble and co-workers60 prepared SAPO-34 membranes on a stainless steel support and found that the CO2 and H2 single gas permeances were not strong functions of the temperature in the range from –20 to 35 1C and the CO2/H2 ideal selectivity was around 2. They measured a separation factor for the mixture of CO2 and H2 over 100 at –20 1C and it decreases with the temperature (Figure 17.11). This result confirms how the carbon dioxide adsorption greatly inhibits the H2 adsorption. Indeed, hydrogen has a higher diffusivity than other molecules through the zeolite materials, but it is weakly adsorbed.61

(a)

Figure 17.11

(b)

(a) Single gas permeances and separation factor for CO2 and H2 as a function of the temperature. (b) CO2/H2 selectivity versus the temperature (CO2–H2 mixture 43–57). Reprinted from M. Hong, S. Li, J. L. Falconer and R. D. Noble, Hydrogen purification using a SAPO-34 membrane, Journal of Membrane Science, 307, 277–283, 2008, with permission from Elsevier.

240

Chapter 17

Figure 17.12

Revised Robeson upper bound for CO2/N2. Reprinted from ‘The upper bound revisited’, Journal of Membrane Science, 320, 390–400, 2008, with permission from Elsevier. Data point of SAPO-34 membrane is also shown for comparison. Reprinted from S. Li and C. Q. Fan, High flux SAPO-34 membrane for CO2/N2 separation, Industrial and Engineering Chemistry Research, 49, 4399–4404, 2010, with permission from ACS.

Membranes with the same topology and synthesized by secondary growth method, showed the highest CO2 permeability and a good CO2/N2 selectivity at 22 1C (Table 17.5).62 Figure 17.12 shows (for the year 2008) the revised upper bound of polymeric membranes for the gas pairs carbon dioxide and nitrogen.63 On the same figure, the data of the SAPO-34 membrane62 are above the Robeson’s upper bound showing the better performance of the zeolite membrane with respect to the polymeric ones.63 For such high permeance and good selectivity values, these membranes meet the requirement for economical industrial operation.34

17.4.2

Hydrogen Separation

Today, over 96% the hydrogen is produced from fossil fuel source and the rest is generated by means of water electrolysis. Most of this H2 is used in different processes in the chemical, petrochemical, pharmaceutical, metallurgical and textile industries.64 However, a large scale production of pure H2 increases significantly the cost of this gas. For separation of hydrogen from gaseous streams, membranes can provide an attractive alternative to PSA and cryogenic distillation. Membrane separation processes consume less energy and there is the possibility of continuous operation.65 Hydrogen selective membranes are separated into two main categories: polymeric and inorganic. Table 17.7

241

Zeolite Membranes for Gas Separations

Table 17.7

Hydrogen separation for different commercial membranes

Membrane

Developer

H2/N2

H2/CO

H2/CH4

Reference

Polysulfone Cellulose acetate Polyimide

Monsanto Separex Ube

39 33 35.4

23 21 30

24 26 —

66 67 68

Table 17.8

Hydrogen permeance and selectivity of different inorganic membranes

Membrane Pd95Au5 Si(400) Silica Carbon molecular sieve MFI–stainless steel tube MFI–alumina tube B-ZSM5–alumina tube SAPO-34–stainless steel tube Silylated by catalytic cracking B-ZSM5– alumina tube SAPO-34–stainless steel tube SAPO-34–stainless steel tube

H2 permeance (nmol m2 H2/ T (1C) s1 Pa1) H2/N2 H2/CO2 H2/CH4 i-C4H10 Reference 400 200 200 150

7200 1850 25 2000

82 000 — — 6.8 — — 100

350

—

2.4

500 450

1730 560

25

47

0.47

450

105

—

25

48

25

15

— 1.7

— —

— 321 23 000

— — —

84 75 76 81 85

— —

70 —

88 86

35

—

86

60

—

—

86

0.60

—

59

—

60

—

17

—

—

60

reports the performance in terms of H2 separation factor of different polymeric commercial membranes. However, some problems such as low resistance in the presence of certain chemicals (hydrochloric acid, sulfur oxide and CO2) make them less attractive. Much effort has been also focused on the synthesis of inorganic membranes such as metal, molecular sieving carbon, zeolite and ceramics for the H2 separation.69 Table 17.8 reports some of the results present in the open literature about inorganic membranes for hydrogen separation. Palladium membranes have been proposed for both production and separation/purification of hydrogen. Pure and CO-free H2 produced might be used to feed the PEM-FC. This membrane type is not discussed here since other chapters in this book deal deeply with hydrogen and palladium-based membranes. Silica is another interesting material for the preparation of inorganic membranes for H2 separation. In particular, these are less expensive and also do not undergo hydrogen embrittlement like palladium membranes. They are generally prepared by chemical vapor deposition (CVD)70,71 or sol–gel72–74

242

Chapter 17 75

techniques. For example, Vos and Verweij prepared silica membranes with very low defect concentration via a TEOS-based sol–gel synthesis. The membranes exhibited high permeances for H2 and much lower permeances for CO2, N2 and O2. These membranes show a minor increase of permeance with the temperature for H2, CH4, N2, O2 and a slight decrease for CO2. Separation factors obtained from gas separation experiments with 50/50 (vol.%) gas mixtures are very similar to permselectivities calculated from single gas permeance experiments. The pore size distribution, determined by physical adsorption, is narrow which is in agreement with a narrow range of kinetic diameters in which permeance is found to decrease drastically.75 Supported silica membrane with very high H2/CH4 separation factors (Table 17.8) were prepared by Prabhu and Oyama.76 However, these membranes present very low hydrothermal stability due to the condensation of Si-OH groups.77 Lee and coworkers78 improved the thermal stability of silica membranes supported on the porous stainless steel disks by the new technique of the soaking–rolling method. the thermal stability and the performance of these membranes can be improved by penetration of the coating layer into the pores of the support and dispersion of colloidal silica particles in the coating layer, which contribute to the minimization of the interface between the coating layer and the stainless steel and the decrease in micro-cracks, respectively.78 The carbon-based membranes show superior permeability–selectivity combination to polymeric ones and are categorized in three classes: carbon membranes, carbon molecular sieve membranes and carbon nanotubes.79,80 For molecular sieve membranes Wang and Hong81 reported a H2/N2 selectivity of about 100 and very high permeance value (Table 17.8). Grainger and Hagg82 prepared carbon molecular sieve membranes derived from cellulose hemicellulose and used for hydrogen separation from light hydrocarbons. High permselectivities were found for hydrogen over methane. For example, the permselectivity was approximately1200–1800 at 90 1C. The permselectivity of hydrogen over carbon dioxide was stable, with average values of 11 at 90 1C and 23 at 25 1C. However, the carbon-based membranes have problems that hinder their introduction on the market. They are very fragile and require more careful handling.83 Moreover, traces of strongly adsorbing vapors can clog the pores of the membranes with consequent lowering of the membranes performance. Zeolite membranes have much higher thermal, chemical and hydrothermal stability when compared with the other inorganic membranes. Bakker et al.85 characterized a metal-supported silicalite-1 zeolite membrane over a broad temperature and pressure range. They found that the molecular sieving and difference in diffusivity and adsorption strength are the key factor determining the separation factor. In particular, the strongest adsorbing specie suppresses the permeation of the other component with consequently high separation factors. An inversion in separation selectivity could be observed increasing the temperature when the diffusivity is the dominant mechanism as reported in Table 17.9. With SAPO-34 membrane Noble and co-workers60 separated H2 from CH4 since this latter gas has its kinetic diameter close to the membrane pore size.

243

Zeolite Membranes for Gas Separations

Table 17.9

Separation selectivities and mechanisms of some binary mixture85

Mixture

Separation mechanism

Feed ratio (kPa/kPa)

T (1C)

Separation selectivity

n-Butane/H2 CO2/H2 H2/CO2

Adsorption Adsorption Diffusion

5/95 50/50 50/50

25 25 350

1.3 12 2.4

(a)

Figure 17.13

(b)

Gas permeances of CH4 and H2 as a function of the temperature (a). CO2/H2 selectivity versus the temperature. (b) (CO2–H2 mixture 54–46; (S) single and (M) mixture). Reprinted from M. Hong, S. Li, J. L. Falconer and R. D. Noble, Hydrogen purification using a SAPO–34 membrane, Journal of Membrane Science, 307, 277–283, 2008, with permission from Elsevier.

Furthermore, although the methane is more strongly adsorbed than hydrogen it diffuses more slowly than H2. The hydrogen and methane permeances present a minimum and a maximum with the temperature, respectively, as Figure 17.13(a) shows for both single gas and mixture feed streams. Thus, both the ideal selectivity and the separation factor present a maximum with the temperature (Figure 17.13(b)). Recently, MFI zeolite membranes were modified for reducing the intracrystalline pore size and minimizing the non-selective pores in order to improve the separation of hydrogen from small gaseous species. In particular, methyldiethoxysilane (MDES) and methyldimethoxysilane are used to reduce the pore of MFI zeolite due to a reaction between the silane and zeolite pore to form SiO2. Hong and co-workers86 modified boron-substituted ZSM-5 and SAPO34 membranes using MDES. For the B-ZSM-5 membranes the authors found that the permeances for the small gases, after the chemical treatment, decrease and the separation factors increase. The selectivity values were more than one order of magnitude higher than the original ones. These results are shown in Figures 17.14 and 17.15. Moreover, the CO2 permeated faster than H2 at 22 1C because this gas is more strongly adsorbed than H2 on the zeolite pores.

244

Chapter 17 (a)

(b)

Figure 17.14

H2 permeance before and after silylation for B-ZSM-5 membrane B2 as a function of temperature in a 50/50 mixture of (a) H2/CH4 and (b) H2/ CO2. Reprinted from M. Hong, J. L. Falconer and R. D. Noble, Modification of zeolite membranes for H2 separation by catalytic cracking of methyldiethoxysilane, Industrial and Engineering Chemistry Research, 44, 4035–4041, 2005, with permission from ACS.

However, after the silylation the CO2 permeance is lower than that of hydrogen due to the smaller pore of the membranes. Regularly, the more adsorbed species is preferentially adsorbed in the zeolite pores, blocking the permeation of the other gases. However, this behaviour has been also found in different studies of smaller-pore zeolite membranes. For example, Aoki et al.87 found that CO2 permeated more slowly than He, H2, N2, O2 and CH4 for NaA membranes. The silylation on SAPO-34 membranes does not change the H2 permeance and the H2/CO2 and H2/N2 selectivities (Table 17.8). This treatment does not modify significantly the size of the zeolite pores. However, during the chemical treatment, the MDES reacted also with the acid sites of the defects, reducing their size. Reducing the defects size the CH4 permeance decreased and consequently the H2/CH4 selectivity increased. MFI zeolite supported membrane has been prepared for the catalytic dehydrogenation of iso-butane.88 For a mixture of H2:iso-butane (50:50, molar%) the separation factor is equal to one at room temperature, whereas at

Zeolite Membranes for Gas Separations

245

(a)

(b)

Figure 17.15

Separation selectivity before and after silylation for B-ZSM-5 membrane B2 as a function of temperature in a 50/50 mixture of (a) H2/CH4 and (b) H2/CO2. Reprinted from M. Hong, J. L. Falconer and R. D. Noble, Modification of zeolite membranes for H2 separation by catalytic cracking of methyldiethoxysilane, Industrial and Engineering Chemistry Research, 44, 4035–4041, 2005, with permission from ACS.

higher temperature (Z500 1C) it is 70 (Figure 17.16). At low temperature the permeation is controlled by the adsorption and, the permeate is enriched in butane. Increasing the temperature the diffusion is the dominant mechanism because the butane is less adsorbed. Figure 17.17 shows the possibility of increasing the conversion of iso-butane of about 30–40% at 510 1C when the MFI membrane is used. The proton exchange membrane fuel cell (PEMFC) has been extensively studied in the last two decades for many applications and especially for low emissions vehicles.89 Pure hydrogen is the ideal fuel for the PEMFC. However, it cannot be stored practically in sufficient quantities on-board a vehicle and, therefore, there is a need for an adequate supply infrastructure. On-board H2 production from liquid fuel such as methanol is considered as a promising method for fuel cell vehicle application. Hydrogen-rich reformed gas presents 1–2% of CO.90 Unfortunately, this CO concentration cannot be tolerated by the PEMFC electrodes at low operating temperature. In order to avoid

246

Chapter 17

Figure 17.16

Separation factor of MFI membranes for H2 : i-C4H10 mixture. Reprinted from U. Illigen, R. Schafer, M. Noack, P. Kolsch, A. Kuhnle and J. Caro, Membrane supported catalytic dehydrogenation of isobutane using an MFI zeolite membrane reactor’’, Catalysis Communications, 2, 339–345, 2001, with permission from Elsevier.

Figure 17.17

Iso-butane conversion with and without removal of H2. Reprinted from U. Illigen, R. Schafer, M. Noack, P. Kolsch, A. Kuhnle and J. Caro, Membrane supported catalytic dehydrogenation of iso-butane using an MFI zeolite membrane reactor’’, Catalysis Communications, 2, 339–345, 2001, with permission from Elsevier.

degradation of the cell performance, the carbon monoxide concentration must be kept below 10 ppm in the hydrogen feed streams.91 CO selective or preferential oxidation is the primary method used for hydrogen deep purification92 and various research groups93,94 are working on this topic. Some research groups studied this process using zeolite-based catalytic membranes.95,96 Bernardo et al.30,97 studied the CO selective oxidation in H2-rich mixtures using catalytic Pt-Y zeolite membranes. The authors showed the possibility of reducing the amount of CO in the presence of large quantities of hydrogen in a

Zeolite Membranes for Gas Separations

247

continuous flow-through membrane reactor. The catalytic membrane reactor succeeded in reducing the outlet CO concentration from 10 000 ppm (1%) down to 10–50 ppm.30 In this case the catalytic zeolite membrane did not have a separative role but provided catalyst nano-sized particles entrapped in the thin zeolite layer (a few microns thick). This resulted in an efficient contact between reagents and catalytic sites reducing by-passing and misdistribution generally shown in a packed bed and low pressure drops (zeolitic layer few microns thick). Recently, modified LTA membranes supported on macroporous carbon discs have been synthesized for H2/CO separation.98 The prepared Na-LTA/ carbon membranes were exchanged with K, Rb and Cs in order to modify the pore size of the Na-A zeolite. The degree of ion exchange achieved is 96% for the K form and 70% for Rb and Cs forms. In order to test the membranes under realistic conditions a gas stream containing both H2 and CO (H2:CO ¼ 50 mL min1:1.25 mL min1, rest He, total flow 100 mL min1) was fed through to the membrane. The NaA membrane does not separate hydrogen

Figure 17.18

H2 permeance (A) and CO permeance (B). Reprinted from F. J. VarelaGandı´ a, A. B. Murcia, D. L. Castello´ and D. C. Amoro´s, Hydrogen purification for PEM fuel cells using membranes prepared by ionexchange of Na-LTA/carbon membranes, Journal of Membrane Science, 351, 123–130, 2010, with permission from Elsevier.

248

Chapter 17

from carbon monoxide. In the case of the KA membrane the pore size of the pore is reduced to 0.28 nm. In this case, the reduction of the pore dimension causes a reduction of the H2 and CO permeances in comparison with the NaA form at room temperature. The H2/CO separation factor is equal to 3.0 at 125 1C and to 0.6 at 150 1C. These latter two results show the impossibility of reaching, at high temperature, the target for purifying the H2 rich gas streams. The CO does not permeate through the membranes loaded with Cs and Rb ions, whereas the H2 permeance value is similar to the NaA membrane, over the whole temperature range investigated. This result is due to the larger size of these ions that partially block the zeolite pores. The results are shown in Figure 17.18.98 These results confirm the good potentiality of the Rb-LTA and Cs-LTA membranes for a deep purification of H2-rich streams, allowing the hydrogen final use, e.g. in fuel cell applications.

17.5 Concluding Remarks In these last 10 years research activity on the zeolite membranes has been enormously extended with a significant progress on the understanding of the separation mechanisms and of the synthesis of thin zeolite layer with high flux and good separation factors. Despite this very intense activity, the industrial application of zeolite membranes is confined to NaA and T-type membranes for pervaporation and vapor permeation processes. The commercialization of these membranes for gas separations is hindered by the impossibility of preparing defect free zeolite membranes because a very low defects concentration in the zeolite layer can destroy the separation performance. In order to improve the quality of the zeolite membranes different preparation methods are reported in the open literature. The most promising is the secondary growth method. This method decoupling the nucleation from crystal growth gives the possibility of optimizing the conditions of each step independently. This overview shows the possibility of separating light gases using zeolite membranes. In particular, FAU and MFI zeolite membranes were used for the CO2/N2 separation despite their large pore size. In these membranes the more adsorbing gas suppresses the permeation of the other component leading to good separation factors. The separation factor decreases with the temperature because the adsorption also decreases. It is possible to have very high CO2/N2 and CO2/ CH4 separation factors by zeolite membranes with small pore size (T, SAPO-34 and DD3R). The separations are a result of a molecular sieving mechanism due to the opening of the pore size combined with the competitive adsorption. The zeolitic membranes with better chemical, thermal and hydrothermal stability than the other inorganic membranes also seem a good candidate for the hydrogen separations. The best zeolite membranes are also those with small pore size. Recently, Rb-LTA and Cs-LTA supported membranes separated hydrogen from CO at all temperatures tested with a separation factor close to infinite, no CO permeance has been observed. Zeolite membranes with this topology are promising materials for a deep purification of H2-rich streams, allowing the hydrogen final use, e.g. in fuel cell applications.

Zeolite Membranes for Gas Separations

249

Great effort is being devoted by researchers to improving the selectivity, the manufacturing methodology and to lowering their production costs. The achievement of these goals might place zeolite membranes in the forefront of separations technology.

17.6 List of Symbols R T g d DKnudsen i dpore Mi Zi Pi Ji DPi xpermeate i xpermeate j xfeed i xfeed j

Constant of the ideal gas law Temperature Geometric factor Membrane thickness Knudsen diffusion coefficient of the i-th gas Pore diameter Molecular weight Viscosity of the i-th gas Average pressure Permeating flux Trans-membrane pressure difference Partial pressure in the permeate the i-th gas Partial pressure in the permeate the j-th gas Partial pressure in the feed the i-th gas Partial pressure in the feed the j-th gas

Acknowledgements The joint bilateral project ‘Functionalized ZnO nano tubular membrane assembly for CO2 capturing: design, fabrication and performance assessment studies’ between ITM-CNR and Sol-gel Nano ceramics section of Materials and Minerals Division of the National Institute for Interdisciplinary Science & Technology (NIIST) of the CSIR (India), is gratefully acknowledged for co-funding this research.

References 1. J. B. Nagy, P. Bodart, I. Hannuis and I. Kiricsi, Synthesis, Characterization and Use of Zeolitic Microporous Materials, DecaGen Ltd, 1998. 2. S. Nair, Z. Lai, V. Nikolakis and G. Xomeritakis, Microporous Mesoporous Mater., 2001, 48, 219. 3. J. Coronas and J. Santamarı´ a, Sep. Purif. Method., 1999, 28, 127. 4. M. Yang, B. D. Crittenden, S. P. Perera, H. Moueddeb and J. A. Dalmon, J. Membr. Sci., 1999, 156, 1. 5. J. Caro, M. Noack and P. Kolsch, Adsorption, 2005, 11, 215. 6. E. McLeary, J. C. Jansen and F. Kapteijn, Microporous Mesoporous Mater., 2006, 90, 198. 7. M. Kondo and H. Kita, J. Membr. Sci., 2010, 361, 223.

250

Chapter 17

8. Y. Morigami, M. Kondo, J. Abe and H. Kita, Sep. Purif. Technol., 2001, 25, 251. 9. Inocermc website: http:\\www.inocermic.de. 10. Kita et al., 6th ICIM Montpellier, 2000.. 11. M. Noack, P. Kolsch, R. Schafer, P. Toussaint and J. Caro, Chem.- Ing.Tech., 2001, 73, 958. 12. H. Richter, Gas Separating Zeolite Membranes. Presentation at the Membrain summer school (Gas separation membranes for zero-emission fossil power plants), Aagen, Germany, 2008. 13. C. Algieri, G. Golemme, S. Kallus and J. D. F. Ramsay, Microporous Mesoporous Mater., 2001, 47, 127. 14. M. Hanebuth, R. Dittmeyer, G. Mabande and W. Schwieger, Chem.- Ing.Tech., 2003, 75, 221. 15. C. Algieri, P. Bernardo, G. Barbieri and E. Drioli, Microporous Mesoporous Mater., 2009, 119, 129. 16. Z. A. E. P. Vroon, K. Keizer, A. J. Burgraaf and H. Verweij, J. Membr. Sci., 1998, 144, 65. 17. K. Kusakabe, T. Kuroda and S. Morooka, J. Memb. Sci., 1998, 148, 13. 18. Y. Hasegawa, K.-I. Sotowa, K. Kusakabe and S. Morroka, Microporous and Mesoporous Mater., 2002, 53, 37. 19. M. C. Lovallo and M. Tsapatsis, AIChE J., 1996, 42, 3020. 20. M. P. Bernal, G. Xomeritakis and M. Tsapatsis, Catal. Today, 2001, 67, 101. 21. S. Mintova and T. Bein, Adv. Mater., 2001, 13, 1880. 22. J. Hedlund, M. Noack, P. Kolsch, D. Creaser, J. Caro and J. Sterte, J. Membr. Sci., 1999, 159, 263. 23. M. Pera Titus, J. Llorens, F. Cunill, R. Mallada and J. Santamaria, Catal. Today, 2005, 104, 281. 24. S. Nakao, I. Kumakiri, T. Yamaguchi and T. Ando, JP 20011097715. 25. A. Huang, Y. S. Lin and W. Yang, J. Membr. Sci., 2004, 245, 41. 26. J. Choi, H.-K. Jeong, M. A. Snyder, J. A. Stoeger, R. I. Masel and M. Tsapatsis, Science, 2009, 325, 590. 27. G. Bonilla, M. Tsapatsis, D. G. Vlachos and G. Xomeritakis, J. Membr. Sci., 2001, 182, 103. 28. M. van der Graaf, E. van der Bijl, A. Stol, F. Kapteijn and J. A. Moulijn, Ind. Eng. Chem. Res., 1998, 37, 77. 29. M. van de Graaf, F. Kapteijn and J. A. Moulijn, J. Membr. Sci., 1998, 144, 87. 30. P. Bernardo, C. Algieri, G. Barbieri and E. Drioli, Sep. Purif. Technol., 2008, 62, 629. 31. M. D. D’Alessandro, B. Smit and J. R. Long, Angew. Chem. Int. Ed., 2010, 49, 6058. 32. Y. Li, H. Zhou, G. Zhu, J. Liu and W. Wang, J. Membr. Sci., 2007, 297, 10. 33. L. Liu, A. Chakma and X. Feng, Ind. Eng. Chem. Res., 2005, 44, 6874. 34. J. Huang, J. Zou and W. S. Winston Ho, Ind. Eng. Chem. Res., 2008, 47, 1261. 35. K. Kusakabe, T. Kuroda, A. Murata and S. Morooka, Ind. Eng. Chem. Res., 1997, 36, 649. 36. R. Krishna and J. M. van Baten, J. Membr. Sci., 2010, 360, 323.

Zeolite Membranes for Gas Separations

251

37. Y. Hasegawa, K. Watanabe, K. Kusakabe and S. Morooka, Sep. Purif. Technol., 2001, 22, 319. 38. X. Gu, J. Dong and T. M. Nenoff, Ind. Eng. Chem. Res., 2005, 44, 937. 39. S. Y. Yang, A. G. Vlessidis and N. P. Evmiridis, Ind. Eng. Chem. Res., 1997, 36, 1622. 40. V. Valtchev, S. Mintova, V. Dimov and A. Toneva, Zeolites, 1995, 15, 193. 41. J. H. Dong and Y. S. Lin, Ind. Eng. Chem. Res., 1998, 37, 2404. 42. K. Weh, M. Noack, I. Sieber and J. Caro, Microporous Mesoporous Mater., 2002, 54, 27. 43. S. Buttefey, A. Boutin, C. Mellot-Draznieks and A. H. Fuchs, J. Phys. Chem. B, 2001, 105, 9569. 44. J. Choi, H.-K. Jeong, M. A. Snyder, J. A. Stoeger, R. I. Masel and M. Tsapatsis, Science, 2009, 325, 590. 45. S. Nair, Z. Lai, V. Nikolakis, G. Xomeritakis, G. Bonilla and M. Tsapatsis, Microporous Mesoporous Mater., 2001, 48, 219. 46. N. Mori and T. Tomita, Microporous Mesoporous Mater., 2008, 112, 88. 47. C. Algieri, P. Bernardo, G. Golemme, G. Barbieri and E. Drioli, J. Membr. Sci., 2003, 222, 181. 48. M. Lovallo, A. Gouzinis and M. Tsapatsis, AIChE J., 1998, 44, 1903. 49. D. W. Shin, S. H. Hyun, C. H. Cho and M. H. Han, Microporous Mesoporous Mater., 2005, 85, 313. 50. W. J. Koros and R. Mahajan, J. Membr. Sci., 2000, 175, 181. 51. W. Zhu, P. Hrabanek, L. Gora, F. Kapteijn and J. Moulijn, Ind. Eng. Chem. Res., 2006, 45, 767. 52. K. Tanaka, R. Yoshikawa, Y. Cui, H. Kita and K. Okamoto, Catal. Today, 2001, 67, 121. 53. Y. Cui, H. Kita and K. Okamoto, Chem. Commun., 2003, 17, 2154. 54. Y. Cui, H. Kita and K. Okamoto, J. Mater. Chem., 2004, 14, 924. 55. K. Nakayama, K. Suzuki, M. Yoshida, K. Yajima and T. Tomita, Method for preparing DDR type zeolite film, and composite DDR type zeolite film. U.S. Pat. 10796890, 2004; Eur. Pat. 2002765562, 25.03.2004. 56. J. van den Bergh, W. Zhu, J. Gascon, J. A. Moulijn and F. Kapteijn, J. Membr. Sci., 2008, 316, 35. 57. T. Tomita, K. Nakayama and H. Sakai, Microporous Mesoporous Mater., 2004, 68, 71. 58. S. Li, J. L. Falconer and R. D. Noble, J. Membr. Sci., 2004, 241, 121. 59. S. Li, J. Martinek, J. L. Falconer, T. Gardner and R. D. Noble, Ind. Eng. Chem. Res., 2005, 44, 3220. 60. M. Hong, S. Li, J. L. Falconer and R. Noble, J. Membr. Sci., 2008, 307, 277. 61. D. W. Breck, Zeolite Molecular Sieve, Wiley, New York, 1974. 62. S. Li and C. Q. Fan, Ind. Eng. Chem. Res., 2010, 49, 4399. 63. L. M. Robeson, J. Membr. Sci., 2008, 320, 390. 64. N. W. Ockwing and T. M. Nenoff, Chem. Rev., 2007, 107, 4078. 65. S. Uemiya, Top. Catal., 2004, 29, 79. 66. J. M. S. Henis, United States Patent, 4, 230, 463, 1980. 67. W. J. Shell and C. D. Houston, Chem. Eng. Prog., 1982, 78, 33.

252

Chapter 17

68. T. Tsuru, K. Yamaguchi, T. Yoshioka and M. Asaeda, AIChE J., 2004, 50, 2794. 69. B. D. Freeman, I. Pinnau, Advanced Materials for Membrane Separations ACS Symposium Series, Volume 876, American Chemical Society, Washington, DC 2004. 70. S. Morooka, S. Yan, K. Kusakabe and Y. Aakiyama, J. Membr. Sci., 1995, 101, 9. 71. T. Nagano, S. Fujisaki, K. Sato, K. Hataya, Y. Iwamoto, M. Nomura and S. I. Nakao, J. Am. Chem. Soc., 2008, 91, 71. 72. R. R. Bhave, Inorganic Membranes: Synthesis, Characterization and Applications. Van Nostrand Reinhold, New York, 1991. 73. A. J. Burggraaf and T. Cot, Fundamentals of Inorganic Membrane Science and Technology. Elsevier, Amsterdam, 1996. 74. T. Tsuru, J. Sol-Gel Technol., 2008, 46, 349. 75. R. M. de Vos and H. Verweij, J. Membr. Sci., 1998, 143, 37. 76. A. K. Prabhu and S. T. Oyama, J. Membr. Sci., 2000, 176, 233. 77. L. J. Wang and F. C. N. Hong, Appl. Surf. Sci., 2005, 240, 161. 78. D.-W. Lee, Y.-G. Lee, B. Sea, S.-K. Ihm and K.-H. Lee, J. Membr. Sci., 2004, 236, 53. 79. A. F. Ismail and L. I. B. David, J. Membr. Sci., 2001, 193, 1. 80. Mita Das, John D. Perry and William J. Koros, Carbo, 2010, 48, 3737. 81. L. J. Wang and F. C. N. Hong, Microporous Mesoporous Mater., 2005, 77, 167. 82. D. Grainger and M.-B. Ha¨gg, J. Membr. Sci., 2007, 306, 307. 83. S. Adhikari and S. Fernando, Ind. Eng. Chem. Res., 2006, 45, 875. 84. O. Hatlevik, S. K. Gade, M. K. Keeling, P. M. Thoen, A. P. Davidson and J. D. Way, Sep. Purif. Technol., 2010, 73, 59. 85. W. Bakker, F. Kapteijn, J. Poppe and J. A. Moulijn, J. Membr. Sci., 1996, 117, 57. 86. M. Hong, J. L. Falconer and R. D. Noble, Ind. Eng. Chem. Res., 2005, 44, 4035. 87. K. Aoki, K. Kusakabe and S. Morooka, J. Membr. Sci., 1998, 141, 197. 88. U. Illgen, R. Schafer, M. Noack, P. Kolsh, A. Kuhnele and J. Caro, Catal. Commun., 2001, 2, 339. 89. I. Rosso, C. Galletti, G. Sarocco, E. Garrone and V. Specchia, Appl. Catal. B: Environ., 2004, 48, 195. 90. S. Ren and X. Hong, Fuel Process. Technol., 2007, 88, 383. 91. B. Rohland and V. Plzak, J. Power Sources, 1999, 84, 183. 92. G. Avgouropoulos and T. Ioannides, Appl. Catal. A, 2003, 224, 155. 93. Y. H. Kim, E. D. Park, H. C. Lee and D. Lee, Appl. Catal. A: Gen., 2009, 366, 363. 94. Y. H. Kim, E. D. Park, H. C. Lee, D. Lee and K. H. Lee, Catal. Today, 2009, 146, 253. 95. Y. Hasegawa, K. Kusakabe and S. Morooka, J. Membr. Sci., 2001, 190, 1. 96. K. I. Sotowa, Y. Hasegawa, K. Kusakabe and S. Morooka, Int. J. Hydrogen Energy, 2002, 27, 339. 97. P. Bernardo, C. Algieri, G. Barbieri and E. Drioli, Catal. Today, 2006, 118, 90. 98. F. J. V. Gandia, A. B. Murcia, D. L. Castello` and D. C. Amoro`s, J. Membr. Sci., 2010, 351, 123.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

CHAPTER 18

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production X. TANa AND K. LI*b a

School of Environmental and Chemical Engineering, Tianjin Polytechnic University, Tianjin 300160, China; b Department of Chemical Engineering and Technology, Imperial College London, South Kensington, London SW7 2AZ, UK

18.1 Introduction Separation of oxygen from air is a big business because oxygen is widely used in many industrial sectors and other special fields such as military, aerospace and medical applications. This market may also be massively expanded in future by use of oxygen instead of air as a feed to reduce carbon emissions in CO2 capture. Currently, oxygen is mainly produced by cryogenic distillation, pressure swing adsorption (PSA) or polymeric membrane separation. These processes are either of high energy consumption or unable to produce highpurity oxygen.1 Coupling a cryogenic air separation unit in an oxyfuel power plant would even reduce its power generation efficiencies from currently around 40% to 30%.2 This energy penalty, thus, promotes the development of more effective oxygen production technologies. In the last 20 years, the mixed oxygen–ionic and electronic conducting (MIEC) ceramic membranes such as La1–xSrxCo1–yFeyO3–a perovskite that exhibit appreciable oxygen ionic and electronic conductivity at elevated Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

253

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

254

Chapter 18

temperature have attracted considerable interests as a potentially economical, clean and effective technology for oxygen production.3–6 It was suggested that this technology enabled to reduce O2 production cost by a 30% or more compared to the cryogenic distillation or PSA process.7 As oxygen transport in MIEC membranes is via diffusion of mobile oxygen vacancies and electronic defects while simultaneously excluding the transport of other gas species, highpurity oxygen can be obtained in one step. It is reported that an industrial-scale MIEC ceramic membrane system for tonnage oxygen production has been developed by Air Products & Chemicals and Praxair but the open report is not available possibly due to industrial secrets.8 The membrane systems are based on the planar or tubular designs which possess limited membrane area per unit of volume. In the last decade, a phase inversion/sintering technique has been developed to produce ceramic hollow fiber membranes including MIEC oxides.9–14 The hollow fiber geometry exhibits many advantages over the planar and tubular ones such as, in particular, higher surface area/volume ratio and facile high-temperature sealing. Furthermore, the phase inversion-induced hollow fiber membranes usually possess an asymmetric structure consisting of a dense layer and porous substrate and thus exhibit noticeably reduced resistance to oxygen permeation. This makes it possible to reduce the membrane system size and operation costs remarkably.15 More importantly, there is no expensive equipment required any longer in the membrane fabrication process and the operation is much simplified as well. Therefore, the hollow fiber membranes have more potential to meet commercial targets in air separation. Although much progress has been made in the development of MIEC ceramic hollow fiber membranes there are still many challenges to face, such as oxygen flux, chemical stability and mechanical reliability. This chapter is mainly concerned with the engineering issues related to the commercialization of MIEC ceramic hollow fiber membranes for oxygen production and highlights the significant development in this field.

18.2 Oxygen Permeation in MIEC Ceramic Membranes 18.2.1

Oxygen Permeation Mechanism

The MIEC ceramic membranes are made of the oxides containing multi-metal ions which usually have perovskite or fluorite crystalline structures. There are oxygen vacancies presented in the membrane bulk due to the partial substitution of the higher-valency metal cations by the lower-valency ones. These oxygen vacancies allow the oxygen ions with enough energy to move from one site to another leading to appreciable oxygen ionic conductivity. In the meantime, the presence of multivalent cations in the membrane compositions ensures a high, often predominating electronic conductivity. In the presence of oxygen gas, the oxygen vacancies on the membrane surfaces tend to be filled by the

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

Figure 18.1

255

Schematic of oxygen permeation across the MIEC membranes ( oxygen vacancy; o-lattice oxygen; h-electron hole).

oxygen atom with the formation of two electron holes as a result of charge compensation: 1 2 O2

 x þ V O , OO þ 2h

ð18:1Þ

where the charged defects are defined using the Kro¨ger–Vink notation. That is,  OxO stands for lattice oxygen, V O for oxygen vacancy and hi for electron hole. When the MIEC membrane is placed under an oxygen partial pressure gradient at elevated temperature the oxygen may permeate from the high oxygen partial pressure side to the low oxygen partial pressure side, yielding a net oxygen flux. Figure 18.1 demonstrates the oxygen permeation process through a MIEC membrane including the following steps in series: (1) oxygen molecular diffusion from the gas stream to the membrane surface (high pressure side); (2) reaction between molecular oxygen and oxygen vacancy on the membrane surface (high pressure side); (3) bulk diffusion of oxygen vacancy across the membrane; (4) reaction between lattice oxygen and electron-hole on the membrane surface (low pressure side); and (5) mass transfer of oxygen from the membrane surface to the gas stream (low pressure side). It should be mentioned that the surface exchange reactions may actually involve many substeps such as oxygen adsorption, dissociation, recombination, and charge transfer.16 In general, the resistances between the gas phase and the membrane (steps 1 and 5) are much smaller than the others and thus may be negligible.17

18.2.2

Permeation Flux

The transfer flux of charged species in a mixed conductor can be described by the Nernst–Planck equation:18 Ji ¼ 

si rmi þ Ci u z2i F 2

ð18:2Þ

View Online

256

Chapter 18

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

where si ; mi ; zi ; Ci are the conductivity, electrochemical potential, charge number and concentration of species i, respectively; u is the local velocity of inert marker; F is the Faraday constant. The electrochemical potential for each charged species consists of a chemical potential or an activity term and a local electrostatic potential term, f: mi ¼ m0i þ RT ln ai þ zi Ff

ð18:3Þ

where mi0 ai, f, R and T are the standard chemical potential, activity and the Galvanic (internal) potential, the gas constant and temperature, respectively. For the ideal state, the activity of defect can be replaced by its concentration (activity coefficient is unit). The conductivity of defect can be correlated to its concentration and diffusivity, which is a measure of the random motion of the species i in the lattice, by the Nernst–Einstein equation: si ¼

z2i F 2 Ci Di RT

ð18:4Þ

where Di is the diffusion coefficient of charged species i. For the oxygen permeation in MIEC membranes, (i) the overall charge P balance is applied or zi Ji ¼ 0; and (ii) the local velocity of inert marker is negligible, u ¼ 0. Accordingly, the transport flux of charged defects in the MIEC membrane at steady state can be derived (one-dimensional model) from eqns (18.2) to (18.4) as:19 " # 1  ti dCi X zi tj dCj  Ji ¼ Di Ci    ð18:5Þ Ci dx z Cj dx j6¼i j where ti is the transport number of defect i: si z2 Di Ci ti ¼ P ¼ P i 2 sj zj D j C j

ð18:6Þ

j

 When the oxygen vacancy, V O , and electron hole, h , are the primary mobile charge carriers, the oxygen vacancy flux may be derived from eqn (18.5) as:

JV ¼ 

ðCh þ 4CV ÞDV Dh dCV  Ch Dh þ 4CV DV dx

ð18:7Þ

where the subscripts h and V represent hole and oxygen vacancy, respectively. In most MIEC perovskite membranes for oxygen permeation, the electronic conductivity usually overwhelms the ionic conductivity, i.e. ChDh c CVDV and Ch c CV.20,21 Therefore, eqn (18.7) may be reduced to: JV ¼ DV 

dCV dx

ð18:8Þ

View Online

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

257

Based on the stoichoimetric relation between oxygen and vacancy, the oxygen bulk diffusion flux in the membrane can be given by:

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

1 DV dCV  JO 2 ¼  JV ¼ 2 2 dx

ð18:9Þ

As DV is generally considered to be a constant at a given temperature, the local oxygen flux in a hollow fiber membrane can, therefore, be written as: JO 2 ¼ or

1 dNO2 DV dCV ¼ 2pr dl 2 dr

dNO2 ¼ pDV dl

00 C Rv

dCV

0

0

CV

R Ro

Rin

dr r

ð18:10aÞ

¼ pDV

CV00  CV Ro ln Rin

ð18:10bÞ

where dNO2 =dl represents the oxygen permeation rate per unit fiber length; Ro and Rin are the outer and inner radius of the hollow fiber, respectively; the superscripts 0 and 00 stand for the high (upstream) and the low oxygen partial pressure (downstream) sides of the membrane. On the other hand, the surface exchange reaction rates (as denoted by eqn (18.1)) integrated with all the sub-steps on the upstream and the downstream sides may be written, respectively, as:    0:5 0 dNO2 0 ¼ 2pRo kf PO2 CV  kr ð18:11aÞ dl   0:5  dNO2 00 ¼ 2pRin kr  kf PO2 CV00 ð18:11bÞ dl where kf and kr are, respectively, the forward and reverse reaction rate 0 constants for the surface reactions; PO2 and P00O2 are the oxygen partial pressures on the upstream and the downstream membrane surfaces, respectively. It is noted that eqn (18.11) is based on the fact that the electron holes are essentially constant on membrane surfaces due to the overwhelming electronic conductivity in MIEC oxides, and thus the exchange reaction rates may be of pseudo zero-order with respect to electron hole concentration at steady state under isothermal condition. Combining eqns (18.9) and (18.10) gives the specific permeation rate through the MIEC membranes in terms of the oxygen partial pressures and membrane dimensions as:22 h 0 i 00 kr ðPO2 Þ0:5  ðPO2 Þ0:5 J^O2 ¼ ð18:12Þ Rm 2kf ðRo  Rin Þ Rm 00 0 0 00  ðPO2 Þ0:5 þ  ðPO2 Þ0:5  ðPO2 PO2 Þ0:5 þ DV Ro Rin

View Online

258

Chapter 18

where the specific permeation flux is defined by:

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

dNO2 J^O2 ¼ 2pRm dl in which Rm is the logarithmic radius of the hollow fiber membrane: Rm ¼

Ro  Rin lnðRo =Rin Þ

As can be seen from eqn (18.12), the oxygen permeation rate through the MIEC hollow fiber membranes is determined not only by the membrane’s inherent properties reflected by the permeation kinetic parameters, DV, kf and kr but also by the membrane thickness (Ro – Rin) and the effective surface areas for exchange reactions (as can be seen from eqn (18.11)). In order to improve the specific permeation rate of the MIEC membranes to a commercial level, which is believed to be at least higher than 1 mL cm–2 min–1,23 considerable efforts have been made on the development of new types of MIEC membranes with improved oxygen permeability.23,24 Besides, the permeation rate can also be promoted by improving the membrane structure, i.e. by decreasing the membrane thickness or increasing the effective membrane surface areas either during or after the preparation process, which will be described in the following sections.

18.2.3

Stability

For practical applications, the MIEC membranes for oxygen separation are required to have sufficiently high structural/chemical stability and sustainable mechanical integrity in addition to high oxygen permeability. Unfortunately, it is usually true that for a MIEC membrane the higher the oxygen permeability the lower is the structural and chemical stability. For example, the perovskite membranes derived from SrCoO3–d (SC) by partial substitution of cobalt with higher valency transition metal cations (Fe, Cr, Ti) have higher oxygen permeability in comparison with other structures because the strontium at Asite induces the formation of oxygen vacancies while the cobalt ions at B-site facilitate fast diffusion of oxygen within the oxide bulk and fast oxygen surface exchange kinetics due to the small binding energy with oxygen.25–30 But they also exhibit low structural/chemical stability due to the high and nonlinear thermal expansion, abrupt phase transition at a temperature as well as interaction with gas species such as carbon dioxide.31 The partial substitution of Sr at A-site or Co at B-site with other cations can lead to improvement of the stability of the resultant membrane, i.e. La0.6Sr0.4Co0.2Fe0.8O3–d (LSCF), but the oxygen permeability can also be subsequently reduced because of the decrease in oxygen vacancy concentration.32–37 It is known that the MIEC membranes inevitably undergo a kinetic demixing/ decomposition process under oxygen separation conditions.38–41 When an oxygen ion conductive membrane is exposed to an oxygen chemical potential

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

Figure 18.2

259

Segregation of La0.3Sr0.7CoO3–d membrane after oxygen permeation38 (a-cross sectional; b-surface; c-EDX scan).

gradient, the metal cations in the membrane material would diffuse in the reverse direction of oxygen vacancies due to the thermodynamic driving force.40 If the mobilities of the cations in the oxide are different, segregation would take place towards the membrane surfaces leading to phase separation. That is, the faster cations would be enriched on the high oxygen surface and the slower cations on the low oxygen surface. As a result, low conductive layers of the segregation phase will be formed on both surfaces of the membrane as shown in Figure 18.2,38 which would suppress the oxygen permeation. Although the oxygen permeation flux may remain unaltered during the time span of measurements because of the highly porous nature of the segregated layer, in the long run, the kinetic demixing could lead to a loss of membrane performance or even to mechanical failure. Once the phase boundary limit is exceeded the membrane may eventually decompose (kinetic decomposition) leading to the membrane collapse. The kinetic demixing/decomposition rate of a MIEC membrane depends on the composition of the membrane material. Unfortunately, few studies have been conducted on the relationship between

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

260

Chapter 18

the kinetic demixing rate and the membrane composition as well as the permeation parameters. It is generally considered that the oxides containing Co exhibit low stability specially under reducing environment because the Co ions diffuse much faster than the others. Accordingly, some efforts have been focused recently on the development of cobalt-free membrane materials of high stability.42–45

18.3 Development of MIEC Hollow Fiber Membranes 18.3.1

Preparation

The MIEC hollow fiber membranes are usually fabricated by a combined phase inversion and sintering process.9–14 The immersion induced phase inversion technique has been widely employed to prepare polymeric hollow fiber membranes. Due to the excellent binding capability of the polymers, a large amount of fine inorganic powder can be mixed inside the polymer solution to form a uniform mixture from which the ceramic hollow fiber precusors can be extruded at room temperature. Subsequent sintering would convert these precursor fibers into the ceramic hollow fiber membrane products. Figure 18.3 shows the preparation procedures of inorganic hollow fiber membranes by the phase inversion and sintering technique. The details are described as follows exemplified for LSCF perovskite hollow fibers. To start with, a calculated quantity of PESf and pyrrolidone (PVP) additive is dissolved in N-methyl-2-pyrrolidone (NMP) solvent to form a polymer solution. A given amount of LSCF powder dried in advance is then added gradually under stirring. The stirring is usually carried out continuously for at least 48 h for complete dispersion of the powder in the polymer solution.

Figure 18.3

Procedures of the preparation of MIEC hollow fiber membranes by the phase inversion-sintering technique.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

261

After degassing, the suspension is moved to a stainless steel reservoir. Under a nitrogen pressure, the suspension in the reservoir is pressurized through a spinneret into a water bath to form hollow fiber precursors where deionized water and tap water are served as the internal and external coagulants, respectively. The hollow fiber precursors are kept in the water bath for more than 24 h to complete the solidification process. After being dried and straightened, the hollow fiber precursors are heated in a vertically positioned tubular furnace at 600800 1C for 2 h to remove the organic polymer binder, followed by sintering at 12001400 1C for 4 h to allow the fusion and bonding to occur. Preparation of other ceramic membranes follows the same procedures as above to form hollow fiber precursors but the sintering temperature and sintering time have to be varied depending on the properties of the membrane material in order to obtain gas-tight hollow fiber membranes. The macrostructure of the ceramic hollow fiber membranes can be controlled by modulating the suspension compositions as well as the spinning conditions, as shown in Figure 18.4.46–49 When water is used as both internal and external coagulants, a sandwich-like structure, i.e. a central dense layer integrated with fingers on both sides is formed (Figure 18.4(a)).

Figure 18.4

MIEC hollow fiber membranes with (A) sandwiched structure; (B) symmetric structure; and (C) highly asymmetric structure prepared using different spinning conditions.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

262

Chapter 18

The formation of such asymmetric structures can be attributed to the rapid precipitation occurred at both the inner and outer walls close to coagulants resulting in short finger pores but the slow precipitation at the center of the fiber giving the sponge-like structure. However, when the suspension viscosity is increased to some extent by adding some non-solvent additive, a symmetric structure without the occurrence of any macrovoids within the fiber wall may be obtained as shown in Figure 18.4(b). On the other hand, when a mixture of solvent and non-solvent (such as deionised water) is used as the internal coagulant while water is still served as external coagulant, a highly asymmetric structure comprised of a dense outside layer integrated with an inside porous layer can be generated as shown in Figure 18.4(c). It is believed that the formation of finger-like voids in asymmetric ceramic hollow fiber membranes is initiated by hydrodynamically unstable viscous fingering developed when a less viscous fluid (non-solvent) is in contact with a higher viscosity fluid (ceramic suspension containing invertible polymer binder). Finger-like void growth occurs only below a critical suspension viscosity, above which a sponge-like structure is observed over the entire hollow fiber cross section.49 Therefore, the length of the finger-like pores can be modulated through varying the suspension viscosity by addition of non-solvent additive. The macrostructure influences greatly on the oxygen permeation flux of the perovskite hollow fiber membranes because not only the effective membrane thickness for permeation but also the effective membrane areas for surface exchange reactions may be changed accordingly. For example, the highly asymmetric LSCF hollow fiber membranes have exhibited 2.6–10.5 times oxygen permeation fluxes higher than the sandwich-structured membranes.46 The sintering is a very important process to obtain high performance MIEC hollow fiber membranes. It not only affects the macro- and microstructures but also the oxygen permeation properties. Compared to the relatively homogeneous ceramics manufactured using normal pressing or casting methods, the sintering process of the ceramic hollow fiber membranes prepared via the phase inversion process is more complex due to the asymmetric structure.50 In order to obtain gas-tight hollow fiber membranes, a sintering temperature higher than that for the homogeneous membranes of the same material is generally required. However, the influence of sintering on the permeation properties of mixed conducting membranes is complicated and cannot be generalized. It depends not only on whether the grain boundaries act as high diffusivity paths or barriers for oxygen transport, which is determined by the type of crystalline solids to adjoin the interfaces, but also on the rate-controlling step of the oxygen permeation process (surface exchange kinetics or bulk diffusion). Nevertheless, if the sintering temperature is too high, some impurity phases may be formed on the grain boundaries, blocking the oxygen transport. Therefore, the sintering process should be carefully controlled in order to obtain both gas-tight and high-flux MIEC hollow fiber membranes. For example, in the preparation of LSCF hollow fiber membranes, the sintering temperature should be controlled at around 1300 1C and the sintering time be limited within 2–4 h.51

View Online

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

18.3.2

263

Surface Modification

As described above, the resistances to oxygen permeation in perovskite membranes may be from the bulk diffusion and the surface exchange reactions. For the MIEC membrane with definite composition, the oxygen permeation rate can be improved by either reducing the effective membrane thickness or increasing the effective membrane surface area. When the membrane is thick, the resistance from bulk diffusion plays a dominant role in oxygen permeation and the permeation flux is related to the reciprocal of membrane thickness. As the membrane thickness decreases, the oxygen permeation rate can be increased, but the relative limiting effect of surface exchange processes also increases. When the membrane thickness is far less than the critical thickness (Lc), the membrane surface exchange reaction becomes the rate-limiting step. Under this circumstance, further increase in the permeation rate can be achieved only by improving the surface exchange kinetics which may be realized (i) by coating the membrane surfaces with porous layers to increase the effective surface area, or (ii) by coating the membrane surfaces with materials of superior oxygen exchange properties as catalyst to promote a faster oxygen exchange.52–55 Two methods have been employed to modify the MIEC hollow fiber membrane surfaces for improving the permeation properties: acid modification and coating modification. The acid modification principally applies the reactivity of the MIEC membrane with some acids such as H2SO4 and HCl. When the hollow fiber membrane surfaces are dense with macrovoids enclosed in the fiber wall, the membrane surface area for oxygen exchange reactions is restricted to the membrane’s external surfaces. By using H2SO4 or HCl to remove the dense skin layers of the hollow fibers the enclosed macrovoids would directly open to the gas phase, as can be seen clearly in Figure 18.5. Therefore, the membrane surface areas for exchange reactions are increased noticeably while the effective membrane thickness for bulk diffusion is also decreased (only limited to the central dense layer). As a consequence, the oxygen permeation flux can be improved greatly, e.g. a maximum improvement factor of 18.6 can be obtained at 800 1C for the H2SO4-modified membranes.56 It is noted that the flux improvement decreases with increasing temperature because the relative limiting effect of the bulk diffusion gradually becomes more noticeable at higher temperatures. The acid modification is easily operational but may lead to the loss of mechanical strength or even of gas-tightness of the hollow fiber membranes. Comparatively, the coating modification will not compromise the mechanical strength of the resultant hollow fibers. It is completed by coating a porous catalyst layer on the membrane surfaces followed by sintering at intermediate temperatures. The catalysts for membrane modification in use may be Ag, Pt or a porous perovskite with superior oxygen exchange activities.57–59 The Ag- and the porous LSCF-modified membranes exhibited 17.8 and 9.3 times the original fluxes in the unmodified hollow fibers at 800 1C, respectively.58

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

264

Chapter 18

Figure 18.5 Effect of the surface modification by 98% sulfuric acid for 40 min on the LSCF perovskite hollow fiber membranes. (A) Fiber wall after modification; (B) surface before modification; (C) surface morphology after modification.

18.3.3

Mechanical Strength

The mechanical strength of the MIEC hollow fiber membranes is a very important parameter for the assembly of hollow fiber membrane modules in practical applications. It is well known that the mechanical strength depends on both their macrostructure (i.e. the enclosed pores) and microstructure (i.e. grain size) of the MIEC hollow fibers membranes. For example, the asymmetric membrane has poorer mechanical strength than the symmetric ones although it exhibits higher oxygen permeation fluxes. In general, the higher sintering temperature facilitates the densification of the MIEC hollow fiber membranes and thus favors the increase in mechanical strength. After certain a densification degree, the mechanical strength of the hollow fibers would be mainly controlled by the binding force between the grain boundaries but not by the grain bulk. Furthermore, the formation of impurity phases especially on the grain boundaries may lead to the loss of mechanical strength.51 Therefore, it cannot be always expected to strengthen the MIEC hollow fiber membranes by increasing sintering temperature. In the case of full densification, that is, the membrane is sintered into gas tight, the mechanical strength of the MIEC

View Online

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

265

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

hollow fiber membranes will be mainly determined by the inherent properties of the membrane material. Unfortunately, few studies have been conducted so far to relate the mechanical strength of the MIEC membranes to their compositions especially at high temperatures.

18.4 Design of Hollow Fiber Membrane Systems 18.4.1

Operation Mode

The oxygen partial pressure gradient across the membrane is essential to oxygen separation from air. Three methods can be employed for the production of pure oxygen: (i) imposing a high pressure on the upstream side (45 bar) (high-pressure mode); (ii) applying a vacuum to the downstream side (pressure o0.2 bar) (vacuum mode); and (iii) passing a sweep gas through the downstream side (sweep-gas mode).60 In the high-pressure operation, pure oxygen is ‘pushed out’ from air and can be easily collected in the downstream side. However, it requires the membrane system to be capable of enduring very high pressure at high temperature, which is hard to achieve for most commonly used materials. In this case, the tubular instead of hollow fiber membranes are considered to be more suitable for practical application because the module preparation, sealing and manifolding are relatively simpler for the tubular membrane system compared with the hollow fiber membranes, which may outweighs the low specific surface area of the tubular system.8 The vacuum operation can also produce pure oxygen directly by ‘pulling out’ from air. The oxygen is collected in the rear outlet of the vacuum pump. Since the pressure difference across membrane is not higher than 1 bar, the requirement for the materials in module fabrication is much lower than that by the high-pressure operation. The sweep-gas operation does not bring the above-mentioned problems, but the downstream phase is a mixture of oxygen and the sweep gas. Accordingly, further separation must be carried out in order to obtain pure oxygen unless the mixture may be directly utilized. Unlike in the measurement of the membrane’s permeation property where a sweep gas such as Ar or He is usually used to yield oxygen concentration gradient, water vapor can be served as the sweep gas because it is not only abundant in nature but also can be easily separated from oxygen via the normal condensation operation.61 Although the sweep-gas operation using other sweeping gases cannot produce directly pure oxygen, it can also find wide applications in CO2 sequestration or catalytic partial oxidations of hydrocarbons.15 For example, in a power plant concept with CO2 sequestration, a part of effluent (CO2) is used as the sweep gas for the perovskite oxygen separator to produce the O2/CO2 mixture free of N2. This mixture can be then fed to fuel combustion, yielding almost pure CO2 effluent for recovery. Nevertheless, this concept requires the development of steam and CO2 resistant MIEC membranes.

View Online

266

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

18.4.2

Chapter 18

Design Equation

The above-mentioned operating modes of the hollow fiber membrane module for oxygen production are shown in Figure 18.6 where the flows in a single hollow fiber membrane are depicted. In practical operation, the air feed is generally introduced into the shell side while the oxygen product flows countercurrently in the lumen side of the hollow fiber membranes. Energy consumption can firstly be ignored in the membrane module design. The following assumptions are adopted for the derivation of design equations:  Plug flows with negligible axial dispersion for both the shell- and lumenside streams  Isothermal operation; also no temperature gradients are present in the membrane module

Figure 18.6

Operation of the hollow fiber membrane modules for oxygen production. (A) High-pressure mode, (B) vacuum mode and (C) sweep-gas mode.

View Online

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

267

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

 Both the diffusion coefficient of oxygen vacancy and the surface exchange rate constants are independent of oxygen partial pressure  Ideal gas law applicable  No leaking in the membranes  Steady state operation. The governing equations for different operations of the module are based on the mass conservations in the lumen- and the shell-side, respectively and are given as follows: High pressure operation mode Lumen side for oxygen 00

d PO2 Vl dl RT

!

¼ 2pmRm  J^O2

ð18:13Þ

!

¼ 2pmRm  J^O2

ð18:14Þ

Shell side for oxygen 0

d PO2 Vs dl RT Shell side for nitrogen

0

ðPs  PO2 ÞVs RT

¼ 0:79FAir

ð18:15Þ

The pressure drop in the fiber lumen is described by the Hagen–Poiseuille equation dPl 8mVl ¼ dl mpR4in

ð18:16Þ

The boundary conditions are l ¼ 0;

0

00

PO2 ¼ 0:21Po ; PO2 ¼ Pl ¼ Pa

ð18:17Þ

Vacuum operation mode The governing equations are the same as eqns (18.13) to (18.16) for the highpressure operation mode but only with different operating pressures on both the shell and the lumen side: l ¼ 0;

0

00

PO2 ¼ 0:21Pa ; PO2 ¼ Pl ¼ Po

ð18:18Þ

View Online

268

Chapter 18

Sweep-gas operation mode Shell side for oxygen

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

0

d PO2 Vs dl RT

!

¼ 2pmRm  J^O2

ð18:19Þ

Shell side for nitrogen 0

ðPs  PO2 ÞVs ¼ 0:79FAir RT

ð18:20Þ

Lumen side for oxygen 00

d PO2 Vl dl RT

!

¼ 2pmRm  J^O2

ð18:21Þ

Lumen side for sweep gas 00

ðPl  PO2 ÞVl ¼ Fw RT

ð18:22Þ

dPl 8mVl ¼ dl mpR4in

ð18:23Þ

Pressure drop in lumen

With the boundary conditions l ¼ 0;

00

PO2 ¼ 0; l ¼ L;

0

PO2 ¼ 0:21Pa ; Pl ¼ Pa

ð18:24Þ

In the above equations, Vl and Vs are the volumetric flow rates of the lumen- and the shell-stream, respectively; FAir and Fw are the molar flow rate of air feed and the sweep gas (water vapor); Pl and Ps are respectively the pressures in the fiber lumen and the shell side, respectively; Po is the operating pressure (high-pressure operation) or the vacuum pressure (vacuum operation); Pa is the atmosphere; m is the number of hollow fibers and m is the viscosity of the lumen gas. Simulation studies were conducted on the membrane module containing 500 LSCF hollow fibers with the dimension of 0.156/0.112 cm in od/id and 50 cm in length to compare the three operation modes.62 The results indicated that the sweep-gas and the vacuum operations have higher efficiency than the highpressure operation but the sweep-gas (water vapor) operation is most energy consuming. The advantages and the disadvantages of these three operation modes are summarized in Table 18.1.

View Online

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Table 18.1

Advantages and disadvantages of the different operation modes

Production capacity Energy consumption Pressure drop Requirements for module assembly (material and sealing)

18.4.3

269

High-pressure mode

Vacuum mode

Sweep-gas mode

Low Moderate Negligible High

High Low Noticeable Low

High High Negligible Low

Hollow Fiber Membrane Systems

In practical applications, the hollow fibers have to be assembled into a membrane module. For the sake of sealing and facile replacement of broken fibers, it is suggested that every seven to eight hollow fibers are firstly bundled together with high-temperature sealant and the fiber bundles are then placed in membrane holder to form a membrane module, as shown in Figure 18.7. Each hollow fiber should be individually tested to be gas tight prior to bundling and every bundle is again tested to be free of leakage in advance of assembly of the module. Figure 18.8 shows the flow chart of the hollow fiber membrane system for oxygen production with the vacuum operating mode.63 The membrane module is placed under a vertically positioned tubular furnace with the sealing points kept out of the bottom inlet of the furnace tube. An oil-free vacuum pump is connected to the lumen of the fibers to yield a negative pressure and to collect the oxygen product.

Figure 18.7

Photos of (a) LSCF hollow fiber membranes; (b) hollow fiber bundles; and (c) hollow fiber membrane module.63

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

270

Chapter 18

Figure 18.8 Flow chart of the hollow fiber membrane system for oxygen production with the vacuum operation mode.

The maximum oxygen production rate at 1070 1C may reach 3.10 L(STP) min–1 with the oxygen concentration of 99.88%. This production rate is far lower than the ideal capacity of the system, 22.9 L(STP) min–1 (or 1.38 m3(STP) h–1) at 1000 1C, which is calculated based on the permeability of each hollow fiber) because most of the membrane areas are not served for oxygen permeation due to the great temperature gradient in the system. Minor leakage (6–15 mL min–1) is present in the system leading to lower oxygen purities than the theoretical value (100%). The system has been operated continuously for 1067 h at around 960 1C with the oxygen purity of 99.4%.

18.5 Energy Consumption and Cost Analysis The final cost of oxygen product for the MIEC membrane systems is mainly determined by the price of the hollow fiber membranes and the energy consumption per volume of oxygen. The exact costs of the MIEC hollow fiber membranes cannot be given presently because up to date there are no such membranes are commercially provided. Furthermore, the degradation data of the MIEC membranes under practical operation conditions are also unknown so far. However, the energy consumption per unit oxygen can be approximately calculated with which the oxygen cost may be estimated. The following calculation for the vacuum operation is based on the assumptions as: (i) negligible pressure drop in fiber lumen; (ii) no concentration and temperature gradients in the membrane module; (iii) pure oxygen product; and (iv) 10% heat loss for the membrane system. Figure 18.9 shows the flow patterns of the membrane system combined with heat exchangers for heat recovery where the parameters for energy calculation are also presented.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

Figure 18.9

271

Diagram of the membrane system with the heat exchange (A) between oxygen product and the air feed; (B) between exhaust gas and the air feed; (C) between both the exhaust and product and the air feed.

The oxygen balance in the membrane system is given by: 0:21FAir ¼ FOx þ ðFAir  FOx Þ  xe

ð18:25Þ

where FAir and FO2 are the air feed flow rate and the oxygen production rate, respectively; xe is the oxygen fraction in the furnace tube and exhaust stream. For the surface modified hollow fiber membrane, the oxygen production rate may be given by:56 h 0 i 00 akr ðPO2 xe Þ0:5 ðPO2 Þ0:5    FO2 ¼ Am   Rm 2kf ðRo Rin Þ 00 0 0 00 0:5 Rm 0:5 0:5 ðPO2 Þ þ ðPO2 xe Þ ðPO2 xe PO2 Þ þ DV Ro Rin ð18:26Þ

View Online

272

Chapter 18

where Am is the membrane area of the hollow fiber module; a the permeation enhancement factor due to surface modification. The heating power of the furnace may be given, respectively by:

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

 Heat exchange between the air feed and the oxygen product only   Q1 ¼ FO2 CpO2 ð5025Þþ ð0:21FAir FO2 ÞCpO2  þ0:79FAir CpN2 ðT 25Þ =0:9

ð18:27aÞ

 Heat exchange between the air feed and the exhaust stream only   Q2 ¼ FO2 CpO2 ðT 25Þþ ð0:21FAir FO2 ÞCpO2  ð18:27bÞ þ 0:79FAir CpN2 ð5025Þ =0:9

 Heat exchange between the air feed and both the exhaust and oxygen streams   Q3 ¼ FO2 CpO2 ðT 0 25Þþ ð0:21FAir FO2 ÞCpO2  ð18:27cÞ þ 0:79FAir CpN2 ð5025Þ =0:9 where Cpi is the specific heat capacity of gas species i; T 0 is the temperature of the air feed after heating by the exhaust stream: T0 ¼

Table 18.2

ð0:21FAir  FO2 ÞCpO2 þ 0:79FAir CpN2  ðT  50Þ þ 25 ð0:21CpO2 þ 0:79CpN2 ÞFAir

ð18:28Þ

Parameters for calculating the energy consumption of oxygen production

Parameter Membrane area Outer diameter of the fiber membrane Inner diameter of the fiber membrane Enhancement factor by surface modification Operating vacuum degree Air feed flow rate Air feed temperature Temperature of exhaust gas Operating temperature Efficiency of vacuum pump Diffusion coefficient of oxygen vacancy Reverse surface exchange reaction rate constant Forward surface exchange reaction rate constant

Am ¼ 1 m 2 0.18 cm 0.12 cm a ¼ 2.0 98.3 kPa 0.2 N m3 min–1 25 1C 50 1C 850–1050 1C 0.8   8852:5 DV ¼ 1:58  102 exp  ; cm2 s1 T   27291 kf ¼ 1:85  104 exp  ; cm Pa0:5 s1 T   29023 ; mol cm2 s1 kr ¼ 2:07  104 exp  T

View Online

273

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

Figure 18.10

Energy consumption of the oxygen product by the LSCF hollow fiber membrane system with heat exchange (A) between the air feed and the oxygen stream; (B) between the air feed and the exhaust stream; and (C) between air feed and both the exhaust and oxygen streams.

The power of the vacuum pump corresponding to the oxygen product rate may be estimated:   Pa VO2  lnðPa =P2 Þ 2:48  103 FO2 Pa ¼  ln W¼ ð18:29Þ Z P2 Z where Z is the efficiency of the vacuum pump.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

274

Chapter 18

Using the parameters listed in Table 18.2, the energy consumption of the membrane system with different membrane areas to produce one molar oxygen product was calculated and plotted against operating temperature in Figure 18.10. For comparison, the oxygen cost by the PSA process is also presented with a dash line in the figure. It indicates that the oxygen production cost of the membrane process can be reduced to be much lower than the PSA cost level when the heat in both the exhaust gas and the oxygen product is recovered. The energy consumption per unit oxygen product by the membrane system decreases with increasing the operating temperature or the membrane area whose effect can be evaluated by the oxygen recovery which is defined as: x¼

FOx  100% 0:21FAir

The calculation also indicates that the oxygen recovery should be limited within 20  40% so as to reduce the overall oxygen production cost of the MIEC membrane systems.

18.6 Concluding Remarks The MIEC membranes in hollow fiber configuration display a lot of advantages over others such as the large membrane area per unit packing volume, the reduced resistance to oxygen permeation as well as the facile assembly into membrane modules. For the practical application in oxygen production, the hollow fiber membranes have to possess high mechanical strength and sufficiently high structural/chemical stability in addition to high specific oxygen permeation rate. A strategy to achieve these requirements is to use the MIEC membrane materials with enough structural/chemical stability but to improve the specific permeation rate by optimizing the macro- and microstructures of the hollow fiber membranes or by post-modifications. The vacuum operation mode is the best for the high energy-efficiency and the low requirement for the system materials. Heat exchangers have to be integrated in the membrane system to recover the heat energy in exhaust gas and oxygen products so as to reduce the oxygen cost to commercial levels. The membrane area and the operating temperature should be optimized so that the oxygen recovery is limited within 2040%. However, there are still many challenges especially in the following aspects have to be faced on the way to commercialization:  The phase inversion/sintering technique has to be further improved to produce in a large scale MIEC hollow fiber membranes with constant macrostructure and performances so as to reduce the membrane costs.  It is difficult and time-consuming to fabricate membrane modules from the hollow fibers due to the poor mechanical strength. Moreover, the membrane modules are also damageable for both transportation and utilization. It would be better if the phase inversion technique can be modified to prepare monolith-structured membranes.

View Online

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

275

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

 In order to integrate the heat exchangers in the membrane system and to eliminate the temperature difference in membrane module by feeding hot air, the high-temperature sealing problem associated with the fabrication of hollow fiber modules and membrane systems has to be solved.

18.7 List of Symbols Symbol

Definition

Units

Am

Effective membrane area for oxygen permeation Activity of charged defect i Concentration of charged defect i Specific heat capacity of gas species Diffusion coefficient of charged defect i Faraday constant Molar flow rate of air feed and the sweep gas Oxygen permeation rate Reverse surface exchange reaction rate constant Forward surface exchange reaction rate constant Length variable of hollow fiber membrane Length of hollow fiber membrane Number of hollow fiber membranes Oxygen permeation flux Atmosphere pressure Pressure in fiber lumen and shell side Oxygen partial pressure in the upstream and downstream side Operating pressure or the vacuum pressure Heating power of the furnace

m2

ai Ci Cp Di F FAir, Fw FO2 kr kf l L m JO2 Pa Pl , Ps P0O2 ; P00O 2 Po Q1, Q2, Q3 R Rm Ro, Rin T T0 ti Vl , Vs W xe zi a d si m mi

Gas constant Algorithmic radius of fiber Outer and inner radius of hollow fiber Temperature Temperature of the air feed after heating by the exhaust stream Transport number of defect i Volumetric flow rate of the lumen- and the shell-stream Power of the vacuum pump Oxygen fraction in the furnace tube and exhaust stream Charge number of defect i Permeation improvement factor due to the surface modification Hollow fiber membrane thickness Conductivity of defect i Gas viscosity Electrochemical potential of defect i

— mol m3 J (mol K)1 m2 s1 9.648104C mol1 mol s1 mol s1 mol (m2 s)1 m (Pa0.5 s)1 m m — mol (m2 s)1 1.013105Pa Pa Pa Pa W 8.314 J (mol K)1 Rm ¼ ðRo  Rin Þ= lnðRo =Rin Þ m K K — m2 s1 W — — — d=Ro – Rin S m1 Pa s J mol1

View Online

276

Chapter 18

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

(Continued ) Symbol

Definition

Units

m0i

Standard chemical potential Oxygen recovery Efficiency of vacuum pump Local electrostatic potential or galvanic (internal) potential Local velocity of inert marker

J mol1 % — V

x Z f u

Subscript h Electron hole V Oxygen vacancy

m s1

— —

References 1. H. J. M. Bouwmeester and A. J. Burggraaf, in Fundamentals of Inorganic Membrane Science and Technology, ed. A. J. Burggraaf and L. Cot, Elsevier Science B.V., Amsterdam, 1996, p. 435. 2. P. N. Dyer, R. E. Richards, S. L. Russek and D. M. Taylor, Solid State Ionics, 2000, 134, 21. 3. S. P. S. Badwal and F. T. Ciacchi, Adv. Mater., 2001, 13, 993. 4. A. Thursfield and I. S. Metcalfe, J. Mater. Chem., 2004, 14, 2475. 5. Y. Liu, X. Tan and K. Li, Catal. Rev., 2006, 48, 145. 6. M. Exter, J. F. Vente, D. Jansen and W. G. Haije, Energy Procedia., 2009, 1, 455. 7. P. A. Armstrong, E. P. Foster, D. A. Horazak, H. T. Morehead and V. E. Stein, in Proceedings of the 22nd International Pittsburgh Coal Conference, Pittsburgh, Pennsylvania, 2005. 8. J. F. Vente, W. G. Haije, R. I. Jpelaan and F. T. Rusting, J. Membr. Sci., 2006, 278, 66. 9. J. Luyten, A. Buekenhoudt, W. Adriansens, J. Cooymans, H. Weyten, F. Servaes and R. Leysen, Solid State Ionics, 2000, 135, 637. 10. X. Tan, S. Liu and K. Li, J. Membr. Sci., 2001, 188, 87. 11. X. Tan, Y. Liu and K. Li, Ind. Eng. Chem. Res., 2005, 44, 61. 12. K. Li, X. Tan and Y. Liu, J Membr. Sci., 2006, 272, 1. 13. S. Liu and G. Gavalas, J. Membr. Sci., 2005, 246, 103. 14. T. Schiestel, M. Kilgus, S. Peter, K. J. Caspary, H. Wang and J. Caro, J. Membr. Sci., 2005, 258, 1. 15. J. Caro, H. Wang, C. Tablet, A. Kleinert, A. Feldhoff, T. Schiestel, M. Kilgus, P. Kolsch and S. Werth, Catal. Today, 2006, 118, 128. 16. B. A. Hassel, T. Kawada, N. Sakai, H. Yokokawa, M. Dokiya and H. J. M. Bouwmeester, Solid State Ionics, 1993, 66, 295. 17. S. J. Xu and W. J. Thomson, Chem. Eng. Sci., 1999, 54, 3839. 18. K. S. Goto, Solid State Electrochemistry and its Applications to Sensors and Electronic Devices, Elsevier, New York, 1988. 19. X. Tan, S. Liu, K. Li and R. Hughes, Solid State Ionics, 2000, 138, 149.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

Engineering Aspects of MIEC Hollow Fiber Membranes for Oxygen Production

277

20. L. Qiu, T. H. Lee, L.-M. Liu, Y. L. Yang and A. J. Jacobson, Solid State Ionics, 1995, 76, 321. 21. V. V. Kharton, E. N. Naumovich and A. V. Nikolaev, J. Membr. Sci., 1996, 111, 149. 22. X. Tan and K. Li, AIChE J., 2002, 48, 1469. 23. A. Leo, S. Liu and J. C. D. Costa, Int. J. Greenhouse Gas Control, 2009, 3, 357. 24. J. Sunarso, S. Baumann, J. M. Serra, W. A. Meulenberg, S. Liu, Y. S. Lin and J. C. D. Costa, J. Membr. Sci., 2008, 320, 13. 25. A. V. Kovalevsky, V. V. Kharton, V. N. Tikhonovich, E. N. Naumovich, A. A. Tonoyan, O. P. Reut and L. S. Boginsky, Mater. Sci. Eng. B, 1998, 52, 105. 26. V. V. Kharton, A. P. Viskup, I. P. Marozau and E. N. Naumovich, Mater. Lett., 2003, 57, 3017. 27. C. Fan, Y. Zuo, J. Li, J. Lu, C. Chen and D. Bae, Sep. Purif. Technol., 2007, 55, 35. 28. T. Nagai, W. Ito and T. Sakon, Solid State Ionics, 2007, 177, 3433. 29. Y. Cheng, H. Zhao, D. Teng, F. Li, X. Lu and W. Ding, J. Membr. Sci., 2008, 322, 484. 30. B. Meng, Z. Wang, X. Tan and S. Liu, J. Eur. Ceram. Soc., 2009, 29, 2815. 31. V. V. Kharton, A. A. Yaremchenko, A. V. Kovalevsky, A. P. Viskup, E. N. Naumovich and P. F. Kerko, J. Membr. Sci., 1999, 163, 307. 32. J. W. Stevenson, I. R. Armstrong, R. D. Carneim, L. R. Pederson and W. J. Weber, J. Electrochem. Soc., 1996, 143, 2722. 33. Z. Shao, G. Xiong, H. Dong, W. Yang and L. Lin, Sep. Purif. Technol., 2001, 25, 97. 34. J. Tong, W. Yang, B. Zhu and R. Cai, J. Membr. Sci., 2002, 203, 175. 35. S. Diethelm, J. van Herle, P. H. Middleton and D. Favrat, J. Power Sources, 2003, 118, 270. 36. L. Tan, L. Yang, X. Gu, W. Jin, L. Zhang and N. Xu, J. Membr. Sci., 2004, 230, 21. 37. H. Lu, J. Tong, Z. Deng, Y. Cong and W. Yang, Mater. Res. Bull., 2006, 41, 683. 38. R. H. E. Doorn, H. J. M. Bouwmeester and A. J. Burggraaf, Solid State Ionics, 1998, 111, 263. 39. H. L. Lein, K. Wiik and T. Grande, Solid State Ionics, 2006, 177, 1587. 40. M. Martin, J. Chem. Thermodyn., 2003, 35, 1291. 41. B. Wang, B. Zydorczak, Z.-T. Wu and K. Li, J. Membr. Sci., 2009, 344, 101. 42. T. Ishihara, T. Yamada, H. Arikawa, H. Nishiguchi and Y. Takita, Solid State Ionics, 2000, 135, 631. 43. X. Zhu, Y. Cong and W. Yang, J. Membr. Sci., 2006, 283, 38. 44. H. Wang, C. Tablet, A. Feldhoff and J. Caro, Adv. Mater., 2005, 17, 1785. 45. Y. Teraoka, H. Shimokawa, Ch. Y. Kang, H. Kusaba and K. Sasaki, Solid State Ionics, 2006, 177, 2245.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00253

278

Chapter 18

46. Z. Wang, N. Yang, B. Meng and X. Tan, Ind. Eng. Chem. Res., 2009, 48, 510. 47. B. Zydorczak, Z. Wu and K. Li, Chem. Eng. Sci., 2009, 64, 4383. 48. W. Yin, B. Meng, X. Meng and X. Tan, J. Alloys Compd., 2009, 476, 566. 49. B. F. K. Kingsbury and K. Li, J. Membr. Sci., 2009, 328, 134. 50. Y. Liu and K. Li, J. Membr. Sci., 2005, 259, 47. 51. X. Tan, Z. Wang and K. Li, Ind. Eng. Chem. Res., 2010, 49, 2895. 52. K. S. Lee, S. Lee, J. W. Kim and S. K. Woo, Desalination, 2002, 147, 439. 53. Y. Teraoka, Y. Honbe, J. Ishii, H. Furukawa and I. Moriguchi, Solid State Ionics, 2002, 152, 681. 54. S. Lee, K. S. Lee, S. K. Woo, J. W. Kim, T. Ishihara and D. K. Kim, Solid State Ionics, 2003, 158, 287. 55. F. M. Figueiredo, V. V. Kharton, A. P. Viskup and J. R. Frade, J. Membr. Sci., 2004, 236, 73. 56. Z. Wang, H. Liu, X. Tan, Y. Jin and S. Liu, J. Membr. Sci., 2009, 345, 65. 57. A. Thursfield and I. S. Metcalfe, J. Membr. Sci., 2007, 288, 175. 58. X. Tan, Z. Wang, H. Liu and S. Liu, J. Membr. Sci., 2008, 324, 128. 59. A. Leo, S. Liu and J. C. D. da Costa, J. Membr. Sci., 2009, 340, 148. 60. X. Tan, Y. Liu and K. Li, AIChE J., 2005, 71, 1991. 61. H. Wang, P. Kolsch, T. Schiestel, C. Tablet, S. Werth and J. Caro, J. Membr. Sci., 2006, 284, 5. 62. X. Tan and K. Li, AIChE J., 2007, 53, 838. 63. X. Tan, Z. Pang and K. Li, J. Membr. Sci., 2008, 310, 550.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

CHAPTER 19

New Metrics in Membrane Gas Separation A. BRUNETTI,*a G. BARBIERIa AND E. DRIOLIa,b a

Institute on Membrane Technology (ITM-CNR), National Research Council, c/o The University of Calabria, Cubo 17C, Via Pietro Bucci, 87036 Rende CS, Italy; b Department of Chemical Engineering and Materials, The University of Calabria, Cubo 44A, Via Pietro Bucci, 87036 Rende CS, Italy

19.1 Introduction Process intensification is a key strategy that the chemical and petrochemical industry is adopting for increasing energy efficiency and profitability.1 In the last few years the potentialities recognized to membrane operations in this field, have contributed to confirm membrane engineering as a powerful tool to fulfil process intensification strategy in the best way. Today membrane technology for gas separation (GS) is a well-consolidated technique, in various cases it competes with traditional operations. Separation of air components, H2 from refinery industrial gases, natural gas dehumidification, separation and recovery of CO2 from biogas and natural gas are some examples in which membrane technology is successfully applied at the industrial level.2,3 For 10 years, membrane operations have been used for the separation of air components or oxygen enriched air for use in several fields, including chemical and related industries, the mechanical field, food packing and so forth. Mixtures containing more than 40% oxygen or 95% nitrogen can be produced. Currently, membranes dominate the fraction of the nitrogen market for applications with

Membrane Engineering for the Treatment of Gases, Volume 2: Gas-separation Problems Combined with Membrane Reactors Edited by Enrico Drioli and Giuseppe Barbieri r Royal Society of Chemistry 2011 Published by the Royal Society of Chemistry, www.rsc.org

279

View Online

280

Chapter 19

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

1

a productive capacity of less than 50 tonne day and relatively low purity (nitrogen 95–99.5%molar). The possibility of using membrane systems in solving problems such as the greenhouse effect related to CO2 emission is also ongoing. Membranes able to separate CO2 from flue gas streams emitted by power plant, kilns, steel mills, etc. with a high CO2/N2 selectivity, might be used at any large-scale industrial CO2 source. The significant positive results achieved in GS membrane systems are, however, still far from realizing the potentialities of this technology. Problems related to pre-treatment of streams, membrane life time, aging and their selectivity and permeability still exist, slowing down the growth of large-scale industrial applications. Together with the investigation of new polymeric, inorganic and hybrid materials able to offer better properties than the materials currently available and to withstand more aggressive environments, the design and optimization of the membrane process will lead to significant innovation toward the large-scale diffusion of the membranes for GS and in this respect, the role of membrane engineering is crucial. To discuss the role of membrane GS for re-designing industrial applications, one of the roles of membrane engineering is the introduction of new metrics for comparing membrane performance with those of traditional operations. These metrics consist in easy and useful indexes able to give an immediate idea of the eventual gain offered by the membrane systems. Moreover, they also allow a better understanding of the application limits of membrane systems for obtaining a determined quality target of a process. These metrics do not replace the existing indicators, referring to other aspects of the production plants. Therefore, it is important to point out that the final evaluation of the ‘sustainability’ of processes must be always carried out by considering the new metrics together with the environmental, economic, and society indicators. The new metrics would represent, in fact, ‘something more to think about’ in performing the overall analysis of processes.4 In this chapter, some metrics specifically focused on GS application will be analyzed and some selection guidelines for specific application processes will be introduced. As a case study, the hydrogen and CO2 separation from refinery streams, will be considered and conventional separation operations, such as adsorption, absorption and cryogenic, will be compared in terms of the productivity/size ratio; mass intensity and energy intensity. As an example to show how to use these metrics for real cases, some experimental results obtained using polyimide membranes will be analyzed from these different points of view. Moreover, some further considerations related to the exergetic aspects to take into account in the process design phase will be highlighted considering the ethylene production cycle as a case study.

19.2 Current Applications of Membranes in Gas Separation In 1950, Weller and Steiner5 proposed the application of membrane processes as feasible solution in the separation of hydrogen from hydrogenation tail gas,

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

New Metrics in Membrane Gas Separation

281

enrichment of refinery gas, and air separation. However, only in 1980 Permea launched the first large industrial application of gas separation membranes6,7 proposing ‘prism membranes’ suitable for hydrogen separation. Since that time many companies such as Cynara-Natco, Separex-UOP, GMS, Generon, Praxair, Air Products, and UBE have become involved in this field8,9 introducing membrane-based operations as substitution for or integration of traditional systems. Table 19.1 lists the main industrial applications in the field of gas separation (GS) where membrane technology is involved, describing also the membrane materials used and the status of membrane technology.10

19.2.1

Case Study: Hydrogen Recovery

Hydrogen recovery is applicable to several processes, divided into three main categories:  Hydrogen recovery from ammonia purge streams  Syngas ratio adjustment  Hydrogen recovery in refineries. The first widespread commercial application of membranes in GS was, as already said, the separation of the hydrogen in the ammonia purge stream, by Permea PrismTM systems. In the ammonia process, the purge stream, almost clean and free of condensable vapors, consists of a mixture of hydrogen, nitrogen, methane and argon, delivered at a high pressure (130–140 bar). It is, thus, an ideal application for membrane technology, since hydrogen is highly permeable with respect to the other gases and the stream already provides the necessary driving force for promoting the permeation. Actually, Permea, Inc. (now owned by Air Products and Chemicals, Inc.) designed a two-step membrane process for this separation which recycles 90% pure H2. Similar membrane systems are already applied also for syngas ratio adjustment (H2/CO ratio) of the streams exiting reformers. Generally, membranes are employed for stripping hydrogen out of the syngas in order to reduce the H2/CO ratio. At the moment, several hundred hydrogen separation plants have been installed.11 The demand for hydrogen recovery in refineries is rapidly increasing also owing to environmental regulations. The hydrogen content in the various refinery purges and off-gases ranges between 30 and 80%, mixed with light hydrocarbons (C1–C5); 90–95% hydrogen purity is required to recycle it into a process unit. A typical refinery operation is the separation of the hydrogen contained in the stream emerging from hydrocrackers. Actually, this process has the disadvantage of removing 4 moles of hydrogen for each mole of hydrocarbon removed. The membranes can be used alone or together with an absorber system, at a reduced capital cost and better process efficiency. At the moment, the PrismTM system (using polysulfone hollow fibers under a thin silicone film) dominates the market for this kind of separation, showing

Process

Ammonia purge gas

Adjustment of H2/CO ratio in syngas plant

Hydrogen recovery in refineries

Ethylene cracker old trains

Air separation

H2/N2

H2/CO

H2/hydrocarbons

H2/light hydrocarbons

O2/N2

Cryogenic distillation

Cryogenic distillation

PSA

PSA

PSA

Traditional technology

Plant installed (Prism by Permea) Lab scale Lab scale Plant installed (Separex) Lab scale Plant installed (Sinopec Zhenhai (China) Prism by Permea Du Pont) Lab scale Pilot plant Lab scale Plant installed (Cynara; Separex; GMS; Air Products) Plant installed (Permea) Plant installed (Medal; DowGeneran; UBE) Plant installed (Aquilo) Plant installed (Air Liquide) Lab scale Lab scale

Status of membrane technology application

282

Polyphenilene Ethyl Cellulose Ion transport (perovskite) Pd-based

Polysulfone Polyimide

Pd-based PTMSP; PMP Pd-based Silicon rubber

Polysulfone Pd-based Silicon rubber Polyimide Pd-based Silicon rubber Polyimide

Membrane material

Main industrial applications of membrane technology for GS

Separation

Table 19.1

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

View Online

Chapter 19

Nitrogen removal

Dehydration

Sweetening of natural gas

Natural gas liquid removal

CO2 capture from flue gas streams

Polyolefin plant resin degassing; Ethylene recovery

N2/CH4

H2O/CH4

CO2/CH4

CO2/ Hydrocarbons

CO2/N2

VOCs/gas Adsorption, Refrigeration and turbo-expander plants

Glycol absorption and cooling inn a propane refrigeration plant (–20 1C) Amine absorption

Amine absorption

Glycol absorption

Cryogenic distillation

Polyimide FSC; PEEKWC; Zeolite, silica based; carbon Silicone rubber; PTMSP

Polyaramide Polyimide Perfluoro-polymers Cellulose acetate; Polyimide polyaramide

Silicon rubber PMP Parel PEBAX Cellulose acetate; Polyimide Polyaramide Cellulose acetate

Plant installed (MTR; OPW VaporsaverTM)

Pilot plant Lab scale

Offshore platforms In Thailand gulf. (Cynara-NATCO); Grace-Separex Plant installed (Medal) Plant installed (Medal) Plant installed (MTR) Early commercial stage. A demonstration system installed

Lab scale Plant installed

Pilot plant

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

View Online

New Metrics in Membrane Gas Separation 283

View Online

284

Chapter 19

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

12

interesting selectivities. Another source of hydrogen from the refinery is the ethylene steam cracking. Picciotti and Kaiser13 estimated that in 2002 ethylene production via steam reforming was more than 100 million tonnes and this production has increased significantly in the last few years. As proposed by Bernardo and co-workers,14,15 the huge amount of H2 produced in the ethylene cycle can be easily upgraded by using membrane separation units as well as membrane reactors. These membrane processes compared with the traditional ones can offer some potential benefits such as lower plant size, lower energy consumption, improved safety and more rational utilization of raw materials.

19.3 Comparison of Membrane Gas Separation and the Other Separation Technologies: Engineering Evaluation 19.3.1

Technologies for Gas Separation

The separation of a gas from the rest of the gaseous stream is traditionally carried out by means of pressure swing adsorption (PSA), cryogenic separation and absorption processes. However, the application of membrane systems is rapidly evolving on the commercial level owing to the advantages related to low capital costs, low energy requirements and modularity. In the following, after a brief description of the conventional separation processes, a comparison among these processes and membrane systems is reported, also introducing some design considerations such as process flexibility, reliability, ease of response to the variations, expansion capability and versatility.16

19.3.1.1

Pressure Swing Adsorption

Pressure swing adsorption is a batch operation that uses multiple vessels to produce a constant product and off-gas flows. The PSA process is based on the capacity of some adsorbents (zeolites, molecular sieve, etc.) to adsorb such gases at high gas-phase partial pressure. Proper selection of the adsorbents is critical to both the performance of the unit and adsorbent life. In hydrogen separation, for example, the hydrogen is adsorbed at higher partial pressure and then desorbed at lower partial pressure. The hydrogen partial pressure is lowered by ‘swinging’ the adsorber pressure from the feed pressure to the tail gas pressure, and by using a high-purity hydrogen purge. In this case the purge is used as a ‘sweep gas’, to carry away the contaminants; however, about 8–20% of the total hydrogen throughput16 is lost in the phase. Multiple adsorbers are used in order to provide constant feed, product and tail gas flows and each adsorber undergoes the same process steps in the same sequence. A minimum pressure ratio of approximately 4:1 between the feed and tail gas pressure is usually required for separations. However, the absolute pressures of the feed and tail gas are also important, particularly to gas recovery.

View Online

New Metrics in Membrane Gas Separation

285

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

The optimum feed pressure range for PSA units in refinery applications is 15–30 bar. The optimum tail gas pressure is as low as possible. Since vacuum is normally avoided, tail gas pressure between 1.1 and 1.4 bar are typically used when high recovery is needed.

19.3.1.2

The Cryogenic Process

The cryogenic process is a low temperature separation process which uses the difference in boiling temperatures (relative volatilities) of the feed components to effect the separation. This process condenses the required amount of feed impurities by cooling the feed stream in multipass heat exchangers. The refrigeration required for the process is obtained by Joule–Thomson refrigeration derived from throttling the condensed liquid hydrocarbons. Additional refrigeration, if required, can be obtained by external refrigeration packages or by turbo-expansion of the hydrogen product. One of the main advantages of the cryogenic process is the ability to produce separated hydrocarbon streams rich in C41, ethane/propane, etc. One of the main drawbacks is the presence of H2O in the gaseous stream can strongly affect the separation performance; therefore the cryogenic separation relies on the assumption that the gaseous stream to be treated is completely dehydrated prior to be cooled.

19.3.1.3

Absorption

Absorption is a process that relies on a solvent’s chemical affinity with a solute to dissolve preferably one species into another. It is widely proposed for CO2 separation where a solvent, generally, monoethanolamine (MEA) or a solid absorbent like lithium zirconate is used to dissolve CO2, but not the other components of a flue gas stream.17 CO2-rich solution is typically pumped to a regeneration column, where CO2 is stripped out from the solution and the solvent recycled for a new batch of flue gas. The absorption equipment should be placed after the flue gas desulfurization-step and before the stack. Optimal conditions for absorption are low temperature and high pressure, making this the best location for absorption to occur. In addition, most solvents are easily degraded by compounds such as fly ash, other particulates, SOx (SO2, SO3) and NOx (NO, NO2), so the absorption step must take place after electrostatic precipitation and desulfurization. In a typical absorption process, the CO2-lean flue gas is either emitted to the atmosphere or possibly used in other applications (e.g. chemical production).

19.3.1.4

Membrane Technology

Membrane technology is most often listed as potential candidate for its application in new plants for GS. The greatest asset to membrane separation is simplicity. Whereas PSA requires the equipment for swinging pressure,

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

286

Chapter 19

cryogenic distillation must endure extreme temperatures and absorption requires huge amount of sorbent, the only equipment necessary for GS is the membrane and fans. There are almost no moving parts, and the construction is fairly simple. The gaseous stream to be separated generally requires a compression, but this is much smaller than that necessary for PSA. Actually, the major drawback of the membrane operation concerns just the low selectivity of some membrane materials and a limitation on the suitable operating conditions. A membrane to be useful for H2 separation from off-gas streams as well as for CO2 capture, should posses a number of properties, such as high permeability,18 high selectivity, thermally and chemically resistance, plasticization resistance, aging resistance, cost effective, ability to be cheaply manufactured into membrane modules.

19.3.2

Selection Guidelines for Gas Separation

The choice of the technology suitable for the specific separation is related to different parameters like economics, stream conditions, product target and, also, to design considerations such as those reported below. These design parameters were introduced by Miller and Sto¨cker16 for H2 separation technologies; however, they can be considered valid, in general, for gas separation.

19.3.2.1

Operating Flexibility

Operating flexibility is the ability to operate under variable feed quality conditions, either on a short-term or long-term basis. The changes in feed composition occur very often for example in refinery applications, particularly when the source of the feed is a catalytic process or when the feedstock to the upstream unit changes. In membrane processes, the increase of the concentrations of feed impurity implies a decrease in product purity, which, however, can be maintained for small feed composition changes by adjusting the feed-to-permeate pressure ratio. Depending on the type of separation to be carried out, for example as with the methane in most refinery membrane applications, the presence of an impurity can be allowed to increase slightly in the product without major downstream impact. The response time of membrane systems is essentially instantaneous, and a corrective action has immediate results. The start-up time required by the process is extremely short. The PSA process shows a great ability to maintain target purity and recovery under changing conditions, requiring a simple cycle time adjustment. The process is self-compensating and even relatively large changes in feed impurity concentrations have little impact on performance. As the concentration of a feed impurity increases, its partial pressure increases, increasing also the amount of the impurity which will be adsorbed. The response time to variations is rapid but not abrupt, generally requiring 5–15 min after a step change in feed quality. The new steady-state upon restart following shut-down is reached in

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

New Metrics in Membrane Gas Separation

287

about 1 h. The cryogenic process has very little flexibility, because changes in the concentration of the lower boiling components of the feed directly affect the product purity. Recovery is not strongly affected. Response time is not as rapid as for PSA or membrane systems. Start-up lasts 8–24 h depending on the used procedure. The flexibility of the absorption process is moderate. This system, in fact, responds quite well to the changes in the feed composition; however, significant differences in the target gas fraction contained in the feed imply higher solvent flow rate that is strictly related to the equipment size apart from the absorber. Furthermore, the feed stream contains oxidizing compounds (oxygen or sulfur dioxide) which induces major problems of amine deactivation and promotes corrosion.19

19.3.2.2

Turndown

Turndown is the capability of a system to operate at reduced capacity. Membrane systems are highly capable of maintaining product purity even though the capacity is reduced down to 10% of the initial design value. This can be done by either reducing feed pressure, increasing permeate pressure, or by system modularity allowing a set number of modules to be isolated. The first two methods can be used for short-term operations, and the latter when operating at significantly reduced capacity for extended periods. No provisions are required in the design phase. PSA units can maintain both recovery and product purity at throughputs of about 30–100% of design by adjustment of the cycle time. Between 0 and 30% of design product rates purity can be maintained, but recovery is reduced unless special provisions are made in the design. The turndown capability of cryogenic systems strongly depends on the design. Condensation units can maintain product purity at slightly reduced recovery at flows down to 30–50% of design. Absorption units can maintain both recovery and product purity at throughputs of about 30–100% of design by adjustment of gas and solvent flow rates.

19.3.2.3

Reliability

Reliability takes into account the on-stream factors which can cause unscheduled shut-downs. Membrane systems are extremely reliable with respect to the on-stream factor. The membrane separation process is continuous and few control components can cause a shut-down. Typically, the response to unscheduled shut-downs is rapid for GS whereas PSA systems are moderately reliable owing to the numerous valves associated with the process which can cause unexpected shut-downs. The new PSA are designed with alternative modes of operation, in which 100% of design capacity can be achieved while by-passing any failed valve or instrument, with only a slight recovery loss. Failures are automatically detected and by-passed by the microprocessor-based control system. However, stronger and periodic control cycles are required. The cryogenic process is considered by refiners to be less reliable than PSA or membrane

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

288

Chapter 19

processes, this is mainly due not to the process itself but to the need of feed pre-treatments. Failure of the pre-treatment system usually results in contaminants freezing in the cold box, leading to shut-down. In this case, a ‘thaw’ is necessary prior to restarting. The pre-treatment system itself is often more complex than the cryogenic system. Absorption systems are moderately reliable. The large equipment associated with the process for avoiding the formation of degradation products and corrosion products and the presence of particles can cause unexpected shut-downs.

19.3.2.4

Ease of Expansion

Ease of expansion takes into account the possibility of and expanding a system and the ease of doing so, out of the design phase. Membrane systems expansion is very easy, since this only requires the addition of identical modules. PSA systems can also be expanded, but this requires additional design considerations and adds cost in the initial phase of the project. Cryogenic units cannot be expanded unless foreseen during the design phase. Generally, they can be over-sized and a capacity increase is often obtained without modification to the cold box itself through addition of a tail gas compressor. Absorption systems can also be expanded, but this requires additional design considerations and adds cost in the initial phase of the project.

19.3.2.5

By-product Value

By-product value is the capability of producing a high-value tail gas. For instance, in many refinery hydrogen upgrading processes, the impurities to be rejected include hydrocarbons which have a value significantly above that of the fuel. This is particularly true for olefin-containing streams. The relative amounts of high-value hydrocarbons and the incremental cost of further separation determine whether by-product recovery is an important parameter to be considered. The cryogenic process is best suited for applications involving by-product hydrocarbon recovery. It is possible to separate hydrocarbon streams containing C2, C3 and C41 components, with recovery of these species generally quite high, with typical overall values of 90% C2 and C3 and up to 100% C41. The membrane process is not capable of providing separate streams rich in specific hydrocarbon fractions; however, it is also used in these applications, since the hydrocarbon-rich stream is available at high pressure. By-product hydrocarbon recovery is normally not economical when the PSA process is used, because the single hydrocarbon-rich tail gas is delivered at too low pressure. The same happens with the absorption process. Combinations of separation processes, such as membrane and cryogenic processes may be applicable to these applications. Table 19.2 summarizes the comparison among the design parameters described above for the four separation processes considered.

View Online

289

New Metrics in Membrane Gas Separation

Table 19.2

Comparison among some important design parameters Membrane system

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

Ease of expansion

Very high (modularity) By-product value Moderate Operating flexibility Moderate Response to Instantaneous variations Start-up after the Extremely short variations (10 min) Turndown Down to 10% Reliability 100% Control requirement Low

19.3.2.6

PSA

Cryogenic Absorption

Moderate

Very low Moderate

Moderate High High Low Rapid (5–15 min) Slow

Moderate Moderate Rapid (5–15 min)

1h

8–24 h

1h

Down to 30%

Down to Down to 30% 30–50% Limited Moderate High High

95% High

Selection of the Separation Process

The selection of the suitable separation process should be driven by specific considerations, strictly related to the output to be obtained. The feed composition and operating conditions (pressure and temperature), the product purity and the final destination of the product strongly affect the choice. 19.3.2.6.1 Composition of the Feed. The composition of the feed and its variability has a large impact on the selection of the separation process because it influences the performance, reliability and pre-treatment required. Membrane systems are suitable for a wide range of feed compositions, owing to the possibility of driving the process acting on different parameters such as feed and permeate pressure, temperature, flow configuration, etc. In the case of hydrogen separation, streams with 75–90%molar hydrogen are most economically upgraded also by PSA with the selection being based on flow, pressure, and pre-treatment requirements. Cryogenic upgrading is applicable to large streams with 30–75%molar hydrogen. 19.3.2.6.2 Pressure and Flow Rates. Feed pressure and product flow rates are best considered together when selecting a hydrogen purification process because the three processes have drastically different economies of scale. The membrane systems have the lowest capital cost alternative for small (less than 30 000 m3(STP) h1 product) flow rates, since it is proportional to the number of modules required. If the feed gas is already at high pressure or if there is downstream recovery from the non-permeate, where the high pressure of this stream can be used as an advantage, the membrane systems can be used also for larger flow rates. For small flow rates at high pressure, such as a purge stream from the high pressure separator of a hydroprocessor, membrane systems are the most economical.

View Online

290

Chapter 19

PSA systems can be economical for flow rates from 10 –10 m (STP) h1 already compressed at 15–30 bar. Differently from membrane systems, PSA units cannot take advantage of high available feed pressures (50–70 bar) and the capital and operating costs associated with feed, product, and/or tail gas compression are almost always a significant portion of the total separation costs and, globally, higher than that required by membrane systems. Cryogenic systems have high capital costs at low product rates, but good economies of scale. In such applications as CO2 separation from flue gas streams the absorption is the direct competitor of membrane systems. It allows the same flow rates of membranes system to be treated and it is ideal when the stream to be separated is not compressed and the concentration of the gas to be recovered is not so high. This is the main advantage with respect to the membrane systems that require a pressure drop between the two membrane sides for performing the separation. The main drawback of this technology is connected to the significant amount of solvent required.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

3

5

3

19.3.2.6.3 Product Purity. Product purity and the levels of specific product impurities are, of course, critical to process selection. In the hydrogen separation process, cryogenic and membrane processes normally produce hydrogen at 90–98%molar, whereas the PSA process normally produces hydrogen at 99 þ %molar. However, depending on the final utilization of the upgraded hydrogen, it can be more conveniently applied to other operations. If the upgraded hydrogen is used as a primary source of make-up hydrogen to a high pressure hydrotreater or hydrocracker, high purity is required and the PSA is often the most used. It is only an incremental portion of the make-up hydrogen to a hydroprocessor, lower product purity is usually required and membrane systems are preferable, owing to the lower capital costs. If the feedstock to be upgraded contains only hydrogen and hydrocarbons, then the principal impurity in the hydrogen product will be methane. When 3–10%molar methane can be tolerated in the hydrogen product, the membrane and cryogenic systems prove to be the most economical. If the feed of the upgrading process contains substantial quantities of CO and CO2, then PSA and the membrane units are most often selected, whereas cryogenic is not suitable. The absorption is not indicated in hydrogen separation, but in CO2 separation is able to obtain a high purity of the stream. Table 19.3 and Table 19.3 Comparison of the three systems utilized in hydrogen separation as a function of feed and product conditions Feed composition, (H2 %molar) Feed pressure and product flow rate Product purity (%molar)

Membrane system

PSA

Cryogenic

30–90

75–90

30–75

410 bar o30 000 m3(STP) h1 90–98

15–30 bar 1000–105 m3(STP) h1 499

420 bar 4105 m3(STP) h1 90–98

View Online

291

New Metrics in Membrane Gas Separation

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

Table 19.4

Comparison of the three systems utilized in CO2 separation from flue gas streams as a function of feed and product conditions

Feed composition, (CO2%molar) CO2 purity (%molar) CO2 recovery

Membrane System

Absorption

Cryogenic

415 80–95(*) 60–80%(*)

45 495 80–95%

45 99.99 99.99%

*Considering a CO2/N2 selectivity of 50 in a one stage process.

Table 19.4 report a comparison among the separation units utilized for hydrogen separation and CO2 capture from flue gas streams respectively as a function of the feed and product conditions. For the process choice, the selection guidelines described above can be helpful, at least in eliminating any inappropriate process. In order to use such guidelines, feed characteristics, contaminant levels, required product purity, allowable impurities, pressure levels and flow rates must be known. These parameters can be used together with experience-based, application-specific guidelines to select the optimum process.

19.3.3

Case Study: Selection Guidelines for the Separation and Recovery of Hydrogen in Refineries

In the choice of the most suitable technology among the three separation processes the specific process from which the hydrogen is produced should be taken into account, also considering the specific requirements that the upgraded H2 must have and its final use as well. Table 19.5 reports some selection guidelines of the most important processes of hydrogen production in refineries.

19.3.3.1

Off-gas from Catalytic Reforming

The largest source of easily recovered hydrogen in a refinery is the off-gas from catalytic reforming that typically contains from 70–90þ %molar hydrogen. Traditionally, the PSA is used for upgrading the large quantities of reformer off-gas to use as a primary source of make-up hydrogen for a hydrocracker or hydrotreater. The high hydrogen content makes the adsorbent requirements low, and the hydrocarbons in the tail gas sent to the fuel is relatively small with respect to the feed gas. The aromatic and HCl content of the feed do not require special pretreatment. For some purposes, such as catalyst regeneration, relatively small quantities of hydrogen at 98þ%molar purity are required. In this case, membrane systems are most economical for these applications requiring less than about 1000 m3(STP) h1 of hydrogen product, owing to low capital costs. For large flows, cryogenic upgrading has sometimes been used. However, the feed hydrogen purity is often too high, and external refrigeration is required. A further drawback is the requirement of a pretreatment section for the aromatic removal. Unless recovered hydrocarbons are of high value, the cryogenic process is not normally used to process reformer off-gas alone.

15–50%vol H2 Hydrocarbon or olefins balance Pressure: 5–20 bar Room Temperature

80–90%vol H2 CO, CO2, CH4 balance Pressure: 15–30 bar

FCC off-gas and other refinery purge streams

Ethylene off-gas

Feed compression: 425 bar H2 purity: 80–90% H2 pressure: 5–20 bar H2 recovery: 50–85% Final destination: low pressure hydrotreaters The tail gas is sent to downstream hydrogen recovery units H2 purity: 95–97% H2 recovery: 80–90% Final destination: make-up hydrogen compressor

H2 purity: 95% H2 pressure: 15–30 bar Final destination: make-up hydrogen compressor

H2 purity: 495% H2 pressure: 15–30 bar Recovery of a mixed stream containing 499 þ % of C2 þ components

292

H2 purity: 99.5% H2 recovery: 80–90% Final destination: primary source of make up hydrogen for hydrocracker

H2 purity: 99% H2 pressure: 1 bar H2 recovery: 80–90% Final destination: primary source of make up hydrogen for hydrocracker Not used because of the low H2 concentration in the feed

Not used due to the low flow rates

High pressure purge gas: 75–90%vol H2 Hydrocarbon balance Pressure: 50–200 bar Room Temperature Low pressure purge gas: 50–75%vol H2 Hydrocarbon balance Pressure: 5–20 bar Room Temperature

Hydroprocessor purge gases H2 purity: 92–98% H2 pressure: 20–40 bar H2 recovery: 85–95% Final destination: make-up hydrogen compressor Not used because of the low feed pressure

Cryogenic Not used due to the necessity of pretreatment for aromatic removals

PSA H2 purity: 99 þ % H2 recovery: 82–90% Final destination: primary source of make up hydrogen for hydrocracker Not used because of the high feed pressure

Membrane system H2 purity: 98%. Final destination: uses such as catalyst regeneration

70–90%vol H2 30–10% C1 to C6 þ ppv of aromatics and HCl Pressure: 15–30 bar Room Temperature

Catalytic reformer off-gas

Selection guidelines for specific application processes in hydrogen production11

Process application Feed conditions

Table 19.5

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

View Online

Chapter 19

View Online

New Metrics in Membrane Gas Separation

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

19.3.3.2

293

Gases from Hydrocrackers and Hydrotreaters

The high pressure and low pressure purge gases from high pressure hydrocrackers and hydrotreaters are good candidates for hydrogen upgrading. The recovered hydrogen is returned to the hydroprocessor with the make-up hydrogen. Many hydroprocessors have both high and low pressure purge streams. The high pressure purge streams are available at 50–200 bar and contain 75–90%molar hydrogen, with the balance being hydrocarbons. Low pressure purge streams are available at much lower pressures, typically 5–20 bar and have hydrogen contents ranging from 50 to 75%molar. The membrane process is the most economical process for high pressure purge gas upgrading, whereas low pressure purge gases are usually upgraded by the PSA process that is better suited than the cryogenic process due to the relatively small flow rates and the highly variable stream composition. However, if the stream containing hydrogen has lower pressure and lower hydrogen content, the membrane systems are more economical than PSA, since they give the highest rate of return on investment, the tail gas being at a high pressure.

19.3.3.3

Hydrogen from Fluid Catalytic Cracking

Hydrogen can be recovered from fluid catalytic cracking off-gas and other low pressure refinery purge streams of low hydrogen content mainly using either the cryogenic or membrane process, depending on flow rate, feed composition and variability, feed pressure and required hydrogen product purity. In almost all cases, the stream is not upgraded for its hydrogen content alone, but also the tail gas is of value. The cryogenic is preferred when there are valuable hydrocarbons, particularly olefins, which can be recovered in addition to hydrogen, or when hydrogen product purity over 90%molar is required. However, feed quality variations and contaminant levels are important considerations in determining whether the cryogenic process is appropriate. The membrane process can recover hydrogen efficiently from these streams at a hydrogen purity 80–90%molar. The low purity hydrogen product can be used effectively in some applications, such as low pressure hydrotreaters, and the tail gas can be sometimes sent to downstream hydrocarbon recovery units, being already compressed.

19.3.3.4

Production of Ethylene

The ethylene production process produces large amounts of hydrogen which can be easily upgraded for refinery use if the ethylene plant is in close proximity to the refinery. Traditionally, the ethylene process uses a series of cryogenic units to separate the products from the ethylene furnace, and a high purity hydrogen stream is produced. Only a small fraction of the available hydrogen is used in the ethylene plant (for acetylene hydrogenation) and the balance is sent to the fuel unless it can be exported. The hydrogen-rich ethylene off-gas normally contains 80–90%molar hydrogen, with CO, CH4, ethylene and nitrogen as impurities. In few cases, the cryogenic system is employed producing hydrogen purity as high

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

294

Chapter 19

as 95%molar with the same impurities through use of external refrigeration. Regardless of the hydrogen purity, the ethylene off-gas must be processed to remove CO down to ppm levels for general refinery use. For off-gas with 80–90%molar hydrogen, the membrane systems are the most suitable for upgrading. The pressure and the high H2 concentration, in fact, allow a high purity product to be produced, even though lower (95–97%molar) than for the PSA system. Furthermore, the retentate gas stream is already compressed, on the contrary to the PSA systems that require tail gas compression, from 1.1 to 1.3 bar to fuel system pressure. The choice between the membrane and PSA processes will depend primarily upon the compression cost of the hydrogen produced by membrane compared to the cost of compressing the PSA tail gas, assuming high hydrogen recovery is desired in both cases. The product purity from the membrane system will be lower (95–97%molar) than for the PSA system.

19.4 New Metrics for Gas Separation To measure progress towards sustainability, many efforts are being made to define indicators of the industrial process and, in particular their impact on three specific areas.20,21 Thus the following indicator were defined:  Environmental indicators, which refer to the resource exploitation and the emission, effluents and waste related to the production  Economic indicators, which take into account the investments profits, values and taxes  Social indicators, which consider the employment situation, the health and safety at work, the benefits/troubles for the community. Most of these indicators are calculated in the form of appropriate ratios which can provide a measure of impact independent of the scale of the operation, or to weigh costs against benefits and, in some cases, they can allow the comparison between different operations.22 As indicated by Sikdar,21 to guarantee the sustainability of a process these three areas (environmental, economic and societal) should be inter-related and the metrics or indicators can be used for capturing the three aspects of sustainability. Two classes of metrics are under development to indicate the state and performance of a system. These metrics are more popularly known as indexes. Those indicating a system state are known as content indexes and those measuring the system behavior, performance indexes. The following discussion is focused on the performance ones only, as we are mostly focused on the means of improving the sustainability characteristics of a membrane system. Membrane operations are well known for their modularity, compactness and flexibility, therefore they can be considered as new operations developed in the logic of process intensification. Recently, new metrics for comparing membrane performances with those of conventional units have been introduced.20,23 These new metrics take into account the size, weight, flexibility, yield and modularity

View Online

295

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

New Metrics in Membrane Gas Separation

of the plants. They are useful for having an immediate indication of the eventual gain that a membrane operation can offer with respect to a traditional one. In this sense, they are useful also for selecting the proper separation technology for a specific process. In order to compare different systems, energy and mass intensity10,24,25 were calculated as reported below, taking into account the power required by the system (eqn (19.1)), the H2 produced and its fraction recovered in the permeate (eqn (19.2)) as valuable product with respect to the total mass fed and the steam and cooling water necessary (eqn (19.3)). Moreover, another index was defined as productivity to footprint ratio (eqn (19.4)). Lower values of these indicators are related to an intensified process. In the ideal situation the mass intensity indicators would approach 1, whereas the energy intensity and the productivity to footprint ratio should be as low as possible: Energy Intensity ¼

Power required H2 product flowrate

ð19:1Þ

Mass Intensity 1 ¼

Total inlet mass H2 mass in permeate

ð19:2Þ

Steam and cooling water required H2 product flow rate

ð19:3Þ

Mass Intensity 2 ¼

H2 product flow rate Productivity= Footprint ¼ Footprint

ð19:4Þ

Already in 1989, Spillman7 reported a first comparison of three different separation technologies in H2 recovery from refinery off-gas, by considering polyimide membranes (Table 19.6). This comparison is here provided by the Table 19.6

H2 recovery from refinery off-gas. (The data reported were elaborated11 from results presented by Spillman8) Membrane Membrane system (80 1C) system (120 1C) PSA Cryogenic

H2 recovery, % H2 purity, % Product flow rate, m3(STP) h1 Power, kW Steam, kg h1 Cooling water, tonne h1 Investment, $ millions Installation area, m2

87 97 3257 220 230 38 1.12 8

91 96 3375 220 400 38 0.91 5

73 98 2643 370 – 64 2.03 60

Metrics Energy intensity, kJ kgH21 Mass intensity_2, kg kgH21 Productivity/footprint, kgH2 (h1 m2)

2808 136 35

2738 133 58

5760 4853 277 342 3.9 2.4

90 96 3375 390 60 99 2.66 120

View Online

296

Chapter 19

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

10

calculation of the indexes. A lower investment cost than PSA or cryogenic separation was estimated for H2 recovery from refinery off gas by polymeric membranes. This comparison is dated 1989, before the decrease in the cost of polymeric membranes. As can be seen, the energy intensity of the membranes system, for both the temperatures considered, is lower than that of the other two processes. This means that the energy required for producing a fixed amount of hydrogen is, in any case, significantly lower than that required by the PSA or cryogenic distillation. Also the addition of the steam and cooling water is halved with respect to that necessary in traditional systems. Moreover, considering the same area occupied by the unit, the membrane systems have more than ten times higher productivity.

19.4.1

Case study: H2 Separation from H2/N2 and H2/CO Mixtures with co-polyimide Hollow Fiber Modules

In the following section, H2 separation from binary mixtures with a hollow fibers membrane module will be considered as a case study for showing an example of possible comparison between the different separation technologies previously listed. In particular, the experimental results obtained by testing a homemade membrane module containing 165 hollow fibers of P84s co-polyimide with H2/N2 and H2/CO mixtures26 will be used as the basis for the comparison between the membrane system and the other separation technologies. The comparison has been made considering PSA and cryogenic systems. The absorption has not been considered because it does not have a wide application in hydrogen separation from refinery streams; therefore no results are available in the literature for comparison. Four different case studies related to the membrane module have been considered in this analysis. The first two refer to results achieved in the experimental study by Choi et al.26 where more details are reported. The last two contain some results such as H2 recovery and stage cut evaluated from these experimental data, but considering a higher feed pressure and vacuum on the permeate side. Table 19.7 lists the operating conditions and membrane properties related to the chosen case studies. Table 19.7

Membrane properties (measured) and corresponding operating conditions (experimentally used: cases 1 and 2; simulated cases 3 and 4) of the case studies chosen for the calculations 3

2

1

H2 Permeance, dm (STP) m h bar Separation factor, Stage cut, % Feed pressure, bar Permeate pressure, bar Temperature, 1C Feed molar composition, molar%

1

Case 1

Case 2

Case 3

Case 4

9.94 6 27

11.8 20 9

9.94 6 35

11.8 20 20

6 1 50 H2:N2=30%:70%

10 0.2

View Online

297

New Metrics in Membrane Gas Separation

Table 19.8

Comparison of membrane operation with traditional separation units. The data reported for PSA and cryogenic refers to8

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

Membrane system Case 1

Case 2

Case 3

Case 4

PSA

Cryogenic

Product flow rate, m h H2 recovery, molar% H2 purity, molar% Installation area, m2

3000 64 72 55

3000 26 90 56

3000 85(a) 72(b) 21

3000 50(a) 90(b) 18

2643 73 98 60

3375 90 96 120

Metrics Mass Intensity_1, kg kgH21 Productivity/footprint, kg H2 m2 h1

52.6 4.8

129.5 4.8

39.6 12.8

67.3 15

46.1 3.9

37.4 2.4

3

1

(a) Value linearly extrapolated considering the correlation reported by Choi et al.26 for a stage cut of 35% and 20%, respectively. (b) H2 purity value assumed to be the same as cases 1 and 2, respectively.

Table 19.8 highlights the comparison among the performances of the membrane gas operation, considering the four different cases studies, with respect to another two traditional separation units. First, comparing the case studies related to the membrane operation. Case 1 (experimental results) considers a higher H2 recovery and a lower H2 purity. As a consequence the installation area required is lower than Case 2 (experimental results), where a lower recovery but a higher H2 purity are obtained. The mass intensity value obtained in both cases is significantly far from 1, which is the ideal value. The mass intensity of case 1 is lower than that of case 2, owing to the lower hydrogen recovery. On the contrary, the productivity footprint are very close to each other, basically owing to the similar value of installation area required. For cases 3 and 4, the same membrane properties and H2 purity values as those measured and used in case 1 and case 2, respectively, were assumed, but considering a higher feed pressure (10 bar) and also vacuum in the permeate side (0.2 bar). The corresponding values of H2 recovery and stage cut were thus extrapolated considering a linear correlation reported in the referred paper. In both cases, the direct consequence of the higher driving force was a significant reduction of the installation area with respect to cases 1 and 2, respectively. As a consequence, a higher productivity footprint was obtained as well as a lower mass intensity. Globally, the process becomes more intensified in proportion to the higher driving force. However, to confirm this idea also the power required for this operation with a higher driving force might be taken into account. The performances of the membrane system (all four cases) compared with the PSA and cryogenic systems show interesting results. In particular, the installation area required by the membrane systems is always lower than the other two separation operations, and it is much more evident considering cases 3 and 4. This is reflected in a significantly higher productivity footprint ratio.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

298

Chapter 19

However, the H2 recovery and purity obtained with the membrane operation are lower or comparable with those of the other two systems; therefore the mass intensity is higher. It must be noticed that the membrane systems considered in these calculations operate at 50 1C. However, at a higher temperature, the membrane permeance increases; therefore higher recovery can be expected. Moreover, the results presented for the PSA and cryogenic systems consider the treatment of a refinery off-gas, where the amount of hydrogen in the feed16 is 50–75% higher than that used in this work (30%).

19.5 Further Evaluations in Membrane Process Design: The Exergetic Aspects An important parameter to take into account in the designing phase of an integrated membrane process is the exergy that quantifies the energy quality. As is known, an energetic analysis, which is based on the first thermodynamic law, does not consider the energy quality. The exergy can be defined as the maximum work obtainable by the system evolution, through reversible transformation from an initial to a final state in equilibrium with the environment27 and the exergetic analysis is one of the most important tools to assess the process efficiency. An interesting exergetic analysis in the ethylene cycle has been proposed by Bernardo and co-workers,14,15 replacing some traditional units with membrane operations, which are (Figure 19.1):  Membrane reactor with oxygen conducting membrane for ethane oxidative dehydrogenation  Single- and multi-stage membrane GS system for H2 recovery Air

Coke

MF for Water Treatment (coke removal)

OEA By Membrane Operation

Pure H2

Water

MR FOR CO CLEAN-UP GAS COMPRESSION

H2

to flares Feedstock

CRACKING FURNACES

Dilution Steam

HOT SECTION

MCs for Acid Gas Removal

SEPARATION

Condensed Acid gas stream

MRs FOR ETHYLENE PRODUCTION

MCs for Water Purification

C2+ Membrane GS ETHYLENE/ ETHANE

C2H4

C4+ CH4 C3+ Membrane GS PROPYLENE/ PROPANE

C3H6

Ethane/Propane recycle

Figure 19.1

Block diagram of a plant redesigned with the integration of some membrane unit operations. Reprinted from P. Bernardo, G. Barbieri, E. Drioli, An exergetic analysis of membrane unit operations integrated in the ethylene production cycle, Chem. Eng. Res. Des, 2006, 84, 405–411, with permission of Elsevier.14

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

New Metrics in Membrane Gas Separation

299

 Membrane contactors for HCl removal and recovery of the water used for producing the ‘dilution stream’ and for acid removal  Micro-filtration unit for the coke removal from the scrubbing water in the decoking phase  Membrane GS unit for oxygen-enriched air to be used for improving the combustion efficiency  Membrane GS units for ethylene/ethane and propylene/propane separation  Membrane reactor for CO clean-up. By means of considerations based on the exergetic calculations, the authors demonstrated how exergetic analysis can help in the process designing phase. They compared traditional and membrane operations, basically considering the H2 recovery stage placed after the traditional reactor where the oxidative dehydrogenation of ethane was carried out. The separation and recovery of the H2, required in the feed for improving the reactor performance,28 was conventionally performed by a cryogenic system. The H2 removal by means of membranes was investigated for compressed gases before the ‘cold train’ and resulted in a reduction of energetic duties related to the refrigeration units, owing to a very different value of the operating temperature and to the successive fractionations operated at lower pressure. The reduction in the compression loads resulted in a capital cost reduction, saving in piping, etc. Moreover, the comparison demonstrated a difference in exergy production rate equal to 760% when H2 was recovered by means of membrane system instead of the traditional H2/CH4 separation. This means that the quality of the energy use in the membrane unit is significantly better than that achievable with the traditional one and the membrane technologies apart from the equipment size reduction and better energy consumption also show reduced exergy losses. Moving with the same approach, in a second study15 the same authors considered ethane oxidative dehydrogenation as a second case study. They proposed to replace the traditional reactor for ethylene production with a membrane reactor containing a dense perovskite oxygen-conducting membrane, which allows only oxygen to permeate the membrane towards the reaction side. This type of solution, consisting of a controlled diffusion of the oxygen versus the reaction side, was proposed in the literature for reducing the over-oxidation of the valuable intermediate products (olefins) and, at the same time, for better controlling the heat management and the explosion risks. Furthermore, with this plant solution the H2 separation unit is no longer necessary, since there is no H2 in the feed. Comparing this improved system with the one proposed by Bernardo et al.,14 where the conventional reactor was followed by a GS unit, it was observed that the membrane reactor use implied an exergy loss 59% lower than the conventional system. In addition the membrane reactor showed better energy efficiency, 1.3-fold more than the traditional reactor þ GS system. These investigations demonstrated how much the exergetic evaluations are important in plant design since they provide further information for the correct estimation of pros and cons of a new operation with respect to those already applied.

View Online

300

Chapter 19

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

19.6 Concluding Remarks To discuss the role of membrane GS for re-designing industrial applications, one of the roles of the membrane engineering is the introduction of new metrics for comparing membrane performance with that of traditional operations. This chapter provides an analysis of some processes for hydrogen production/ separation referring to some new indexes introduced recently also by process intensification strategy. These metrics, consisting in easy and useful indexes able to give an immediate idea of the eventual gain offered by a membrane system, allow also a better understanding of the application limits of membrane systems to obtain a determined quality target of a process. Together with the new metrics an exergetic analysis is also presented as fundamental aspect to take into account in the design phase of an integrated membrane process. All the metrics and the exergetic calculation approaches reported above for GS applications can be easily extended to the other types of unitary operation as a useful tool for the evaluation of pros and cons during the design phase of a new plant where the membrane operation would replace the traditional ones.

Acknowledgement The Italian Ministry for Foreign affairs, Direzione generale per la promozione e la Cooperazione Culturale, Rome, Italy, is gratefully acknowledged for co-funding this research.

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

F. Dautzenberg and M. Mukherjee, Chem. Eng. Sci., 2001, 56, 251. A. Stankiewicz and J. A. Moulijn, Ind. Eng. Chem. Res., 2002, 41, 1919. J. C. Charpentier, Ing. Quim. (Madrid), 2006, 38(434), 16. A. Criscuoli and E. Drioli, Ind. Chem. Eng. Res., 2007, 46, 2268. S. M. Weller and W. A. Steiner, J. Appl. Phys., 1950, 21, 279. R. W. Baker, Ind. Eng. Chem. Res., 2002, 41, 1393. R. W. Spillman, Chem. Eng. Prog., 1989, 85, 41. E. Sanders, D. O. Clark, J. A. Jensvold, H. N. Beck, G. G. Lipscomb and F. L. Coan, U.S. Pat. 4,772,392. 1988. O. M. Ekiner, R. A. Hayes and P. Manos, U.S. Pat. 5,085,676. 1992. A. Brunetti, P. Bernardo, E. Drioli and G. Barbieri, in Membrane Gas Separation, ed. Y. Yampolskii and B. Freeman, Wiley and Sons, ch. 14, p. 281. N. W. Ockwig and T. M. Nenoff, Chem. Rev., 2007, 107, 4078. R. R. Zolandz and G. K. Fleming in Membrane Handbook, ed. W. S. W. Ho and K. K. Sirkar, Chapman and Hall, New York, 1002, p. 78. M. Picciotti and V. Kaiser, Hydrocarbon Process, 1979, 58(5), 99. P. Bernardo, G. Barbieri and E. Drioli, Chem. Eng. Res. Des., 2006, 84, 405. P. Bernardo, A. Criscuoli, G. Clarizia, G. Barbieri, E. Drioli, G. Fleres and M. Picciotti, Clean Tech. Environ. Policy, 2004, 6, 78.

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00279

New Metrics in Membrane Gas Separation

301

16. G. Q. Miller and J. Sto¨cker, Selection of a Hydrogen Separation Process, NPRA Annual Meeting, San Francisco, California, 1989. 17. D. Aaron and C. Tsouris, Sep. Sci. Technol., 2005, 40, 321. 18. C. E. Powell and G. G. Qiao, J. Membr. Sci., 2006, 279, 1. 19. A. Veawab, P. Tontiwachwuthikul and A. Chakma, Ind. Eng. Chem. Res., 1999, 38, 3917. 20. A. Criscuoli and E. Drioli, Ind. Chem. Eng. Res., 2007, 46, 2268. 21. S. K. Sikdar, AIChE J., 2003, 49, 1928. 22. IChemE. Sustainable Development Progress Metrics: Recommended for Use in the Process Industries, Institution of Chemical Engineers, Rugby, 2006, p. 1. Available: http://www.icheme.org/sustainability/metrics.pdf. 23. A. Brunetti, A. Caravella, G. Barbieri and E. Drioli, J. Membr. Sci., 2007, 306, 329. 24. A. D. Curzons, D. J. C. Constable, D. N. Mortimera and v. L. Cunningham, Green Chem., 2001, 3, 1. 25. A. A. Martins, T. M. Mata, C. A. V. Costa and S. K. Sikdar, Ind. Eng. Chem. Res., 2007, 46, 2962. 26. S.-H. Choi, A. Brunetti, E. Drioli and G. Barbieri, Sep. Sci. Technol., 2011, 46, 1. 27. R. Molinari, R. Gagliardi and E. Drioli, Des., 1995, 100, 125. 28. S. S. Bharadway, J. J. Maj, J. H. Siddall, M. D. Beardnen, C. B. Murchison and G. E. Lazaruk, U.S. Pat. 6,566,573. 2003.

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index Page numbers in italics refer to entries in Membrane Engineering for the Treatment of Gases, Volume 2.

absorption 285 amines 196, 211 vacuum swing 226, 230 acid gas removal 152–60, 216, 234–7 capital expenses 155 from natural gases 234–7 membrane system 235–7 operating expenses 155 AF 1600 91 AF 2400 91 aging of membranes see membrane aging air compression 226–7 purification see air separation systems air drying systems 230–2 applications 232 high pressure 232 membrane system 232 Air Liquide 85, 151, 228, 231, 233 Air Products 43, 44, 151, 169, 225, 228, 231, 233, 254 air separation systems 184–5, 226–30 nitrogen production 226–9 oxygen production 229–30 oxygen/nitrogen simultaneous production 230 alcohol reforming 57–60, 61

hydrogen yield 58, 62 results 59 steam 95 alkali metals, as functionalization additives 166 alkanes, oxi-dehydrogenation 207–9 alumina membranes 92 amines absorption 196, 211 washing 52 ammonia decomposition 62 dehydrogenation 93 synthesis loop 233 ammonium chloride 168 Amoco 157 amorphous cells gas transport properties 19–20 packings 17–18 realistic selection 18–19 amorphous fluoropolymer membranes 107–10 properties 109 aquifer CO2 storage 199–200 Aquillo 228 Archimedes number 32 argon, pipeline flow effects 198 Arrhenius behavior 9 Asahi Glass 231 Aspen Plus 197, 222 atomic layer deposition 47

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

atomistic simulation 8–16, 21 molecular dynamics 3, 9–12 Monte Carlo simulations 3, 12–13, 22 quantitative structure property relationships 14–16 transition state theory 3, 12, 13–14, 21 autothermal reforming 42, 52 Barrer, R.M. 85 benzene 173, 233 Research Octane Number 170 binary gas mixtures 88–9 CO2 separation 90, 91, 262 hydrogen separation 91, 261, 296–8 biogas 238 purification 157–8, 159 upgrading 181–2 see also natural gas 2,2-bis(4-methacryloxypolyethoxy phenyl) propane 96 bisphenol A ethoxylate diacrylate 95 block copolymer membranes 97–103 hydrogen sulfide removal 161, 162 PEBAX 91, 97–8, 99, 161 structure-property relationships 100–1 blocked fibers 146–7 Blue Membranes GmbH 180 BOC 95 Boltzmann statistics 9 Bondi method 67 boundary conditions PBMRs 14, 15 Danckwerts’ 9 Brownmillerite 194 bubble phase component mass balances, FBMRs 19, 23, 33 butane 154, 165 butyl rubber 93 by-product value 288–9 capital expenses (CAPEX) acid gas removal 155 carbon capture 30, 31, 41, 156

303

carbon adsorption 48 carbon capture 29–57, 85 capital expenses 30, 31, 41, 156 flue gases 183–4, 239–40 zeolite membranes 232–3 and global warming 29–31 high performance membranes 84–124 hybrid systems 52–3 membrane module simulation 32–40 capture step boundary conditions 32–4 classical methodology 38–40 simplified framework 34–8 membrane processes 31–2 operating expenses 30, 31, 41, 156 oxycombustion 30 postcombustion 30–1, 196–214 boundary conditions 197–201 competing technologies 211–12 driving force 201–3 energy requirement 44–5 membrane area trade-off 45–7 membranes 200–1 multi-stage processes 47–8 selectivity 42–4 simulation studies 40–8 techno-economic analysis 203–11 precombustion 30, 86 processes for 86 role of water in 51–2 selectivity and productivity 48–50 and sequestration 44 trade-off relationships 90–1 carbon dioxide capture see carbon capture effect on membrane aging 74–5 emissions 85–6 plasticization by 160 quality requirements for pipeline transport 199 reduced emissions 122 removal from natural gas 152–60 biogas purification 157–8, 159

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

304

carbon dioxide (continued) enhanced oil recovery 153–7 future directions 158–60 storage 199–200 transport 33, 198–9 carbon dioxide separation 122 binary gas mixtures 90, 91, 262 membrane systems 239–40 zeolite membranes 231–40 carbon membranes 178–9 Carbon Membranes Ltd 180 carbon molecular sieve membranes 162–91 aging and regeneration 177–8 characterization 170–8 Fourier transform infrared spectroscopy 170–1 gas permeation 173–7 gas sorption 173 scanning electron microscopy 170 X-ray diffraction 171–2 X-ray photoelectron spectroscopy 171 hydrogen separation 242 industrial applications 181–8 air separation 184–5 biogas upgrading 181–2 flue gas carbon capture 183–4 high-temperature 186–7 natural gas purification 182–3 petrochemical industry 185–6 module construction 180–1 permeability, Robeson upper boundary 162–3 production 164–70 carbonization 168–9 material functionalization 165–6 material selection 164–5 post-treatment 169–70 precursor preparation 166 pretreatment 166–8 transport mechanisms 178–81 Knudsen diffusion 179 molecular sieving 180 selective surface flow 179–80

Subject Index

carbon monoxide 43, 92 adsorption of 48 clean-up 93, 99–100 conversion 97, 98 Selox 99 carbon monoxide inhibition 137–61 and concentration polarization 142–57 concentration polarization coefficient 143–8, 150 inhibition coefficient 148–9, 150 permeance and flux effects 151–3 permeation reduction coefficient 149–51 permeation reduction maps 153–7 carbon nanotubes, as functionalization additives 166 carbonization 168–9 cascaded membrane systems 222–4 catalyst phase mass balance, PBMRs 10–11 catalytic partial oxidation 129–30 catalytic reforming 291–2 cellulose acetate 85, 165 permeators 150–1 chemical cross-linking 177 chemical vapor deposition 45, 47, 115, 241 Chevron 157 Claus process 130, 132 co-current flow 136–7 co-polyimide hollow fiber modules 296–8 coal bed methane 238 cohesive energy density 107 cold rolling 115 commercial applications 215–44, 280–4 air drying systems 232 carbon molecular sieve membranes 181–8 gas separation 216 hydrogen 92–100, 232–4 nitrogen 228–9

View Online

305

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

high-temperature 186–7 hollow fiber modules 133 see also individual applications COMPASS 16–17 component mass balance for PBMRs one-dimensional models 8–9, 10 two-dimensional models 14 compressor purge treatment 31 computational fluid dynamics 6, 7, 22, 25, 28 computational methods 5–16 atomistic simulation 8–16, 21 DDFT 6 dissipative particle dynamics 5–6 fluid dynamics 7, 22 gas separation 7–8 lattice Boltzmann simulation 8 concentration polarization 140–1 and carbon monoxide inhibition 142–57 concentration polarization coefficient 143–7, 150, 157 definition 143–5 and hydrogen permeation driving force 145–8 and inhibition 147–8 limit values 146 concentration profile 140–1 concurrent modeling 6 condensed-phase optimized molecular potentials for atomistic simulation studies see COMPASS continuity equation for PBMRs one-dimensional models 8 two-dimensional models 13 conversion index 100, 104 and feed pressure 104–5 copolymerization 114 copolymers with intrinsic microporosity (CoPIMs) 111 corrosion 198 counter-current flow 137 CRI-Criterion 64–5, 66, 67, 68 cross-flow 137–8 filtration 226

cryogenic air distillation 215–17 cryogenic gas separation 32, 52, 216–17, 285 cyclic steady state 249 cyclohexane, Research Octane Number 170 Cynara 157, 225 Cytop 91, 109 Dallas Production Inc. 157 Damko¨hler’s number 97, 98 Danckwerts’ boundary conditions 9 Daynes, H.A. 85 DD3R membranes 236, 237–9 DDFT 6 dead-end filtration 226–7 hollow fibers 201 Debye-Waller factor 14 dehydration of natural gas 237 dehydrogenation 60–2 ammonia 93 high-temperature 186–7 light hydrocarbons 93 oxidative 93, 210–12 propane 128–9 Delaunay simplices 4 dendrimer membranes 103–5 dense gas-solid systems, multi-scale modeling 25–9 desalination 32 di-isopropyl dimethyl PEEK-WC 16, 17 selectivity/solubility 20 DIDMPEEK 16, 17 selectivity/solubility 20 diffusion 87–8 diffusion-enhanced membranes 105–19 amorphous fluoropolymers 107–10 intrinsic microporosity 110–13 substituted polyacetylenebased 105–6 thermally rearranged 113–19 diffusion-selective membranes 89 digital image analysis 29

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

306

dimethyl PEEK-WC 16, 17 selectivity/solubility 20 dimethylbutanes, Research Octane Number 170 dimethylterephthalate 173 discrete bubble model 26 discrete particle models 26–7, 28 displacement purge adsorption 247–8 dissipative particle dynamics 5–6 DMF 168 DMPEEK 16, 17, 19 selectivity/solubility 20 Doolittle relationship 75 driving force 145–7, 201–3 feed compression 201–3 suction at permeate side 202–3 sweep operation 203 dry nitrogen seals 228 dry reforming 52 DuPont 151, 246 Einstein formulation 10, 14 electroless plating 45–7, 115 electroplating 115 ellipsometry 61, 69 emulsion phase component mass balances, FBMRs 19, 23, 33 energy effective dispersion coefficients 30 requirement 44–5 energy balance 141–2 FBMRs 20 PBMRs 14–15 Energy Centre of the Netherlands 47, 69 hydrogen separation modules 116 enhanced oil recovery 199 CO2 238–9 nitrogen facilitated 229 technique 153–7 Environgenics 150, 158 ethane 154 ethanol reforming see alcohol reforming ethylene production 293–4

Subject Index

thermal water splitting with oxidative dehydrogenation 210–12 exergetic analysis 298–9 extraction index 100, 105–6 and feed pressure 106 ExxonMobil 99, 241 ISOSIEVE process 171 facilitated transport membranes 177–8 performance 180–1 FAU zeolite membranes 223, 224, 234, 235, 248 faujasite 179 FBMRs see fluidized bed membrane reactors 6FDA-IPDA polyimide 61 properties 62 feed composition 37, 289 feed compression 46, 201–2 permeate side 46, 202–3 feed pressure 289–90 and conversion index 104–5 and extraction index 106 and volume index 101, 102 fibers blocked 146–7 diameter 143 length 144 see also hollow fiber membranes; membrane modules Fick, A. 84 Fick’s law 10, 86, 179 film thickness, and membrane aging 59–60, 61–3, 77, 78 films see membranes filtration cross-flow 226 dead-end 201, 226–7 Fischer-Tropsch synthesis 42, 205 fixed site carrier membranes 31, 50 selectivity and permeability 50 flow rate 289–90 flue gases 34

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

carbon capture 183–4, 239–40 zeolite membranes 232–3 CO2 separation 86 composition 33 desulfurization unit 197, 198 drying 31 pressure 211 separation 86 fluid catalytic cracking 173, 293 fluidized bed membrane reactors 2, 3, 17–29 advantages/disadvantages 17 autothermal methane reforming 53–4 constitutive equations 32–3 modeling of 18–25 multi-scale modeling of dense gassolid systems 25–9 fluorescence confocal optical microscopy 228 fluorinated polyimides 89 fluorite 194 fluoropolymers 107–10 flux hydrogen, carbon monoxide inhibition 151–3 oxygen 192–3, 195 MIEC membranes 255–8 and partial pressure gradient 203 perovskite membranes 192–3, 195 food storage/preservation 228 Formusa Plastics 241 Fourier transform infrared spectroscopy 170–1 fractional free volume 67, 106 and aging time 76 superglassy polymers 58 free volume elements 91 friction coefficient for PBMRs one-dimensional models 8 two-dimensional models 14 Galden HT 55 108 gases

307

critical temperature 92 diffusion coefficients 10 diffusivity 246 see also individual gases gas exchange coefficient 33 gas permeation see permeation gas permeation unit 49 gas separation 7–8, 245–6 adsorption 217 air drying 230–2 applications 232 high pressure 232 membrane system 232 air separation systems 226–30 nitrogen applications 228–9 nitrogen generation 226–8 oxygen generation 229–30 oxygen/nitrogen simultaneous production 230 biogas systems 238 carbon molecular sieve membranes 162–91 CO2 84–124 coal bed methane 238 commercial applications 216 concentration polarization 140–1 and carbon monoxide inhibition 142–57 cryogenic 32, 52, 216–17, 285 enhanced oil recovery 238–9 high performance membranes 89–91 hydrogen see hydrogen separation membrane modules 220–1 hollow fiber 126, 131–2, 133–8 non-ideal construction 142–7 performance 138–42 plate-and-frame 126, 127–9 simulation 32–40 spiral wound 126, 129–31 membranes 166, 215–16, 217–24 fabrication 219–20 hybrid separation 224 materials 218–19 modules see membrane modules producers 225

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

308

gas separation (continued) selectivity and permeability 50 systems 221–4 natural gas upgrading systems 234–9 acid gas removal 234–7 dehydration 237–8 new metrics 294–8 organic vapor systems 240–1 oxygen see oxygen separation polymeric membrane systems 217–24 pressure-swing adsorption see pressure-swing adsorption processes 216–17 selection guidelines 286–91 by-product value 288–9 ease of expansion 288 operating flexibility 286–7 reliability 287–8 separation process 289–91 turndown 287 substitute natural gas 238 technologies 284–6 trade-off relationship 3, 88–9 zeolite-based membranes 231–48 gas sorption 173 gas transport amorphous cells 19–20 CO2 33, 198–9 glassy polymer membranes 3–4 see also individual membrane types gas-surface interactions 139–40 gasoline isomerate fractionation 170–2 General Electric Company 99 Generon 225, 228 GKSS membranes 151 modules 129 glassy polymer membranes 3, 58 aging see membrane aging diffusion in 10 fluorinated 166 gas separation 166 gas transport 3–4

Subject Index

hydrogen sulfide removal 161–3 olefin/paraffin separation 174–7 permeability 90–1 plasticization 51 QSPR correlation 4 see also membranes global warming 29–31 see also greenhouse gases Grace Membrane Systems 151 Graham, Thomas 85 grain storage 228 granular flow, kinetic theory 26 Green-Kubo formulation 10 greenhouse gases 41, 85, 124, 196 removal of 216 see also carbon dioxide Gusev-Suter method 13–14, 19 Hagen-Poiseuille law 143 Hamiltonians 9 helium recovery 216 Heniffin 228 Henry sorption 140 Henry’s law 87, 179, 180 hexane, Research Octane Number 170 HFE 7100 108 high performance membranes 84–124 dendrimer 103–5 diffusion-enhanced 105–19 poly(ethylene oxide) 91, 93–7 high-temperature applications 186–7 dehydrogenation 186–7 steam methane reforming 187 higher hydrocarbons, removal of 164–5 highly permeable polymers 118 hollow fiber modules 126, 131–3 applications 133 co-polyimide 296–7 configuration 134 dead-end filtration 201 MIEC see MIEC hollow fiber membranes operation of 133–8

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

flow 133–5 flow patterns 136–8 mode 135–6 perovskite 198 properties 132 see also individual membrane types humidity, and membrane aging 177 hybrid membrane/PSA processes 251–73 co-current blowdown 257, 268 configurations 255 counter-current blowdown 257, 268 more permeable component least adsorbed 260–8 more adsorbed 268–73 purging 257, 268 Sips isotherm 260, 268, 269 stream composition/recovery 267 hybrid molecular-continuum modeling 6–7 hybrid separation processes 224 hydrazine 168 hydro-desulfurization 168 hydrocarbon catalytic partial oxidation 129–30, 204 Hydrocarbon Operating Inc. 157 hydrocarbon processing 150–95 natural and biogas membranes 150–68 acid gas removal 152–60 higher hydrocarbon removal 164–5 hydrogen sulfide removal 160–3 inert gas removal 165–8 suppliers 150–2 water removal 163–4 petrochemicals 173–85 monomer recovery 185 separation of light olefins/ paraffins 173–84 xylene isomer separation 184–5 petroleum refining 168–73 gasoline isomerate fractionation 170–2 hydrogen purification 168–70

309

hydrocracking 232 gases from 293 hydrogen applications 44 production/purification see hydrogen separation recovery in refineries 233 separation 91 hydrogen chloride 168 hydrogen embrittlement phenomenon 113 hydrogen flux, carbon monoxide inhibition 151–3 hydrogen permeation 101 carbon monoxide inhibition 142–57 and concentration polarization 142–57 concentration polarization coefficient 143–8, 150 inhibition coefficient 148–9, 150 permeance and flux effects 151–3 permeation reduction coefficient 149–51 permeation reduction maps 153–7 driving force 145–7 inorganic membranes 241 hydrogen permeators 168 hydrogen separation 41–4, 168–70, 216, 281–4 alcohol reforming 57–60 ammonia decomposition 62 autothermal reforming 42 binary gas mixtures 91, 261, 296–8 chemical reactions 42 dehydrogenation 60–2 from fossil fuels 88 integrated membrane process 89–90 liquid carriers 42–3 membrane reactors 1–39, 87–109 applications 92–100 membranes 232–4 applications 233–4

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

310

Subject Index

hydrogen separation (continued) carbon molecular sieve 242 cellulose acetate 241 palladium-based 40–86, 110–36, 137–9, 241 polyimide 241 polysulfone 241 silica 241–2 zeolite 240–8 methane reforming 41–2, 50–4 proton exchange membrane fuel cells 245–6, 247 thermal water splitting 210–12 with oxidative dehydrogenation 210–12 with partial methane oxidation 210 traditional reactors 91 up-scaled 63–79 CRI-Criterion 64–5, 66, 67, 68 Energy Centre of the Netherlands 69 Membrane Reactor Technologies 70 Pall Corporation 65, 67, 69 water gas shift 42, 43, 89, 93, 95–9 hydrogen sorbents 43 hydrogen sulfide catalytic decomposition 130–2 removal 160–3 block polymer membranes 161 glassy polymer membranes 161–3 rubbery polymer membranes 160–1 hydrotreaters, gases from 293 Hyflon AD 91, 108–9 properties 109 Hysep modules 69, 116 HYSIS system 222

inhibition coefficient 148–9, 150, 157 Inocermic GmbH 225–6 integrated gasification combined cycle process 55, 86, 127 integrated membrane reactors 119, 132–3 intrinsic microporosity 91, 105, 110–13 ion transport membranes 30, 219 IPSORB process 171 isopentane, Research Octane Number 170

ideal selectivity 36, 49 immersion precipitation 7 inert gases on-board generator systems 229 removal from natural gas 165–8

magnetic suspension balance 173 magnetron sputtering 115 Markov chains 12 mass, effective dispersion coefficients 30

Joule-Thomson coefficient 141, 236 KAPOF project 163 key performance indicators 204–6 Knudsen diffusion 10, 125, 179, 230 Knudsen flow 145 Kroeger-Vink notation 192, 255 Kvaerner Membrane Systems 151, 157, 225 Kvaerner process 240 lattice Boltzmann simulation 8, 22, 26 Lennard-Jones radius 20–1, 180, 185 light hydrocarbons dehydrogenation 93 oxidative 93 hydrogen separation from 88, 89 oxidative dehydrogenation 93 partial oxidation 93 steam reforming 93–5 limit conversion in membrane reactors 3–6 liquid membranes 31 selectivity and permeability 50 liquid ring pumps 202 Lorentz-Lorentz equation 67

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

mass balance 35–6 bubble phase 19, 23, 33 emulsion phase 19, 23, 33 FBMRs 19 mass transfer 35–6 resistance model 145 material functionalization 165–6 additives 166 Matrimid 60, 95, 179 oxygen permeability 65–6 properties 62 Matrimid-PEO 97 matrix phase swelling 247 Medal (MEmbrane Dupont Air Liquide) membranes 85, 151, 157, 169, 225 Meirovitch’s scanning method 18 Membrain project 197 membranes 1–28, 50 aging see membrane aging alumina 92 amorphous fluoropolymer 107–10 block copolymer 97–103 carbon 178–9 carbon capture 29–57 postcombustion 200–1 carbon molecular sieve 162–91 cellulose acetate 85, 165 chemical sensitivity 158–9 classification 111 commercial applications 215–44, 280–4 compaction 159 DD3R 236, 237–9 defects 144–5 dendrimer 103–5 diffusion-enhanced 105–19 amorphous fluoropolymers 107–10 intrinsic microporosity 110–13 substituted polyacetylenebased 105–6 thermally rearranged 113–19 diffusion-selective 89 fabrication 45–7, 114–16, 219–20 facilitated transport 177–8, 180–1

311

film thickness 59–60, 61–3 fixed site carrier 31, 50 gas separation 166, 215–16, 217–24 fabrication 219–20 hybrid separation 224 materials 218–19 modules see membrane modules producers 225 systems 221–4 glassy polymer see glassy polymer membranes high performance see high performance membranes hollow fiber see hollow fiber membranes hydrogen separation 232–4 applications 233–4 carbon molecular sieve membranes 242 cellulose acetate membranes 241 palladium-based membranes 40–86, 110–36, 137–9, 241 polyimide membranes 241 polysulfone membranes 241 silica membranes 241–2 zeolite membranes 240–8 ion transport 219 liquid 31, 50 microporous 92 MIEC 253–78 natural/biogas see natural/biogas membranes numerical simulation 16–20 amorphous cell packings 17–18 force field and choice of ensemble 16–17 gas tranport through amorphous cells 19–20 realistic amorphous cell selection 18–19 palladium alloy 40, 47–9, 90, 113–14, 218

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

312

membranes (continued) palladium-based see palladiumbased membranes PEBAX 161 performance 139–40 permeability see permeability permeance 49 perovskite 192–222 plasticization of 160, 218 PMMA 59 and permeability 60 poly(ethylene oxide) 91, 93–7 polyimide 162, 165, 241 polysulfone 115 postcombustion carbon capture 200–1 pretreatment 166–8 Prism 168, 246 pyrolyzed 182 regeneration 178 rubbery polymer see rubbery polymer membranes SAPO-34 236, 239–42 selective 2, 3, 111–14 selective layers block polymers 161 compaction of 159 rubbery polymers 160–1 silica 91, 92, 241–2 silicon-based 92 sorption-enhanced 92–105 block copolymer membranes 97–103 dendrimer membranes 103–5 poly(ethylene oxide) 91, 93–7 staged/cascaded systems 222–4 tantalum 112 technology 285–6 thermally rearranged see thermally rearranged membranes thickness 59–60, 143–4 titania 45, 92 trade-off relationships 3, 88–91 zeolite see zeolite membranes membrane aging 58–83, 177 aging time 62, 64, 73

Subject Index

CO2 exposure 74–5 experimental conditions 72–5 film thickness 59–60, 61–3, 77, 78 humidity effect 177 modeling 75–8 o-Ps lifetime 70–1 optical properties 67 oxygen chemisorption 177 and oxygen permeability 63–6, 73 PALS 68–72 previous history 72–5 regeneration 178 relative density 68 thin and ultra-thin films 60–7 membrane contactors 31 membrane micro-reactors 2 membrane modules 127–33, 220 carbon molecular sieve membranes 180–1 design 125–49 hollow fiber 126, 131–2, 133–8, 220 plate-and-frame 126, 127–9 spiral wound 126, 129–31, 220 honeycomb configuration 180 hydrogen separation 116, 117 Hysep 69, 116 nitrogen generation 227 non-ideal construction 142–7 blocked fibers 146–7 fiber diameter variation 143 fiber length variation 144 membrane defects 144–5 membrane thickness variation 143–4 performance 138–42 concentration, pressure and temperature profiles 140–1 energy balance 141–2 membrane 139–40 pressure losses 142 purity 138 recovery 138–9 selectivity 138 stage cut and pressure ratio 139

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

reformer and membrane module systems 63–4, 124, 126 simulation 32–40 capture step boundary conditions 32–4 classical methodology 38–40 simplified framework 34–8 see also membranes Membrane Reactor Technologies 53, 70 membrane reactors architecture 52–4, 118–21 benefits and drawbacks 123 integrated 119, 132–3 staged 119–21, 132–3 benefits and drawbacks 121–4 better CO2 purification 122 increased process efficiency 121 methane economy 121–2 reaction temperature reduction 121 reduced CO2 emissions 122 reduced dependence on natural gas cost 122 catalytic 90 configuration 103 dense gas-solid systems 25–9 features of 110–18 feed pressure and conversion index 104–5 and extraction index 106 and volume index 101, 102 fluidized bed 2, 3, 17–29, 32–3 advantages/disadvantages 17 autothermal methane reforming 53–4 constitutive equations 32–3 modeling of 18–25 multi-scale modeling of dense gas-solid systems 25–9 hydrogen separation 1–39, 87–109 applications 92–100 limit conversion in 3–6 methane reforming 52–4 modeling of 1–39 operating temperature 102

313

oxygen separation 197 packed bed 2–3, 6–17, 29–33 catalyst configuration 2 constitutive equations 29–32 drawbacks 3 one-dimensional models, heterogeneous 9–13 two-dimensional models 13–17 permselective separation 90 reactant distribution 90 reaction volume and membrane type 104 removal of product from reaction volume 91 vs traditional reactors 100–6 water gas shift 55–7, 100 membrane reformers 63 see also alcohol reforming; methane reforming membrane selectivity see selectivity membrane systems 216, 221–2 staged/cascaded 222–4 see also membranes Membrane Technology & Research 240–1 membrane-distillation coupling 179, 182–4 Membratek-Envig 158 MemfoACT 182 mesoscale modeling 5, 21 metal nitrates, as functionalization additives 166 metal oxides, as functionalization additives 166 methane adsorption of 48 coal bed 238 economy 121–2 oxidative coupling 205–7 partial oxidation to syngas 204–5, 210 saving of 121–2 separation 90, 261 methane reforming 41–2, 50–4, 91, 93–5 autothermal reforming 42, 52

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

314

methane reforming (continued) combined reforming 52 dry reforming 52 membrane reactor design 52–4 results 51 steam 41–2, 50–1, 91, 124, 187–8 methanol steam reforming 95 methyl rubber 93 methyldiethoxysilane 243 methyldimethoxysilane 243 methylpentanes, Research Octane Number 170 Metropolis algorithm 9 MFI zeolite membranes 171, 223, 224, 229, 234, 236, 243–5, 246, 248 microporosity, intrinsic 91, 105, 110–13 microporous membranes 92 MIEC hollow fiber membranes 197, 253–78 design 265–70 energy consumption and cost analysis 270–4 mechanical strength 264–5 operating equations 266–9 high pressure mode 267 sweep-gas 268 operation mode 265 oxygen permeation 254–60 mechanism 254–5 permeation flux 255–8 stability 258–60 preparation 260–2 surface modifications 263–4 see also perovskite membranes mining, nitrogen gas in 229 mixed oxygen-ionic and electronic conducting ceramic membranes see MIEC hollow fiber membranes Mobil 157 molecular dynamics 3, 9–12 Molecular Gates 165 molecular mechanical methods 8 molecular sieving 125, 180 mono-diethanol ammine 122 monomer recovery 185

Subject Index

Monsanto 85, 157, 246 Monte Carlo simulations 3, 12–13, 22 Moore’s law 21 MRT 95 hydrogen separation modules 116 MTR Inc. 165 mullite 226 multi-scale molecular modeling 1–28, 25–9 computational methods 5–16 numerical simulation of polymer membranes 16–20 multi-stage evaporation 32 multi-stage membrane systems 47–8 multiphase reactors see membrane reactors NanoGLOWA project 197, 200, 205, 212 Natco Group Inc. 85, 235 National Energy Technology Laboratory 184 natural gas 150–68 acid gas removal 152–60, 216, 234–7 CO2 removal 152–60 biogas purification 157–8, 159 enhanced oil recovery 153–7 future directions 158–60 cost of 122 dehydration 237 higher hydrocarbon removal 164–5 hydrogen sulfide removal 160–3 block polymer membranes 161 glassy polymer membranes 161–3 rubbery polymer membranes 160–1 inert gas removal 165–8 membrane suppliers 150–2 nitrogen removal 167, 238 purification 182–3 reforming 233–4 steam 124–7

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

sweetening 86 upgrading systems 234–9 acid gas removal 234–7 dehydration 237–8 water removal 163–4 pressurization of permeate stream 164 reduction of water vapor pressure 163–4 see also biogas natural gas liquids 240–1 Navier-Stokes equations 13, 26, 27 Nernst-Einstein equation 256 Nernst-Planck equation 255–6 Newton’s second law 9, 27 NGK Insulators 171 nickel, as functionalization additive 166 niobium membranes 112 nitric oxide decomposition in syngas production 213–14 nitrogen applications 228–9 dry nitrogen seal 228 food storage/preservation 228 grain storage 228 mining 229 oil recovery 229 on-board inert gas generator systems 229 tire inflation 228–9 generation 226–8 air compression 226–7 air purification 227 membrane modules 227 pipeline flow effects 198 production 184 removal from natural gas 167, 238 nitrogen separation 90, 216 nitrogen-CO2 separation 90, 262 Nitrosep units 168 nitrous oxide decomposition in syngas production 212–13 Nitto Denko 175 numerical simulation 16–20 amorphous cell packings 17–18

315

force field and choice of ensemble 16–17 gas transport through amorphous cells 19–20 realistic amorphous cell selection 18–19 olefins 173 polymerization 216 production 207–9 separation 173–84 carbon membranes 178–9 facilitated transport membranes 177–8 glassy polymer membranes 174–7 membrane-distillation couplings 179, 182–4 zeolite membranes 179 on-board inert gas generator systems 229 operating expenses (OPEX) acid gas removal 155 carbon capture 30, 31, 41, 156 organic vapor separation systems 240–1 orthogonal collocation methods 38 oxidation catalytic partial 129–30, 204 light hydrocarbons 93 methane 204–5, 210 oxidative coupling of methane 205–7 oxycombustion carbon capture 30 oxygen and accelerated corrosion 199 chemisorption 177 pipeline flow effects 198 oxygen enriched air 229–30 oxygen flux 192–3, 195 and partial pressure gradient 203 oxygen permeation and membrane aging 63–4 MIEC membranes 254–60 mechanism 254–5 permeation flux 255–8

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

316

oxygen permeation (continued) stability 258–60 perovskite membranes 196, 197 oxygen separation 184, 196, 198–204, 216, 229–30 cryogenic air distillation 215–17 elevated pressure on permeate side 202–3 evacuation on permeate side 201–2 with elevated pressure on feed side 203–4 from air consumption in partial oxidation 204–9 low temperature 200 from oxygen-containing gases 209–14 decomposition of N2O and NO 212–14 thermal water splitting 210–12 membrane reactors 197 MIEC membranes 253–78 organic vs inorganic membranes 215–17 perovskite membranes 192–222 pressure swing adsorption 216–17, 253 sweep gases in 196, 198–201 Wagner equation 192, 193, 203 packed bed membrane reactors 2–3, 6–17, 29–33 catalyst configuration 2 constitutive equations 29–32 drawbacks 3 one-dimensional models 7–13 heterogeneous 9–13 pseudo-homogeneous 7–9 two-dimensional models 13–17 Pall Corporation 65, 67, 69 palladium alloy membranes 40, 47–9, 90, 113–14, 218 hydrogen permeation 101 see also palladium-based membranes

Subject Index

palladium-based membranes 40–86, 91, 137–61 advantages and drawbacks 92 alcohol reforming 57–60, 61 hydrogen yield 62 ammonia decomposition 62 carbon monoxide inhibition 142–57 concentration polarization 140–1 cost analysis 117–18 dehydrogenation reactions 60–2 development 45 endothermic/exothermic reactions 60–2 Energy Centre of the Netherlands 116 fabrication methods 45–7, 114–16 gas-surface interactions 139–40 hydrogen embrittlement phenomenon 113 hydrogen separation 40–86, 110–36, 137–9, 241 hydrogen transport 112–13 Japanese 116–17 methane reforming 50–4 autothermal reforming 42, 52 steam reforming 41–2, 50–1 water gas shift reaction 42, 43, 55–7 MRT 116 poisoning of 92 selective 110–36 SINTEF 47, 117 structural stability 49 up-scaled technology 64–79 CRI-Criterion 64–5, 66, 67, 68 Energy Centre of the Netherlands 69 Membrane Reactor Technologies 70 Pall Corporation 65, 67, 69 palladium-copper alloy 49 palladium-gold alloy 49 palladium-silver alloy 48 PALS see positron annihilation lifetime spectroscopy

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

PAMAM dendrimers 103–5 paraffin separation 173–84 carbon membranes 178–9 facilitated transport membranes 177–8 glassy polymer membranes 174–7 membrane-distillation couplings 179, 182–4 zeolite membranes 179 Parker 225, 228 particle image velocimetry 29 PBMRs see packed bed membrane reactors PEBAX membranes 91, 97–8, 99, 161 PEEKs 16, 17 selectivity/solubility 20 PEEK-WC 16, 17, 20 selectivity/solubility 20 PEGDA 94, 95, 96 PEGDMA 96, 104 PEGMEA 94, 95, 96 pentane, Research Octane Number 170 PEO see poly(ethylene oxide) Permea Inc. 151, 158 permeability 2–3, 246 fixed site carrier membranes 50 glassy polymer membranes 90–1 liquid membranes 50 mass transfer expression 35 Matrimid 60 and membrane aging 63–6, 73 PIMs 112 poly(ethylene oxide) membranes 94, 95 polysulfone membranes 64 QSPR correlation 4 thermally rearranged membranes 114, 116 see also permeation permeability coefficient 2–3 PEO membranes 94, 95 permeance 49, 151–3 and carbon monoxide inhibition 151–3

317

permeate 139–40, 141 composition 35, 37 concentration 145 water vapor pressure reduction 163–4 permeate stream 126 feed compression 202–3 suction 202–3 water-concentrated, pressurization of 164 permeation 32, 165–7, 174–7 carbon molecular sieve membranes 173–7 driving force 145–7, 201–3 feed compression 201–3 suction at permeate side 202–3 sweep operation 203 hydrogen carbon monoxide inhibition 142–57 inorganic membranes 241 Knudsen diffusion 10, 179, 230 membrane module simulation 32–40 capture step boundary conditions 32–4 classical methodology 38–40 simplified framework 34–8 membrane performance 166 mixed gas measurements 176–7 oxygen MIEC membranes 254–60 perovskite membranes 196, 197 single gas tests 174–6 solution-diffusion 86–7, 89, 125 zeolite membranes 229–31 see also permeability permeation flux 256–8 permeation reduction coefficient 139, 149–51, 158 permeation reduction maps 153–7, 158 permeative-stage membrane reactors 53

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

318

permselectivity 2, 90 amorphous fluoropolymers 109 PIMs 112 poly(ethylene oxide) membranes 95 thermally rearranged membranes 116 perovskite membranes 192–222 disadvantages of 217 hollow fiber 198 dead-end 201 materials 194–6 oxidative coupling of methane 207–8 oxygen flux 192–3, 195 oxygen permeation 196, 197 stability 195 thickness 193 see also MIEC membranes perturbation methods 38 pervaporation 184–5, 224–5 PETEDA dendrimers 105 petrochemical industry 173–85 hydrocarbon recovery 185–6 monomer recovery 185 separation of light olefins/ paraffins 173–84 xylene isomer separation 184–5 petroleum refining 168–73 gasoline isomerate fractionation 170–2 hydrogen purification 168–70 physical vapor deposition 45, 115 PIMs see polymers with intrinsic microporosities pinhole defects 145 plasticization 160, 218 plate-and-frame modules 126, 127–9 plug flow 135 PMMA 59 point plating 46 poly-ether-ketones 16, 17, 19–20 poly(1-p-tert-butylphenyl-2phenylacetylene) (PTBPPA) 106 poly(1-triethylsilyl-1-propyne) (PTESP) 106

Subject Index

poly[1-(trimethyl-silyl)propyne] see PTMSP poly(1-trimethylgermyl-1-propyne) (PTMGP) 106 poly(4-methyl-2-pentene) (PMP) 106 polyacrylonitrile 162, 165 poly(amide-b-ethylene oxide) (PEBAX) 97–8 structure-property relationships 99 poly(amidoamine) (PAMAM) dendrimers 103–5 polyaramides 233, 246 polybenzimidazole 113 polybenzothiazole 113 polybenzoxazole (PBO) 113 polybutadiene 93 poly(2,6-dimethyl-1,4-phenylene oxide) 62 polydimethylsiloxane 10–11, 65, 85, 92, 93 polyether-ether-ketones see PEEK polyethersulfone 85 polyethylene 93 poly(ethylene glycol) acrylate 96 poly(ethylene glycol) diacrylate 96 poly(ethylene glycol) dimethacrylate 96 poly(ethylene glycol) ethyl acrylate 96 poly(ethylene glycol) methyl ether acrylate 96 poly(ethylene glycol) methyl ether methacrylate 96 poly(ethylene oxide) membranes 91, 93–7 block copolymer 97–103 cross-linking 94–7 permeability 94, 95 permselectivity 95 postcombustion carbon capture 200–1 poly(ethylene oxide)–poly(butylene terephthalate) copolymer 102 polyethylene terephthalate 173 polyimide 165, 233, 246

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

polyimide membranes 85, 98, 162, 166, 241 aromatic polyimides 219 gas separation 166, 168, 169, 246 polymer membranes see membranes polymers with intrinsic microporosities (PIMs) 91, 105, 110–13 properties 112 synthesis 110 poly(methyl methacrylate) see PMMA poly(oxa-p-phenylene-3,3-phthalidop-phenylenxoxa-p-phenylenexoxip-phenylene) see PEEK-WC poly(phenylene oxide) 162, 165 poly(phthalazinone ether sulfone ketone) 162, 165 polypyrrolone (PPL) 113 polysulfone 233 polysulfone membranes 63, 85, 115 oxygen permeability 64 properties 62 polytetrafluoroethylene (PTFE) 107 polyurethanes 98 polyvinylpyrrolidone 165 porosity profile 30–2 positron annihilation lifetime spectroscopy (PALS) 68–72, 105, 111 postcombustion carbon capture 30–1, 196–214 boundary conditions 197–201 CO2 storage 199–200 CO2 transport 198–9 power plant 197–8 competing technologies 211–12 driving force 201–3 feed compression 201–3 suction at permeate side 202–3 sweep operation 203 energy requirement 44–5 membrane area trade-off 45–7 membranes 200–1 multi-stage processes 47–8 selectivity 42–4

319

simulation studies 40–8 energy requirement 44–5 energy requirement/membrane area trade-off 45–7 multi-stage process 47–8 selectivity 42–4 techno-economic analysis 203–11 key performance indicators 204–6 process configurations 203–4 with retentate recycling 208–11 without retentate recycling 206–8 power plants 197–8 Praxair 254 precombustion carbon capture 30, 86 decarbonization 44 pressure losses 142 pressure profile 140–1 pressure ratio 36, 139 pressure vacuum swing adsorption 226 pressure-swing adsorption 165, 169, 226, 247–75, 284–5 adsorbent productivity 251 blowdown 250 cyclic steady state 249 energy requirements 251 high-pressure adsorption 248–9, 250 hybrid membrane processes 251–73 co-current blowdown 257, 268 configurations 255 counter-current blowdown 257, 268 more permeable component least adsorbed 260–8 more permeable component more adsorbed 268–73 purging 257, 268 Sips isotherm 260, 268, 269 stream composition/ recovery 267 hydrogen production 88 oxygen separation 216–17, 253

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

320

pressure-swing adsorption (continued) permeate pressure 254 pressure equalization 249 pressure range 251 pressurization 250 process scaling 251 product purity/recovery 251 purging 251 Skarstrom process 249 pretreatment of membranes 166–8 chemical 167–8 physical 167 Prism membrane systems 164, 168, 246, 281 product purity 290–1 product/permeate recycling 203 productivity 48–50 propane 154, 165 dehydrogenation 128–9 propylene adsorption 48 proton exchange membrane fuel cells 137, 245–6 PSA see pressure-swing adsorption PTMSP 60, 68, 91, 106 aging studies 79 fractional free volume 58 properties 62 purity 138 and recovery 139 PVDF 109 pyrolyzed membranes 182 QSAR 14–16 QSPR 4, 14–16 quantitative structure activity relationships see QSAR quantitative structure property relationships see QSPR quantum mechanics 8 recovery 138–9 purity trade-off 139 recycling product/permeate 203 retentate 203

Subject Index

process with 208–11 process without 206–8 reformer and membrane module systems 63–4, 124, 126 reforming alcohol 57–60, 61, 62, 95 autothermal 42, 52 catalytic 291–2 dry 52 methane 41–2, 50–4, 91, 93–5 natural gas 233–4 regeneration of membranes 178 relative density 68 Research Octane Number 170 residue recycling see recycling retentate recycling 203–4 process with 208–11 process without 206–8 retentate stream 126 retrofit 30 reverse osmosis 32, 140 Richardson equation 9, 11, 12 Robeson plot 184 root pumps 202 rotary lobe blowers 202 rubber 93 rubbery polymer membranes 166 apolar with polar solvent 161 silicones 161 diffusion in 10 hydrogen sulfide removal 160–1, 162 polar 161 properties 93 SABIC 241 SAPO-34 membranes CO2 separation 236, 239–40 hydrogen separation 242–3 scanning electron microscopy 61, 170 seeding 226, 227 selective membranes 111–14 selective surface flow 179–80 selectivity 2, 3, 138

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

and aging time 66 carbon capture 42–4, 48–50 CO2/N2 50, 51, 52 fixed site carrier membranes 50 H2O/CO2 51 ideal 36, 49 liquid membranes 50 Matrimid 60 membrane modules 138 PEO membranes 95 Selox process 93, 99–100 Separex 85, 151, 158 Shell Global Solutions 160 Shell Oil Company 95, 157 Sieverts driving force 147 Sievert’s law 90–1, 139, 147, 148, 158 Sieverts-Fick law 113 Sieverts-Langmuir model 9, 156 silica 45 silica membranes 91, 92 hydrogen separation 241–2 silica nanoparticles, as functionalization additives 166 silicones 92, 161 silver, as functionalization additive 166 Sinopec 241 SINTEF hydrogen separation modules 47, 117 Sips isotherm 260, 268, 269 Skarstrom process 249 solution-diffusion 86–7, 89, 125 Solvay Solexis 108 sorption 87 sorption-enhanced membranes 92–105 block copolymer membranes 97–103 dendrimer membranes 103–5 poly(ethylene oxide) 91, 93–7 ‘spacer’ molecules 18 spatial multi-scale calculations 6 spiral wound modules 126, 129–41 applications 133 spray pyrolysis 115 sputtering 45, 115

321

stage cut 139 staged membrane reactors 119–21, 132–3, 222–4 steam ejectors 202 steam reforming light hydrocarbons 93–5 methane 41–2, 50–1, 91, 93–5, 124, 187 natural gas 124–7 steam-to-carbon ratio 42 Struik model 76, 77, 78 substitute natural gas 238 substituted polyacetylene-based membranes 105–6 suction 202–3 Sulfatreat process 161 sulfur 92 Sun Explo 157 superglassy polymers 106 suppliers gas separation membranes 225 natural/biogas membranes 150–2 sweep gases 203, 268 oxygen separation 196, 198–201 syngas 216 production 42, 129 methane partial oxidation 204–5, 210 N2O/NO decomposition 212–14 tantalum membranes 112 Teflon 107, 108 Teflon AF 109 temperature profile 140–1 temperature-swing adsorption 169, 247 temporal multi-scale calculations 6 terephthalic acid 173 tetrafluoroethylene 108 tetrafluoroterephthalonitrile (TFTPN) 111 5,5 0 ,6,6 0 -tetrahydroxy-3,3,30,30tetramethylspirobisindane (TTSBI) 111

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

322

tetramethyl PEEK-WC 16, 17 selectivity/solubility 20 tetramethylhexafluoropolycarbonate (TMHFPC) 61 Theodorou-Suter chain-generation approach 18 thermal polarization 51 thermally rearranged membranes 91 diffusion-enhanced 113–19 permeability 114, 116 structure-property relationships 115 thin films aging behavior 60–7 coating technique 65 time-lag method 85 tire inflation, nitrogen pressure 228–9 titania membranes 45, 92 TMPEEK 16, 17 selectivity/solubility 20 Tokyo Gas 63, 95 toluene 173, 233 total mass balance, FBMRs 19 total momentum balance equation for PBMRs one-dimensional models 8 two-dimensional models 14 trade-off relationships 3, 45–7 gas separation 3, 88–9 binary gas mmixtures 90, 91 traditional reactors hydrogen production 91 vs membrane reactors 100–6 transfer term, FBMRs 20 transition metals, as functionalization additives 166 transition state theory 3, 12, 13–14, 21 see also Gusev-Suter method 19 2,2,4-trifluoro-5-trifluoromethoxy-1, 3-dioxole (TTD) 108 turndown 287 Ube Industries 151, 169, 225, 228, 231, 233 ultra-thin films 3 aging behavior 60–7

Subject Index

ultrafiltration 140 United Technologies Corporation 98 UOP 233, 235 MOLEX process 171 vacuum pumps 35–7 vacuum swing absorption 226, 230 Van der Waals radii 180 vanadium membranes 112 VaporSep 241 volatile organic compounds 85, 91 volume index 100, 101 and feed pressure 101, 102 and reactor configuration 103 Voronoi polyhedra 4 Wagner equation 192, 193, 203 waste gas shift recovery 86 water in carbon capture 51–2 thermal splitting 210–12 with oxidative dehydrogenation 210–12 with partial methane oxidation 210 water gas shift 42, 43, 55–7, 86, 124, 127–8 Damko¨hler’s number 97, 98 hydrogen separation 89, 93, 95–9 membrane reactors 100 performance of 56, 96 water removal 163–4 pressurization of permeate stream 164 reduction of water vapor pressure 163–4 Weir Envig 158 Widom particle insertion method 12 X-ray diffraction 171–2 X-ray photoelectron spectroscopy 171 xylenes 173, 233 isomer separation 184–5

View Online

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

Subject Index

zeolite Sil-1 45 zeolite T membranes 236–7 zeolite-based membranes 49, 91, 92, 171–3, 219, 223–52, 247 cation exchange 233, 234 characteristics 224 DD3R 236, 237–9 FAU 223, 224, 234, 235, 248 gas separations 231–48 CO2 231–40 hydrogen 240–8 mass transport in 229–31 MFI 171, 223, 224, 229, 234, 236, 243–5, 246, 248

323

olefin/paraffin separation 179 preparation 225–9 dead-end filtration 226–7 rapid thermal processing 227 seeding 226, 227 supports 225–6 in pressure swing adsorption 216 SAPO-34 236, 239–40, 242–3 selectivity 49 topology 227–8, 229 zeolites as functionalization additives 166 zirconia membranes 45, 92 yttria-stabilized 45

Published on 06 July 2011 on http://pubs.rsc.org | doi:10.1039/9781849733489-00302

View Online