CALCIUM LOOPING PROCESSES FOR CARBON CAPTURE

Transcription

1 CALCIUM LOOPING PROCESSES FOR CARBON CAPTURE DISSERTATION Presented in Partial Fulfillment of the Requirements for the Degree Doctor of Philosophy in the Graduate School of The Ohio State University By Shwetha Ramkumar, B.Tech. Graduate Program in Chemical and Biomolecular Engineering The Ohio State University 2010 Dissertation Committee: Dr. Liang-Shih Fan, Adviser Dr. Bhavik R. Bakshi Dr. Andre F. Palmer xi

2 Copyright by Shwetha Ramkumar 2010 i i

3 ABSTRACT A growing need for the reduction in anthropogenic carbon dioxide (CO 2 ) emission has led to a global push toward the development of efficient, economical, and reliable carbon capture and sequestration technologies (CCS) for application to fossil fuel based power plants. Several options are being investigated for the implementation of CCS on pre-combustion and post-combustion systems including using solvents, sorbents, membranes and chemical looping processes. The calcium looping process (CLP) which is a calcium sorbent based chemical looping process, has the potential to reduce the cost and increase the efficiency of CCS implementation on post-combustion and pre-combustion systems. In the CLP, a regenerable calcium-based sorbent is used to chemically absorb CO 2, sulfur, and halide impurities from synthesis gas or hydrocarbon feedstock during the production of hydrogen(h 2 ) and electricity or only electricity. The removal of CO 2 drives the water-gas shift reaction and hydrocarbon reforming reaction forward via Le Chatelier s principle enabling the production of high-purity H 2. The process operates at high temperature (e.g., C), eliminating the need for a water gas shift catalyst and allowing the exothermic heat of the CO 2 absorption reaction to be recovered for use in generating steam. This significantly reduces the energy penalty associated with ii

4 CO 2 capture. The spent sorbent consisting mostly of calcium carbonate (CaCO 3 ) is heated in a calciner to regenerate calcium oxide (CaO) for reuse in the process and to release a concentrated CO 2 stream that can be dried and sequestered. Overall CO 2 emissions from the process are essentially zero. The regenerated sorbent is reactivated in a hydrator, to eliminate sintering and improve the recyclability of the sorbent, before being reintroduced into the H 2 production reactor. Among various reaction and process factors that are of importance to the CLP, the reactivity and recyclability of the calcium based sorbent are vital. The nature of calcium sorbent sintering that has been observed during multicyclic operation could pose a severe limitation to the commercialization of the process. In realistic calcination conditions, the sorbent loses one third to half of its original reactivity in a single cycle due to calcination at 950 ºC and 1000 ºC respectively. Several methods of improving the recyclability of CaO sorbents have been investigated including sorbent pretreatment, modification by addition of supports and reactivation. Hydration of the sorbent as a reactivation method after every calcination cycle was found to be very effective in improving sorbent performance. The Wt% capture of the sorbent was found to be constant at 50% during multicyclic CO 2 capture with sorbent hydration in every cycle in both bench scale and subpilot scale tests. The CLP for production of H 2 from syngas was investigated and very high purity H 2 was produced with less than 1ppm of hydrogen sulfide (H 2 S) at high temperatures and pressures. For near stoichiometric steam addition, high carbon iii

5 monoxide (CO) conversion and H 2 purity can be obtained at high pressures and an optimal temperature of 600 ºC. At atmospheric pressure, the presence of a water gas shift catalyst with CaO sorbent improves the purity of H 2. At high pressures, typical of commercial deployment, the absence of the catalyst and the reduction of excess steam addition do not have any effect on CO conversion and high H 2 purity is obtained. For a hydrocarbon feed, the steam reforming of the hydrocarbon is integrated with the water gas shift and carbonation reaction in a single reactor. In addition to improving the conversion of the hydrocarbon to H 2, the CLP also provides an efficient mode of internal heat integration where the endothermic energy for the reforming reaction is obtained from the exothermic energy released by the combined water gas shift and carbonation reaction. Single cycle tests have shown that the conversion of methane (CH 4 ) is improved to a large extent by the addition of CaO sorbent at 650 ºC. High purity H 2 is obtained at low steam to carbon(s:c) ratios of 3:1 for various pressures ranging from 1 to 11 atms. The effect of calcination conditions on the extent of CH 4 reforming was determined. The reactivity of the sorbent was found to decrease over multiple cycles due to calcination in both pure nitrogen and in a mixture of steam and CO 2. Hydration was found to be effective in reducing the sintering of the sorbent. System analysis using ASPEN Plus has shown that the CLP has a high efficiency for conversion of both coal as well as natural gas to H 2 and electricity. The CLP is being scaled up to a 25 KW subpilot unit demonstration at the Ohio State University and the unit is currently under construction. The subpilot scale unit design is iv

6 based on the thermodynamic, kinetic and sorbent reactivity studies and cold flow tests. This unit will be used to conduct continuous testing for the production of H 2 from a simulated syngas stream and a mixture of hydrocarbons. v

7 Dedicated to my parents and sister for their love and support. vi

8 ACKNOWLEDGMENTS I would like to sincerely express my gratitude toward my adviser, Professor Liang-Shih Fan, for his invaluable guidance, support and encouragement throughout my graduate education at The Ohio State University. His perseverance, enthusiasm and quest for knowledge has been a constant source of motivation. I am indebted to him for the trust he confided in me and the time he spent with me, enriching me with his experiences that help me mature as a professional and an individual. I am also grateful to Professors Bhavik R. Bakshi, Kurt W. Koelling, David L. Tomasko and Andre F. Palmer for serving in my qualifier, candidacy, and dissertation committees, and thereby providing valuable suggestions and comments in this research study. I would also like to take this opportunity to thank all faculty in this department, especially for their support and encouragement. I would like to specifically thank Dr. Mahesh Iyer who served as my mentor throughout this study. The insightful discussions with him formed an invaluable part of this study. I would like to thank Dr. Robert Statnick, an invaluable member of our extended research team, for his guidance, support and motivation. His vast industry experience was immensely helpful in developing this study. I would like to thank Dan vii

9 Connell from CONSOL Energy for his support with the techno-economic evaluations. I would like to thank the members of my research team: Danny Wong, William Wang for all their support with all the post combustion work. I would also like to thank Nihar Phalak and Niranjani Deshpande for all their support during the later part of the study. It was the tremendous team effort of this group that helped in the development of the overall calcium looping process. I would like to thank Dr. Alissa Park for her mentoring and support. She has been a great friend and source of inspiration. I would also like to thank the members of our research group Dr. Songgeng Li., Fuchen Yu, Zhenchao Sun, Siddharth Gumuluru, Fanxing Li, Zhao Yu, Fei Wang, Deepak Sridhar, Ray Kim, Andrew Tong, Liang Zeng, Dr. Puneet Gupta, and Dr. Luis Velazquez-Vargas for their support and friendship. Finally, I greatly enjoyed working with the undergraduate research assistants: Brittany Valentine, Jessica Huber, Theresa Vonder-Haar, Eric Sacia, and Brian Stelzer and cherish their friendship. My sincere thanks to Lynn Flanagan, Amy Dudley, Susan Tesfai, Angela Jones, Kari Uhl, Bill Cory and Paul Green for helping me in administrative and other fronts. I would like to thank my friends at OSU, for all their warmth and hospitality that made my stay in Columbus a wonderful experience to cherish forever. I am indebted to the Ohio Coal Development Office (OCDO) of the Ohio Air Quality Development Authority (OAQDA) and the US Department of Energy for viii

10 providing financial assistance throughout this study. My special gratitude goes to Mr. Bob Brown and Mr. Dan Cicero for providing useful suggestions. Most importantly, I would like to thank my parents for their unconditional love, and support and for their faith in me. Without their constant encouragement and motivation this study would not have been possible. I would like to thank my sister, Shmita Ramkumar, who is my biggest source of strength and inspiration, for her love and support. ix

11 VITA June 21, Born Chennai, India July 2001 June B.S. Chemical Engineering Anna University, Chennai, India September 2005 present.. Graduate Research Associate Chemical and Biomolecular Engineering The Ohio State University Columbus, OH, USA PUBLICATIONS 1. Wang W., Ramkumar S., Li S., Wong D., Iyer M.V., Gumuluru S., Sakadjian B, Statnick R.M., and Fan L-S Sub-Pilot Demonstration of the Carbonation- Calcination Reaction (CCR) Process: High Temperature CO 2 and sulfur capture from Coal Fired Power Plants, Ind. Eng. Chem. Res. (2010). 2. Fan L-S., Li F., and Ramkumar S. Utilization of chemical looping strategy in coal gasification processes. Particuology, 6(3), (2008) 3. Ramkumar S., Li S., Wang W., Gumuluru S., Sun Z., Phalak N., and Fan L.-S. Results from the Carbonation-Calcination Reaction (CCR) Process, Proc. 26 th Intl. Pittsburgh Coal Conf., Pittsburgh, PA, September (2009) 4. Ramkumar S., and Fan L.-S. Calcium Looping Process for Clean Fossil Fuel Conversion, Proc. 26 th Intl. Pittsburgh Coal Conf., Pittsburgh, PA, September (2009) x

12 5. Ramkumar S., Phalak N., Sun Z., and Fan L.-S. Calcium Looping Process Enhanced Coal to Liquid Technology, Proc. 26 th Intl. Pittsburgh Coal Conf., Pittsburgh, PA, September (2009) 6. Ramkumar S., Connell D., and Fan L-S. Calcium Looping process for clean fossil fuel conversion 1st Meeting of the High Temperature Solid Looping Cycles Network, Oviedo, Spain, September (2009) 7. Ramkumar S., Wang W., Li S., Gumuluru S., Sun Z., Phalak N., Wong D., Iyer M., Statnick R.M., Fan L-S., Sakadjian B., and Sarv H. Carbonation-Calcination Reaction(CCR) Process for High Temperature CO2 and Sulfur Removal 1st Meeting of the High Temperature Solid Looping Cycles Network, Oviedo, Spain, September (2009) 8. Ramkumar S., and Fan L-S., Calcium looping process for clean fossil fuel conversion, 8th World Congress of Chemical Engineering, Montreal, Canada, August (2009) 9. Ramkumar S., Wang W., Li S., Wong D., Iyer M.V., Gumuluru S., Statnick R.M., and Fan L-S., Carbonation-Calcination Reaction Process for High Temperature CO 2 and sulfur Removal, 8th World Congress of Chemical Engineering, Montreal, Canada, August (2009) 10. Sakadjian B., Wang W., Li S., Ramkumar S., Gumuluru S., Fan L-S., and Statnick R.M., Sub-Pilot Demonstration of the CCR Process: High Temperature CO 2 Capture Sorbents for Coal Fired Power Plants Proc. Int. Tech. Conf. Coal Utilization & Fuel Systems, Clearwater, FL (2009) 11. Fan L.-S., Li F., Velazquez-Vargas L.G., and Ramkumar S. Chemical Looping Gasification. Proc. 9th International Conference on Circulating Fluidized Beds. Hamburg, Germany. May (2008). FIELDS OF STUDY Major Field: Chemical Engineering xi

13 TABLE OF CONTENTS Page Abstract...ii Dedication... vi Acknowledgments... vii Vita......ix List of Tables..xvi List of Figures....xviii Chapters: Chapter Introduction Chapter Literature Review: Processes for Enhanced H 2 Production with CO 2 capture Introduction CO 2 Acceptor Process HyPr-RING Process Zero Emission Coal Alliance (ZECA) Process ALSTOM Hybrid Combustion-Gasification Process Fuel-Flexible Advanced Gasification-Combustion Process Chapter Reactivity and Recyclability of Calcium Based Sorbents for CO 2 Capture Introduction Sorbent Reactivity Over Multicyclic Reactions Synthesis of High Reactivity Precipitated Calcium Carbonate (PCC) Sorbent Pretreatment of Calcium Based Sorbents and Addition of Supports Reactivity Testing of Ca-based Sorbents for CO 2 Capture xii

14 3.4.2 Recyclability of Natural, Pretreated and Supported Sorbents Effect of Realistic Calcination Conditions on Sorbent Reactivity Experimental Methods Results and Discussion Sorbent Reactivation by Hydration Lab Scale Testing Experimental Methods Results and Discussion Sub-Pilot Scale Demonstration of Reactivation of Calcium Sorbent by Hydration Experimental Methods for the 120 KWth Subpilot Scale Testing Results and Discussion Conclusions.. 58 Chapter Enhanced Catalytic H 2 Production from Syngas Introduction Calcium Looping Process(CLP) Configuration and Thermodynamics The Carbonation Reactor The Calciner Sorbent Reactivation by Hydration Materials and methods Chemicals, Sorbents, and Gases Fixed Bed Reactor Unit Setup Water Gas Shift Reaction Testing Simultaneous Water Gas Shift and Carbonation Catalyst Pretreatment Combined H 2 Production with H 2 S Removal Results and Discussion Effect of Process Parameters on the Extent of Water Gas Shift Reaction using HTS Catalyst Enhancing the Water Gas Shift Reaction by In-situ CO 2 Removal (HTS Catalyst and CaO Sorbent) Simultaneous Water Gas Shift, Carbonation and Sulfidation Reaction Testing Effect of Catalyst Type on the Water Gas Shift Reaction Conclusions.. 99 Chapter Enhanced Non-Catalytic H 2 production from Syngas Introduction Materials and Methods Chemicals, Sorbents, and Gases Experimental Setup: Fixed Bed Reactor. 130 xiii

15 5.2.3 Water Gas Shift Reaction in the Presence and Absence of HTS Catalyst Simultaneous Water Gas Shift and CO 2 Removal Combined H 2 Production with H 2 S Removal Results and Discussion Baseline Water Gas Shift Reaction Testing Water Gas Shift Reaction in the Presence of Only CaO Sorbent H 2 Production in the Presence of CaO Sorbent Only and a Mixture of CaO Sorbent and Catalyst Multicyclic Investigation of H 2 Production in the Presence of CaO Sorbent Only Enhanced H 2 Production With CO 2 and Sulfur Capture H 2 Production From Coal Gasification Derived Syngas Process Overview System Thermodynamics Analysis H 2 Production From Syngas Derived From Natural Gas Feedstocks Syngas from Steam Reforming of Natural Gas Syngas from Autothermal Reforming of Natural Gas Syngas from Partial Oxidation of Natural Gas Addressing The Issue of Sulfur in the Feedstock Experimental Analysis of the Regeneration of CaS Conclusion. 158 Chapter Process Simulation and Economics of the Calcium Looping Process (CLP) for production of h 2 from Coal Introduction Production of Fuel Cell Grade H 2 With a PSA Cogeneration of H 2 and Electricity Production of Only H 2 With Internal Heat Integration Production of H 2 Having a Purity of 94 98% Without a PSA Cogeneration of H 2 and Electricity Production of H 2 With Internal Heat Integration Comparison of the Process Efficiencies for Different Gasifiers Effect of Process Parameters on CLP Performance Using Syngas From a GE Gasifier Approach Sensitivity Analysis for the Yield and Purity of H 2 Produced Sensitivity Analysis for the Extent of Contaminant Removal from the Product H Sensitivity Analysis for the Cold Gas Efficiency and Overall Process Efficiency Effect of Addition of Sorbent Hydration to the CLP Process 219 xiv

16 6.7 Techno-Economic Analysis of H 2 Production From Coal Conclusions Chapter Enhanced Reforming of Hydrocarbons Introduction Process Configuration and Thermodynamics The Carbonation Reactor System Calciner or Sorbent Regeneration Reactor Hydrator or Sorbent Reactivation Reactor Experimental Methods Chemicals, Sorbents, and Gases Bench Scale Experiment Setup Steam Methane Reforming in the Presence of a Ni-based Catalyst Simultaneous Steam Methane Reforming, Water Gas Shift and Carbonation Multicyclic Steam Methane Reforming and Spent Sorbent Calcination Results and Discussion Base-line Steam Methane Reforming Testing Simultaneous Reforming with In-situ CO 2 Removal (Catalyst with CaO Sorbent) Effect of Sorbent Calcination Conditions on the Extent of Steam Reforming Calcination in N 2 with Sorbent Hydration Realistic Sorbent Calcination in a Steam/CO 2 Atmosphere with Sorbent Hydration Applications of CLP in Hydrocarbon Reforming Steam Reforming of Natural Gas and Other Hydrocarbons for H 2 and Electricity Generation Implementation of Carbon Capture in Liquid Fuels Production From Coal Conclusions 293 Chapter Subpilot scale testing and recommendations for future work Introduction Cold Flow Testing Design of the Subpilot Scale Unit Conclusions Recommendations for Future Work Appendix - A xv

17 LCA Analysis - Comparison of the conventional coal to H 2 process with the CLP process References xvi

18 LIST OF TABLES Table Page Table 2.1: A typical composition of the H 2 rich synthesis gas from the gasifier Table 5.1: Typical fuel gas compositions obtained from different gasifiers (Stultz and Kitto, 1992) Table 5.2: Fuel gas composition entering the water gas shift reactor after steam addition (S:C ratio =1:1) (adapted from Stultz and Kitto, 1992) Table 5.3: Fuel gas composition entering the water gas shift reactor after steam addition (S:C ratio =3:1) (adapted from Stultz and Kitto, 1992) Table 5.4: Extent of equilibrium CO conversion and CO 2 capture in the CLP from Steam Methane Reforming (SMR) derived syngas Table 5.5: Extent of equilibrium CO conversion and CO 2 capture in the CLP from Auto Thermal Reforming (ATR) derived syngas Table 5.6: Extent of equilibrium CO conversion and CO 2 capture in the CLP from partial oxidation (POX) derived syngas Table 6.1: Properties of Illinois # 6 coal Table 6.2: Composition of the syngas exiting from the Shell gasifier Table 6.3: Intermediated pressures for compression of the CO 2 for sequestration Table 6.4: Components list for the ASPEN Plus simulation Table 6.5: ASPEN Plus models used for the simulation of the CLP Table 6.7: Power balance in the CLP process Table 6.8: Process simulation results for the CLP process xvii

19 Table 6.9 Summary of the schemes investigated for the production of H 2 alone and for the coproduction of H 2 and electricity with a PSA Table 6.10 Summary of the schemes investigated for the production of H 2 alone and for the coproduction of H 2 and electricity without a PSA Table 6.11 Comparison of the efficiency of the H 2 production process for different gasifiers Table 6.12: Levelized annual costs and levelized cost of H 2 for the conventional coal to H 2 plant (adapted from DOE, 2010) Table 6.13: Levelized annual costs and levelized cost of H 2 for the CLP plant Table 7.1: Thermodynamic extent of the various reactions occurring in the carbonator Table 7.2: Stream data for the integration of the CLP in a steam methane reforming process Table 7.3: Energy balance for the production of H 2 and electricity from natural gas using the CLP Table 7.4: Heat required for the production of steam at 650 ºC from water at 15 atms Table 7.5: Heat required for preheating the natural gas at 15 atms to 650 ºC Table 7.6: Heat released by the solids from the calciner in the H 2 production reactor Table 7.7: Heat generated from the H 2 production reactor Table 7.8: Heat released on cooled the H 2 from 650 ºC to ambient temperature Table 7.9: Heat required for preheating the PSA tail gas from 650 ºC to 900 ºC Table 7.10: Heat required for preheating the oxygen from ambient temperature to 900 ºC Table 7.11: Heat absorbed by the solids from the H 2 production reactor in the calciner Table 7.12: Heat released from the calciner Table A.1: Energy balance for the conventional process xviii

20 Table A.2: Water balance for the conventional process Table A.3: Quantification of the inputs for the conventional process Table A.4: Quantification of outputs for the conventional process Table A.5: Energy balance for the CLP Table A.6: Water balance for the CLP Table A.7: Quantification of the inputs for the CLP Table A.8: Quantification of the outputs for the CLP Table A.9: Quantification of the Global Warming Potential(GWP) for the inputs and outputs for the conventional process Table A.10: Quantification of the Global Warming Potential(GWP) for the inputs and outputs for the CLP xix

21 LIST OF FIGURES Table Page Figure 1.1: Historical data and projections of the world energy consumptions till 2030 (EIA, 2009) Figure 1.2: Projections of the world energy supply by different fuel types including fossils fuels and renewable (EIA, 2009) Figure 1.3: Implementation of Carbon Capture and Sequestration (CCS) in fossil fuel based power plants Figure 2.1: Schematic diagram of the reactor system in the gas synthesis block of the CO 2 Acceptor process (Dobbyn et al., 1978) Figure 2.2: Schematic diagram of the HyPr-RING process Figure 2.3: Schematic of the ZECA process Figure 2.4: Schematic of the ALSTOM process Figure 2.5: Schematic of the GE process Figure 3.1: Comparison in the CO 2 capture capacity of CaO sorbents obtained from different precursors. (Calcination conditions: T = 700 ºC, P = 1 atm, pure N 2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) Figure 3.2: Comparison in the multicyclic conversion of PCC powder sorbent PCC pelletized and broken sorbent (Calcination conditions: T = 700 ºC, P = 1 atm, pure N 2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) xx

22 Figure 3.3: CO 2 capture capacity of pretreated and supported Ca-based sorbents over multiple carbonation calcination cycles (Calcination conditions: T = 700 ºC, P = 1 atm, pure N 2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) Figure 3.4: Effect of steam concentration in the calcination carrier gas on the CO 2 capture capacity of CaO sorbent (Calcination conditions: T = 900 ºC, P = 1atm) Figure 3.6: Effect of steam calcination on multicyclic carbonation and calcination of CaO sorbent (Calcination conditions: T = 900 ºC, P = 1 atm, carrier gas = 50%H2O/50% CO 2 ; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) Figure 3.7: Effect of hydration conditions on sorbent reactivity Figure 3.8: Effect of hydration pressure on sorbent reactivity (Hydration temperature = 600 ºC) Figure 3.9: Effect of steam hydration on sorbent reactivity over multiple calcinationhydration-carbonation cycles (Calcination conditions: T = 900 ºC, P = 1 atm, carrier gas = pure CO 2 ; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas, Hydration conditions: T = 500 ºC, P ~ 1 atm, 90% H 2 O/10% N 2 feed gas) Figure 3.10: Process flow diagram of the CLP for CO 2 and SO 2 removal from combustion flue gas Figure 3.11: Snapshot of the sub-pilot scale facility of the CLP integrated with a coal fired combustor Figure 3.12: Effect of hydration on the % CO 2 removed from the flue gas over multiple cycles Figure 3.13: Wt.% CO 2 capture achieved by the hydrated sorbent over multiple cycles Figure 4.1: Schematic of the CLP xxi

23 Figure 4.2: Thermodynamic data illustrating the equilibrium constants of the water gas shift reaction and the combined water gas shift and carbonation reaction Figure 4.3: Thermodynamic data for the hydration and carbonation of CaO sorbent. 102 Figure 4.4: Equilibrium H 2 purity in the carbonator at varying temperatures, pressures and S: C ratios. ( Feed gas: 10% CO and balance nitrogen) Figure 4.5: Thermodynamic data for the sulfidation (H 2 S) of CaO with varying steam partial pressures. (P Total = 30 atm) Figure 4.6: Thermodynamic data for predicting the equilibrium COS concentration for CaO sulfidation with varying CO 2 concentration (P Total = 30 atm) Figure 4.7: Thermodynamic data for predicting the equilibrium HCl concentration for CaO reaction with HCl with varying steam concentration (P Total = 30 atm) 106 Figure 4.8: Thermodynamic data for the carbonation of CaO Figure 4.9: Thermodynamic data for the hydration of CaO Figure 4.10: Simplified flow sheet of the bench scale experimental setup Figure 4.11: X-ray diffraction patters of the HTS catalyst before pretreatment (hematite) Figure 4.12: X-ray diffraction patters of the HTS catalyst after pretreatment (magnetite) Figure 4.13: Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction in the presence of HTS catalyst at (a) 1 atm (b) Figure 4.14: Effect of reaction temperature and pressure on the observed partial pressure ratio for the water gas shift reaction in the presence of HTS catalyst at a S:C ratio of (a)1:1 (b)3: Figure 4.15: Typical curves for the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst depicting (a) Gas composition (mol%) and (b) CO conversion (650 ºC, 1 atm, S:C ratio of 3:1) xxii

24 Figure 4.16: Effect of pressure on purity of H 2 produced during the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at a S:C ratio of (a) 3:1 (b) 1:1 (650 ºC) Figure 4.17: Effect of S:C ratio on the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 650 ºC (a) CO conversion at 1 atm (b) H 2 gas composition at 1 atm (c) CO conversion at 21 atm (d)h 2 gas composition at 21 atm Figure 4.18: Effect of temperature on CO conversion by the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 1 atm and S:C ratio of 3: Figure 4.19: Effect of temperature on CO conversion by the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 21 atm and S:C ratio of (a) 3:1 (b) 1: Figure 4.20: Effect of S:C ratio on (a) the composition of H 2 S in the H 2 stream and (b) CO conversion in the presence of the catalyst and sorbent during the simultaneous water gas shift, carbonation and sulfidation reaction(600 ºC, 1 atm) Figure 4.21: Effect of S:C ratio on the composition of H 2 S in the H 2 stream during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent and HTS catalyst (600 ºC, 1 atm) Figure 4.22 : Effect of S:C ratio and temperature on CO conversion during the water gas shift reaction in the presence of STC and HTS catalyst Figure 4.23: Effect of reaction temperature on CO conversions for various pressures at an S:C ratio of 1:1 for the STC (0.25g STC, Total flow = slpm) Figure 4.24: Effect of reaction temperature on CO conversions for the HTS and STC at 11 atms and S:C ratio of 1:1(Total flow = slpm) Figure 4.25: Effect of reaction temperature on CO conversions for the HTS and STC at 21 atms and S:C ratio of 1:1(Total flow = slpm) xxiii

25 Figure 4.26: Effect of S:C ratio, type of catalyst and effect of H 2 S on CO conversion during the water gas shift reaction(650 ºC, 1atm) Figure 4.27: Effect of temperature on CO conversion (Temperature=650 C, Pressure = 1 atm, S:C ratio= 1:1) Figure 4.28: Comparison in the CO conversion obtained at different S:C ratios for different sorbent and catalyst mixtures (650 ºC, 1atm) Figure 5.1: Simplified flow sheet of the bench scale experimental setup Figure 5.2: Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction at 1 atm Figure 5.3: Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction at 21 atm Figure 5.4: Typical breakthrough curves for the production of H 2 in the presence of CaO sorbent without catalyst (a) Gas composition (mole%) and (b) CO conversion (600 C, 21 atm, S:C ratio of 3:1) Figure 5.5: Effect of pressure on CO conversion obtained in the presence of CaO sorbent without catalyst (650 C, S:C ratio of 3:1) Figure 5.6: Effect of S:C ratio on CO conversion obtained in the presence of CaO sorbent without catalyst at (a) 1 atm, (b) 11 atm, (c) 21 atm (650 C) Figure 5.7: Effect of temperature on CO conversion obtained in the presence of CaO sorbent without catalyst at a S:C ratio of (a) 1:1 and (b) 3:1 (1 atm) Figure 5.8: Effect of CO concentration in the feed on the (a) CO conversion and (b) purity of H 2 produced in the presence on CaO sorbent without catalyst (11 atm, 600 C, S:C ratio of 3:1) Figure 5.9: SEM image of the (a) initial CaCO 3 sorbent (b) CaO sorbent obtained from the calcination of CaCO Figure 5.10: SEM image of sorbent at the end of the water gas shift and carbonation reaction in the absence of a catalyst at (a) 1 atm (b) 21 atm (S:C ratio of 3:1, 600 C) xxiv

26 Figure 5.11: Comparison in the product H 2 purity in the presence of the sorbent and in the presence of the sorbent and catalyst mixture at 1 atm (650 C, S:C ratio of 1:1) Figure 5.12: Comparison in the product H 2 purity in the presence of the sorbent and in the presence of the sorbent and catalyst mixture (650 C, 21 atm) Figure 5.13: Product H 2 purity obtained over multiple reaction and regeneration cycles in the presence of CaO sorbent without catalyst at 4.5 atms. (600 C, S:C ratio of 3:1) Figure 5.14: Product H 2 purity obtained over multiple reaction - regeneration cycles in the presence of CaO sorbent without catalyst at 21 atms (600 C, S:C ratio of 3:1) Figure 5.15: Effect of S:C ratio on the (a) extent of H 2 S removal and (b) the purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent (1atm, 600 o C) Figure 5.16: Effect of temperature on the (a)extent of H 2 S removal and (b) purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent (1 atm, S:C ratio of 1:1) Figure 5.17: Effect of pressure on the (a) extent of H 2 S removal (b) purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent (S:C ratio of 1:1, 600 o C) Figure 5.18: SEM image of the (a) initial CaCO 3 sorbent (b) CaO sorbent obtained from the calcination of CaCO 3 (c) sorbent at the end of the water gas shift, carbonation and sulfidation reaction at 1 atm (c) CaO sorbent obtained from the calcination of CaCO 3 (600 o C, S:C ratio of 1:1) (c) sorbent at the end of the water gas shift, carbonation and sulfidation reaction at 21 atm (600 o C, S:C ratio of 1:1) Figure 5.19: (a) Conventional process for H 2 production from coal (b) Integration of the CLP in a conventional process for H 2 production from coal xxv

27 Figure 5.20: Integration of the CLP in a coal gasification system for the production of electricity, H 2 and liquid fuels Figure 5.21: Comparison of the PCO 2 in the carbonator with the equilibrium PCO 2 for the carbonation of CaO for a S:C ratio of (a)1:1 (b)3: Figure 5.22: Comparison of the PH 2 O in the carbonator with the equilibrium PH 2 O for the hydration of CaO for a S:C ratio of (a)1:1 (b)3: Figure 5.23: Effect of temperature on equilibrium CO conversion in the water gas shift reactor at a S:C ratio of (a) 1:1 (b) 3: Figure 5.24: Effect of temperature on equilibrium CO conversion in the presence of CaO in the carbonation reactor of the CLP at a S:C ratio of (a) 1:1 (b) 3: Figure 5.25: Effect of temperature on equilibrium H 2 purity in the presence of CaO at a S:C ratio of (a) 1:1 (b) 3: Figure 5.26: Effect of temperature and S:C ratio on the % of carbon captured in the CLP using syngas from different gasifiers as the feed Figure 5.27: Conventional steam reforming of natural gas for H 2 production with a methanator Figure 5.28: Conventional steam reforming of natural gas for H 2 production with a PSA Figure 5.29: CLP integrated in the conventional steam reforming of natural gas process Figure 5.30: Conventional partial oxidation process for conversion of natural gas to H Figure 5.31: CLP integrated in the partial oxidation of natural gas for H 2 production 200 Figure 5.32: Effect of the change in temperature and steam composition on the regeneration of CaS with H 2 O Figure 5.33: Effect of the change in steam and CO 2 composition on the regeneration of CaS in the presence of H 2 O and CO Figure 5.34: H 2 S evolved in the presence of H 2 O and CO 2 from spent sorbent produced during combined CO 2 and H 2 S removal at 1 and 21 atms xxvi

28 Figure 6.1: The CLP for coproduction of fuel cell grade H 2 and electricity from coal Figure 6.2: ASPEN simulation flow diagram for the CLP process with a PSA Figure 6.3 Aspen simulation for the production of H 2 using the CLP without a PSA. 246 Figure 6.4: Aspen model used for sensitivity analysis of the combined reactions occurring in the H 2 production reactor of the CLP Figure 6.5: Effect of temperature on the H 2 purity produced at the outlet of the carbonation reactor (S:C ratio = 3, Pressure = 10 atms) Figure 6.6: Effect of pressure on the H 2 purity produced at the outlet of the carbonation reactor( S:C ratio = 3, Temperature = 600 ºC) Figure 6.7: Effect of S:C ratio on the H 2 purity produced at the outlet of the carbonation reactor ( Pressure = 10 atms, Temperature = 600 ºC) Figure 6.8: Effect of temperature and S:C ratio on the extent of H 2 S removal Figure 6.9: Effect of temperature and S:C ratio on the extent of COS removal Figure 6.10: Effect of temperature and S:C ratio on the amount of CO impurity present in the H 2 stream Figure 6.11: Effect of temperature and S:C ratio on the extent of CO 2 removal Figure 6.12: Effect of temperature and S:C ratio on the amount of CH 4 impurity present in the H 2 product stream Figure 6.13: Effect of pressure on the cold gas efficiency, process efficiency and H 2 purity obtained from the H 2 production reactor at various S:C ratios Figure 6.14: Effect of S:C ratio on H 2 purity, cold gas efficiency and process efficiency Figure 6.15: Effect of temperature on H 2 purity, cold gas efficiency and process efficiency (1:1, 10 atms) Figure 6.16: Effect of Ca:C ratio on H 2 purity, cold gas efficiency and process efficiency (600 ºC, 1:1, 10 atms) Figure 6.17: Effect of the addition of sorbent hydration to the CLP xxvii

29 Figure 6.18: Process flow diagram of the conventional coal to H 2 plant used for the economical analysis ( DOE, 2010) Figure 6.19: Process flow diagram of the CLP plant used for the economical analysis Figure 7.1: Schematic of the CLP for the conversion of hydrocarbons to H Figure 7.2: Thermodynamic data illustrating the equilibrium constants of the steam reforming of CH 4, water gas shift and carbonation reaction Figure 7.3: Simplified schematic of the bench scale experimental setup Figure 7.4: Effect of temperature and S:C ratio on (a)h 2 purity and (b) the amount of CO, CO 2 and CH 4 remaining in the product gas for the steam methane reforming reaction in the presence of Ni-based catalyst ( P = 1 atm) Figure 7.5: Breakthrough curve in the composition of the product gases obtained during the simultaneous reforming, water gas shift and carbonation reaction. (T = 650 ºC, P = 1 atm) Figure 7.6: CH 4 conversion obtained during the simultaneous reforming, water gas shift and carbonation reaction. (T = 650 ºC, P = 1 atm) Figure 7.7: Effect of temperature and S:C ratio on (a) H 2 purity (b) conversion of CH 4 (P = 1atm) Figure 7.8: Effect of temperature and S:C ratio on the amount of (a) CO and (b) CO 2 remaining in the product gas for H 2 production from methane with/without sorbent. ( P = 1 atm) Figure 7.9: Effect of pressure on (a) H 2 purity and (b) CH 4 concentration in the product stream. (T = 650 ºC, S:C ratio = 3) Figure 7.10: Effect of pressure on (a) CO 2 and (b) CO concentration in the product stream. (T = 650 ºC, S:C ratio = 3) Figure 7.11: Effect of pressure on the prebreakthrough and postbreakthrough concentration of CH 4, CO and CO 2 in the product stream. (T = 650 ºC, S:C ratio = 3) xxviii

30 Figure 7.12: Effect of calcination conditions on (a) H 2 purity and (b) CH 4 composition in the product gas for cycles 1,2,3 and 4. [(Reforming reaction conditions :T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2 and 3 are calcined in pure N 2 at 950 and the sorbent for cycle 4 is calcined in a 50:50 CO 2 /H 2 O atmosphere at 950 ºC.)] Figure 7.13: Effect of hydration on (a) H 2 purity and (b) CH 4 composition in the product gas for cycles 1, 2, 3 and 4. [(Reforming reaction conditions :T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in pure N 2, T = 950, P = 1 atm)(hydration conditions: hydration of calcined sorbent from the 3 rd cycle in a 80:20 H 2 O/N 2 atmosphere, T = 600, P = 11 atm)] Figure 7.14: Effect of hydration on H 2 purity for cycles 1,2,3 and 4. [(Reforming reaction Conditions: T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in pure N 2, T = 950, P = 1 atm)(hydration conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20 H 2 O/N 2 atmosphere, T = 600, P = 11 atm)] Figure 7.15: Effect of hydration on (a) H 2 purity and (b) CH 4 content in the product gas for cycles 1,2,3 and 4. [(Reforming reaction Conditions: T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in a 50:50 CO 2 /H 2 O atmosphere, T = 950, P = 1 atm) (Hydration conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20 H 2 O/N 2 atmosphere, T = 600, P = 11 atm)] Figure 7.16: Integration of the CLP in a natural gas reforming system Figure 7.17: Detailed schematic for H 2 production from natural gas Figure 7.18: Conventional CTL plant Figure 7.19: Integration of the CLP in a CTL plant in two configurations Figure 7.20: Integration of the CLP in a CTL plant configuration Figure 7.21: Integration of the CLP in a CTL plant configuration Figure 8.1: Standard deviation of pressure in the fluidized bed xxix

31 Figure 8.2 (a): Schematic diagram of the cold flow model for the CLP Figure 8.2 (b): Snapshot of the cold flow model for the CLP Figure 8.3: Cold flow model for the hydrator Figure 8.4: Schematic of the subpilot scale unit being constructed at OSU for testing the Calcium Looping Concept for H 2 production Figure 8.5: Sorbent hopper and screw feeder Figure 8.6: Water cooled heat exchanger Figure A.1: Schematic of a conventional gasification plant for the cogeneration of H 2 and electricity Figure A.2: Input-output diagram for the conventional coal to H 2 process Figure A.3: Schematic of a CLP plant for the cogeneration of H 2 and electricity Figure A.4: Flow sheet developed for the CLP using ASPEN plus simulator Figure A.5: Input-output diagram for the coal to H 2 process using the CLP xxx

32 CHAPTER 1 INTRODUCTION The world energy demand, as shown in Figure 1.1, is projected to increase by 40% at a rate of 1.5% per year from 2007 to 2030 (EIA, 2009). Although the energy generation from renewable resources is projected to grow, as illustrated in Figure 1.2, fossil fuels are still projected to contribute a major portion of the energy needs in the near future (EIA, 2009). A growing need for the reduction in anthropogenic carbon dioxide (CO 2 ) emission has led to a global push towards the development of efficient, economical, and reliable carbon capture and sequestration technologies (CCS) for application to fossil fuel based power plants. The implementation of CO 2 capture in fossil fuel based systems could be through post-combustion capture, oxy-combustion and pre-combustion capture as illustrated in Figure 1.3. Post-combustion capture technology involves the combustion of coal or natural gas to produce hot flue gas which is used to generate steam. The CO 2 from the flue gas is then captured. The capture of CO 2 from flue gas results in a large increase in parasitic energy and cost of electricity (COE) due to the large volumes of flue gas and the low concentration of CO 2 (13-14%) for coal combustion and 3-4% for natural gas 1

33 combustion). In oxy-combustion, the fuel is burnt in oxygen and recycled flue gas, to produce a concentrated stream containing CO 2 and steam which is then dried, compressed and transported for sequestration. Although oxy-combustion obviates the need for a separate CO 2 capture stage, it requires an Air Separation Unit (ASU) which is energy intensive and expensive. Pre-combustion capture involves the gasification of coal or the reforming of natural gas to produce syngas. The syngas is then cleaned and sent to shift reactors to convert the carbon monoxide (CO) to hydrogen (H 2 ) and CO 2 in the presence of steam. Downstream of the shift reactors, the CO 2 is removed using solvents like amines, rectisol, selexol, etc., and the H 2 stream is further purified in a Pressure Swing Adsorber (PSA) for high H 2 purity applications. The application of CCS to gasification systems has been found to be more efficient and economical when compared to CCS for post-combustion systems. It has been estimated that with the implementation of CCS using solvent based systems, the increase in the COE for an Integrated Gasification Combined Cycle (IGCC) will be 25 to 40 % while that for Pulverized Coal (PC) boilers will be 60 to 85%.(MIT, 2007) In a carbon constrained scenario, it has been estimated that the cost of a super critical PC boiler will be $2140/KWe while that of an IGCC will be $1890/KWe (MIT, 2007). In addition to being more economical and efficient, gasification is also very versatile and capable of producing H 2 and liquid fuels in addition to electricity. Several options are being investigated for the implementation of CCS on precombustion systems including using solvents, sorbents, membrane and chemical 2

34 looping processes. The calcium looping process (CLP) which is a calcium sorbent based chemical looping process, has the potential to reduce the cost and increase the efficiency of CCS implementation. (Abanades et al, 2007, Ramkumar et al, 2009) In this study, the CLP concept and application to various feed streams (syngas, natural gas and other hydrocarbons) has been studied using thermodynamic analysis, lab, bench and subpilot scale experimental studies and system analysis and preliminary process economics. Chapter 2 gives an overview of the calcium sorbent based processes that have been developed for the enhanced conversion of fossil fuels to hydrogen and electricity with simultaneous CO 2 capture. In Chapter 3, the study conducted at the Ohio State University on calcium sorbent reactivity and recyclability is detailed. Sorbent modification, pretreatment and reactivation methods to improve recyclability have been described. Sorbent reactivation by hydration has been found to be very effective in maintaining the sorbent reactivity over multiple cycles in the lab, bench and subpilot scale investigations. Chapter 4 describes the thermodynamics and experimental analysis of the application of the CLP to catalytic H 2 production in the presence of a water gas shift catalyst with insitu CO 2 and sulfur capture from syngas. The effect of different water gas shift catalysts and process conditions on the purity of H 2 is discussed. 3

35 Chapter 5 describes the non catalytic production of H 2 by the CLP. The process conditions for which the water gas shift catalyst can be eliminated without a decrease in H 2 purity have been identified. The application of the CLP to syngas produced from a gasification system, steam methane reforming, partial oxidation and autothermal reforming process has been discussed. Chapter 6 discusses the system analysis and techno-economical analysis conducted for the production of H 2 from gasifier derived syngas. Sensitivity analysis for the effect of various process parameters on system efficiency has been conducted. Techno-economic analysis predict that the CLP has a potential to reduce the cost of H 2 and electricity production from gasifier syngas. Chapter 7 describes the application of the CLP to H 2 and electricity production from hydrocarbons. The CLP aids in combining several unit operations including steam reforming of the hydrocarbons, water gas shift reaction and CO 2 capture in a single reactor. The integration of the CLP enhanced steam reforming process to H 2 and electricity production from natural gas and to a coal to liquid fuel (CTL) process has been described. Chapter 8 describes the scaleup of the CLP for H 2 production from bench to subpilot scale at the Ohio State University. Cold flow tests have been conducted and a subpilot scale unit has been designed based on the information discussed in the 4

36 previous chapters. The subpilot scale unit is currently under construction. Chapter 8 also provides recommendations for future work. 5

37 Figure 1.1: Historical data and projections of the world energy consumptions till 2030 (EIA, 2009). 6

38 Figure 1.2: Projections of the world energy supply by different fuel types including fossils fuels and renewable (EIA, 2009). 7

39 Post Combustion Electricity CO 2 free Flue Gas Coal Or Natural Gas Boiler Steam Generator CO 2 Capture Oxy Combustion Coal Or Natural Gas Air ASU Boiler Electricity Steam Generator CO 2 Compression, Transportation and Sequestration Coal or Natural Gas Pre Combustion Gas Gasifier Cleanup Reformer Electricity CO WGSR 2 capture Gas PSA FT Reactor Turbine Steam Hydrogen Turbine Ammonia Synthesis Liquid Hydrogenation Fuels Other chemicals N 2 and Steam synthesis Figure 1.3: Implementation of Carbon Capture and Sequestration (CCS) in fossil fuel based power plants. 8

40 CHAPTER 2 LITERATURE REVIEW: PROCESSES FOR ENHANCED H 2 PRODUCTION WITH CO 2 CAPTURE 2.1 INTRODUCTION Hydrogen can be produced conventionally from coal by the gasification process, natural gas by the steam methane reforming process and higher hydrocarbons by the partial oxidation process. In a typical coal gasification system, the coal is fed along with steam and/or oxygen to the gasifier to produce syngas. The syngas is then cooled using a gas cooler or a water quench. The quench system also provides the excess steam required for the water gas shift reaction.(holt, 2005, MIT, 2007)) While higher temperatures enhance the kinetics of the water gas shift reaction, the equilibrium limitation of the water gas shift reaction adversely affects H 2 production and the H 2 yield falls with rising temperature. Hence, a high steam: CO (S:C) ratio is required to enhance CO conversion and the consequent H 2 yield. The S:C ratio required at 550 o C can be as high as 50 in a single-stage operation or 7.5 for a more expensive dual-stage process to obtain 99.5 % pure H 2. (Haussinger et al, 2000) Numerous research studies have focused on the development of low temperature catalysts to improve H 2 9

41 production. (Haussinger et al, 2000) Commercially, the dual stage sweet water gas shift reaction is carried out in series, with a HTS ( o C) stage containing iron oxide catalyst and a LTS ( o C) stage containing copper catalyst. (Loyd et al, 1996) The commercial iron oxide catalyst has a sulfur tolerance of about several hundred ppms while the copper catalyst has a lower tolerance to sulfur and chloride impurities. (Haussinger et al) Hence syngas clean up is required upstream of the shift reactors which is achieved in conventional scrubbing towers using physical solvents like selexol or, rectisol or chemical solvents like amine based solvents. This low temperature syngas cleanup process is energy intensive due to the gas cooling and reheating requirements. In a sour gas shift system, where the sulfur content of synthesis gas is greater than 1000 ppm, a sulfided catalyst is used in a series of reactors at a temperature of C and the desulfurization unit is located downstream of the water gas shift reactors.(loyd et al, 1996, Hiller et al, 2007) After the shift reaction, the syngas is subjected to scrubbing using solvents to remove the CO 2 and is sent to the PSA unit to produce a pure stream of H 2. The tail gas from the PSA unit is then used as fuel for power generation. Several methods to enhance the purity of H 2 with the simultaneous separation of CO 2 have been cited in literature. A slight advancement in the commercial method of H 2 production has been to remove the CO 2 from the reaction mixture between the two stages of the shift reaction. However solvents operate at ambient temperatures and this method involves severe energy penalties due to cooling and reheating of the 10

42 reaction gas mixture. An effective technique to shift the water gas shift reaction to the right for enhanced H 2 generation has been to remove H 2 from the reaction mixture. This concept has led to the development of H 2 separation membranes. Kreutz et al, 2002 have described the integration of these membranes in a commercial coal gasification unit. (Kreutz et al, 2002) The syngas produced from the gasifier is shifted at a high temperature over a STC followed by a water gas shift H 2 membrane reactor which aids in producing more H 2 and separating it from the gas mixture. (Kreutz et al, 2002) However, ceramic membranes have a very low H 2 permeability and the intermediate temperature composites inspite of having a high H 2 flux are difficult to fabricate and are very susceptible to poisoning. The cermet membranes are superior to the other two classes of membranes but again they are susceptible to poisoning and are expensive. (Roark et al, 2002) Donghao Ma and Carl R. F. Lund (2003) have reported the investigation of a Pd membrane reactor system packed with HTS catalyst. (Ma and Lund, 2003) For optimum performance these reactors require 2 stages and a S:C ratio of 3. These reactors also suffer from inhibition effects of CO 2, which reduces the yield of H 2 from 90% to 50%. (Ma and Lund, 2003) In addition, membranes cannot completely remove H 2 from the mixture and suffer from a considerable pressure drop across them. (Roark et al, 2002) Any remaining H 2 in the main stream would dilute the CO 2 and would lead to poor process economics. High temperature CO 2 membranes have been developed which operate in the 11

43 same temperature range as that of the water gas shift reaction. Although polymeric membranes for the removal of CO 2 from H 2 have been found to have several advantages like simplicity of operation, high energy efficiency and lower cost, most polymers have a poor H 2 /CO selectivity. Hence they are not very effective in shifting the equilibrium of the water gas shift reaction and producing high purity H 2. (Chung et al, 2006) An alternative concept to drive the water gas shift reaction forward has been to remove the CO 2 from the reaction mixture using solid sorbents which either physisorb or react with the CO 2 in the water gas shift reactor. The separation of CO 2 from the reaction mixture at high temperatures removes the equilibrium constraint of the water gas shift reaction and enhances H 2 production. Sorbents that operate at higher temperatures are beneficial to the process as the water gas shift reaction has superior kinetics due to the high temperature and enhanced thermodynamic extent due to the removal of CO 2 from the product gas stream. Faster rates of reaction and larger CO 2 capture capacities allow the use of smaller reactors and require a smaller amount of solids circulation through the system. CaO has a high CO 2 capture capacity and removes CO 2 to ppm levels at a high temperature of 600 ºC making it one of the most suitable sorbents for this application.(gupta and Fan, 2002) The concept of utilizing CaO for CO 2 capture has existed for well over a century. It was first introduced by DuMotay and Marechal in 1869 for enhancing the gasification of coal (Squires, 1967) and followed by CONSOL s CO 2 Acceptor process 12

44 (Curran et al, 1967) a century later when this concept was tested in a 40 t/day plant. A variation of this process, the HyPr-RING process, (Lin et al, 2005, Lin et al, 2002) was developed in Japan for the production of H 2 at high pressures. Several other processes have also been developed to enhance H 2 production using calcium based sorbents such as the ZECA (Ziock et al, 2001), Alstom (Andrus, 2006) and GE process (Rizeq et al, 2001). A detailed description of these processes is provided in the following sections. 2.2 CO 2 ACCEPTOR PROCESS The CO 2 Acceptor process was developed by the Consolidation Coal Company and later Conoco Coal Development Company (Dobbyn et al., 1978). The American Gas Association, U. S. Department of Interior, and Energy Research and Development Administration were among the major sponsors for the development of this process. The CO 2 Acceptor process was designed to produce synthetic pipeline gas from pulverized lignite or sub-bituminous coal. There are five main operational blocks comprising this process, i.e. the feedstock preparation block, the gas synthesis block, the gas cleanup block, the methanation block, and the utility block. Figure 2.1 shows the schematic of the gas synthesis block, where Ca-based sorbent circulating between two fluidized bed reactors a gasifier and a regenerator operating at high pressures of ~11 atms(150 psig) aids in the conversion of coal to H 2. (Curran et al, 1967) In the reactor system of the gas synthesis block shown in Figure 2.1, preheated coal is ground to ~150 μm before it is fed to the gasifier. The gasifier is a fluidized bed 13

45 with steam as the fluidizing gas. It is operated at 800 ~ 850 ºC and 10 atm. The relatively low temperatures in the gasifier enables the use of high sodium coal which tends to fuse forming silica and alumina solid aggregates at high temperatures (>870 ºC). The acceptor is the calcined limestone or dolomite sorbent, which is fed at the top of the gasifier. The sorbent reacts with CO 2 produced from the combined steamcarbon reaction and water-gas shift reaction. The exothermic carbonation reaction of the sorbent provides the heat for the endothermic steam-carbon reaction and also drives the water-gas shift reaction towards the forward direction, thereby increasing the H 2 content in the product gas. The major reactions that occur in the gasifier include: Steam Carbon Reaction: C + H 2 O CO+H 2 (2.1) Water Gas Shift Reaction: CO + H 2 O CO 2 + H 2 (2.2) Carbonation: CaO + CO 2 CaCO 3 (2.3) CaO.MgO + CO 2 CaCO 3.MgO (2.4) The H 2 rich gas obtained from the gasifier, containing ~20% CO and CO 2, is subsequently quenched, purified, methanated, and transported by pipelines. The spent sorbent discharged from the bottom of the gasifier is then fed to the regenerator where the spent sorbent is regenerated at 1010 ºC and 10 atm through the calcination reaction: Calcination: CaCO 3 CaO + CO 2 (2.5) The heat for calcination is provided by combusting the residual char, which is discharged from the gasifier to the regenerator. The regenerator is also a fluidized bed. 14

46 Air is used as both fluidizing and oxidizing gas for the regenerator. The regenerated sorbent is sent back to the gasifier to complete the loop. The CO 2 containing flue gas is generated in the regenerator. The CO 2 Acceptor process was proved to be technically feasible after being successfully tested in a pilot plant facility located at Rapid City, South Dakota. The designed capacity of the pilot plant was 40 ton/day. Over a period of six years that ended in 1977, an accumulative operation of 13,000 hours was achieved with the longest continuous run of ~2300 hours. A total of 6500 tons of dry coal of various ranks including three North Dakota lignites, one Texas lignite and three subbituminous coals were tested in the facility. As noted, the gas synthesis reactor system, which produces H 2 rich gas from coal with the aid of the calcium based CO 2 sorbent or acceptor, characterizes the key innovation of the CO 2 Acceptor process. The detailed description of a pilot scale gas synthesis system, which includes the regenerator and the gasifier, is given below. Regenerator The regenerator decomposes the spent sorbent from the gasifier via coal char combustion. It was estimated that MW th was released by char combustion in the pilot regenerator. The fluidized bed regenerator provides some, though not complete, mixing of the char and the spent sorbent. The heat released by char reactions coupled with endothermic calcination leads to ~17ºC temperature gradients throughout 15

47 the fluidized bed regenerator. The feed to the regenerator contains calcium sulfide (CaS), part of which will be oxidized to Calcium Sulfate (CaSO 4 ) in the presence of CO 2 and O 2 as shown by: 1/4 CaS +CO 2 1/4 CaSO 4 +CO (2.6) 1/4 CaS +1/2O 2 1/4 CaSO 4 (2.7) The concurrent presence of CaS and CaSO 4 in the regenerator generates CaO and SO 2 as shown by: 1/4 CaS +3/4 CaSO 4 CaO + SO 2 (2.8) This reaction occurs through a series of intermediate steps. At temperatures above 955 o C, a transient liquid of a eutectic mixture of CaSO 4 and CaS is formed as an intermediate which solidifies and is deposited on the regenerator walls. Thus, a reducing environment with a CO concentration of 1 5% is maintained in the regenerator in order to prevent the formation of the transient liquid. Consequently, the gas exiting from the regenerator will contain a small amount of CO in the CO 2 stream. Further treatment of the CO containing exit stream will be needed. The temperature in the regenerator is controlled by adjusting the air flow rate. During the regenerator operation, nearly all the char fed to the regenerator is consumed. The coal ash is entrained with the spent air and collected in the external cyclone-lock hopper located at the gaseous product outlet of the regenerator. It was determined that 99% of the carbon in coal was converted by the gasifier and the regenerator. 16

48 A minimum temperature of 471 ºC in the regenerator is required for immediate combustion of char. At the start up, part of the flue gas from the regenerator, re-heated at natural gas furnaces, was used to increase the temperature of the regenerator to ~538 C in order to initiate char combustion. After the initiation of combustion, the heat released from combustion increased the temperature of the regenerator. At steady state, the regenerator was operated at 1010 C. Under such a high temperature, spent sorbent is decomposed, releasing CO 2 in the calcination reaction. The regenerated sorbent is discharged from the regenerator through two outlets: one outlet purges a predetermined amount of sorbent while the other outlet discharges the remainder into the gasifier where the regenerated hot sorbent is used for coal gasification. The heat from the hot sorbent is partially used to balance the heat requirement of the endothermic steam gasification reaction in the gasifier. Gasifier The gasifier is operated using a fluidized bed with continuous solids feeding. The hot sorbent particles from the regenerator are fed to the upper part of the gasifier while coal and steam are injected to the middle part and the lower part of the gasifier, respectively. In the gasifier, the CaO sorbent is converted to CaCO 3. The exothermic heat of the carbonation reaction is used to compensate for the endothermic gasification reaction. Examining the converted sorbent or spent sorbent settled at the bottom of the gasifier before its transport by gravity to the engager pot, it is found that the extent of conversion of the sorbent in the gasifier from CaO to CaCO 3 is high. The spent sorbent 17

49 in the engager pot is then pneumatically transported back to the regenerator by air. The fresh makeup sorbent is also provided to the engager pot. The presence of the reaction products such as H 2, CO, CO 2, and methane (CH 4 ) was found to limit the rate of gasification (Dobbyn et al., 1978) as noted by the rate of gasification initially at the lower part of the gasifier to be much larger than that at the higher part due to the difference in the gas composition. For example, 62% of the char was gasified at the bottom section of the bed while only 12% additional char was gasified in the middle section. In order to create a more uniform reaction rate throughout the gasifier, a portion of the product gas from the gasifier was recycled to the bottom of the gasifier. Such a product gas recycling step assists in moderating the rate of the reaction in the gasifier where solids are not well mixed. Further, the presence of the recycled gases decreases the partial pressure of the steam and hence, decreases the formation of calcium hydroxide (Ca(OH) 2 ) in the gasifier. Thus, the formation of the eutectic mixture of CaO- Ca(OH) 2 -CaCO 3 is minimized. It was found that lignite char was distinctively more reactive than sub-bituminous char. Synthesis gas and synthetic natural gas are continuously produced in this process. A typical synthesis gas composition obtained from the gasifier is given in Table 2.1. The synthesis gas produced from the gasifier is used in the methanation system, which is not shown in Figure 2.1, for the production of synthetic natural gas. A typical composition of the synthetic natural gas obtained from the CO 2 Acceptor process is given in Table 2.2. As can be seen in Table 2.2 the heating value of the 18

50 synthetic natural gas obtained from the pilot plant exceeds 900 Btu/SCF (33.5MJ/NM 3 ). The pilot plant studies also examined the factors that affected the activity of the sorbent and its environmental impact. Some Findings from the Process Testing The activity of the CO 2 sorbent, which is expressed by the ratio of the weight of CO 2 absorbed to the weight of the fresh (unreacted) sorbent, was found to be the key parameter to determining the technical feasibility of this process. The average acceptor activity must exceed a certain level in order for the process to be in heat balance since the system heat requirements are met by the sensible and reaction heat released by the acceptor at a given CO 2 removal rate. The minimum activity for the CO 2 acceptor process was found to be 0.26 for dolomite sorbent and 0.14 for limestone sorbent. Through the pilot plant testing, an activity of 0.35 was achieved using the dolomite sorbent, which exceeds the minimum activity requirement. In order for the process to be economically attractive, the sorbent needs to maintain a high activity with a minimal purge rate. It was observed that the activity of the acceptor decreases as the number of carbonation-calcination cycles increase. Other important factors that affect the activity of the acceptor include the acceptor residence time in the gasifier and the reactor operating temperatures. The decrease in the acceptor reactivity was attributed to the CaO crystalline growth. It was hypothesized that the calcium atom is relatively mobile at the gasifier/regenerator operating conditions, 19

51 especially when the operating temperature is high (Dobbyn et al., 1978). The high mobility of the calcium atom leads to fast CaO crystalline growth which forms a bridge between closely placed CaO crystals during the carbonation reaction. The crystals formed in this bridging effect are highly stable and have slow carbonation reaction kinetics. Such CaO crystals remain intact during the calcination reaction and tend to continuously grow in size and hence, reduce the gas diffusion rate. Thus, the reactivity of the acceptor decreases over time, especially at high temperatures. The crystalline growth effect was found to be more significant in the limestone acceptor than in the dolomite acceptor. To increase the recyclability of the acceptor, two approaches, i.e. acceptor reactivation and acceptor structure modifications, were tested in the pilot scale facility. Satisfactory results for both approaches were reported. Besides these attempts in improving the acceptor reactivity and recyclability, several strategies were adopted to enhance the energy conversion efficiency of the CO 2 Acceptor process. These strategies include the utilization of high pressure exhaust gas for air compression and the recovery of heat from exhaust gas for steam generation. Through the pilot testing, the metallurgical aspect of the reactor materials and its feasibility in usage were determined HYPR-RING PROCESS The H 2 Production by Reaction Integrated Novel Gasification Process (HyPr- RING) currently under development in Japan, is similar to the CO 2 Acceptor process. Both processes promote fuel conversion using CaO and/or Ca(OH) 2 sorbents. While 20

52 the CO 2 Acceptor process was aimed at synthetic natural gas production, the goal of the HyPr-RING process is the production of high purity H 2. (Lin, et al. 2004) The HyPr-RING process comprises principally two units, i.e., gasifier and regenerator, as shown in Figure 2.2. Coal is introduced, along with CaO and steam, in the gasifier where the following reactions take place: CaO + H 2 O Ca(OH) 2 (2.9) C + H 2 O CO + H 2 (2.1) CO + H 2 O CO 2 + H 2 (2.2) CaO + CO 2 CaCO 3 (2.3) Ca(OH) 2 + CO 2 CaCO 3 + H 2 O (2.10) The solids mixture from the gasifier, which contains unreacted coal char, CaCO 3, Ca(OH) 2 and ash, are fed into the regenerator where the following reactions take place: C + O 2 CO 2 (2.11) CaCO 3 CaO + CO 2 (2.5) CaO sorbent and ash from the regenerator are then recycled back to the gasifier. Before re-entering the gasifier, a portion of the solids mixture is discharged and the fresh makeup is added. This purge step helps prevent the ash accumulation and maintain the sorbent reactivity. 21

53 The concentration of H 2 S, NH 3 and HCN in the syngas stream were reported to be 2.2 ppm, ~0 ppm and 3.2 ppm respectively in the pilot unit located at Japan s Coal Energy Center with a coal feeding rate of 3.5 kg/hr. The HyPr-RING process has been extensively studied for H 2 production. A comparison of the quantity of the synthesis gas produced from the pyrolysis of coal mixed with CaO and Ca(OH) 2 revealed that the extent of pyrolysis of coal is improved in the order, from high to low, of coal/ca(oh) 2 offering the best extent of reaction, followed by a coal/cao mixture, and lastly with pure coal at pressures of atm. As this demonstrates, there are advantages of using Ca(OH) 2 as a sorbent. Water is supplied at a high temperature so that calcium is in the form of Ca(OH) 2 and aids in the reforming reaction. Further, CaO from the decomposition of Ca(OH) 2 enhances the pyrolysis reaction by the removal of CO 2 from the product gases. CaO also has a catalytic effect on the decomposition of tar, which further increases the gaseous product yield. The composition of H 2 in the gaseous products was found to be the highest in the temperature range of C. This temperature range conforms to the optimal temperature for the combined water gas shift and carbonation reaction. An increase in pressure results in an increase in the H 2 purity since the water gas shift and carbonation reactions are kinetically favored at higher pressures (Lin et al, 2003). Studies of H 2 generation from a mixture of pulverized coal and CaO with high pressure steam in a fixed bed reactor revealed that at a temperature of 700 C, the hydration of CaO occurs at a steam partial pressure higher than 30 atm. The yield of H 2 22

54 was found to be doubled with an increase in temperatures from 650 to 700 C and the yield of H 2 increases by 1.5 times with an increase in the total pressures from atm and the steam partial pressures from 7-42 atm.(lin et al, 2002). Gasification using pellets containing a mixture of coal and CaO in a fixed bed reactor revealed that although there is a decrease in the volume of the product gas, the composition of the product gas from pellet gasification is similar to that from gasification using the pulverized coal and CaO mixture. Further, in the gasifier, pellets retain their size and morphology at a gasification temperature of 650 C. At 700 C, however, the pellets are separated into two distinct parts: a dark part containing carbon and a white part containing a mixture of CaO, Ca(OH) 2 and CaCO 3 which form a eutectic melting mixture of solids. Recycling the CaO pellets between the reaction and the regeneration results in a constant H 2 yield over 4 cycles beyond which CaO is significantly deactivated due to the deposition of ash and inerts on the surface of the pellet. (Lin et al, 2004). Studies of the HyPr-RING process in a fluidized bed reactor at 650 C and 50 atm revealed that the hydration of CaO and the carbonation of the Ca(OH) 2 occurred in series, resulting in a gaseous product containing 76% H 2, 17% CH 4, 2%C 2 H 4, 3% C 2 H 6 and 2% CO 2. As the time scale for the combined hydration, water gas shift and carbonation reactions is 1-2 sec, which is much shorter than that for the gasification reaction, CO in the product gases is completely converted to CO 2 and almost all CO 2 is removed by CaO or Ca(OH) 2. In the continuous flow gasifier, the increase in the rate 23

55 of the combined water gas shift, reforming, and carbonation reactions is higher than the increase in the rate of the methanation reaction when the total and steam partial pressures are increased, as noted earlier. This behavior leads to the enhancement of H 2 production and the inhibition of CH 4 formation.(lin et al, 2004). The carbon conversion was found to be 60% near the entrance area of the fluidized bed reactor and 80% at the outlet of the reactor. The eutectic melting of the Ca(OH) 2, CaCO 3 and CaO mixture, which occurs at 700 C in the fixed bed experiments with the pelletized coal and CaO, was not present in the fluidized bed at 650 C. However, in the fluidized bed reactor, even at a low temperature of 650 C, particle growth occurs due to crystallization and cohesion of calcium compounds.(lin et al, 2006). Studies of the effect of various sorbents including CaCO 3, CaOSiO 2, MgO, SnO and Fe 2 O 3 on H 2 production indicate that high purity H 2 is obtained only with CaCO 3 and CaOSiO 2 sorbents and CO 2 cannot completely be removed from the product gas using the other sorbents (Lin et al, 2005). For different Ca-based sorbents, the rate of hydration was found to decrease with an increase in the CaO content. Further, the initial rate of hydration increases with an increase in the surface area of the sorbent while the final rate increases with an increase in the porosity. (Lin et al, 2008). Studies of the regeneration of the spent calcium sorbent in a 100% CO 2 environment and the reactivity of the calcined sorbent for the hydration and carbonation reactions reveal that for a residence time of 70 min for the calcination sorbent in a fluidized bed reactor, 73% of CaCO 3 calcined at 920 C, 95% calcined at 24

56 1020 C and almost 100% above 1020 C. The rates of the hydration and carbonation reactions decrease with an increase in the calcination temperature. Further, the extent of carbonation of CaO decreases from 60% at 950 C to 52% at 1000 C, and 40% at 1020 C. Thus, to improve the extent of carbonation, the hydration of CaO is desired in order to improve the porosity of the sorbent (Yin et al, 2007). Calcination in the presence of steam yields a sorbent that requires only half the time for hydration, compared to a sorbent obtained from calcination in the presence of 100% CO 2. The extent of carbonation of completely calcined CaO is also increased from 40% for 100% CO 2 calcination to 70% for steam calcination (0.4 atm CO 2 partial pressure and 30 atm total pressure) (Yin et al, 2008). Thus, by the combination of steam calcination and hydration, the sorbent loading in the process can be significantly reduced. For generating H 2 with high purity for fuel cell applications, extensive cleaning of H 2 S, CH 4, and other pollutants or byproducts from the H 2 stream will be necessary. The energy efficiency, defined as the high heating value (HHV) of the H 2 produced divided by the HHV of the coal converted, for this process was reported to be 77% (Lin et al, 2005). It can be noted that the difference between the CO 2 Acceptor process and the HyPr-RING process lies in the gasifier operating conditions. Comparing the operating conditions of the gasifier used in the CO 2 Acceptor process, i.e., ºC and 10 atm, the operating conditions of the gasifier used in the HyPr-RING process have a lower operating temperature (650 ºC) and a higher operating pressure (30 atm). The lower temperature and higher pressure in the HyPr-RING process gasifier 25

57 thermodynamically favors the carbonation reaction, thereby further enhancing the H 2 production. Moreover, an excess of steam is used in the HyPr-RING process to enhance the reactivity of the CaO sorbent by refreshing the pore structure of the particles ZERO EMISSION COAL ALLIANCE (ZECA) PROCESS The Zero Emission Coal Alliance Process, or ZECA process, was proposed by Klaus Lackner and H. Ziock (ZECA Corporation, 2002). Figure 2.3 shows the schematic diagram of the process. In this process, coal is first converted to methane by reacting with H 2 in a gasifier: C + 2H 2 CH 4 (2.12) This hydrogasification step also produces light hydrocarbons. The CH 4 and light hydrocarbons are then sent to the reformer. Steam and CaO sorbent are introduced to the reformer to convert the hydrocarbons into H 2 via sorbent enhanced reforming reactions: CH 4 + H 2 O CO + 3H 2 (2.13) CO + H 2 O CO 2 + H 2 (2.2) CO 2 + CaO CaCO 3 (2.3) The H 2 gas generated in the reformer is split into two streams: one stream is recycled to the hydrogasifier and the other is sent to a solid oxide fuel cell (SOFC) for 26

58 power generation. The spent sorbent, consisting mainly of CaCO 3, is regenerated in a calciner. CO 2 is readily separated in this step: CaCO 3 CaO + CO 2 (2.5) The heat required for the calcination reaction is provided by the waste heat from the solid oxide fuel cell system, which is operated using H 2 from the reformer. The fuel cell also generates steam, which is recycled back to the gasifier and the reformer for CH 4 and H 2 generation. Stoichiometrically, to convert one mole of carbon, 2 moles of H 2 gas are consumed in the gasifier and 4 moles of H 2 gas are generated in the reformer. Therefore, there is a net gain of 2 moles of H 2 gas per mole of carbon converted. The excess H 2 stream is used to meet the process heat requirement and to generate electricity ALSTOM HYBRID COMBUSTION-GASIFICATION PROCESS In a typical configuration of the ALSTOM chemical looping process for H 2 production, calcium based sorbents and bauxite ore are used to carry oxygen, CO 2, and heat in three loops. The first loop is the CaSO 4 -CaS loop in which coal is gasified using CaSO 4, an oxygen carrying agent, to produce CO. CO is then converted to CO 2 and H 2 by the water gas shift reaction. The CaS produced in this process is regenerated in air to produce CaSO 4 through an exothermic oxidation reaction. The second loop consists of the CaO-CaCO 3 loop in which the CaO sorbent is used to remove CO 2 during the water gas shift reaction, forming CaCO 3 while producing a pure stream of H 2. The 27

59 third loop is a heat transfer loop in which hot CaSO 4 or bauxite is used to transfer the heat from the exothermic CaS oxidation reaction to the calciner to support the endothermic calcination of CaCO 3. (Andrus et al, 2006) 2.6. FUEL-FLEXIBLE ADVANCED GASIFICATION-COMBUSTION PROCESS The GE process comprises two loops, an oxygen transfer loop and a carbon transfer loop, and involves three reactors. In the first reactor, coal is gasified to produce CO and H 2 along with CO 2, which is constantly removed by the CaO sorbent. The reacted CaCO 3 product along with the unconverted char is then routed to the second reactor where hot oxygen transfer material from the third reactor is reduced while converting the char to CO 2. The hot solids also provide heat for the calcination of CaCO 3. In the third reactor, the reduced oxygen transfer material is reoxidized, releasing a considerable amount of heat that heats up the solids and generates steam for power production. The GE process obtains a H 2 concentration of only 80%. (Rizeq et al, 2001) In the processes discussed above, CO 2 is removed in the gasifier by the CaO sorbent. Brun-Tsekhovoi et al., 1988, Fan et al., 2007, Ortiz and Harrison, 2001, Han and Harrison, 1994, Johnsen et al., 2006, Balasubramanian et al., 1999, Hufton et al., 1999, and Akiti et al., 2004, have also applied CO 2 removal by CaO to the removal of CO 2 and the production of H 2 from syngas through the water-gas shift reaction and from CH 4 through the sorption-enhanced steam methane reforming reaction. Chapter 4, 28

60 5 and 6 describe the conversion of syngas to H 2 in the presence of CaO sorbent. Chapter 4 discusses the production of H 2 in the presence of CaO sorbent and a water gas shift catalyst while Chapter 5 discusses the non-catalytic H 2 production in the presence of CaO. Chapter 6 is a system analysis of the process. The conversion of CH 4 to H 2 in the presence of CaO sorbent is described in Chapter 7. 29

61 Gas Type CH 4 H 2 CO CO 2 N 2 H 2 O HHV Btu/SCF, (MJ/NM 3 ) Percentage (%) (14.1) Table 2.1: A typical composition of the H 2 rich synthesis gas from the gasifier 30

62 Gas Type CH 4 H 2 CO CO 2 N 2 HHV Btu/SCF, (MJ/NM 3 ) Percentage (%) > 900 (33.5) Table 2.2: A typical composition of the synthetic natural gas from the methanation system 31

63 Flue Gas Product Gas Ash Spent Sorbent CaO 1010ºC Regenerator 823ºC Gasifier Coal Fuel char Lift gas Sorbent makeup CaCO 3 Engager Pot Air Steam Figure 2.1: Schematic diagram of the reactor system in the gas synthesis block of the CO 2 Acceptor process (Dobbyn et al., 1978) 32

64 Water H 2 Water Regenerator CO 2 Gasifier O 2 C/CaCO 3 /Ca(OH) 2 Steam CaO Coal Ash/ CaO Steam Steam Turbine A.S.U. N 2 Figure 2.2: Schematic diagram of the HyPr-RING process 33

65 Figure 2.3: Schematic of the ZECA process 34

66 H 2 O CaCO 3 Coal CaO CO + H 2 O + CaO CaCO 3 + H 2 Loop 2 CaCO 3 CaO + CO 2 CaCO 3 CO Hot Cold Loop 3 Bauxite Bauxite CaSO 4 4C + CaSO 4 4CO + CaS Loop 1 CaS + 2O 2 CaSO 4 CaS Figure 2.4: Schematic of the ALSTOM process 35

67 H 2 CO 2 N 2 Coal Gasification C + 2H 2 O CO 2 + 2H 2 CaO + CO 2 CaCO 3 CaCO 3 Regeneration FeO Loop 1 C + Fe 2 O 3 CO 2 + FeO Loop 2 CaO CaCO 3 CaO + CO 2 Fe 2 O 3 Oxidation 4FeO + O 2 2Fe 2 O 3 Steam Steam Air Figure 2.5: Schematic of the GE process 36

68 CHAPTER 3 REACTIVITY AND RECYCLABILITY OF CALCIUM BASED SORBENTS FOR CO 2 CAPTURE 3.1 INTRODUCTION The successful operation of the CLP is highly dependent on the performance of the CaO particles for CO 2 and sulfur capture. In the CLP, the sorbent participates in several reactions in at least two reactors: the carbonation reactor and the calciner. The carbonation reaction occurs in the temperature range of 500 to 750 ºC and the calcination reaction occurs at higher temperatures. The reactivity of the sorbent over multiple cycles is very important for the economics of the process since it affects the size of the reactors and the amount of solid circulation and sorbent makeup. Some of the major factors that affect the solid circulation and makeup are the reactivity of the sorbent, the recyclability, which depends on the temperature and gas atmosphere of calcination, the amount of sulfur in the feed gas and the extent of attrition of the sorbent. In this chapter, the reactivity and recyclability of natural and synthetic sorbents is investigated. In addition, the effect of realistic calcination conditions on sorbent 37

69 reactivity is explored and the effectiveness of sorbent reactivation by hydration is determined on the bench and subpilot scale. 3.2 SORBENT REACTIVITY OVER MULTICYCLIC REACTIONS The reaction between CaO and CO 2 occurs in two distinct stages. The first stage occurs rapidly and is kinetically controlled, while the second stage is slower and diffusion controlled. For any commercial application, only the first stage of the reaction should be considered in order to use a compact reactor for the removal of CO 2. Abanades et al studied the rate and the extent of the carbonation reaction and the variation of these parameters with multiple carbonation and calcination cycles.(abanades and Alvarez, 2003) The CaO conversion at the end of the rapid kinetically controlled regime is found to decay sharply for naturally occurring limestone with an increase in the number of cycles. Although the initial decay is smoother for dolomite and other modified sorbents, it is intrinsic to most sorbents used in the CLP. In addition to the decay in CO 2 capture capacity, dolomite and other supported sorbents also have the disadvantage of carrying more inert material in the loop thereby increasing the parasitic energy requirement of the regeneration process. Since the cost of the supported and modified sorbents is also higher, their performance over multiple cycles also needs to be significantly higher in order to compete with natural limestone. The decay in lime conversion over multiple cycles has been reported by numerous researchers including Curran et al, Shimizu et al, Silaban and Harrison, Barker, and Aihara et al (Curran et al, 1967, Shimizu et al, 1999, Saliban and 38

70 Harrison, 1995, Barker, 1973, Aihara et al, 2001). Using these data, Abanades and Alvarez concluded that the decay in conversion is dependent only on the number of cycles and independent of the reaction times and conditions (Abanades and Alvarez, 2003). Using a simple relationship given in Equation (3.1), Abanades related the conversion of lime for any given cycle number (x c,n ) to fitted constants (f, b) and the cycle number (N) as given by: x c,n = f N+1 + b (3.1) where the fitted parameters f and b have a numerical value of.782 and.184, respectively (Abanades, 2002). Taking into consideration the sorbent conversion decay over multiple cycles, the kinetics of the reaction, and mass and energy flows, Abanades developed Equation (3.2) to determine the maximum capture efficiency of CO 2 in a system containing a continuous purge of solids and a make up of fresh sorbent. (Abanades, 2002) E CO F F 1+ 0 f 0 FR FR = + b (3.2) 2 F F F 0 CO f F F F R R R where E CO is the maximum obtainable efficiency, F 2 0 is the fresh feed added to the system (mol CaO/s); F R is the total amount of sorbent required to react with the CO 2 in the system (mol CaO/s); FCO2 is the flow of CO 2 (mol/s); and f and b are constants as defined in Equation (3.1). b is the residual carbonation conversion due to the formation 39

71 of a product layer of carbonate inside the macropores in highly sintered sorbents. This residual carbonation of the lime sorbent is beneficial as it aids in reducing the amount of fresh sorbent to be added. From an economic standpoint, it is desirable to minimize the ratios F R / F CO and F 2 0 /F R in order to minimize the energy required for calcination and the amount of fresh sorbent required (Abanades, 2002). For F O and F R to be low, the sorbent should have a high resistance to sintering. The CaCO 3 product layer formation and pore pluggage during carbonation and the sintering of CaO during calcination are both attributed to the decay and irreversibility of limestone. Abanades et al concluded that micropores contribute to the fast stage of the carbonation reaction (Abanades and Alvarez, 2003). The fast reaction stage ceases when the micropores connecting the crystal grains are plugged due to the increase in the molar volume during the formation of CaCO 3 from CaO, where CaCO 3 has greater than twice the molar volume as CaO. In the larger pores (mesopores and micropores), CaCO 3 forms a layer on the CaO wall (Alvarez and Abanades, 2005). Although the pore is sufficiently large to handle the increase in pore volume, the resistance of CO 2 diffusion through the CaCO 3 layer dramatically increases. The increased resistance forms the boundary between the two stages of carbonation. Sintering of CaO during calcination over multiple cycles results in grain growth which drastically reduces the CaO microporosity while increasing the mesoporosity. This leads to a reduced fast carbonation reaction zone, and therefore, a decrease in CO 2 capture capacity over multiple cycles (Abanades and Alvarez, 2002). 40

72 Sun et al also investigated the sintering mechanism of limestone with increasing number of cycles and attributed sintering to be due to CO 2 released during the calcination process (Sun et al, 2007). They showed that the increase in the carbonation time did not have any effect on the structure of the calcine as the calcination process eliminates the changes caused by carbonation. However, an increase in the calcination time resulted in a decrease in the pore volume for pores <220 nm (Sun et al, 2007). Similar to the observation made by Abanades et al with the increase in the number of cycles the pore volume decreased for pores < 220nm and consequently increased for pores >220 nm (Abanadez and Alvarez, 2003). A sintering model has been developed by Sun et al based on the packed bed model, shrinking core model and a modified sintering kinetic model and the average CO 2 conversion is given below ( Sun et al, 2007): X carb = 1.07 (n+1) (3.3) To be commercially viable, the CaO sorbent must maintain its reactivity towards CO 2 over multiple cycles. Additives and processed sorbents have been investigated, but these techniques undermine the main advantage of using natural limestone, which is its low cost. Using natural limestone has its challenges, which must be overcome. The effect of doping CaO with NaCl and Na 2 CO 3 has also been investigated in a Thermo Gravimetric Analyzer (TGA) (Salvador et al, 2003). The addition of NaCl 41

73 increased the CO 2 removal capacity of the sorbent to 40% over 13 cycles due to favorable changes in the pore structure and surface area of the sorbent while the addition of Na 2 CO 3 did not have any effect on the extent of carbonation. When the doped sorbents were tested in the fluidized bed, both NaCl and Na 2 CO 3 caused a decrease in the CO 2 removal capacity of the CaO sorbent which might be attributed to the coating of the surface of the sorbent, leading to pore blockage during the calcination stage ( Salvador et al, 2003). This chapter focuses on modification and reactivation methods that could be used to improve the reactivity of Ca-based sorbents over multiple cycles. 3.3 SYNTHESIS OF HIGH REACTIVITY PRECIPITATED CALCIUM CARBONATE (PCC) SORBENT One method of improving the recyclability of Ca-based sorbents is to modify the pore structure of the sorbent to increase the pore volume and surface area. Fan et al. have developed a wet precipitation process to synthesize a high surface area Precipitated Calcium Carbonate (PCC) (Fan et al, 1998, Fan and Gupta, 2006) The PCC - CaO sorbent can achieve almost complete conversions (> 95%) due to presence of mesopores (5-30 nm). PCC is synthesized by bubbling CO 2 through a slurry of Ca(OH) 2. The surface properties of the sorbent are tailored by the addition of anionic surfactants (Agnihotri et al, 1999, Ghosh-Dastidar et al, 1996, Wei et al., 1997). The system reaches an optimum only when the zeta potential equals zero. The sorbent 42

74 optimization process results in production of a sorbent with a surface area of 60 m 2 /g and a pore volume of 0.18 cc/g. CaCO 3 primarily occurs in three different polymorphs, each of which may have multiple morphologies depending on the arrangement of the atoms and ions in the crystal structure. These polymorphs are all present in nature as well as in synthesized PCC and can be classified as calcite, aragonite and vaterite. Calcite is the most stable polymorph and typically occurs in the triagonalrhombohedral (acute to obtuse), scalenohedral, tabular and prismatic morphologies. Calcite crystals also display intergrowth or twinning to form fibrous, granular, lamellar and compact structures. The rhombohedral and prismatic forms find applications in paper coating and in polymer strength enhancing agents while the scalenohedral form is used in paper filling due to its light scattering ability. Calcite exhibits a unique property by which its solubility in water decreases with increasing temperature. The aragonite polymorph has an orthorhombic morphology with needle shaped or acicular crystals. Twinning of these crystals results in the formation of pseudohexagonal structures which could be in a columnar or fibrous matrix. Aragonite is unstable at standard temperatures and pressures and eventually gets converted to calcite over geological timescales. Aragonite also exhibits a higher density and solubility than calcite. The needle shaped morphology of aragonite is beneficial for high gloss 43

75 paper coating applications as well as for strength enhancing additives in polymeric materials. Vaterite is the most unstable form of CaCO 3 at ambient conditions and readily gets converted to calcite (at lower temperatures) and aragonite (at higher temperatures of 60 ºC) on exposure to water. Vaterite is usually spherical in shape and has a higher solubility in water than the other polymorphs. The transformation of aragonite and vaterite to calcite is accelerated with temperature (Yamaguchi and Murakawa, 1981) Although PCC predominantly contains calcite, various factors in the synthesis procedure like the extent of saturation of the Ca(OH) 2 solution, ph of the solution, concentration of CO 2, etc dictate the type and size of its morphology. For example, PCC synthesized from highly saturated aqueous Ca(OH) 2 solutions contains aragonite at 70 ºC and vaterite at 30 ºC (Wary and Daniels, 1957). Cizer et al (2008) have shown that rhombohedral calcite crystals formed by the exposure of Ca(OH) 2 to 100% CO 2 are micrometer sized while that precipitated with 20% CO 2 are submicrometer sized. In addition, it was also found that during the initial stages of carbonation, when the concentration of Ca 2+ ions in the solution is greater than the concentration of CO - 3 ions, a scalenohedral calcite is precipitated. The scalenohedral morphology gets transformed into the rhombohedral form during the later stages of precipitation, when the CO 3 - concentration in the solution is high. 44

76 3.4 PRETREATMENT OF CALCIUM BASED SORBENTS AND ADDITION OF SUPPORTS The CO 2 capture capacity of CaO obtained from several precursors was determined for a single cycle and for multiple cycles in a TGA. The experimental procedure and the details of the TGA setup can be obtained elsewhere (Iyer et. al, 2004) Reactivity Testing of Ca-based Sorbents for CO 2 Capture Figure 3.1 illustrates the comparison in the CO 2 capture capacity of the CaO sorbent obtained from different precursors. The CO 2 capture capacity has been defined by the weight % capture which is the grams of CO 2 removed/ gram of the CaO sorbent. The Wt% capture of CaO obtained from limestone is 58%. It can be seen that the weight % capture attained by the sorbent obtained from PCC powder is 74% when compared to that of 60% attained by the Ca(OH) 2 hydroxide sorbent and 20% attained by the ground lime sorbent. In order to improve the strength of the PCC particles, the PCC powder was pelletized into 2mm pellets and then ground to a size of 150 microns. The CO 2 capture capacity of the PCC pellets as well as the pelletized and broken sorbent was also determined. The CO 2 capture capacity of the pelletized and broken PCC is almost the same (71%) as the PCC powder as shown in Figure 3.1. The PCC pellet requires a very large residence time due to mass transfer resistance but reaches 45

77 the same final CO 2 capture capacity of 71% as that of the PCC pelletized and broken sorbent Recyclability of Natural, Pretreated and Supported Sorbents Since the PCC powder as well as the PCC pelletized and broken sorbents have very high CO 2 capture capacities and require almost the same residence time for carbonation, a multicyclic calcination and carbonation experiment was conducted on the two sorbents. Figure 3.2 illustrates the comparison in the conversion attained by the PCC powder and PCC pelletized and broken sorbents over 5 calcination and carbonation cycles. It can be seen that during the first cycle the PCC powder and PCC pelletized and broken sorbent both achieve the same CaO conversion. As the number of cycles increases, the conversion falls for both sorbents due to sintering but the conversion for the PCC powder sorbent falls more than the pelletized and broken PCC sorbent. This shows that the sintering of the PCC sorbent could be reduced by pelletizing the PCC sorbent and grinding it to the size range of 150 microns. This not only improves the multicyclic conversion but also improves the strength of the sorbent. To improve the recyclability of CaO for CO 2 capture, several pretreatment methods as well as the addition of metal oxide supports were evaluated in a TGA as shown in Figure 3.3. The calcination of the sorbent precursor during the first cycle as well as calcination of the spent sorbent every cycle was conducted at 700 ºC in pure nitrogen. The wt% capture which is the weight of CO 2 captured/ unit weight of the 46

78 sorbent was determined by reacting the calcined sorbent with CO 2 in a feed gas containing 10% CO 2 and 90% N 2 at 650 ºC. Linwood Carbonate (LC) is naturally occurring limestone and its reactivity decreases from 59 wt% capture to 30% in 18 cycles. As the simplest method of pretreatment, a freshly calcined sample of LC was hydrated with water to produce Linwood Hydrate (LH) at ambient temperature which was then tested in the TGA for 18 carbonation and calcination cycles. As shown in Figure 3.3, the first cycle reactivity of LH is lower than LC but LH performs better in the cyclic tests. The reactivity of LH only decreases from 53% to 43% and LH has a 13% higher reactivity at the end of 18 cycles than LC. This shown that pretreatment of the sorbent by hydration has the potential to improve the recyclability of calcium sorbents. As described in the section above, PCC was synthesized by the addition of a surfactant and the first cycle reactivity as well recyclability of PCC was found to be higher than most of the other sorbents. PCC powder has a first cycle reactivity of 67% and a wt% capture of 49% at the end of 18cycles. The effect of pretreatment of the LC sorbent with formic and acetic acid was also investigated. The pretreatment of the LC sorbent with formic and acetic acid prior to the first calcination step aids in increasing the pore volume of the sorbents due to the formation of calcium acetate and calcium formate which have a higher molar volume than CaCO 3. The calcium formate precursor has the highest first cycle reactivity of 70% but its reactivity decreases steeply to 27% at the end of 18 cycles. Calcium acetate precursor has a lower first cycle wt% capture than the calcium formate precursor of 63% but it has good recyclability over multiple cycles similar to the LH precursor. In addition to sorbent 47

79 pretreatment, the effect of addition of metal oxides like MgO, SiO 2 and Al 2 O 3 to CaO sorbent was also investigated. The synthetic sorbents were prepared in the laboratory from LC. Calcined LC was mixed with water to make a slurry of Ca(OH) 2. The metal oxide was added to the slurry and CO 2 was bubbled through the mixture to precipitate out a mixture of CaCO 3 and the metal oxide. The slurry was filtered and the solid sorbent was then dried in an oven. The modified sorbent which is a mixture of CaCO 3 and the metal oxide was then subjected to 18 carbonation and calcination cycles in the TGA. All the sorbents with the metal oxide supports have a low first cycle wt% capture due to the presence of the metal oxide which behaves as an inert during the carbonation reaction. The wt% capture of the sorbent with MgO decreases from 51% to 41% while that with SiO 2 decreases from 47% to 41% over 18 cycles. The sorbent with Al 2 O 3 has a lower wt% capture over 18 cycles than the other two supported sorbents which decreases from 44% to 34% over 18 cycles. 3.5 EFFECT OF REALISTIC CALCINATION CONDITIONS ON SORBENT REACTIVITY In the previous sections, the multicyclic reactivity of sorbents was investigated in a TGA with calcination conducted in ideal conditions (at a low temperature of 700 ºC in a pure stream of N 2 carrier gas). In a CO 2 constrained scenario, carrier gases like N 2 and air cannot be used in the calciner as they will mix with the CO 2 produced by the calcination of the spent sorbent. Hence in the absence of these carrier gases the temperature of calcination is increased significantly. From thermodynamic analysis, a minimum temperature of 890 ºC is required to calcine CaCO 3 in a pure CO 2 48

80 atmosphere. The temperature of calcination can be reduced by using a condensable gas like steam as a carrier gas in the calciner. This will result in the production of a wet CO 2 stream which can then be dried and compressed for sequestration. The following section describes the effect of realistic calcination conditions on the reactivity of the sorbent in a bench scale calciner Experimental Methods A detailed description of the bench scale rotary bed calciner is provided elsewhere and consists of a stainless steel reactor tube rotating within a horizontal furnace ( Sakadjian et al, 2007). The carrier gas consisting of pure CO 2 or a mixture of steam and CO 2 was fed to the reactor and the outlet of the reactor was connected to a CO 2 analyzer. The sorbent was loaded in the reactor tube and the temperature was increased to the calcination temperature. At the end of calcination, the CO 2 capture capacity of the sorbent was determined in a TGA apparatus procured from PerkinElmer Corp. In the TGA a small sample of the sorbent (15-20 mg) was placed in a quartz boat suspended from a platinum wire. The sorbent was brought to a reaction temperature of 650 C in flowing nitrogen. Subsequently, the flow was switched to the reaction gas stream which contained 10% CO 2 and balance N 2. The TGA records the increase in the sample weight with respect to time, which signifies the CO 2 capture by the sorbent. The Wt% CO 2 capture capacity of the sorbent was then determined as the grams of CO 2 captured *100 /gram of CaO sorbent. 49

81 3.5.2 Results and Discussion Figure 3.4 illustrates the effect of realistic calcination conditions on the reactivity of limestone sorbent. The Wt% CO 2 capture of the original limestone sorbent calcined in ideal conditions in a 100% nitrogen stream at 700 ºC in the TGA is 50%. The realistic calcination of the limestone sorbent in the bench scale rotary bed calciner at 900 ºC in a pure CO 2 atmosphere produced CaO sorbent with a Wt% CO 2 capture of 28%. Hence, the sorbent only retains half of its original reactivity to CO 2 after a single cycle of realistic calcination at 900 ºC. The effect of calcination in the presence of a mixture of steam and CO 2 at 900 o C was also determined on the reactivity of the sorbent. Almost complete calcination of the sorbent was obtained in every case. It is found that on calcination of the limestone sorbent in an atmosphere of 33% steam and 67% CO 2, a CaO sorbent with a Wt% CO 2 capture of 35% is obtained. A further increase in the steam concentration to 50% resulted in the production of a more reactive sorbent with a Wt% CO 2 capture of 45%. Hence, the addition of steam aids in reducing the extent of sintering of the sorbent and results in the production of a CaO sorbent that is more reactive to CO 2. To determine the recyclability of the sorbent with steam calcination, a multicyclic carbonation- calcination test was conducted and the results are shown in Figure 3.6. The Wt% capture of the original sorbent calcined in ideal conditions at a temperature of 700 ºC in pure nitrogen is 50%. During the cyclic testing, the sorbent calcination was conducted in the bench scale rotary bed calciner while the carbonation 50

82 was conducted in a bench scale fixed bed reactor shown in Figure 3.5. The sorbent was packed in the fixed bed reactor and carbonation was conducted in a 10% CO 2 /90% N 2 stream at 650 ºC. On calcining the limestone sorbent at 900 ºC in a 50%/50% H 2 O/CO 2 atmosphere, the Wt% capture of the sorbent reduces to 45%. During the second and third cycles the reactivity further decreases to 30% and 25%. Hence although steam calcination reduces the extent of sintering, the sorbent reactivity is not maintained a constant and it continues to fall over multiple cycles SORBENT REACTIVATION BY HYDRATION LAB SCALE TESTING The use of sorbent hydration as a pretreatment method was investigated in the TGA and found to improve the recyclability of the sorbent as shown in Figure 3.3. In order to maintain the reactivity of the sorbent a constant over multiple cycles, sorbent hydration was also investigated as a reactivation process. Sorbent reactivation by hydration was included as a step in every carbonation- calcination cycle Experimental Methods The effect of sorbent hydration with water at 25 o C, and steam at high temperatures was investigated on sorbent reactivity. Limestone sorbent was calcined in realistic calcination conditions at 1000 o C in a 100% CO 2 atmosphere in the bench scale rotary bed calciner described earlier (Sakadjian et al, 2007). At the end of calcination, the sorbent was hydrated. Water hydration was conducted by spraying water at 25 o C on the sorbent with vigorous stirring. Steam hydration was conducted in 51

83 the bench scale reactor shown in Figure 3.5 in a 20% nitrogen and 80% steam atmosphere. At the end of hydration, the CO 2 capture capacity of the sorbent was determined in the TGA. A small sample of the sorbent (15-20 mg) was placed in the quartz boat suspended from the platinum wire. The hydrated sorbent was brought to a reaction temperature of 650 C in flowing nitrogen. Complete dehydration of the sorbent occurred by the time the sample was heated to 650 o C. Subsequently, the flow was switched to the reaction gas stream containing 10% CO 2 and balance N 2. The TGA records the increase in the sample weight with respect to time, which signifies the CO 2 capture by the sorbent. The Wt% CO 2 capture capacity of the sorbent was then determined as the grams of CO 2 captured *100 /gram of CaO sorbent Results and Discussion Figure 3.7 illustrates the effect of water hydration (at 25 o C) and steam hydration (at 150 o C and 500 ºC) on the CO 2 capture capacity of the sorbent. The original limestone sorbent calcined in the TGA in the presence of 100% nitrogen at 700 o C has a Wt% CO 2 capture capacity of 52%. Calcination of the limestone at 1000 o C in pure CO 2 in the bench scale calciner reduces its Wt% CO 2 capture capacity to 20%. On hydration of the sorbent with water at ambient temperature, the Wt% CO 2 capture capacity increases to >55%. Another method of hydration at atmospheric pressure, in the presence of steam at 150 o C yielded in the production of a sorbent with 52 Wt% CO 2 capture. Steam hydration at atmospheric pressure and 500 ºC yielded in a sorbent with a Wt% capture of 45%. While the extent of hydration obtained with water and 52

84 with steam at 150 ºC is greater than 95%, it is only 80% when the sorbent is hydrated at 500 ºC. The reduction in extent of hydration might have resulted in the lower CO 2 capture capacity observed for the sorbent hydrated at 500 ºC. Figure 3.8 shows the effect of hydration at a high temperature of 600 o C for total pressures ranging from 8 atms to 21 atms. The Wt% CO 2 capture of the sorbent calcined in 100% CO 2 at 1000 o C increases from 20% to 45% by pressure hydration at 600 o C and 8 atms. The reactivity of the sorbent is found to decrease to a small extent on increasing the pressure of hydration at 600 o C. Further investigation is required to determine if this decrease in reactivity is due to an increase in the sintering of the sorbent at high pressures. Multicyclic calcination hydration carbonation tests were also conducted as shown in Figure 3.9. The calcination was conducted at a temperature of 950 ºC. The calcined sorbent was hydrated at 500 ºC and carbonated in a 10%CO 2 / 90%N 2 stream at 650 ºC. The Wt% capture was calculated on the basis of the Ca(OH) 2 in the sorbent sample and not on the basis of entire solid sample weight as before. As illustrated in Figure 3.9, the Wt% capture of the sorbent is maintained a constant over multiple cycles. Hence sorbent hydration is a promising method of completely reactivating the sorbent and improving its recyclability. 53

85 3.7 SUB-PILOT SCALE DEMONSTRATION OF REACTIVATION OF CALCIUM SORBENT BY HYDRATION The effectiveness of hydration on improving the recyclability of the sorbent was tested in a subpilot scale demonstration for CO 2 and SO 2 capture from combustion flue gas using the CLP process. Figure 3.10 illustrates the process flow diagram of the calcium based CO 2 and SO 2 capture process from combustion flue gas. The CaO sorbent or Ca(OH) 2 sorbent is injected into the carbonator, which is an entrained bed reactor, where it reacts with the CO 2 and SO 2 to form CaCO 3 and CaSO 4 at a high temperature between 450 C and 650 C. Thermodynamic limitations prevent greater than 90% CO 2 removal from a coal combustion flue gas stream at temperatures greater than 650 C. The CaO sorbent could be obtained from such precursors as natural limestone, hydrated lime, and reengineered and supported sorbents. The spent sorbent mixture is then regenerated by calcining it at a high temperature between 850 C-1300 C where the CaCO 3 decomposes to yield CaO and a pure, dry stream of CO 2 when calcined. The calciner could be a flash or entrained bed calciner, a fluidized bed or a rotary kiln. While energy has to be provided for the calcination reaction, the carbonation reaction is exothermic and releases high quality heat. Hence, a good indirect heat integration strategy aids in reducing the parasitic energy consumption of the process. With Ca(OH) 2 as the sorbent, the CaO is further reactivated by hydration and re-circulated to the carbonator, while the CO 2 is compressed and transported for sequestration. Since CaSO 4 begins to decompose only at temperatures greater than 54

86 1450 C, under the conditions experienced in the calciner, CaSO 4 is stable and a small amount of solids must be continuously purged out of the system to prevent complete conversion of sorbent to CaSO 4.The amount of solid purge from the CLP will depend on the amount of sulfur and flyash that are fed to the carbonator to prevent the accumulation of inert solids in the process. Based on a preliminary economic analysis, the purge percentage will be in the range of 2% to 10%.Thus the CLP process captures CO 2 in the flue gas stream and converts it into a concentrated sequestration ready CO 2 stream. The CLP process is capable of capturing CO 2 from flue gas streams produced from various fuels including coal, oil, natural gas, biomass, etc, Experimental Methods for the 120 KWth Subpilot Scale Testing The effectiveness of sorbent reactivation by hydration was tested in a 120KWth subpilot scale demonstration of the calcium based CO 2 capture process at the Ohio State University (Wang et al, 2009). Coal was stored in a coal hopper, which is connected to an underfeed stoker, provided by Babcock & Wilcox Co., Barberton, OH,. The underfeed stoker has two Forced Draft (FD) fans that provide combustion air to the stoker. Natural gas is connected to the inlet of the stoker for start-up and to maintain gas temperature. The flue gas stream is transported through the ductwork via an Induced Draft (ID) fan. Connected to the ductwork are a hopper and screw-feeder, two sets of gas analyzers, multiple temperature monitoring ports, multiple pressure measurement ports, a cyclone and a baghouse. Figure 3.11 illustrates a snapshot of the sub-pilot scale facility. 55

87 A Schenck-Accurate mid-range volumetric hopper, is the main sorbent feeder and is connected to the calciner feed inlet. An electrically-heated rotary calciner manufactured by FEECO that has a maximum operating temperature of 980 C is used to calcine the spent calcium sorbent. The calcined sorbent was hydrated offline for the data reported in this study and injected into the flue gas duct. Once injected into the ductwork, the sorbent is entrained by the flue gas, and it simultaneously reacts with the CO 2 and SO 2 present in the flue gas. At the end of the process, a Donaldson Torit downflow baghouse is used to separate the solid sorbent from the CO 2 /SO 2 free flue gas which is emitted to the outside atmosphere. To monitor the gas composition, two sets of gas analyzers are employed. One set of gas analyzers is located upstream of the sorbent injection port and is used as the baseline. The other set of gas analyzers is located downstream of the sorbent injection. The difference, after correcting for air in leakage and other factors, between the two measurements determines the percent removal. The gas analyzers are CAI 600 analyzers and continuously monitor the concentrations of CO 2, SO 2, and CO. In addition, a CAI NOxygen analyzer monitors the upstream oxygen and nitrogen oxides concentrations, while a Teledyne Analytical 3000P analyzer monitors the downstream oxygen concentration. All data are continuously recorded via a data acquisition system. Multiple Type K thermocouples continuously monitor the temperature throughout the entire system to determine the proper operating temperature for the carbonation 56

88 reaction, which occurs at a reasonable rate between 450 C and 650 C (Gupta and Fan, 2002, Koji et al, 2003, Abanadez et al, 2003, Lee, 2004, Wong, 2007) Prior to each experimental run, all analyzers were calibrated. The stoker was heated and operated according to the start-up procedures. Once the flue gas temperature at the sorbent injection location reached approximately 650 C, which is sufficiently high to allow both the carbonation and sulfation reaction to proceed at a high rate and achieve greater than 90% removal of the CO 2. The flowrate of the sorbent was set via the control panel. After the sorbent reacted with the CO 2 and SO 2 in the carbonator, the gas temperature was lowered and the spent sorbent was collected in the bag house. To calcine the spent sorbent, the calciner temperature was set to 950 C and the calcined solids were then reactivated using offline hydration. The carbonationcalcination-hydration cycle was repeated. At the completion of each experiment, solids from the baghouse were collected and analyzed via a TGA Results and Discussion The effect of sorbent reactivation by hydration was investigated on the %CO 2 removal from the flue gas and on the Wt% capture of the sorbent over multiple cycles. Figure 3.12 illustrates the carbonation -calcination cycles for pulverised lime over 3 cycles and the calcination-carbonation-hydration cycles for Ca(OH) 2 over 4 cycles. A maximum of only 50% CO 2 removal is achieved in all the tests shown in Figure

89 since a substoichiometric calcium to carbon (Ca:C) mole ratio of 0.75 was used for testing. Greater than 90% CO 2 removal is achieved for a Ca:C ratio of 1.3 (Wang et al, 2009). For the pulverized lime sorbent the CO 2 and SO 2 capture in all the cycles was conducted using commercially available pulverized lime sorbent obtained from Greymont. As illustrated in Figure 3.12 the % CO 2 removal decreases from 50% in the first cycle to 20% in the second cycle. No CO 2 was captured by the sorbent in the third cycle. This shows that the sorbent sinters to a large extent in the system and loses all its reactivity in three cycles. For the cycles with Ca(OH) 2, the CO 2 and SO 2 capture in the first cycle was conducted using commercially available Ca(OH) 2 from Graymont. The CO 2 and SO 2 capture in the remaining cycles was achieved using sorbent that was reactivated by hydration. As can be seen from Figure 3.12, uniform CO 2 removal was achieved at a temperature of 625 ºC, a Ca:C ratio of 0.75 and a constant residence time. Figure 3.13 illustrates the Wt% CO 2 capture achieved by the Ca(OH) 2 sorbent during the 4 cycles shown in Figure The Wt% capture of the sorbent is maintained a constant at 52% over the 4 cycles. Hence hydration is very effective in maintaining sorbent reactivity over multiple cycles even in the subpilot scale facility. 3.8 CONCLUSIONS Among various reaction and process factors that are of importance to the CLP, the reactivity and recyclability of the calcium based sorbent are vital. The nature of CaO/CaCO 3 sintering that has been observed during multicyclic operation could pose a severe limitation to the commercialization of the process. In this Chapter, several 58

90 methods of improving the recyclability of CaO sorbents have been investigated including sorbent pretreatment, modification by addition of supports and reactivation. Reengineering the sorbent morphology by increasing the pore volume and surface area of the precursor has been found to be effective in improving the reactivity and recyclability of the sorbent. PCC sorbent that is synthesized from natural limestone has an improved performance due to presence of a greater surface area and pore volume. A similar improvement was observed by pretreating natural limestone with acetic acid which also increases the pore volume and surface area. Hydration of the sorbent also showed an improvement in sorbent performance. Although the addition of metal oxide supports to natural CaO sorbent improves the recyclability of the sorbent it reduces the amount of CO 2 that can be captured by a certain amount of sorbent due to the presence of the inert metal oxide. The effect calcination conditions of sorbent reactivity was also investigated and the sorbent loses one third to half of it original reactivity in a single cycle due to calcination at 950 ºC and 1000 ºC respectively. Modification of calcination conditions by the addition of steam in the calciner was found to improve the reactivity of the sorbent although a loss in reactivity was still observed over multiple cycles. Hydration of the sorbent as a reactivation method after every calcination cycle was found to be very effective in improving sorbent performance. The Wt% capture of the sorbent was found to be constant at 50% during multicyclic CO 2 capture with sorbent hydration in every cycle in both bench scale and subpilot scale tests. 59

91 Weight% Capture PCC PCC-Pelletised and Broken PCC-Whole Pellet Calcium Hydroxide Ground Lime Limestone Time(sec) Figure 3.1: Comparison in the CO 2 capture capacity of CaO sorbents obtained from different precursors. (Calcination conditions: T = 700 ºC, P = 1 atm, pure N 2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) 60

92 Comparison in the multicyclic conversion of the PCC powder sorbent and the PCC pelletised and crushed sorbent 1.0 PCC PCC pelletised and crushed 0.8 Conversion Time (sec) Figure 3.2: Comparison in the multicyclic conversion of PCC powder sorbent PCC pelletized and broken sorbent (Calcination conditions: T = 700 ºC, P = 1 atm, pure N 2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) 61

93 80 CO 2 Capture Capacity (%) Number of Cycles Formic Acid Acetic Acid MgO SiO 2 Al 2 O 3 PCC LC LH Figure 3.3: CO 2 capture capacity of pretreated and supported Ca-based sorbents over multiple carbonation calcination cycles (Calcination conditions: T = 700 ºC, P = 1 atm, pure N 2 carrier gas; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) 62

94 60 50 Wt% Capture Original Sorbent 0% 33% 50% Steam Concentration in Carrier Gas Figure 3.4: Effect of steam concentration in the calcination carrier gas on the CO 2 capture capacity of CaO sorbent (Calcination conditions: T = 900 ºC, P = 1atm) 63

95 Thermocouple And Pressure Guage Steam & Gas Mixture Steam Generator Water In Gas Gas Mixture Mixture Sorbent MFC MFC MFC MFC 64 Hydrocarbon Analyzer Back Pressure Regulator Heated Steel Tube Reactor Water Syringe Pump H 2 CO CO 2 Hydro carbons Analyzers (CO, CO 2, H 2, H 2 S) Heat Exchanger Water Trap Figure 3.5: Simplified flow sheet of the bench scale fixed bed experimental setup 64

96 60 50 Wt% Capture Original Sorbent Number of Cycles Figure 3.6: Effect of steam calcination on multicyclic carbonation and calcination of CaO sorbent (Calcination conditions: T = 900 ºC, P = 1 atm, carrier gas = 50%H2O/50% CO 2 ; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas) 65

97 C 500C Wt % Capture Original Sorbent Calcined Sorbent Water Hydration Steam Hydration Steam Hydration Figure 3.7: Effect of hydration conditions on sorbent reactivity 66

98 60 Wt% Capture C 600C 600C 10 0 Calcined Sorbent atms psig atms psig atms psig Hydration Pressure Figure 3.8: Effect of hydration pressure on sorbent reactivity (Hydration temperature = 600 ºC) 67

99 60 50 Wt% Capture Cycles Figure 3.9: Effect of steam hydration on sorbent reactivity over multiple calcinationhydration-carbonation cycles (Calcination conditions: T = 900 ºC, P = 1 atm, carrier gas = pure CO 2 ; Carbonation conditions: T = 650 ºC, P = 1 atm, 10% CO 2 /90% N 2 feed gas, Hydration conditions: T = 500 ºC, P ~ 1 atm, 90% H 2 O/10% N 2 feed gas) 68

100 Carbonator Co 2 Free Flue Gas Fresh Sorbent Spent Sorbent Purge Pure CO 2 69 Calciner Flue Gas Calcined sorbent Bag House Steam Hydrator Boiler Coal Air Figure 3.10: Process flow diagram of the CLP for CO 2 and SO 2 removal from combustion flue gas 69

101 Carbonation Reactor Figure 3.11: Snapshot of the sub-pilot scale facility of the CLP integrated with a coal fired combustor. 70

102 60 50 Cyclic CO 2 Removals Calcium Hydroxide Pulverized Ground Lime % CO 2 Removal Hydration Hydration Hydration Cycle Number Figure 3.12: Effect of hydration on the % CO 2 removed from the flue gas over multiple cycles 71

103 50 Wt % Capture Cycle Number Figure 3.13: Wt.% CO 2 capture achieved by the hydrated sorbent over multiple cycles 72

104 CHAPTER 4 ENHANCED CATALYTIC H 2 PRODUCTION FROM SYNGAS 4.1 INTRODUCTION This chapter describes the CLP for high purity H 2 production in the presence of a water gas shift catalyst from syngas. The CLP combines H 2 production with CO 2, sulfur and chloride capture from the syngas stream in a single stage reactor. Most H 2 production processes reported in literature require a separate sulfur clean up unit to prevent poisoning of the sorbent used for CO 2 capture. Sulfur is present in syngas in the form of H 2 S and carbonyl sulfide (COS). According to equilibrium calculations, at temperatures below 1027C (1300K) which exists in the gasifier, all sulfur radicals combine to form predominantly H 2 S which is close to 95% of the total sulfur content and COS forms the other 5%. (Jazbec et al, 2004) There have been studies conducted on the simultaneous calcination and sulfidation of calcium based sorbents at temperatures higher than 600 ºC. ( De Diego et al, 2004) There have also been studies on the sulfidation of CaCO 3 in the presence of CO 2 but the CO 2 was used only to maintain a high enough partial pressure to prevent the calcination of CaCO 3 (Fenouil et al, 1994, Fenouil, 1995, Zevenhoven et al, 1998, De Diego et al, 1999). However there is no mention of studies conducted on simultaneous CO 2 and sulfur 73

105 capture integrated with H 2 production in the literature. In the CLP described in the sections below, simultaneous CO 2 and H 2 S capture is achieved during the production of H CALCIUM LOOPING PROCESS (CLP) CONFIGURATION AND THERMODYNAMICS Several options are being investigated for the implementation of CCS on coal gasification systems including using solvents, sorbents, membrane and chemical looping processes. The CLP which is a calcium sorbent based chemical looping process, has the potential to reduce the cost and increase the efficiency of H 2 and/or electricity production from coal derived syngas by implementing the principles of process intensification (Fan et al, 2007, Fan et al, 2008, Ramkumar et al, 2009, Ramkumar et al, 2010). The CLP integrates the water-gas shift reaction with in-situ CO 2, sulfur, and halide removal at high temperatures in a single stage reactor. It utilizes a high temperature regenerable CaO sorbent which in addition to capturing the CO 2, enhances the yield of H 2 and simultaneously captures sulfur and halide impurities. The advantages of the CLP include: 1) The simplification of the coal to H 2 process by integration of the reaction and separation steps. This results in a decrease in the number of process units and combines the two staged water gas shift reactors (HTS and LTS), the CO 2, sulfur and halide capture units into a single stage reactor. 2) The enhancement in H 2 yield at high temperatures due to elimination of the 74

106 equilibrium limitation of the water gas shift reaction. 3) The potential to reduce excess steam requirement for the water gas shift reaction due to the enhanced thermodynamics of H 2 production by the combined water gas shift and carbonation reactions. 4) The potential to eliminate the requirement for water gas shift reaction catalyst due to H 2 production at high temperatures. 5) Although energy needs to be supplied for the endothermic calcination reaction, the carbonation reaction is exothermic at high temperatures of o C resulting in the production of high quality heat. By using a good strategy of heat integration it is possible to achieve high process efficiencies. 6) The calcination reaction results in the production of a pure sequestration ready CO 2 stream. A schematic of the CLP is shown in Figure 4.1. The CLP comprises the carbonation reactor, the calciner and the hydrator. In the carbonation reactor highpurity H 2 is produced while contaminant removal is achieved, in the calciner the calcium sorbent is regenerated and a sequestration-ready CO 2 stream is produced and in the hydrator the sorbent is reactivated. Thermodynamic analyses are conducted for the reactions occurring in each reactor using HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland) software. All the reactions shown in Figure 4.1 are found to be thermodynamically spontaneous but reversible and the extent of each of these reactions depends on the partial pressure of the respective gas species and the reaction 75

107 temperature. The following sections give a description of the three reactors using thermodynamic analyses The Carbonation Reactor The carbonation reactor comprises either a fluidized bed or an entrained flow reactor that operates at pressures ranging from 1 to 30 atm and temperatures of o C. The exothermic heat released from the carbonation reactor can be used to generate steam or electricity. In the carbonation reactor, the thermodynamic constraint of the water gas shift reaction is overcome by the incessant removal of the CO 2 product from the reaction mixture, which enhances H 2 production and obviates the need for excess steam addition. This is achieved by the concurrent water gas shift reaction and carbonation reaction of CaO to form CaCO 3 thereby removing the CO 2 product from the reaction mixture. In addition, the CaO sorbent is also capable of reducing the concentration of sulfur and halides in the outlet stream to ppm levels. The in-situ removal of CO 2 removes the equilibrium limitation of the water gas shift reaction thereby obviating the need for excess steam addition. Thermodynamic analysis, presented subsequently, predicts that the removal of H 2 S using CaO is inhibited by the presence of steam. Since almost all the steam is consumed in the enhanced water gas shift reaction, the removal of H 2 S is favored in the system. The reactions occurring in the carbonation reactor are as follows: Water gas shift reaction: CO + H 2 O H 2 +CO 2 (ΔH = -41 kj/mol) (4.1) 76

108 Carbonation: CaO + CO 2 CaCO 3 (ΔH = -178 kj/mol) (4.2) Sulfur capture (H 2 S) : CaO + H 2 S CaS + H 2 O (4.3) Sulfur capture (COS) : CaO + COS CaS + CO 2 (4.4) Halide capture(hcl) : CaO + 2HCl CaCl 2 +H 2 O (4.5) Thermodynamic analysis of reactions occurring in the carbonation reactor The equilibrium constants for the water gas shift reaction and the combined water gas shift and carbonation reaction for various temperatures are shown in Figure 4.2. The equilibrium constants are obtained using HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland). The equilibrium constant for the water gas shift reaction can be defined as shown below: H2 C 2 P H 2 P CO2 K eq1 = P CO P H2 O (4.6) where P CO2, P H2, P CO, P H2O are the partial pressures of CO 2, H 2, CO and H 2 O at equilibrium. The combined water gas shift and carbonation reaction is as follows: Combined water gas shift and carbonation: CO + H 2 O + CaO H 2 + CaCO 3 (6) The equilibrium constant for the combined water gas shift and carbonation reaction is defined as shown below: 77

109 K eq2 = P H2 P CO P H2 O (4.7) where K eq2 = K eq1 * K carb and K carb is the equilibrium constant of the carbonation reaction. The equilibrium of the water gas shift reaction decreases with an increase in the temperature resulting in low H 2 yields at higher temperatures. Hence, in the conventional water gas shift system, a LTS is used after the HTS to convert the CO slip and to increase the yield of H 2 in the presence of a LTS catalyst. The equilibrium constant of the combined water gas shift and carbonation reaction is significantly higher than the equilibrium constant of the water gas shift reaction alone, in the desired temperature of operation ranging from 500 to 750 o C. Hence, the CLP is capable of producing a much higher H 2 yield, and hence, purity due to almost complete CO conversion, when compared to the conventional H 2 production process. Equilibrium curves for the partial pressures of H 2 O (P H2O ), CO 2 (P CO2 ) and H 2 S (P H2S ) as a function of temperature, for the hydration, carbonation and sulfidation reactions with CaO were also obtained using HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland). The relationship between reaction temperature and equilibrium partial pressure of CO 2 and H 2 O for the carbonation and hydration reaction with CaO sorbent is shown in Figure 4.3. Hydration : CaO + H 2 O Ca(OH) 2 (4.8) 78

110 Carbonation and hydration of CaO are reversible reactions which occur depending on the conditions of temperature and partial pressures of CO 2 and H 2 O respectively. Carbonation of CaO occurs at conditions above the equilibrium PCO 2 curve while calcination of CaCO 3 occurs at conditions below the curve. Similarly hydration of CaO occurs above the PH 2 O curve while dehydration occurs at conditions below the curve. Figure 4.4 illustrates the equilibrium H 2 purity that can be obtained in the carbonation reactor for a feed gas containing 10% CO and balance nitrogen. Pure H 2 can be obtained even for stoichiometric S:C ratios and atmospheric pressure at temperatures below 500 ºC. The purity of H 2 begins to decrease with an increase in the temperature. Increase in the S:C ratio and pressure favor the production of pure H 2 at all temperatures. Bench scale experiments at various temperatures, pressures and S:C ratios have been conducted for a feed gas stream containing 10% CO and balance nitrogen. The results of these tests have been discussed in sections 4.3 and 4.4 of this chapter. For the reversible sulfidation of CaO, the extent of H 2 S removal will depend on the temperature and P H2O in the carbonation reactor. Figure 4.5 depicts the equilibrium H 2 S concentrations in the product H 2 stream, in ppm, for varying moisture concentrations (P H2O ) at 30 atm total system pressure. It can be seen that the equilibrium H 2 S concentration in the product H 2 stream increases with the increase in P H2O. At a temperature of 600 o C, the H 2 S concentration is 0.1 ppm for a P H2O of

111 atm and 1ppm for a P H2O of 0.2 atm. By operating the carbonation reactor at nearstoichiometric steam requirement, it is possible to obtain low concentrations of steam in the reactor system leading to low H 2 S concentrations of less than 1 ppm in the product stream. It can also be seen that the reactor system will favor H 2 S removal using CaO at around o C, which is a suitable temperature for the carbonation reaction as well. Similarly, the removal of COS is dependent of temperature and partial pressure of CO 2 in the carbonation reactor. Figure 4.6 illustrates the equilibrium COS concentrations in the product H 2 stream for varying CO 2 concentrations at 30 atm total system pressure. The equilibrium COS concentration in the product H 2 stream increases with the increase in temperature and partial pressure of CO 2. Since CO 2 in the carbonation reactor can be reduced to very low concentrations in the product stream by the CaO sorbent, COS capture by the CaO sorbent will occur to large extents. The equilibrium partial pressure of CO 2 for the carbonation reaction with CaO sorbent at 600 ºC is less than 0.01 atms as shown in Figure 4.3. Hence the equilibrium COS concentration in the product H 2 stream can be predicted from Figure 4.6 to be lower than atms. CaO sorbent is also capable of capturing HCl in the carbonation reactor. Similar to the removal of H 2 S by CaO sorbent, the extent of HCl capture also dependent on the temperature and the partial pressure of steam in the carbonator. The equilibrium HCl concentration in the product H 2 stream from the carbonation reactor for varying 80

112 temperatures and steam partial pressures is shown in Figure 4.7. Hence the reduction of S:C ratio in the carbonation reactor will improve the extent of HCl removal The Calciner The spent sorbent at the exit of the carbonation reactor is a mixture consisting of CaCO 3, CaO, CaS and calcium chloride (CaCl 2 ). The CaCO 3 in the spent sorbent mixture is regenerated back to CaO in the calciner. The calciner is operated at atmospheric pressure in a rotary or a fluidized bed system. The heat can be supplied directly or indirectly using a mixture of fuel and oxidant. From the thermodynamic curve for CaO and CO 2 shown in Figure 4.8, calcination is found to occur at temperatures above 890 o C in the presence of 1 atm of CO 2. Dilution of CO 2 in an indirectly fired calciner with steam or oxy-combustion of a fuel (syngas, natural gas, coal, etc) in a direct fired calciner will permit the calcination reaction to be conducted at temperatures lower than 890 o C. The reaction occurring in the calciner is: Calcination: CaCO 3 CaO + CO 2 (4.9) The regenerability of CaO sorbents over multiple cycles has been the major drawback of high temperature calcium based CO 2 capture processes. CaO sorbents are prone to sintering during the high-temperature calcination step. There is a decrease in sorbent reactivity even when steam is present in the calcination atmosphere. Over multiple cycles, the percentage of sintered CaO progressively increases and reduces the CO 2 capture capacity of the sorbent (Curran et al, 1967, Iyer et al, 2004, Sun et al, 81

113 2008, Abanades and Alvarez, 2003, Barker, 1973, Bhatia and Perlmutter, 1983, Saliban et al, 1996, Koji et al, 2003, Wang et al, 2005, Sun et al, 2007). Due to sintering, higher solid circulation or make-up rates need to be used to maintain a high level of CO 2 removal ( Romeo et al, 2009). Pretreatment methods have been developed to reduce the decay in reactivity, which involve hydration of the sorbent (Koji et al, 2003, Iyer, 2003, Manovic et al, 2007,Fennell et al, 2007, Sun et al, 2008), preheating and grinding of the sorbent (Manovic and Anthony, 2008) and synthesis of novel sorbents by physical or chemical modification of the precursor (Sun et al, 2008, Gupta and Fan, 2002, Sakadjian, 2004, Salvador, 2003, Reddy and Smirniotis, 2004, Lu et al, In the CLP process, the addition of a sorbent reactivation step by hydration, as part of the carbonation-calcination cycle is used to reverse the effect of sintering during each cycle and thus maintain the sorbent reactivity. (Fan et al, 2008) Sorbent hydration has been found to be effective in maintaining sorbent reactivity as shown in Chapter Sorbent Reactivation by Hydration The calcination process causes sintering of the sorbent which results in a reduction in its reactivity and hence, the overall CO 2 capture capacity. The hydration process reverses this effect by increasing the pore volume and surface area available for reaction with the gas mixture. Figure 4.9 shows the partial pressure of steam required for hydration of the sorbent at various temperatures. Hydration occurs at atmospheric pressure at temperatures below 500 o C. At temperatures of 600 o C and above hydration occurs at steam partial pressures of above 4 atms. Operation of the 82

114 hydrator at high temperatures reduces the extent of cooling and reheating of the solids required between the calciner and the carbonation reactor. This aids in reducing the parasitic energy consumption of the process. Hydration at higher temperatures also produces high quality heat which can be used to produce steam or electricity. Depending of the reactivity of the calcined sorbent, a fraction of the calcined sorbent or the entire stream of sorbent could be hydrated. The reactivity of the calcined sorbent will depend on a variety of reasons including the type of calciner (direct or indirect), mode of calcination (rotary kiln, fluidized bed or entrained bed), the temperature of calcination and the gas atmosphere within the calciner. The reaction occurring in the hydrator is shown below: Hydration: CaO + H 2 O = Ca(OH) 2 (4.8) The Ca(OH) 2 from the hydrator is conveyed to the carbonation reactor where it dehydrates to produce high reactivity CaO and steam. The steam obtained from the dehydration reaction is consumed in the water gas shift reaction. The advantage of this reactivation process is that no excess steam is required for hydration. Part or all of the steam required for the water gas shift reaction is supplied to the hydrator depending on the fraction of the calcined sorbent that is sent to the hydrator for reactivation. 83

115 4.3 MATERIALS AND METHODS Chemicals, Sorbents, and Gases The HTS and STC catalyst were procured from Süd-Chemie Inc., Louisville, KY. The HTS catalyst consists of iron (III) oxide supported on chromium oxide while the STC catalyst consists of cobalt-molybdenum on alumina support. The CaO precursor for the tests conducted in this chapter was PCC. PCC was synthesized from Ca(OH) 2 obtained from Fisher Scientific (Pittsburgh, PA). The high surface area PCC (BET analysis; SA 49.2 m 2 /g; PV 0.17 cm 3 /g) was synthesized using a dispersant modified wet precipitation technique. The anionic dispersant used in this process was N40V, supplied by Ciba Specialty Chemicals (Basel, Switzerland). PCC was synthesized by bubbling CO 2 through a slurry of hydrated lime. The neutralization of the positive surface charges on the CaCO 3 nuclei by negatively charged N40V molecules forms CaCO 3 particles characterized by a higher surface area/pore volume and a predominantly mesoporous structure. Details of this synthesis procedure have been reported elsewhere (Fan and Gupta, 2006, Fan et al, 1998). The feed gas for all the H 2 production tests was a mixture of 10%CO and 90% Nitrogen (N 2 ) Fixed Bed Reactor Unit Setup Figure 4.10 shows the bench scale, fixed bed reactor system, used for studying H 2 production at various process conditions. The bench scale reactor is coupled with a set of continuous gas analyzers which detect concentrations of CO, CO 2, H 2 S, CH 4 and 84

116 H 2 in the product stream. The reactor setup is capable of handling high pressures and temperatures of up to 21 atms and 900 ºC respectively, which are representative of the conditions in a commercial syngas to H 2 system. The mixture of gases from the cylinders is regulated and sent into the fixed bed reactor by means of mass flow controllers that can handle pressures of about 21 atms. From the mass flow controllers the reactant gases flow to the steam generating unit. The steam generating unit is maintained at a temperature of 200 o C and contains a packing of quartz chips which provide a large surface area of contact and mixing between the reactant gases and steam. The steam generating unit not only facilitates the complete evaporation on the water being pumped into the steam generating unit but it also serves to preheat the reactant gases entering the reactor. The reactor, which is heated by a tube furnace, is provided with a pressure gauge and a thermocouple to monitor the pressure and temperature within. The rector consists of two concentric sections; the inner section is filled with the catalyst or sorbent-catalyst mixture and the outer section provides a preheating zone for the gases before they come in contact with the bed of solids. The sorbent and catalyst loading section of the reactor is detachable which enables easy removal and loading of the sorbent. The reactant gases leaving the reactor enter a back pressure regulator which builds pressure by regulating the flow rate of the gases and is capable of building pressures of up to 68.9 atm. The back pressure regulator is very sensitive and the pressure within the reactor can be changed quickly without any fluctuations. In addition, the back pressure regulator is also 85

117 capable of maintaining a constant pressure for a long period of time. The valve seat material of the regulator is made of PEEK which is corrosion resistant to acidic H 2 S vapors, which makes it suitable for conducting sulfur removal experiments. As shown in Figure 4.10, the inlet of the back pressure regulator is connected to the reactor rod and the outlet is connected to a heat exchanger. Since the entire section of the equipment setup upstream of the backpressure regulator will be exposed to high pressures, flexible stainless steel lines are used to withstand the pressure and the reactor is constructed from inconel which is resistant to corrosion due the high pressure high temperature steam and H 2 S gas. The product gas mixture exiting the back pressure regulator is then cooled in a heat exchanger using chilled ethylene glycol-water mixture to condense the unconverted steam. The product gas at the exit of the heat exchanger is dried in a desiccant bed and is sent to a set of continuous analyzers capable of determining the concentrations of CO, CO 2, H 2 S, CH 4 and H 2 in the gas stream Water Gas Shift Reaction Testing The water gas shift reaction was conducted using the catalysts obtained from Süd-Chemie. These experiments were conducted as base line experiments to determine the conditions for maximum water gas shift catalytic activity at different ranges of temperatures ( o C), S:C ratios and pressures, which are beyond the commercial mode of operation, but are of interest for the CLP. Catalyst particles were used in a 86

118 fixed bed reactor setup for all the experiments. The total flow rate of the gases through the reactor was maintained a constant at 725 sccm for all the experiments and the concentration of CO in the reaction mixture was maintained at 10.3 % g of the catalyst was loaded into the reactor and the pressure, temperature and gas flow rates were adjusted for each run. The dry gas compositions at the outlet of the reactor were monitored continuously using the CO, CO 2, H 2 S, CH 4 and H 2 gas analyzers Simultaneous Water Gas Shift and Carbonation The combined water gas shift and carbonation reaction was conducted using a sorbent (CaO) to catalyst ratio of 10:1 by weight. The CaCO 3 sorbent was calcined by heating the sorbent-catalyst mixture to 700 o C in a stream of N 2 until the CO 2 analyzer confirmed the absence of CO 2 in the outlet stream. At the end of calcination, the feed gas was switched from nitrogen to a mixture of 10% CO and 90% nitrogen for the combined water gas shift and carbonation reaction. The combined water gas shift and carbonation reaction experiments were conducted at 600, 650, and 700 C with a S:C ratio of 3:1, 2:1, 1:1 at various pressures ranging from 1-21 atm Catalyst Pretreatment It is imperative to understand the HTS catalyst composition during calcination of the sorbent which occurs in the presence of a CO 2 atmosphere at high temperature. Iron oxide occurs in three different phases: Hematite (Fe 2 O 3 ), magnetite (Fe 3 O 4 ) and wustite (FeO). The active phase of the HTS catalyst is magnetite. However, in the 87

119 presence of an oxidizing atmosphere, like CO 2 or steam, the magnetite phase gets oxidized to hematite which is likely during the calcination step. This is evident from the iron oxide phase diagram for a CO-CO 2 system (Ross, 1980). Thus, a pretreatment procedure was developed which consists of treating the oxidized catalyst in a 20%/80% of H 2 /H 2 O atmosphere at 600 o C which reduces the hematite to magnetite. The effectiveness of the pretreatment procedure was confirmed by X-ray diffraction analyses of the HTS catalyst before and after the pretreatment procedure. The HTS catalyst as obtained contains hematite phase as shown in Figure The catalyst is subsequently subjected to the pretreatment procedure which changes its phase to the active magnetite form as shown in Figure In the commercial deployment of the CLP, pretreatment of the catalyst can be avoided by using a fixed fluidized bed reactor for the carbonation reactor in which the catalyst remains in the carbonation reactor while the CaO sorbent is looped between the carbonation reactor and the calciner. In this configuration the, catalyst is never exposed to oxidizing gases in the calciner. No deactivation of the STC catalyst was observed during calcination Combined H 2 Production with H 2 S Removal To study the effect of sulfur on the CLP, 5000 ppm of H 2 S was mixed with CO, N 2 and steam before being sent to the reactor. The H 2 production tests were conducted in the presence of the catalyst and CaO sorbent. 88

120 4.4 RESULTS AND DISCUSSION Effect of Process Parameters on the Extent of Water Gas Shift Reaction using HTS Catalyst An investigation of the water gas shift reaction in the presence of a HTS catalyst was conducted in the bench scale fixed bed reactor to determine the effect of temperature, pressure and S:C ratio on the extent of reaction. Figure 4.13a shows the CO conversion profiles for increasing reaction temperatures and S:C ratios at ambient pressures. The CO conversion increases with increasing temperature as it approaches the equilibrium value at an optimal temperature ( o C) beyond which it begins decreasing monotonically. At a pressure of 1 atm and a S:C ratio of 3:1, the conversion increases from 45.8 % at 450 ºC to 83.2 % at 600 o C. Beyond 600 o C, the conversion decreases and at 800 ºC, it is 69.4%. This decrease in conversion with the increase in temperature is observed due to the thermodynamic limitation of the water gas shift reaction. Thus at lower temperatures although the equilibrium constant is high, the reaction rate is low. At high temperatures, although the reaction is very fast, the equilibrium constant is low. Consequently maximum conversion is reached at an optimum temperature at which both the kinetics and the reaction equilibrium are favorable. From Figure 4.13a, it can also be seen, as expected, that the conversion increases with the increase in the S:C ratio for all temperatures. At a temperature of 650 o C, the conversion is 63.5% for a S:C ratio of 1:1, 71.6% for 2:1 and 80.28% for 3:1. As can be seen in Figure 4.13b, the effects of reaction temperatures and S:C ratios 89

121 on CO conversion at 21 atm follow the same trend as that at 1 atm. In addition, below ºC, the CO conversion at 21 atm is greater than at 1 atm due to an increase in the rate of the reaction with increase in pressure. The observed partial pressure ratios were computed for different S:C ratios, temperatures and pressures and were compared with the equilibrium values obtained from HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland). The observed partial pressure ratio (K obs ) was computed from the experimental data and is defined as the ratio of the product of partial pressures of the products to that of the reactants as given below: Kobs = P H2 P CO2 P CO P H2 O (4.6) As shown in Figures 4.14a and 4.14b it was found that each value of the observed partial pressure ratio (K obs ) was within the equilibrium value. From the figures it can be seen that the partial pressure ratio increases with an increase in the temperature till it approaches equilibrium and then decreases along the equilibrium curve. Besides, as the pressure increases, the system is closer to equilibrium for both S:C ratios of 1:1 and 3:1. This can be explained by the increase in the rate of the reaction with increase in pressure. 90

122 4.4.2 Enhancing the Water Gas Shift Reaction by In-situ CO 2 Removal (HTS Catalyst and CaO Sorbent) From Figures 4.13a and b, it can be observed that the CO conversion achieved in the presence of the HTS catalyst is only 80-90% even at a high pressure of 21 atms and a high S:C ratio of 3:1. At atmospheric pressure and a stoichiometric S:C ratio, a low CO conversion of 20-60% is obtained. In order to enhance the H 2 yield, CaO sorbent could be introduced into the H 2 production reactor for in-situ CO 2 removal from the reaction zone. This increase in H 2 yield can be explained by the LeChatlier s principle where the simultaneous CO 2 removal drives the equilibrium limited water gas shift reaction forward. This concept was demonstrated by conducting the combined water gas shift and carbonation reaction in the presence of the calcined PCC sorbent and HTS catalyst in the fixed bed reactor. Figures 4.15a illustrates the typical breakthrough curves obtained during the combined water gas shift reaction and carbonation reaction for the N 2 free dry product gas compositions. High purity H 2 is produced in the pre-breakthrough region due to in-situ CO 2 removal by the sorbent. As the sorbent gets exhausted the breakthrough region occurs followed by the postbreakthrough region in which all the sorbent has been converted to CaCO 3 and H 2 production occurs in the presence of the HTS catalyst. The concentrations of CO and CO 2 in the product gas mixture are very low in the pre-breakthrough region and increase in the breakthrough region due to the depletion of the sorbent. Figure 4.15b illustrates the typical breakthrough curve obtained for CO conversion. 91

123 Effect of Pressure The effect of pressure on the combined water gas shift and carbonation reaction for S:C ratios of 3:1 and 1:1 are shown in Figures 4.16a and 4.16b. The purity of H 2 produced in the pre-breakthrough region of the curves increases with the increase in pressure. From Figure 4.16a it can be observed that during the initial pre-breakthrough period for a S:C ratio of 3:1, 95.6 % H 2 is produced at 1 atm, 99.7% pure H 2 is obtained at 11 atm, and 99.8% pure H 2 is produced at 21 atm. The extent of the prebreakthrough region, which signifies the extent of conversion of the CaO sorbent, also increases with the increase in pressure. A similar observation is made for a lower S:C ratio of 1:1 as shown in Figure 4.16b. It can be inferred that higher pressure results in increased partial pressure of CO 2 which enhances the extent and rate of carbonation due to higher driving force. This consequently results in enhanced CO conversion, sorbent conversion and H 2 yield Effect of S:C Ratio The effect of S:C ratio on the CO conversion and H 2 purity for the combined water gas shift and carbonation reaction are shown in Figures 4.17a, 4.17b, 4.17c and 4.17d. The effect of S:C ratio at atmospheric pressure is shown in Figures 4.17a/b while that at 21 atm is shown in Figures 4.17c/d. It can be seen that at atmospheric pressure a reduction in the S:C ratio results in a decrease in the CO conversion and associated H 2 purity. However, at a higher pressure of 21 atm, almost 100 % CO 92

124 conversion and H 2 purity is achieved for all the three S:C ratios in the pre-breakthrough region. This can again be attributed to the higher partial pressure of CO 2 contributing to enhanced carbonation kinetics which plays a key role in driving the water gas shift reaction to completion. Besides, from a process design and cost perspective, operation at high pressures clearly illustrates the benefit of using a smaller amount of steam for a high CO conversion, resulting in cost savings Effect of Temperature The effect of temperature on the CO conversion achieved in the presence of the CaO sorbent and HTS catalyst at atmospheric pressure and a S:C ratio of 3:1 is depicted in Figure In the pre-breakthrough region, the CO conversion decreases with the increase in temperature due to equilibrium limitations of the combined water gas shift and carbonation reaction. In the initial pre-breakthrough region, a CO conversion of 95% is obtained at 600 and 650 ºC and it decreases to 90% at 700 ºC. Figures 4.19(a) and (b) illustrate the effect of temperature on the combined reactions at 21 atm and S:C ratios of 3:1 and 1:1 respectively. At a high S:C ratio of 3:1, there is almost no change in the CO conversion with the change in temperature as can be seen in Figure 4.19a. On decreasing the S:C ratio to the stoichiometric amount, it is observed in Figure 4.19b, that temperature plays a significant role in the extent of CO conversion and a temperature of 600 ºC is optimum for achieving high CO conversions of 99.7%. Thus, from a process design perspective this defines the 93

125 operating temperature for achieving high CO conversions and H 2 yield while maintaining low steam requirements Simultaneous Water Gas Shift, Carbonation and Sulfidation Reaction Testing Since syngas obtained from the gasifier contains 0.5 to 4% sulfur mostly in the form of H 2 S, the effect of sulfur and the extent of its removal by the CaO sorbent were determined on the combined water gas shift and carbonation reaction. Integrated H 2 production, CO 2 and H 2 S removal using calcium sorbent and HTS catalyst was investigated by the addition of 5000 ppm of H 2 S to the fixed bed reactor feed. The calcium sorbent was used to simultaneously capture H 2 S and CO 2 while enhancing H 2 production in the presence of the HTS catalyst. As illustrated in Figure 4.20(a), it was found that H 2 S concentration in the outlet H 2 stream is reduced to a few ppm in the pre-breakthrough region by the reaction of H 2 S with the CaO sorbent. In the thermodynamics section of the sulfidation of CaO, illustrated in Figure 4.5, it was observed that the extent of H 2 S removal is inhibited by the presence of a high partial pressure of steam in the system. This concept is demonstrated in the experimental results depicted in Figures 4.20a and Figure 4.20a illustrates the entire breakthrough curve of H 2 S concentration in the product H 2 stream with the pre-breakthrough and breakthrough regions. Figure 4.21 is a magnified image of the pre-breakthrough region in Figure 4.20a and it shows that with the increase in S:C ratio, the H 2 S concentration in the H 2 product increases. At a lower S:C 94

126 ratio of 1:1, the H 2 S in the outlet stream is lower than 1 ppm while at an S:C ratio to 3:1, the H 2 S concentration increases from 2 ppm to 30 ppm in 750 secs during the prebreakthrough region. At a S:C ratio of 1:1, in the pre-breakthrough region, the carbonation reaction enhances the water gas shift reaction which results in the consumption of most of the steam. Hence H 2 S removal by the calcium sorbent is enhanced and the H 2 S composition in the outlet stream is low. As the reaction proceeds, the CaO sorbent gets consumed to form CaCO 3 and CaS resulting in the breakthrough curve seen in Figure 4.20a. Since the steam composition in the system is higher for an S:C ratio of 3:1 the H 2 S concentration in the product stream is higher. During the breakthrough region, H 2 S reacts with both CaO and CaCO 3. The post-breakthrough region is not visible in Figure 4.20a as the H 2 S concentration in the product will keep increasing with time until all the CaCO 3 is also converted to CaS. In the post-breakthrough region the H 2 S concentration in the product will be equal to the H 2 S concentration in the feed stream. Figure 4.20b illustrates the change in CO conversion with respect to time for S:C ratios of 3:1 and 1:1. In the pre-breakthrough region, the CO conversion for an S:C ratio of 3:1 is slightly higher than that for 1: Effect of Catalyst Type on the Water Gas Shift Reaction A STC procured from Sud Chemie was also tested for its suitability in the CLP. The water gas shift reaction was conducted in the presence of the STC at a range of 95

127 temperatures ( ºC), pressures (1-21 atms) and S:C ratios (1:1-3:1). The performance of the HTS catalyst was compared with that of the STC catalyst. As illustrated in Figure 4.22, it was found that there is an increase in the CO conversion with an increase in the S:C ratio for both the STC as well as the HTS catalyst. It was also found that at temperatures below 650 ºC the CO conversion in the presence of the HTS catalyst is higher that the CO conversion obtained in the presence of the STC. 550 ºC-650 ºC is found to be the optimum temperature of operation in the presence of the HTS catalyst and ºC is found to the optimum temperature of operation for the STC. The water gas shift reaction was conducted in the presence of the STC at a range of temperatures (400 ºC to 800 ºC) and a range of pressures (1-21 atms) as shown in Figure It was found that the conversion increases with increase in temperature due to improved kinetics and beyond a particular temperature decreases since the reaction is exothermic. The conversion increases with an increase in pressure and at each pressure the conversion reaches a maximum value at a particular temperature. As shown in Figure 4.23, with increase in pressure the temperature for maximum conversion decreased from 750 ºC at 1 atm, 700 ºC at 11 atms to 600 ºC at 21 atms. As shown earlier, at atmospheric pressures, the HTS catalyst gives higher conversion than the STC at all temperatures. As shown in Figure 4.24, at a pressure of 11 atms the HTS catalyst gives maximum CO conversion at a temperature of 550 ºC 96

128 while the STC gives maximum conversion at a temperature of 700 and 750 ºC. It was found that with the increase in temperature above 675 ºC the conversion in the presence of the STC increases above the conversion obtained in the presence of the HTS catalyst. At a temperature of 600 ºC the conversion in the presence of STC is 50% while in the presence of the HTS the conversion is 64%. At a temperature of 700 ºC, the conversion in the presence of the sulfur tolerant is 60% while in the presence of HTS is 58%. Hence in this case it is very important to determine the rate of carbonation at different temperatures as equilibrium conversion for carbonation decreases at temperatures above 650 ºC. The temperature of operation of the sorbent and the catalyst should be similar for production of the highest purity H 2. Hence combined water gas shift and carbonation reactions need to be conducted in the presence of the catalyst and CaO sorbent to determine the catalyst best suited for the reaction. A similar trend was observed at 21 atms and as the temperature was increased the conversion obtained in the presence of the STC increased and was equal to the conversion in the presence of the HTS catalyst at temperatures above 650 ºC. As shown in Figure 4.25, at a temperature of 600 ºC, conversion in the presence of the HTS catalyst is 70% while that in the presence of the STC is 65%. But at a temperature of 700 ºC both catalysts give the same conversion of 65%. The effect of H 2 S on the activity of the HTS and STC catalyst was investigated. Figure 4.26 depicts the comparison in CO conversion achieved at atmospheric pressure in the presence and absence of H 2 S in the inlet gas steam. It was found that at 650 ºC, 97

129 the CO conversion decreases in the presence of H 2 S for both the HTS catalyst and the STC catalyst. It has been shown in literature, that the HTS catalyst still retains half its original activity in its sulfided form and the same inference is obtained from Figure 4.26 at both S:C ratios of 3:1 and 1:1 (Hla et al, 2009). Although the decrease in the conversion obtained in the presence of the STC catalyst is very low when compared to that in the HTS catalyst, it was found that even in the presence of H 2 S, the HTS catalyst shows higher CO conversion at a temperature of 650 ºC. The combined water gas shift and carbonation reaction was conducted at different temperatures at atmospheric pressure in the presence of the STC and CaO sorbent. As shown in Figure 4.27 it can be seen that the conversion decreases as the temperature is increased and it is highest at 650 ºC. In the 650 to 750 ºC temperature range, although conversion of CO in the presence of STC increases with temperature till 750 ºC (in Figure 4.22), the combined reaction conversion decreases with increase in temperature (Figure 4.27). This is because maximum CO 2 removal occurs at 650 ºC and at temperatures higher than 650 ºC, the equilibrium conversion for the carbonation reaction decreases. The enhancement in CO conversion on the addition of CaO sorbent to the STC catalyst is illustrated in Figure At both ratios of 3:1 and 1:1, the CO conversion was found to be the highest in the presence of the HTS catalyst and CaO sorbent. Although the CO conversion is increased by the addition of CaO to the STC, it is still 98

130 lower than the conversion obtained in the presence of the mixture of HTS catalyst and CaO sorbent. 4.5 CONCLUSIONS Enhancement in the production of high purity H 2 from syngas can be achieved using CaO sorbent that can drive the equilibrium limited water gas shift forward by insitu removal of CO 2. Thermodynamic analyses for the reactions occurring in the carbonation reactor, calciner and hydrator were conducted to determine the operating window for various process parameters. Operating at near stoichiometric steam conditions is advantageous for simultaneous sulfur removal to low levels in the product H 2 stream. Bench scale experimental data demonstrate that greater than 99% pure H 2 can be produced at high temperatures and pressures. For near stoichiometric conditions, high CO conversion and H 2 purity can be obtained at high pressures and an optimal temperature of 600 ºC. This operating temperature was also found to be favorable for simultaneous H 2 S removal to <1ppm in the product H 2 stream. At atmospheric pressure, a water gas shift catalyst which has high activity in the optimal temperature range of the carbonation reaction ( ºC) in the presence of sulfur is beneficial. The HTS catalyst with CaO sorbent results in the production of a high purity H 2 stream at atmospheric pressure. This is important in situations where the conversion of fuel gas to H 2 at atmospheric pressure is beneficial. Further investigation conducted to determine whether a catalyst is required for the production of H 2 at high pressures within a short residence time is described in Chapter 5. 99

131 Reaction Hydrogen Purge Stream Pure CO 2 gas Fresh Sorbent Regeneration Integrated Hydrogen reactor Syngas Net Heat Output Heat Input Calciner Dehydration : Ca(OH) 2 CaO + H 2 O WGSR : CO 2 removal : CO + H 2 O CO 2 + H 2 CaO + CO2 CaCO3 Sulfur : CaO + H2S CaS + H 2O Halide : CaO + 2HX CaX 2 + H 2 O Calcination: CaCO3 CaO + CO2 100 Reactivation Hydrator Heat Output H 2 O Hydration : CaO + H 2 O Ca(OH) 2 Figure 4.1: Schematic of the CLP 100

132 K(Equilibrium Constant) WGSR WGSR + Carbonation Temperature (ºC) Figure 4.2: Thermodynamic data illustrating the equilibrium constants of the water gas shift reaction and the combined water gas shift and carbonation reaction 101

133 Equlibrium Partial Pressure (atm) CaO + CO 2 CaCO 3 CaO + H 2 O Ca(OH) 2 Hydration Carbonation Dehydration Calcination P H2O 2 P CO Temperature (C) Figure 4.3: Thermodynamic data for the hydration and carbonation of CaO sorbent 102

134 Hydrogen purity (%) 60 S/C = 1: atm psig atm psig 01 psig atm S/C = 3: atm psig atm psig 01 psig atm Temperature (C) Figure 4.4: Equilibrium H 2 purity in the carbonator at varying temperatures, pressures and S: C ratios. (Feed gas: 10% CO and balance nitrogen) 103

135 Equilibrium H 2 S Conc (ppm) with 30 atm total pressure CaO+ H 2 S CaS + H 2 O 20 atm (P H2O ) 2 atm (P H2O ) 0.2 atm (P H2O ) 0.02 atm (P H2O ) CLP Typical Gasifier Temperature ( o C) Figure 4.5: Thermodynamic data for the sulfidation (H 2 S) of CaO with varying steam partial pressures. (P Total = 30 atm) 104

136 Equilibrium COS Conc (ppm) with 30 atm total pressure 10 1 atm (P CO2 ) 0.1 atm (P CO2 ) atm (P CO2 ) atm (P CO2 ) Temperature ( o C) Figure 4.6: Thermodynamic data for predicting the equilibrium COS concentration for CaO sulfidation with varying CO 2 concentration (P Total = 30 atm) 105

137 Equilibrium HCl Conc (ppm) with 30 atm total pressure atm 2 atm 0.2 atm 0.02 atm Temperature ( o C) Figure 4.7: Thermodynamic data for predicting the equilibrium HCl concentration for CaO reaction with HCl with varying steam concentration (P Total = 30 atm) 106

138 Equlibrium Partial Pressure (atm) P CO2 Carbonation Calcination Temperature ( o C) Figure 4.8: Thermodynamic data for the carbonation of CaO 107

139 10 2 Equlibrium Partial Pressure (atm) Hydration Dehydration Temperature ( o C) Figure 4.9: Thermodynamic data for the hydration of CaO 108

140 Thermocouple And Pressure Guage Steam & Gas Mixture Steam Generator Water In Gas Gas Mixture Mixture Vent Sorbent & Catalyst Powder MFC MFC MFC MFC 109 Hydrocarbon Analyzer Back Pressure Regulator Heated Steel Tube Reactor Water Syringe Pump H 2 CO CO 2 Hydro H 2 S carbons Analyzers (CO, CO 2, H 2, H 2 S) Heat Exchanger Water Trap Figure 4.10: Simplified flow sheet of the bench scale experimental setup 109

141 FE2O3HTS Lin (Counts) Theta - Scale FE2O3HTS - File: HTS.RAW - Type: 2Th/Th unlocked - Start: End: Step: Step time: 1.8 s - Temp.: 25 C (Room)- Time Started: 0 s - 2-Theta: Theta: Chi: Operations: Smooth Background 1.000,1.000 Background 1.000,1.000 Import (C) - Iron Oxide - Fe2O3 - Y: % - d x by: 1. - WL: I/Ic PDF Figure 4.11: X-ray diffraction patters of the HTS catalyst before pretreatment (hematite) 110

142 Lin (Counts) HTS6001HR Theta - Scale HTS6001HR - File: HTS600~1.RAW - Type: 2Th/Th unlocked - Start: End: Step: Step time: 1.5 s - Temp.: 25 C (Room) - Time Started: 0 s - 2-Theta: Theta: Chi: Operations: Smooth Smooth Import (C) - Magnetite - Fe.99Fe1.97Cr.03Ni.01O4 - Y: % - d x by: 1. - WL: Cubic - a b c alpha beta gamma Face-centred - Fd-3m (227) Figure 4.12: X-ray diffraction patters of the HTS catalyst after pretreatment (magnetite) 111

143 CO Conversion :1 2:1 3: Temperature ( o C) (a) CO Conversion :1 2:1 3: Temperature ( o C) Figure 4.13: Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction in the presence of HTS catalyst at (a) 1 atm (b) (b)

144 Partial Pressure Ratio atm 11 atm 21 atm Theoretical Temperature ( o C) (a) Partial Pressure Ratio atm 11 atm 21 atm Theoretical Temperature ( o C) Figure 4.14: Effect of reaction temperature and pressure on the observed partial pressure ratio for the water gas shift reaction in the presence of HTS catalyst at a S:C ratio of (a)1:1 (b)3:1 (b) 113

145 CO CO2 H2 Gas Compositions Time (sec) (a) CO Conversion Time (sec) Figure 4.15: Typical curves for the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst depicting (a) Gas composition (mol%) and (b) CO conversion (650 ºC, 1 atm, S:C ratio of 3:1) (b) 114

146 100 H2 Gas Composition (%) atm 11 atm 21 atm Time(sec) (a) 100 H2 Gas Composition (%) atm 11 atm 21 atm Time (sec) Figure 4.16: Effect of pressure on purity of H 2 produced during the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at a S:C ratio of (a) 3:1 (b) 1:1 (650 ºC) (b) 115

147 CO Conversion :1 2:1 3: Time (sec) (a) 100 H2 Gas Composition(%) Time (sec) (b) 1:1 3:1 Continued Figure 4.17: Effect of S:C ratio on the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 650 ºC (a) CO conversion at 1 atm (b) H 2 gas composition at 1 atm (c) CO conversion at 21 atm (d)h 2 gas composition at 21 atm 116

148 Figure 4.17 continued CO Conversion Time (sec) 1:1 2:1 3:1 (c) 100 H2 Gas Composition (%) :1 2:1 3: TIme (sec) (d) 117

149 CO Conversion C 650C 700C Time(sec) Figure 4.18: Effect of temperature on CO conversion by the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 1 atm and S:C ratio of 3:1 118

150 CO Conversion C 650 C 700 C Time (sec) (a) CO Conversion C 650 C 700 C Time(sec) Figure 4.19: Effect of temperature on CO conversion by the combined water gas shift and carbonation reaction in the presence of CaO sorbent and HTS catalyst at 21 atm and S:C ratio of (a) 3:1 (b) 1:1 (b) 119

151 H2S Concentration (ppm) :1 3: Time(sec) (a) :1 3:1 CO Conversion Time(sec) Figure 4.20: Effect of S:C ratio on (a) the composition of H 2 S in the H 2 stream and (b) CO conversion in the presence of the catalyst and sorbent during the simultaneous water gas shift, carbonation and sulfidation reaction (600 ºC, 1 atm) (b) 120

152 30 H2S concentration (ppm) Time(sec) Figure 4.21: Effect of S:C ratio on the composition of H 2 S in the H 2 stream during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent and HTS catalyst (600 ºC, 1 atm) 121

153 Conversion STC- 3:1(S:C) STC - 1:1(S:C) HTS - 3:1(S:C) HTS - 1:1(S:C) Temperature (C) Figure 4.22: Effect of S:C ratio and temperature on CO conversion during the water gas shift reaction in the presence of STC and HTS catalyst 122

154 70 60 CO Conversion psig 1 atm atm psig atm psig Temperature (C) Figure 4.23: Effect of reaction temperature on CO conversions for various pressures at an S:C ratio of 1:1 for the STC (0.25g STC, Total flow = slpm) 123

155 CO Conversion Sulfur Tolerant Catalyst HTS Temperature (C) Figure 4.24: Effect of reaction temperature on CO conversions for the HTS and STC at 11 atms and S:C ratio of 1:1(Total flow = slpm) 124

156 CO Conversion Sulfur Tolerant Catalyst HTS Temperature (C) Figure 4.25: Effect of reaction temperature on CO conversions for the HTS and STC at 21 atms and S:C ratio of 1:1(Total flow = slpm) 125

157 CO Conversion Time (secs) STC - 1:1(S:C) - In the presence of H2S STC - 1:1(S:C) - In the absence of H2S STC - 3:1(S:C) - In the presence of H2S STC - 3:1(S:C) - In the absence of H2S HTS - 1:1(S:C) - In the presence of H2S HTS - 1:1(S:C) - In the absence of H2S HTS - 3:1(S:C) - In the presence of H2S HTS - 3:1(S:C) - In the absence of H2S Figure 4.26: Effect of S:C ratio, type of catalyst and effect of H 2 S on CO conversion during the water gas shift reaction (650 ºC, 1atm) 126

158 Conversion C 700C 750C Time (sec) Figure 4.27: Effect of temperature on CO conversion (Temperature=650 C, Pressure = 1 atm, S:C ratio= 1:1) 127

159 CO Conversion Time (sec) 3:1(S:C)-STC with CaO 1:1(S:C)-STC with CaO 3:1(S:C)-STC 1:1(S:C)-STC 3:1(S:C)-HTS with CaO 1:1(S:C)-HTS with CaO Figure 4.28: Comparison in the CO conversion obtained at different S:C ratios for different sorbent and catalyst mixtures (650 ºC, 1atm) 128

160 CHAPTER 5 ENHANCED NON-CATALYTIC H 2 PRODUCTION FROM SYNGAS 5.1 INTRODUCTION In Chapter 4, H 2 production with contaminant removal in the presence of CaO sorbent and a water gas shift catalyst was investigated. The presence of the sorbent and catalyst in the carbonation reactor results in the production of high purity H 2 with low levels of CO, CO 2 and sulfur but introduces issues and costs associated with the separation of the sorbent and catalyst prior to calcination or pretreatment of the catalyst to the active form after its deactivation in the presence of CO 2 in the calciner at high temperatures, replacement of the spent catalyst, deactivation of the catalyst in the presence of sulfur impurities (H 2 S) and the use of expensive STC. In an attempt to further simplify the process, the non catalytic CLP was investigated (Iyer et al, 2006, Iyer et al, 2006, Ramkumar et al, 2008). The feasibility of enhancing the purity of H 2 and the optimum process conditions for H 2 production in the absence of a water gas shift catalyst were determined. 129

161 5.2 MATERIALS AND METHODS Chemicals, Sorbents, and Gases The HTS catalyst was procured from Süd-Chemie Inc., Louisville, KY and consists of iron (III) oxide supported on chromium oxide. The CaO sorbent was obtained from a PCC precursor which was synthesized from Ca(OH) 2 obtained from Fisher Scientific (Pittsburgh, PA). The high surface area PCC (BET analysis; SA 49.2 m 2 /g; PV 0.17 cm 3 /g) was synthesized using a dispersant modified wet precipitation technique. The anionic dispersant used in this process was N40V, supplied by Ciba Specialty Chemicals (Basel, Switzerland). PCC was synthesized by bubbling CO 2 through a slurry of hydrated lime. The neutralization of the positive surface charges on the CaCO 3 nuclei by negatively charged N40V molecules forms CaCO 3 particles characterized by a higher surface area/pore volume and a predominantly mesoporous structure. Details of this synthesis procedure have been reported elsewhere (Fan et al, 1998, Gupta and Fan, 2002). The feed gas for all the H 2 production tests was a mixture of 10%CO and 90% Nitrogen (N 2 ) Experimental Setup: Fixed Bed Reactor Figure 5.1 shows the integrated experimental setup, used for the bench scale studies of the non-catalytic CLP for H 2 production from syngas. The setup is similar to the one used earlier to study the catalytic CLP system, The bench scale reactor is coupled with a set of continuous gas analyzers which detect concentrations of CO, 130

162 CO 2, H 2 S, CH 4 and H 2 in the product stream. The reactor setup is capable of handling high pressures and temperatures of up to 21 atms and 900 o C respectively, which are representative of the conditions in a commercial syngas to H 2 system Water Gas Shift Reaction in the Presence and Absence of HTS Catalyst The extent of the water gas shift reaction was determined at different temperatures in an empty stainless steel reactor. The reactant gases were made to flow through the empty heated reactor and the product gases were analyzed by means of continuous analyzers. The extent of the water gas shift reaction was also determined in the presence of the HTS catalyst obtained from Süd-Chemie g of the catalyst was loaded into the reactor and the pressure, temperature and gas flow rates were adjusted for each run. The dry gas compositions at the outlet of the reactor were monitored continuously using the CO, CO 2, H 2 S, CH 4 and H 2 gas analyzers. The total flow rate of the gases through the reactor was maintained a constant at 725 sccm for all the experiments and the concentration of CO in the reaction mixture was maintained at 10.3 % Simultaneous Water Gas Shift and CO 2 Removal The combined water gas shift and carbonation reaction was conducted either using a catalyst-sorbent mixture or non-catalytically, using only the sorbent without the water gas shift catalyst. The combined experiments were conducted using a sorbent (CaO) to catalyst ratio of 10:1 by weight or only CaO sorbent. The effect of various 131

163 temperatures (600, 650, and 700 C), S:C ratios (3:1, 2:1, 1:1) and pressures (1-21 atm) was investigated. The CaCO 3 sorbent was calcined by heating the sorbent-catalyst mixture or only the sorbent to 700 o C in a stream of N 2 until the CO 2 analyzer confirmed the absence of CO 2 in the outlet stream. Multicyclic experiments were conduct in the fixed bed reactor with only CaO sorbent by alternating the carbonation and calcination steps and switching between the above mentioned temperatures and feed gas streams. The total flow rate of the gases through the reactor was maintained a constant at 725 sccm for all the experiments and the concentration of CO in the reaction mixture was maintained at 10.3 % Combined H 2 Production with H 2 S Removal: To study the effect of sulfur on the CLP, 5000 ppm of H 2 S is mixed with the CO, N 2 and steam before being sent to the reactor. The H 2 production tests are conducted in the presence CaO sorbent as described in the section above. 5.3 RESULTS AND DISCUSSION Baseline Water Gas Shift Reaction Testing Base line experiments in an empty stainless steel reactor and in the presence of a HTS catalyst were conducted to study the kinetics of the water gas shift reaction. A comparison of the extent of the water gas shift reaction in the presence and absence of a catalyst gives a perspective of the feasibility of eliminating the need for the water gas 132

164 shift catalyst in the carbonation reactor of the CLP. Figure 5.2 shows the CO conversion obtained when a 10%CO and 90% N 2 feed stream is reacted with steam at different temperatures in an empty stainless steel reactor and in a stainless steel reactor with HTS catalyst at atmospheric pressure. The CO conversion in the presence of HTS catalyst was higher than in the empty reactor at temperatures lower than 800 C. In both the presence and absence of the catalyst, the CO conversion increases with the increase in temperature due to higher kinetics of the water gas shift reaction. Beyond a particular optimum temperature, the CO conversion decreases with the increase in temperature due to the thermodynamic limitation of the water gas shift reaction. It can be seen that, as expected, the CO conversion increases with the increase in S:C ratio. The effects of reaction temperatures and S:C ratios on CO conversion at 21 atm, shown in Figure 5.3, follow the same trend as that at 1 atm. These baseline experiments show that CO conversion occurs even in the absence of a catalyst due to rapid kinetics of the water gas shift reaction in the temperature range of 500 to 750 o C which is the temperature range at which CO 2 removal occurs with CaO sorbent. Hence this CO conversion achieved in the empty reactor can be further improved by the addition of CaO sorbent to the reaction system and removing the thermodynamic constraint of the water gas shift reaction Water Gas Shift Reaction in the Presence of Only CaO Sorbent The results obtained above lead to the conclusion that the water gas shift reaction takes place to a considerable extent even in the absence of a catalyst at 133

165 relatively higher temperatures than the conventional water gas shift reaction. Hence it is possible to increase the yield of H 2 from the water gas shift reactor by shifting the equilibrium of the reaction in the forward direction by removing the CO 2 product formed. The CO 2 formed by the water gas shift reaction is removed using CaO sorbent. Figure 5.4 (a) shows the N 2 and steam free gas concentration at the outlet of the reactor due to the combined water gas shift and carbonation reaction at 600 o C and 21 atm. High purity H 2 is produced with very low levels of CO and CO 2 during the prebreakthrough region of the curve when the CaO sorbent is active. As the CaO sorbent gets consumed, the purity of H 2 reduces and the concentration of CO and CO 2 increase in the breakthrough region of the curves. In the post-breakthrough region of the curve, the CaO sorbent is completely consumed and the composition of the outlet gas is similar to the composition at the outlet of the non catalytic water gas shift reaction. Figure 5.4 (b) illustrates the CO conversion obtained with time for the gas compositions obtained in Figure 5.4 (a). Almost complete conversion of CO is obtained in the pre-breakthrough region of the curve where the combined water gas shift and carbonation reaction takes place Effect of Pressure and S:C Ratio Pressure has been found to have an important role in increasing the purity of H 2 by the combined water gas shift and carbonation reaction in the presence of CaO sorbent. Figure 5.5 shows the effect of the change in pressure on CO conversion at a temperature of 650 o C and S:C ratio of 3:1. The CO conversion is found to increase 134

166 with the increase in pressure. At 1 atm, a clear pre-breakthrough region is not obtained and a 90 to 95% CO conversion is obtained in the initial part of the breakthrough curve. As the pressure is increased to 4.5 atms, a pre-breakthrough CO conversion of greater than 98% is observed and at a pressure of 21 atms, almost 100% CO conversion is observed in the pre-breakthrough region. Since pressure has been found to be an important variable, the combined effect of pressure and S:C ratio was investigated to determine conditions where the S:C ratio can be decreased without causing a large decrease in CO conversion or H 2 purity. Combined water gas shift and carbonation experiments were conducted in the absence of a catalyst for various S:C ratios and pressures ranging from 1 to 21 atm. When the S:C ratio is decreased from 3:1 to 1:1 at ambient pressure, the CO conversion decreases in the breakthrough curve as shown in Figure 5.6 (a). At higher pressures of 11 and 21 atms, there is almost no decrease in the initial pre-breakthrough CO conversion with the decrease in S:C ratio. As illustrated in Figure 5.6 (b) at a pressure of 11 atms, the CO conversion remains at 98 to near 100% for both S:C ratios of 3:1 and 1:1. At 21 atms, a near 100% CO conversion is obtained in the pre-breakthrough curve for all S:C ratios of 3:1, 2:1 and 1:1 as shown in Figure 5.6(c). Hence by operating the carbonation reactor at high pressures it is possible to reduce the excess steam addition without causing a decrease in the CO conversion and the corresponding H 2 purity. 135

167 Effect of Temperature The effect of temperature was investigated at various S:C ratios for the combined water gas shift and carbonation reaction. Figures 5.7 (a) and (b) illustrate the change in CO conversion obtained when the temperature is varied from 600 to 700 o C at two S:C ratios of 3:1 and 1:1 at atmospheric pressure. At both S:C ratios, it can be seen that the highest CO conversion in the pre-breakthrough region is obtained at 600 o C and the CO conversion decreases with the increase in temperature due to the highly exothermic nature of the combined water gas shift and carbonation reaction. A reverse trend is obtained in the post-breakthrough region where the water gas shift reaction occurs in the absence of both a sorbent and a catalyst and its rate increases with the increase in temperature Effect of CO Concentration in the Feed Gas The effect of CO concentration in the reactant gas was investigated at a pressure of 11 atms on the CO conversion and purity of H 2 produced for the same amount of sorbent loaded. As shown in Figures 5.8 (a) and (b), near 100% CO conversion and high purity H 2 was produced for both 10% and 15% CO in the feed stream. With an increase in the CO concentration, the pre-breakthrough region of the curve becomes shorter. This is due to the higher flow rate of CO 2 produced from the 136

168 CO in the feed by the water gas shift reaction which results in the faster conversion of the CaO bed to CaCO Sorbent Characterization and Morphology Analysis Scanning Electron Microscopy (SEM) analysis was conducted on the sorbent samples, to visualize the changes in the physical structure of the sorbent. PCC sorbent was examined using SEM as shown in Figure 5.9 (a). It can be seen that the surface of PCC is rough and the structure is porous and not dense like the structure of the limestone sample observed by Abanades and Alvarez (Abanadez and Alvarez, 2003). It can be clearly observed that the structure of PCC has been modified to improve the porosity by introducing mesopores in the structure using surface modifying agents. The PCC was then calcined to form PCC- CaO which was also examined under the SEM. Figure 5.9 (b) is the image of a freshly calcined sample of PCC and it shows smaller sized clusters than the PCC precursor. The calcined sorbent is then used in the water gas shift reactor at atmospheric pressure to remove the CO 2 produced and to shift the equilibrium of the water gas shift reaction in the forward direction, thereby increasing the yield and purity of H 2. Figure 5.10 (a) shows the surface characteristics and pore structure of the PCC sorbent which has undergone carbonation during the water gas shift reaction at atmospheric pressure. During H 2 production, ~ 70% conversion of CaO to CaCO 3 was obtained. It can be seen that the structure and surface of the first carbonated sample is different from the fresh PCC sample shown in Figure 5.9 (a). Elongated structures can be observed on the surface of the first carbonated sample. 137

169 Figure 5.10 (b) shows the surface structure for calcium sorbent which has undergone carbonation during H 2 production at 21 atms. At 21 atms it is found that ~85% conversion of CaO to CaCO 3 is obtained. It can be seen that the surface structure formed during carbonation at 21 atms is similar to that formed during carbonation at 1 atm, but is denser due to the compaction at higher pressures H 2 Production in the Presence of CaO Sorbent Only and a Mixture of CaO Sorbent and Catalyst The effect of the presence of water gas shift catalyst in the carbonation reactor was investigated at atmospheric pressure and a high pressure of 21 atms. Figure 5.11 depicts the H 2 purity obtained in the presence and absence of the catalyst at atmospheric pressure. It can be seen that the H 2 purity obtained in the presence of the catalyst is 90% while it is 70% in the absence of the catalyst. In addition, a clear prebreakthrough region is observed in the presence of the catalyst for H 2 purity while there is almost no pre-breakthrough region in the absence of the catalyst. In contrast, at a high pressure of 21 atm, there is no difference in the purity of H 2 produced in the absence and presence of the catalyst. Almost 100% pure H 2 is produced in both cases. The same effect is observed at both S:C ratios of 3:1 and 2:1 as shown in Figure Hence although the catalyst can be eliminated without causing a decrease in H 2 purity at high pressures the same is not true at atmospheric pressure. However, in commercial facilities most of the H 2 production applications are typically deployed at high pressures. 138

170 5.3.4 Multicyclic Investigation of H 2 Production in the Presence of CaO Sorbent Only Multicyclic reaction and regeneration of the calcium sorbent was conducted to determine the effect of the number of cycles on the purity of H 2 produced in the fixed bed reactor. During the reaction step, H 2 was produced from a 10% CO/90% N 2 feed stream in the presence of CaO sorbent. The gas compositions for CO, CO 2, H 2 and hydrocarbons were recorded using continuous analyzers connected to a computer program. At the end of the reaction step, the sorbent was calcined at 750 o C in N 2. Figure 5.13 illustrates the purity of H 2 obtained when the reaction step is conducted at a pressure of 4.5 atms for 10 cycles. The purity of H 2 in the product stream is found to decrease with sorbent cycling from near 100% to 97% at the end of 10 cycles. In addition, it can be observed that for each additional cycle, the pre-breakthrough region is shorter than the previous one. This trend might be due to the reduction in useful porosity available for the carbonation of CaO due to sintering of the sorbent. Figure 5.14 illustrates the N 2 and steam free H 2 purity obtained from a 10% CO/90% N 2 feed stream in the presence of CaO sorbent at an operating pressure of 21 atm. The purity of H 2 in the pre-breakthrough region remains almost constant for 10 cycles. However, the time for which the pre-breakthrough region lasted decreased with the increase in the cycle number but to a lower extent than at 4.5 atms. The shortening of the pre-breakthrough region with each progressive cycle again might be attributed to sorbent sintering. 139

171 5.3.5 Enhanced H 2 Production With CO 2 and Sulfur Capture In the CLP, the CaO sorbent assumes the role of a multipollutant capture sorbent, in addition to enhancing the water gas shift reaction. Hence, the influence of various process variables like temperature, S:C ratio and pressure on the purity of H 2 produced and the extent of H 2 S removed during the combined water gas shift, carbonation and sulfidation reaction of CaO was determined Effect of S:C Ratio Figures 5.15 (a) and (b) illustrate the effect of varying S:C ratio on the extent of H 2 S removal and the purity of H 2 produced in the combined water gas shift, carbonation and sulfidation reaction. The extent of H 2 S removal by the CaO sorbent is found to increase with the decrease in S:C ratio in the carbonation reactor. As shown in Figure 5.15 (a), the concentration of H 2 S in the product H 2 stream decreases from 100 ppm to <1ppm with the decrease in S:C ratio from 3:1 to 0.75:1. This decrease in H 2 S concentration is due to a reduction in the inhibiting effect of steam on the reaction between H 2 S and CaO. The same inference was drawn from the thermodynamic analysis shown in Chapter 4. The effect of the change in S:C ratio on the purity of H 2 is illustrated in Figure 5.15 (b). Similar to observations made earlier in this chapter, at atmospheric pressure the purity of H 2 is found to decrease with the decrease in S:C ratio during the breakthrough region of the curves. 140

172 Effect of Temperature Figures 5.16 (a) and (b) illustrate the effect of temperature on the extent of H 2 S removal and the purity of H 2 produced respectively via breakthrough curves. A low concentration of H 2 S in the order of ~10 ppm is detected in the outlet H 2 stream at temperatures ranging from 560 to 600 o C at atmospheric pressure. With the increase in the temperature above 600 o C, the H 2 S concentration in the outlet is found to increase to 50 ppm at 650 o C and 90 ppm at 700 o C. The effect of temperature on the purity of H 2 has been illustrated in Figure 5.16 (b). The H 2 purity is found to be the highest (70%) within the temperature range of o C. The purity of H 2 is found to decrease to 60% with the decrease in temperature to 560 o C. A similar effect is observed with the increase in temperature to 700 o C. Hence, from Figures 5.16 (a) and (b), it can be inferred that the optimum temperature of operation for simultaneous H 2 production and H 2 S removal is ~600 o C Effect of Pressure Pressure has been found to be a very important variable for the non catalytic production of high purity H 2 at low S:C ratios. The effect of the increase in pressure on the extent of H 2 S removal and the purity of H 2 produced is illustrated in Figures 5.17 (a) and (b). The concentration of H 2 S in the product H 2 stream is found to decrease from 10 ppm to <1ppm when the pressure is increased from 1 atm to 21 atm. Hence, the combined effect of operating at a low S:C ratio and high pressure results in the 141

173 production of a H 2 stream with <1 ppm of sulfur impurities. Figure 5.17 (b) illustrates the effect of the increase in pressure on the purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction. At a temperature of 600 o C and a stoichiometric S:C ratio, the purity of H 2 is found to increase from 70% to >99% with the increase in pressure of 1 atm to 21 atm. Hence, the CLP is capable of producing high purity H 2 (>99%) with <1 ppm of sulfur impurities in it at stoichiometric S:C ratios Sorbent Characterization and Morphology Analysis The scanning electron microscopy (SEM) analysis is conducted on the sorbent samples, to visualize the changes in the physical structure of the sorbent. PCC sorbent is examined using SEM as shown in Figure 5.18 (a). It can be seen that the surface of PCC is rough and the structure is porous. The PCC is then calcined to form PCC- CaO which is also examined under the SEM. Figure 5.18 (b) is the image of a freshly calcined sample of PCC and it shows smaller sized clusters than the PCC precursor. The calcined sorbent is then used in the carbonation reactor at atmospheric pressure to remove the CO 2 and H 2 S and shift the equilibrium of the water gas shift reaction in the forward direction, thereby increasing the yield and purity of H 2. Figure 5.18 (c) shows the surface characteristics and pore structure of the PCC sorbent which has undergone carbonation and sulfidation during the water gas shift reaction at atmospheric pressure. It can be seen that the structure and surface of the first carbonated sample is different from the fresh PCC sample shown in Figure 5.18 (a). Elongated structures can be 142

174 observed on the surface of the first carbonated sample. Figure 5.18 (d) shows the surface structure for calcium sorbent which has undergone carbonation and sulfidation during H 2 production at 21 atms. It can be seen that the surface structure formed during carbonation at 21 atms is similar to that formed during carbonation at 1 atm, but is denser due to the compaction at higher pressures H 2 PRODUCTION FROM COAL GASIFICATION DERIVED SYNGAS Process Overview In the conventional coal gasification process, H 2 can be produced from coal through the sweet shift or the sour shift route (Stiegel and Ramezan, 2006). Figure 5.19(a) illustrates the conventional coal to H 2 process in which coal is fed along with steam and/or oxygen to the gasifier to produce syngas. In the sweet shift route, the syngas is cooled in a radiant cooler. The ash is then separated from the cool syngas which is fed to a syngas scrubber for ammonia and HCl removal. Following this, sulfur is removed using a solvent based system as the commercial HTS catalyst has a sulfur tolerance of about several hundred ppms while the LTS catalyst has a lower tolerance to sulfur and chloride impurities. Ash, ammonia, HCl and sulfur removal is conducted at low temperatures of 40 to 200 o C which is energy intensive due to the gas cooling and reheating requirements (Haussinger et al, 2000). The syngas temperature is then raised for the water gas shift reaction. Higher temperatures enhance the kinetics of the water gas shift reaction. However, as shown in Chapter 4, the equilibrium limitation of 143

175 the water gas shift reaction adversely affects H 2 production, with the H 2 yield falling with rising temperature. Hence, a high S:C ratio is required to enhance CO conversion and the consequent H 2 yield. The S:C ratio required at 550 o C can be as high as 50 in a single-stage operation or 7.5 for a more expensive dual-stage process to obtain 99.5 % pure H 2 (David, 1980). Numerous research studies have focused on the development of low temperature catalysts to improve H 2 production (David, 1980). Commercially, the dual stage sweet water gas shift reaction is carried out in series, with a HTS ( o C) stage containing iron oxide catalyst to convert bulk of the CO and a LTS ( o C) stage containing copper catalyst ( Haussinger et al, 2000). Following the shift reactors, the syngas is fed to a mercury removal unit and a CO 2 capture unit based on physical solvents like selexol or, rectisol or chemical solvents like amine based solvents. For high purity H 2 production, a PSA is used as the final step and the tail gas from the PSA is combusted to produce electricity. In a sour gas shift system, syngas is cooled using a water quench which provides the excess steam required for the water gas shift reaction and removes impurities like ash, HCl and ammonia (Holt, 2005, MIT, 2007). Since the sulfur content of synthesis gas is greater than 1000 ppm, a sulfided catalyst is used in a series of reactors at a temperature of C (Lloyd et al, 1996, Hiller et al, 2007). CO 2 and sulfur removal is achieved in a dual stage acid gas removal system and the H 2 is finally purified in a PSA. Figures 5.19(b) and 5.20 show the integration of the CLP in a typical coal or biomass gasification system with the cogeneration of electricity and H 2. The syngas 144

176 from the gasifier is cooled in a radiant heater and fed along with steam and CaO to the carbonation reactor in the CLP. The water gas shift reaction almost goes to completion in the presence of the CaO sorbent. The CaO sorbent reacts with the CO 2, sulfur and halide impurities and removes them from the product stream. The product gas stream from the reactor contains predominantly H 2 which is purified further in a PSA for ultra pure applications (eg. Fuel cells). The H 2 stream upstream of the PSA could also be converted to electricity in a combined cycle system for the generation of electricity or used for the production of fuels and chemicals. The spent sorbent from the carbonation reactor is then regenerated in the calciner where a sequestration ready CO 2 stream is produced. When calcination is conducted in the presence of steam, a CO 2 stream containing a small concentration of H 2 S is produced from the calciner, which can then be sequestered as is (Smith et al, 2007). Since CaS and CaCl 2 cannot be regenerated completely, a portion of the sorbent mixture is purged at the exit of the carbonation reactor. Fresh sorbent make up is added upstream of the calciner. The amount of purge and makeup will depend on the sulfur and chloride content of the coal syngas and the extent of sintering of the sorbent. The sorbent makeup and purge will result in the production of a H 2 stream with constant purity and will prevent the accumulation of inert material (CaCl 2 and CaS) in the circulating sorbent mixture. On comparison of Figures 5.19(a) and (b) it can be seen that by using the CLP, the unit operations in the coal to H 2 process can be significantly reduced. 145

177 5.4.2 System Thermodynamics Analysis Thermodynamic analysis was conducted for the reactions occurring in the carbonation reactor with syngas feed compositions from different gasifiers. The equilibrium constants for the reactions occurring in the carbonator were obtained using HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland) software and are shown in Chapter 4. Table 5.1 shows the syngas compositions from different gasifiers. It is noted that air is used as the oxidant in the moving bed, dry gasifier while oxygen is used in all the other gasifiers. Using the composition of the syngas, the feasibility of the simultaneous water gas shift, carbonation and sulfidation reaction in the temperature range of ºC was determined. The temperature range of ºC was chosen as the preferred operating range primarily because the kinetics of CO 2 capture by CaO is high and the equilibrium partial pressure of CO 2 is low as shown in Chapter 4. Steam is added to the syngas before it is fed to the carbonation reactor to adjust the S:C ratio for the water gas shift reaction. Tables 5.2 and 5.3 list the compositions of the syngassteam mixture for S:C ratios of 1:1 and 3:1 respectively. Depending on the thermodynamic extent of the water gas shift reaction occurring in the carbonator, the PCO 2 in the carbonator can be determined. Figure 5.21(a) illustrates the PCO 2 in the carbonator after the water gas shift reaction has occurred in the syngas-steam mixture with a S:C ratio of 1:1 from various gasifiers. It also shows the equilibrium PCO 2 required for the carbonation reaction with CaO. It can be seen that the PCO 2 in the carbonator for various gasifiers is higher than the equilibrium PCO 2 for the carbonation 146

178 of CaO and hence it can be inferred that in the temperature range of ºC, CO 2 removal is achieved by CaO to greater than 90%. Figure 5.21(b) illustrates the PCO 2 in the shifted syngas-steam mixture for a S:C ratio of 3:1. It can be seen that Figure 5.21(b) is very similar to Figure 5.21(a) and CO 2 removal is achieved by CaO to high extents for a S:C ratio of 3:1. In order to determine whether the CaO sorbent will undergo hydration, the PH 2 O in the carbonation reactor was determined after the combined water gas shift and carbonation reaction has occurred to equilibrium for S:C ratios of 1:1 and 3:1. Figure 5.22(a) shows the PH 2 O in the carbonator for a S:C ratio of 1:1 and since the PH 2 O in the carbonator is lower than equilibrium PH 2 O for the hydration of CaO, hydration will not occur for any syngas composition in the temperature range under consideration. Figure 5.22(b) shows the PH 2 O in the carbonator for a S:C ratio of 3:1 and it can be seen that hydration will occur at lower temperatures. Hydration of CaO will occur at temperatures below 600 ºC for fluidized bed and moving bed ( dry), below 670 ºC for entrained flow (dry), below 680 ºC for moving bed (slagging) and below 700 ºC for entrained flow ( slurry) gasifier syngas. Figures 5.23 (a) and (b) depict the equilibrium CO conversion obtained in a conventional water gas shift reactor. The CO conversion for all syngas compositions and S:C ratios decreases with the increase in temperature due to the equilibrium limitation of the exothermic water gas shift reaction. The CO conversion increases with the increase in S:C ratio from 1:1 to 3:1 for all compositions of syngas. Figures

179 (a) and (b) illustrate the CO conversion obtained in the carbonation reactor of the CLP for different syngas feed compositions. It can be seen that the CO conversion is enhanced in the presence of the CaO sorbent in the CLP in comparison to the conventional water gas shift reactor. It can be seen that greater than 95% CO conversion can be obtained for syngas from all the gasifiers by operating in the temperature range of 550 to 650 C. Although greater CO conversions can be obtained at temperatures lower than 550 C, the kinetic of the water gas shift reaction and CO 2 removal by CaO will decrease, resulting in the need for the use of larger reactors. Figures 5.25 (a) and (b) illustrate the purity of H 2 produced by the carbonation reactor in the CLP at S:C ratios of 1:1 and 3:1. High H 2 purities can be obtained at both S:C ratios of 3:1 and 1:1 in all the gasifiers where oxygen is used as the oxidant. In the moving bed, dry gasifier, lower H 2 purities are obtained due to dilution by nitrogen since air is used as the oxidant in the gasifier. Figure 5.26 depicts the total equilibrium carbon capture obtained in the CLP at S:C ratios of 1:1 and 3:1. The % carbon captured is defined as the total moles of carbon in the form of CO, CO 2 and CH 4 that is removed in the process. The % carbon captured for all syngas compositions and S:C ratios decreases with the increase in temperature. This decrease is due to the decrease in not only CO conversion but also CO 2 capture by CaO at high temperatures. It can be seen that greater than 95% carbon capture can be obtained from all the gasifiers by operating in the temperature range of 550 to 650 o C. Although greater CO conversions can be obtained at temperatures lower than 550 o C, 148

180 the kinetic of the water gas shift reaction and CO 2 removal by CaO will decrease, resulting in the need for the use of larger reactors. The % carbon captured increases with the increase in S:C ratio from 1:1 to 3:1 for all compositions of syngas. 5.5 H 2 PRODUCTION FROM SYNGAS DERIVED FROM NATURAL GAS FEEDSTOCKS Syngas from Steam Reforming of Natural Gas Steam reforming has been used commercially since 1931 for the conversion of hydrocarbons to H 2 as shown in Figure 5.27 (Gunardson, 1998). Catalytic steam reforming is conducted at high temperatures of 800 to 1000 ºC and pressures of 7.9 to 24.7 atms ( 8 25 bars)in the presence of a catalyst. Nickel based catalyst is most commonly used and is resistant to sulfur poisoning in hydrocarbon feeds containing less than 0.1 ppm of sulfur. Since the steam reforming process is endothermic, heat is supplied by burning natural gas in air which produces a flue gas. The hydrocarbon feed to the reformer is preheated by the product syngas and the reformer flue gas. As a first step in the reforming process the feed is hydrogenated at ºC to convert all the organic sulfur to H 2 S. For hydrocarbon feeds containing high concentrations of organic sulfur, the feed is scrubbed with amine solution to reduce the concentration of H 2 S to 25 ppm. The feed is then passed over a ZnO catalyst at ºC to further reduce the H 2 S concentration to 0.01 ppm and enters the reformer. At the exit of the reformer the hot gases are cooled to ºC and sent to the HTS reactor. For maximizing the H 2 yield from this process, the gases from the exit of the HTS reactor are further 149

181 shifted in a LTS stage. The gases are then cooled and fed to a CO 2 removal unit where the CO 2 is scrubbed using amine based solvents or selexol, rectisol, etc. Since CO and CO 2 are poisons for applications like NH 3 synthesis, refinery processes and petrochemical processes, CO and CO 2 need to be removed to less than 5 ppm in the H 2 stream. This is achieved in the methanation reactor at 310 ºC. (Gunardson, 1998) High purity H 2 is also produced by using a Pressure Swing Absorption unit which replaces the CO 2 removal units and the methanator. Figure 5.28 illustrates the process flow for the production of H 2 using a steam reformer and PSA assembly. The syngas from the reformer is shifted in a HTS stage since the performance and efficiency of the PSA is dependent on the purity of H 2 fed to it. Addition of a LTS stage will improve the recovery of H 2 and the efficiency of the PSA while increasing the cost of H 2 production. (Gunardson, 1998) Figure 5.29 illustrates the integration of the CLP in an existing steam methane reformer for the conversion of syngas to H 2 with simultaneous CO 2 capture. The syngas produced in the reformer consists of a mixture of CO, CO 2, H 2, CH 4 and steam. The dry gas composition of the syngas is shown in Table 5.4. (Gunardson, 1998) The syngas is fed to the carbonation reactor in the CLP along with CaO sorbent and steam. The CO in the syngas is converted to H 2 and the CO 2 product is simultaneous removed by the CaO sorbent. The equilibrium product gas composition at the exit of the carbonator at a temperature of 650 ºC and a pressure of 25 atms is also shown in Table 150

182 5.4. From thermodynamic analysis, about 99% of CO conversion to H 2 and 99% CO 2 removal by CaO sorbent is achieved in the carbonation reactor. In addition to converting syngas feed to H 2, the CLP can also be used to convert hydrocarbons like natural gas directly to H 2 in a single stage reactor. This configuration of the CLP is elaborated in Chapter Syngas from Autothermal Reforming of Natural Gas In autothermal reforming, a portion of the natural gas is combusted in the reformer with pure oxygen from an ASU. The hot gases ( ºC) are then passed through a catalyst bed where the steam methane reforming and water gas shift reactions take place to equilibrium. Table 5.5 shows the composition of syngas from the exit of the autothermal reformer. (Gunardson, 1998) If the CLP is added to the autothermal reformer for the production of H 2 in a carbon constrained scenario, then 99% of CO conversion and CO 2 removal can be obtained in the process. The equilibrium product gas from the carbonation reactor of the CLP is shown in Figure Syngas from Partial Oxidation of Natural Gas: The partial oxidation of natural gas is an uncatalyzed reaction with steam and oxygen to produce syngas. A schematic of the H 2 production process from syngas derived from the partial oxidation of natural gas is shown in Figure The syngas is 151

183 then cooled in a radiant heater or using a water quench system. The water quench in addition to cooling the gas also saturates the gas with water which is used for shifting the CO in the syngas. The syngas is then scrubbed with water to remove carbon and is fed to a high and low temperature shift system. The CO 2 in the shifted syngas is then removed using amine solvents, rectisol, selexol, or other solvents. (Gunardson, 1998) The HTS, LTS and CO 2 capture unit operations can be replaced by the CLP. Table 5.6 illustrates the composition of syngas obtained from the partial oxidation of natural gas. (Gunardson, 1998) When this syngas is fed to the carbonation reactor, 99% CO conversion and 97% CO 2 removal is achieved in the presence of steam of the CaO sorbent. Figure 5.31 illustrates the integration of the CLP with a partial oxidation system for the production of high purity H ADDRESSING THE ISSUE OF SULFUR IN THE FEEDSTOCK The effect of sulfur on the CLP is significant when the CLP is applied to precombustion systems. During the production of H 2 and electricity from syngas containing H 2 S and COS, as in coal gasification derived syngas, the CaO reacts with H 2 S and COS to form CaS. The CaS is stable in the carbonation reactor under reducing conditions. In a direct oxyfired calciner, the CaS is stable if the calciner is maintained under reducing conditions. This method was used in the CO 2 acceptor process where the air fired coal calciner was maintained under reducing conditions by the addition of 5% CO. If the 152

184 calciner atmosphere is oxidizing, then the CaS is converted to CaSO 4 in the presence of O 2, CO 2 and H 2 O, as shown below, depending on the residence time of the solids in the calciner. CaS oxidation: CaS + 2O 2 CaSO 4 (5.1) CaS + 4CO 2 CaSO 4 + 4CO (5.2) CaS + 4H 2 O CaSO H 2 (5.3) In a direct oxyfired coal calciner, CaSO 4 may also be produced from the direct reaction of CaO with SO 2 in the presence of O 2 as shown below: CaO + SO O 2 CaSO 4 (5.4) In an oxidizing atmosphere in the calciner, the CaS and CaSO 4 may form a eutectic mixture producing CaO and SO 2 as shown below: CaS + 3CaSO 4 4CaO + 4SO 2 (5.5) The SO 2 produced in the calciner exits with the CO 2 and the entrained sorbent mixture. The sorbent and gas mixture is cooled at the exit of the calciner before it is fed to a particle capture device. The SO 2 in the gas is captured by the sorbent when the sorbent and gas mixture is cooled before the particle capture device. Hence almost no SO 2 is present in the CO 2 stream that is sent for sequestration. 153

185 Thermodynamic analyses predict the complete conversion of CaS to CaO in the calciner and the mixture of solids entering the hydrator contains only CaO, CaCO 3, CaSO 4 and inerts (flyash from the calciner and inerts in the limestone sorbent). During actual operation, the extent of CaS conversion to CaO will depend on the residence time in the calciner. The residence time in the calciner can be as short as 2 secs in a commercial flash calciner. Hence if CaS is present in the solid mixture entering the hydrator then the CaS may be converted to CaO and H 2 S in the presence of the steam as shown below. CaS + H 2 O CaO + H 2 S (5.6) The solids at the exit of the hydrator will be a mixture of Ca(OH) 2, CaO, CaSO 4, CaCO 3 and CaS. When these solids are fed into the carbonator, the CaSO 4 will get reduced in the presence of CO and H 2 as shown below. CaSO 4 + 4CO CaS + 4CO 2 (5.7) CaSO 4 + 4H 2 CaS + 4H 2 O (5.8) The reduction of CaSO 4 results in a decrease of about 10% of H 2 yield from the carbonation reactor due to the consumption of CO and H 2. Although H 2 yield is reduced in the carbonation reactor, the subsequent oxidation of the CaS in the calciner is exothermic and the energy released in the calciner aids in the calcination of CaCO 3. This will help in reducing the amount of coal that needs to be added to the oxyfired 154

186 calciner. Hence the efficiency of the overall process will not be changed significantly due to the reactions involving sulfur in the CLP. Depending on whether the calciner is operated in an oxidizing or reducing mode, the composition of the sorbent mixture circulating through the CLP and the composition of the purged solids will vary. If CaS is present in the purged solids it will have to be converted to CaSO 4 or CaO before disposal. It has been shown in the CO 2 acceptor process and the HyPr-RING process that the complete oxidation of CaS does not occur in one stage in the calciner and hence CaS is always present in the solids mixture. The oxidation of CaS does not go to completion since the CaSO 4 formed has a higher molar volume than CaS. The oxidation reaction is slowed down by diffusional resistance and CaSO 4 forms an outer layer leaving CaS in the core. Squires et al and Keairns et al suggested using a mixture of H 2 O and CO 2 to convert CaS to CaCO 3 and H 2 S (Squires et al,1971, Keairns et al, 1976). They found that the rate of the regeneration reaction increased with the increase in temperature and complete regeneration could be achieved at 650 ºC. An investigation of CaS regeneration with H 2 O and CO 2 is described in the following sections Experimental Analysis of the Regeneration of CaS: The regeneration of a mixture of CaS and CaCO 3 was investigated in a fixed bed reactor setup. The samples for the regeneration experiments were obtained from 155

187 the combined water gas shift, carbonation and sulfidation experiments conducted for H 2 production at 600 ºC and at different pressures illustrated in Chapter 5. The regeneration was studied in the presence of steam alone and in the presence of steam and CO g of the spent sorbent containing a mixture of CaCO 3 and CaS was packed in the fixed bed reactor. High pressure steam was produced by pumping water into a heated tube and the steam produced was carried into the preheating section of the reactor by nitrogen in the case of regeneration in the presence of steam alone or by a mixture of nitrogen and CO 2 in the case of regeneration in the presence of steam and CO 2. The preheated gas mixture was then fed into the heated fixed bed reactor containing the spent sorbent. At the exit of the reactor, the mixture of gases was cooled and conditioned and fed into a continuous analyzer system for recording the concentration of H 2 S, CO 2 and CO. These concentrations were then recorded continuously for every second to yield the breakthrough curve Regeneration in the Presence of Steam reaction: The regeneration of CaS in the presence of steam is achieved by the following CaS + H 2 O CaO + H 2 S (5.6) A small amount of CaS is also oxidized by steam to give CaSO 4 as given below: CaS + 4H 2 O CaSO4 + 4H 2 (5.3) 156

188 Figure 5.32 illustrates the H 2 S evolved from CaS at different temperatures of 650 and 700 ºC and steam compositions of 31% and 15%. The spent sorbent samples for the three sets of data given below were obtained by conducting simultaneous water gas shift, carbonation and sulfidation reaction at 600 ºC and 21 atms in the presence of 5000ppm of H 2 S. As shown in Figure 5.32, it was found that when the regeneration is conducted in the presence of 31% steam and 69% nitrogen, a larger concentration of H 2 S is obtained at the outlet of the reactor at 700 ºC than at 650 ºC. In addition it was also found that when the regeneration reaction is conducted at 700 ºC the H 2 S concentration at the outlet of the reactor is higher when a 31 % steam concentration is used than when the 15% steam concentration is used. This suggests that the increase in steam concentration increases the conversion of CaS to CaO Regeneration in the Presence of Steam and CO 2 occur: During the regeneration of CaS with H 2 O and CO 2, the following reactions CaS + CO 2 + H 2 O CaCO 3 + H 2 S (5.9) CaS + H 2 O CaSO 4 + H 2 (5.3) CaS + H 2 O CaO +H 2 S (5.6) The regeneration of CaS was conducted in the presence of a mixture of steam and CO 2 and the effect of the change in concentration of steam and CO 2 was studied on the conversion of CaS to CaO. As shown in Figure 5.33, it was found that when a 15% 157

189 steam and 15% CO 2 mixture is used the extent of conversion of CaS is higher than when a mixture of 31% steam and 31% CO 2 is used. In contrast it can be seen that the rate of removal of H 2 S is higher when a 31% steam and CO 2 mixture is used than when the 15% steam and CO 2 mixture is used. This suggests that with the 31 % steam and CO 2 mixture the initial rate of decomposition of H 2 S is very high but with time this reaction is limited. This might be due to the sintering of the sorbent due to the presence of steam or CO 2. Figure 5.34 illustrates the comparison in the regenerability of the sorbent which has undergone carbonation and sulfidation at a high pressure of 21 atms and that which has undergone carbonation and sulfidation at atmospheric pressure. Since the sulfidation reaction is favored by the increase in pressure the sample obtained from the carbonation and sulfidation experiment at 21 atms releases more H 2 S than the sample obtained from the carbonation and sulfidation experiment conducted at atmospheric pressure. The rate of release of H 2 S from the sample is also higher for the sample treated at 21 atms CONCLUSION The feasibility and optimum process conditions for the production of H 2 in the absence of a water gas shift catalyst were determined. Experimental analysis revealed that CaO sorbent was found to enhance the thermodynamics of the water gas shift reaction and H 2 purity at a high reaction rate in the absence of the catalyst. Pressure 158

190 was found to have a large effect on H 2 purity. At high pressures, typical of commercial deployment, the absence of the catalyst and the reduction of excess steam addition did not have any effect on CO conversion and high H 2 purity (>99%) was obtained. A greater enhancement in H 2 purity was found to occur at lower temperatures of 600 and 650 o C and the effect of CaO sorbent was found to diminish with the increase in temperature. The effects of sintering of the CaO sorbent were observed on H 2 purity during multiple reaction and regeneration cycles without hydration. The effect of S:C ratio, temperature, and pressure was also studied on H 2 purity and the extent of H 2 S removal by CaO sorbent. Lowering the S:C ratio in the carbonator was found to improve the extent of H 2 S removal by the CaO sorbent. Greater than 99% H 2 purity with less than 1 ppm of H 2 S was obtained at a stoichiometric S:C ratio at high pressures. The integration of the CLP in coal gasification and natural gas reforming systems is also discussed and the advantages of the CLP process have been highlighted. 159

191 Moving Bed, dry Moving Bed slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Oxidant Air Oxygen Oxygen Oxygen Oxygen Fuel Sub Bituminous Bituminous Lignite Bituminous Bituminous Pressure (atm) CO (mole %) H2 (mole %) CO2 (mole %) H2O (mole %) N2 (mole %) CH4+ HCs (mole %) H2S + COS (mole %) Table 5.1: Typical fuel gas compositions obtained from different gasifiers (Stultz and Kitto, 1992). 160

192 Moving Bed, dry Moving Bed slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Oxidant air Oxygen Oxygen Oxygen Oxygen Fuel Sub Bituminous Bituminous Lignite Bituminous Bituminous Total Pressure (atm) CO (atm) H2 (atm) CO2 (atm) H2O (atm) N2 (atm) CH4+ HCs (atm) H2S + COS (atm) Table 5.2: Fuel gas composition entering the water gas shift reactor after steam addition (S:C ratio =1:1) (adapted from Stultz and Kitto, 1992) 161

193 Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Oxidant air Oxygen Oxygen Oxygen Oxygen Fuel Sub Bituminous Bituminous Lignite Bituminous Bituminous Pressure (atm) CO (atm) H2 (atm) CO2 (atm) H2O (atm) N2 (atm) CH4+ HCs (atm) H2S + COS (atm) Table 5.3: Fuel gas composition entering the water gas shift reactor after steam addition (S:C ratio =3:1) (adapted from Stultz and Kitto, 1992) 162

194 Syngas from SMR(%) CLP Product Gas(%) % removal CH CO CO H Table 5.4: Extent of equilibrium CO conversion and CO 2 capture in the CLP from Steam Methane Reforming (SMR) derived syngas 163

195 Syngas from ATR (%) CLP Product Gas (%) % removal CH CO CO H N Table 5.5: Extent of equilibrium CO conversion and CO 2 capture in the CLP from Auto Thermal Reforming (ATR) derived syngas 164

196 Syngas from POX (%) CLP Product Gas (%) % removal CH CO CO H N Table 5.6: Extent of equilibrium CO conversion and CO 2 capture in the CLP from partial oxidation (POX) derived syngas 165

197 Thermocouple And Pressure Guage Steam & Gas Mixture Steam Generator Water In Gas Gas Mixture Mixture Sorbent MFC MFC MFC MFC Hydrocarbon Analyzer Back Pressure Regulator Heated Steel Tube Reactor Water Syringe Pump H 2 S H 2 CO CO 2 Hydro carbons Analyzers (CO, CO 2, H 2, H 2 S) Heat Exchanger Water Trap Figure 5.1: Simplified flow sheet of the bench scale experimental setup 166

198 CO Conversion Temperature ( o C) Empty Reactor, S:C - 3:1 With HTS Catalyst, S:C - 3:1 Empty Reactor, S:C - 2:1 With HTS Catalyst, S:C - 2:1 Empty Reactor, S:C - 1:1 With HTS Catalyst, S:C - 1:1 Figure 5.2: Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction at 1 atm 167

199 CO Conversion Empty reactor, S:C - 3:1 With HTS catalyst, S:C - 3:1 Empty reactor, S:C - 2:1 0.2 With HTS catalyst, S:C - 2:1 Empty reactor, S:C - 1:1 With HTS catalyst, S:C - 1: Temperature (C) Figure 5.3: Effect of reaction temperature and S:C ratio on the conversion of CO by the water gas shift reaction at 21 atm 168

200 Outlet Gas Compositions (%) H2 CO2 CO Time (sec) (a) CO Conversion Time (sec) Figure 5.4: Typical breakthrough curves for the production of H 2 in the presence of CaO sorbent without catalyst (a) Gas composition (mole%) and (b) CO conversion (600 C, 21 atm, S:C ratio of 3:1) (b) 169

201 CO Conversion Time(sec) 1 atm 4.5 atm 21 atm Figure 5.5: Effect of pressure on CO conversion obtained in the presence of CaO sorbent without catalyst (650 C, S:C ratio of 3:1) 170

202 CO Conversion Time (sec) 3:1 2:1 1:1 (a) CO Conversion Time (sec) (b) 3:1 1:1 Continued Figure 5.6: Effect of S:C ratio on CO conversion obtained in the presence of CaO sorbent without catalyst at (a) 1 atm, (b) 11 atm, (c) 21 atm (650 C) 171

203 Figure 5.6 continued CO Conversion Time (sec) 3:1 2:1 1:1 (c) 172

204 CO Conversion C 650C 700C Time(sec) (a) CO Conversion C 650C 700C Time(sec) (b) Figure 5.7: Effect of temperature on CO conversion obtained in the presence of CaO sorbent without catalyst at a S:C ratio of (a) 1:1 and (b) 3:1 (1 atm) 173

205 CO Conversion TIme (sec) 10% CO 15 % CO (a) % CO 15% CO H2 Purity (%) TIme(sec) (b) Figure 5.8: Effect of CO concentration in the feed on the (a) CO conversion and (b) purity of H 2 produced in the presence on CaO sorbent without catalyst (11 atm, 600 C, S:C ratio of 3:1) 174

206 (a) (b) Figure 5.9: SEM image of the (a) initial CaCO 3 sorbent (b) CaO sorbent obtained from the calcination of CaCO 3 175

207 (a) (b) Figure 5.10: SEM image of sorbent at the end of the water gas shift and carbonation reaction in the absence of a catalyst at (a) 1 atm (b) 21 atm (S:C ratio of 3:1, 600 C) 176

208 CaO and HTS catalyst CaO H2 Purity (%) Time (sec) Figure 5.11: Comparison in the product H 2 purity in the presence of the sorbent and in the presence of the sorbent and catalyst mixture at 1 atm (650 C, S:C ratio of 1:1) 177

209 100 H2 Gas H 2 Purity Composition (%) Without catalyst Without catalyst With catalyst - 3:1 With catalyst - 2: Time(sec) Figure 5.12: Comparison in the product H 2 purity in the presence of the sorbent and in the presence of the sorbent and catalyst mixture (650 C, 21 atm) 178

210 H2 Purity (%) Cycle 1 Cycle 3 Cycle 5 Cycle 7 Cycle Time (sec) Figure 5.13: Product H 2 purity obtained over multiple reaction and regeneration cycles in the presence of CaO sorbent without catalyst at 4.5 atms. (600 C, S:C ratio of 3:1) 179

211 H2 Purity (%) Cycle 1 Cycle 3 Cycle 5 Cycle 7 Cycle Time (sec) Figure 5.14: Product H 2 purity obtained over multiple reaction - regeneration cycles in the presence of CaO sorbent without catalyst at 21 atms (600 C, S:C ratio of 3:1) 180

212 H2S concentration (ppm) :1 1:1 3: Time(sec) (a) :1 1:1 3:1 H2 Purity (%) Time (sec) (b) Figure 5.15: Effect of S:C ratio on the (a) extent of H 2 S removal and (b) the purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent (1atm, 600 o C) 181

213 H2S concentration (ppm) o C 600 o C 650 o C 700 o C Time(sec) (a) H2 Purity (%) o C 600 o C 650 o C 700 o C Time(sec) (b) Figure 5.16: Effect of temperature on the (a)extent of H 2 S removal and (b) purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent (1 atm, S:C ratio of 1:1) 182

214 H 2 S concentration (ppm) < 1 ppm 300 psig 0 psig 1 atm 21 atm Time (sec) (a) 100 H2 Gas H Composition (%) 2 Purity (%) High purity H 2 21 atm 1 atm 0 psig 300 psig Time(sec) (b) Figure 5.17: Effect of pressure on the (a) extent of H 2 S removal (b) purity of H 2 produced during the combined water gas shift, carbonation and sulfidation reaction in the presence of CaO sorbent (S:C ratio of 1:1, 600 o C) 183

215 PCC (a) (b) PCC- Calcined PCC Carbonated and sulfidedat 0 psig (c) PCC Carbonated and sulfidedat 300 psig (d) Figure 5.18: SEM image of the (a) initial CaCO 3 sorbent (b) CaO sorbent obtained from the calcination of CaCO 3 (c) sorbent at the end of the water gas shift, carbonation and sulfidation reaction at 1 atm (c) CaO sorbent obtained from the calcination of CaCO 3 (600 o C, S:C ratio of 1:1) (c) sorbent at the end of the water gas shift, carbonation and sulfidation reaction at 21 atm (600 o C, S:C ratio of 1:1) 184

216 CO 2 Sequestration CO 2 Compression Water Sour Shift Sulfuric Acid Plant H 2 S Removal 185 Air ASU Coal Feed Gasifier Radiant Cooler Raw Quench Scrubber Syngas Shift Reactors Ash/HCl/ Ammonia/ HTS/LTS Sulfur Reactors Removal Mercury Removal Sweet Shift Steam Mercury Removal CO 2 Removal Air Dual Stage Acid Gas Removal PSA Boiler H 2 Flue Gas Coal Water Slag Steam Turbine (a) Continued Figure 5.19: (a) Conventional process for H 2 production from coal (b) Integration of the CLP in a conventional process for H 2 production from coal 185

217 Figure 5.19 continued CO 2 Sequestration CO 2 Compression Solids Waste Calcium Radiant Cooler Calciner Looping Carbonator Mercury Removal Air Process ASU Water Limestone PSA H 2 Gasifier Coal Feed Coal Water Slag Steam Turbine (b) 186

218 Steam INTEGRATED WGS +H 2 S +COS + HCL CAPTURE Hydrogen Air To Steam Turbine Coal/Biomass Gasifier CaO BFW Air Oxygen CaCO 3 CO 2 Steam H 2 +O 2 Rotary Calciner Fuel Cell Air Compressor Gas Turbine Generator HRSG Stack Slag Air Separation Fuels & Chemicals Steam Turbine Figure 5.20: Integration of the CLP in a coal gasification system for the production of electricity, H 2 and liquid fuels 187

219 100 Partial Pressure of CO 2, atm Temperature(C) Equilibrium PCO2 for Carbonation of CaO Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry (a) Continued Figure 5.21: Comparison of the PCO 2 in the carbonator with the equilibrium PCO 2 for the carbonation of CaO for a S:C ratio of (a)1:1 (b)3:1 188

220 Figure 5.21 continued 100 Partial Pressure of CO 2, atm Temperature(C) Equilibrium PCO2 for Carbonation of CaO Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry (b) 189

221 100 Partial Pressure of H 2 O, atm Temperature(C) Equilibrium PH 2 O for Hydration of CaO Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry (a) Continued Figure 5.22: Comparison of the PH 2 O in the carbonator with the equilibrium PH 2 O for the hydration of CaO for a S:C ratio of (a)1:1 (b)3:1 190

222 Figure 5.22 continued 100 Partial Pressure of H 2 O, atm Temperature(C) Equilibrium PH 2 O for Hydration of CaO Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry (b) 191

223 CO Conversion Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Temperature (C) (a) CO Conversion Moving Bed, dry Moving Bed slagging 0.2 Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Temperature (C) (b) Figure 5.23: Effect of temperature on equilibrium CO conversion in the water gas shift reactor at a S:C ratio of (a) 1:1 (b) 3:1 192

224 CO Conversion Moving Bed, dry Moving Bed, slagging 0.80 Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Temperature (C) (a) CO Conversion Moving Bed, dry Moving Bed slagging 0.96 Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Temperature (C) (b) Figure 5.24: Effect of temperature on equilibrium CO conversion in the presence of CaO in the carbonation reactor of the CLP at a S:C ratio of (a) 1:1 (b) 3:1 193

225 H2 purity (%) Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Temperature (C) (a) H2 Purity (%) Moving Bed, dry Moving Bed, slagging Fluidized Bed Entrained Flow, slurry Entrained Flow, dry Temperature(C) (b) Figure 5.25: Effect of temperature on equilibrium H 2 purity in the presence of CaO at a S:C ratio of (a) 1:1 (b) 3:1 194

226 100 % Carbon Captured (mole %) :1, Moving Bed, dry 1:1, Moving Bed, slagging 1:1, Fluidized Bed 1:1, Entrained Flow, slurry 1:1, Entrained Flow, dry 3:1, Moving Bed, dry 3:1, Moving Bed slagging 3:1, Fluidized Bed 3:1, Entrained Flow, slurry 3:1, Entrained Flow, dry Temperature ( C) Figure 5.26: Effect of temperature and S:C ratio on the % of carbon captured in the CLP using syngas from different gasifiers as the feed 195

227 Steam Reformer Sulfur Removal High Temperature Shift Reactor Low Temperature Shift Reactor CO 2 Absorber Methanator H Natural Gas Reformer Flue Gas CO 2 Figure 5.27: Conventional steam reforming of natural gas for H 2 production with a methanator 196

228 Steam Reformer Sulfur Removal High Temperature Shift Reactor PSA H 2 Natural Gas Reformer Flue Gas Purge Gas used as reformer fuel Figure 5.28: Conventional steam reforming of natural gas for H 2 production with a PSA 197

229 Steam Reformer Sulfur Removal Calcium Looping Process Polishing PSA H 2 Natural Gas Reformer Flue Gas Figure 5.29: CLP integrated in the conventional steam reforming of natural gas process 198

230 Oxygen Steam Reformer Sulfur Removal Water Quench Scrubber High Temperature Shift Reactor Low Temperature Shift Reactor CO 2 Absorber H Natural Gas CO 2 Figure 5.30: Conventional partial oxidation process for conversion of natural gas to H 2 199

231 Oxygen Steam Reformer Sulfur Removal Water Quench Scrubber Calcium Looping Process H 2 Natural Gas CO 2 Figure 5.31: CLP integrated in the partial oxidation of natural gas for H 2 production 200

232 CaS_700C_15%H2O CaS_650C_31%H2O CaS_700C_31%H2O H2S (ppm) Time(sec) Figure 5.32: Effect of the change in temperature and steam composition on the regeneration of CaS with H 2 O 201

233 CaS_700C_15%H2O_15%CO2 CaS_700C_31%H2O_31%CO2 35 H 2 S (ppm) Time (sec) Figure 5.33: Effect of the change in steam and CO 2 composition on the regeneration of CaS in the presence of H 2 O and CO 2 202

234 H 2 S (ppm) CaS_700C_31%H2O_0psig CaS_700C_31%H2O_300psig Spent sorbent from carbonation and sulfidation at 21 atms Spent sorbent from carbonation and sulfidation at 1 atms Time (sec) Figure 5.34: H 2 S evolved in the presence of H 2 O and CO 2 from spent sorbent produced during combined CO 2 and H 2 S removal at 1 and 21 atms 203

235 CHAPTER 6 PROCESS SIMULATION AND ECONOMICS OF THE CALCIUM LOOPING PROCESS (CLP) FOR PRODUCTION OF H 2 FROM COAL 6.1 INTRODUCTION The CLP for high purity H 2 production from coal, described in Chapters 4 and 5, has been analyzed in this Chapter based on the Aspen Plus process simulation through two schemes. The first scheme is based on the cogeneration of H 2 and electricity in the same facility. In the second scheme, the only product is H 2 and all the energy produced in the process is used internally for the parasitic energy requirement. Both simulations are conducted for the production of 280 t/day of H 2 from Illinois #6 coal. 6.2 PRODUCTION OF FUEL CELL GRADE H 2 WITH A PSA Cogeneration of H 2 and Electricity The CLP produces high purity H 2 and electricity with high efficiency by integrating various unit operations like the water gas shift reaction, and CO 2, sulfur and hydrogen halide removal from synthesis gas into a single stage reactor at high 204

236 temperatures. The CLP-assisted coal to H 2 process comprises of six major unit operations: Shell gasifier ASU Calcium looping reactor Calciner Pressure Swing Absorber (PSA) Heat recovery and steam generation block Process Configuration The process flow diagram for the CLP for co-production of fuel cell grade H 2 and electricity is given in Figure 6.1. A Shell gasifier is used to gasify 2,405 tonnes/day of Illinois # 6 coal in the presence of oxygen supplied by the ASU. The properties of the coal are given in Table 6.1. The Shell gasifier produces 951,624 m 3 /day of syngas at a temperature of 1538 ºC and a pressure of 36 atm. The composition of syngas produced at the outlet of the gasifier is given in Table % of the syngas produced at the outlet of the gasifier is fed to the H 2 production reactor or carbonator for the production of high purity H 2 while 21% of the syngas is combusted in the calciner to provide the energy required for the endothermic calcination reaction. 205

237 79% of the hot syngas from the gasifier is cooled in a radiant heater and is fed to the H 2 production reactor or carbonator along with CaO sorbent and steam. In the carbonation reactor, H 2 production and purification are achieved by the integrated water gas shift reaction, carbonation and sulfidation of the CaO sorbent at a temperature of 600 C and pressure of 21 atm. A Ca:C mole ratio of 1.32 and S:C ratio of 3 are used to achieve high CO conversions and almost 100% carbon and sulfur capture. The heat produced in the carbonation reactor through the exothermic water gas shift and carbonation reactions is used to produce high temperature and high pressure steam which is used to generate electricity. The H 2 rich product stream is further purified in a PSA to produce % H 2 which can be used either in H 2 fuel cells or for the production of fuels and chemicals. Since the purity of the dry H 2 feed stream to the PSA is very high (~94%) the energy consumption in the PSA is considerably reduced. The spent sorbent, which is separated from the H 2 product in a particulate capture device (PCD), is regenerated in the calciner at 850 C to produce a sequestration ready CO 2 stream. All the energy required for the calcination of the sorbent is supplied by the combustion of the syngas and PSA tail gas with oxygen in the calciner. At this stage, 6.2 wt% of the spent sorbent is purged and an equivalent moles of calcium in the form of CaCO 3 is added as make up to maintain the high reactivity of the sorbent mixture towards CO 2 and sulfur capture. In this process, the pure H 2 stream is produced at 21 atm which is compressed to 60 atm for transportation and the CO 2 is compressed to 150 atm for sequestration. The intermediate pressures used for the multistage CO 2 compression are shown in Table

238 In this scenario, 280 tonnes/day of H 2 is produced with an efficiency of 56.5% (HHV) and 81MW electricity is produced with an efficiency of 10% Components and Physical Properties Table 6.4 lists the components used in the simulation. Syngas obtained from a Shell gasifier using Illinois #6 coal is selected as the specific feedstock. RKS-BM is selected as the property method for conventional components Operating Conditions and Premises The basis for the conceptual process evaluation of the CLP is given below: Analysis has been conducted for a production rate of 280 tonnes/day of H 2 and 81 MWe of electricity Illinois #6 coal has been used at a feed rate of 2,405 tonnes/day. A Shell gasifier is used to generate the syngas at a temperature of 1538C and pressure of 36 atm A Ca:C ratio of 1.32 is used to achieve almost 100% CO 2 and sulfur capture. A S:C ratio of 3:1 is used to achieve high CO conversion in the CLP. A solids purge of 6.2 wt% is used for the calcium sorbent and an equivalent moles of calcium in the form of CaCO 3 is added as make up. 100% calcination occurs in the calciner and the heat is supplied by the complete combustion of syngas and PSA tail gas with oxygen. 207

239 Mercury removal is conducted using an activated carbon bed. The temperature is maintained constant during the individual unit operations. H 2 purity of greater than % is obtained by using a PSA. H 2 is produced from the calcium looping reactor at 21 atm and is compressed to 60 atm for transportation while the sequestration ready CO 2 is compressed to 150 atm. The mechanical efficiency of pressure changers such as compressors, turbines, and expanders is 1 while their isentropic efficiency is A part of the solid waste which includes fly ash, bottom ash, gasifier slag, etc is reused in the plant while the rest of it is disposed in the coal mine. Waste water is treated before discharge to meet effluent guidelines Table 6.5 shows the Aspen Plus models used for the various important unit operations in the CLP Results and Discussions Based on the description of the process and the assumptions given above, the ASPEN Plus simulation for the production of H 2 and electricity using the CLP is shown in Figure 6.2. Table 6.6 shows the material and energy balance for the entire process. Table 6.7 illustrates the power balance in the CLP process. A total of MWe of power is produced in the CLP process out of which 61MWe is consumed in the process for coal 208

240 preparation, pumps and compressors for the ASU, CO 2 and H 2..Net electricity of 81 MWe is produced which can be exported. Table 6.8 illustrates the results of the ASPEN simulation conducted for the conversion of coal to H 2 and electricity. The CLP process produces 280 tonnes/day of H 2 and 81 MWe of net electricity with 100% CO 2 and sulfur capture. The overall process efficiency from coal to H 2 and electricity is 66.5%, which is much higher than the conventional coal to H 2 process. The CLP has a higher efficiency than the conventional H 2 production process from coal using solvents due to the integration of various unit operations in a single stage reactor and the high temperature gas clean up achieved in the system Production of Only H 2 With Internal Heat Integration For only H 2 production, a Shell gasifier is used to gasify 2,180 t/day of Illinois #6 coal in the presence of oxygen supplied by the ASU. The properties of the coal are given in Table 6.1. The Shell gasifier produces 847,200 m 3 /day of syngas at the temperature of 1538 C and the pressure of 36 atm. Due to the high content of sulfur in the coal, the syngas contains 1.15% of H 2 S and 848 ppm COS. Since the CLP is capable of in situ sulfur capture during the production of H 2, it can handle high sulfur coal effectively. The composition of syngas produced at the outlet of the gasifier used in the simulation is given in Table % of the syngas produced at the outlet of the gasifier is fed to 209

241 the calcium looping reactor for the production of high purity H 2 while 11.3% of the syngas is combusted in the calciner to provide the energy required for the endothermic calcination reaction. The hot syngas is cooled in a radiant heater and is fed to the calcium looping reactor along with high temperature and high pressure steam and PCC-CaO sorbent. In the carbonation reactor, H 2 production, purification and sulfur removal are achieved by the integrated water gas shift reaction, carbonation, and sulfidation of the CaO sorbent at a temperature of 600 C and pressure of 21 atm. The H 2 rich product stream is then further purified in a PSA to produce up to % pure H 2 which can be used either in H 2 fuel cells or for the production of fuels and chemicals. Since the purity of the H 2 feed stream to the PSA is very high (94 98%) the energy consumption in the PSA is considerably reduced. The spent sorbent, which is separated from the H 2 product in a cyclone, is regenerated in the calciner at 850 C to produce a sequestration ready CO 2 stream. At this stage, 8% of the spent sorbent is purged and a makeup of PCC sorbent is added to maintain the high reactivity of the sorbent mixture toward CO 2 and sulfur capture. In this process, the pure H 2 stream is produced at a high pressure of 20 atm and the CO 2 is compressed to a pressure of 150 atm for transportation to the sequestration site. A Ca:C ratio of 1.3 is used to achieve almost 100% carbon and sulfur capture and sequestration from coal. This process leads to the production of 280 t/day of H 2 from coal with an efficiency of 62.3% (HHV) as shown in Table

242 6.3 PRODUCTION OF H 2 HAVING A PURITY OF 94 98% WITHOUT A PSA Cogeneration of H 2 and Electricity The Aspen Plus flow sheet for the CLP for the production of 94 98% pure H 2 without a PSA is given in Figure 6.3. The syngas obtained from the Shell gasifier is split into two streams, one fed to the calcium looping reactor and the other to the calciner. In this scenario, for the cogeneration of H 2 and electricity in the absence of a PSA, 2,350 t/day of coal is used for the production of 280 t/day of H 2 at an efficiency of 57.8% (HHV). In addition to the H 2, MW of electricity is produced with an efficiency of 8.5% from coal Production of H 2 With Internal Heat Integration In the case of H 2 production without a PSA, 19% of the syngas is required to supply the energy for the endothermic calcination reaction and the remaining 81% is fed to the calcium looping reactor for the production of high purity H 2. This process in the absence of a PSA also leads to the production of 280 t/day of H 2 with an efficiency of 63% from coal (HHV). The two scenarios for the production of H 2 or cogeneration of H 2 and electricity from coal by the CLP in the absence of a PSA are summarized in Table Comparing the processes with and without the PSA unit, it can be seen that the H 2 generation efficiencies for these processes with internal heat integration are almost the 211

243 same. For the cogeneration of H 2 and electricity, the overall efficiencies of the processes with and without the PSA unit are also similar. Although, with the PSA unit, the H 2 generation efficiency (56.5%) is lower than without the PSA unit (57.8%), more electricity is produced with the PSA unit (81 MW) than without the PSA unit (67.5 MW). 6.4 COMPARISON OF THE PROCESS EFFICIENCIES FOR DIFFERENT GASIFIERS The CLP is optimized for high purity H 2 production using syngas obtained from three different gasifiers, the Shell, Lurgi, and GE gasifiers (Zheng and Furinsky, 2005). A comparison of the efficiencies obtained for the different gasifiers is given in Table The type of gasifier used has a large effect on the process efficiency due to the composition of the syngas obtained and the inherent efficiency of the gasifier for the conversion of coal to syngas. It is seen in Table 6.11 that the CLP in combination with the Shell gasifier has the highest efficiency due to the high efficiency of the dry feed Shell gasifier. The co-generation of H 2 and electricity yields a higher efficiency than the case optimized for the production of H 2 alone. 6.5 EFFECT OF PROCESS PARAMETERS ON CLP PERFORMANCE USING SYNGAS FROM A GE GASIFIER As shown in the previous section the CLP integrated with the Shell gasifier has the highest efficiency closely followed by the GE gasifier. However from an economic standpoint, the cost of a GE gasifier is lower than a Shell gasifier and hence the 212

244 sensitivity analysis conducted in this section is based on the simulation for the CLP with the GE gasifier, Approach The thermodynamics of the combined reactions occurring in the H 2 production reactor or carbonation reactor has been investigated using ASPEN Plus software. The effect of temperature, pressure and S:C ratio has been investigated on the combined water gas shift, carbonation and sulfidation reaction. The analysis has been conducted using syngas obtained from a GE or Texaco gasifier. In the ASPEN model shown in Figure 6.4, the syngas from the GE gasifier is fed to the H 2 production reactor along with steam and CaO (from the calciner). At the outlet of the H 2 production reactor, the H 2 product is separated from the solids and the mixture of solids containing CaO, CaCO 3, Ca(OH) 2 and CaS is regenerated in the calciner. 8% of the solids obtained at the exit of the H 2 production reactor is purged and a makeup stream containing an equivalent quantity of CaCO 3 is added to maintain the fraction of CaS in the circulating solids stream at equilibrium Sensitivity Analysis for the Yield and Purity of H 2 Produced Sensitivity analyses have been conducted for the H 2 production reactor, to investigate the effect of temperature, pressure and amount of steam addition on the purity and yield of H 2 produced. 213

245 Effect of Temperature As it can be seen in Figure 6.5, the purity of H 2 is very high in the temperature range of 550 to 650 ºC. In this temperature interval, the thermodynamic limitation of the water gas shift reaction is removed due to the incessant removal of the CO 2 product and a very high yield of H 2 is produced. At temperatures of 650 ºC and above, the purity of H 2 decreases as the equilibrium conversion of the carbonation reaction decreases. At temperatures of 880 ºC and above, the carbonation reaction does not occur because at these temperatures equilibrium favors the reverse or calcination reaction. Hence at temperatures above 880 ºC, H 2 is produced only due to the water gas shift reaction Effect of Pressure The effect of pressure in the range of 1 to 40 atms was investigated and it was found that the change in pressure results in a comparatively smaller change in H 2 purity when compared to the change in temperature. Figure 6.6 illustrates the effect of pressure on H 2 purity. Pressure influences the combined reaction for H 2 production in two ways. Although high pressure has a positive effect on the thermodynamics of the carbonation reaction which increases the equilibrium conversion of the water gas shift reaction, it also favors the methanation reaction which results in a decrease in the overall H 2 yield. The formation of one mole of CH 4 results in the loss of 3 moles of H 2 and hence the yield of H 2 decreases with the increase in pressure. 214

246 From Figure 6.6 it can be seen that the purity of H 2 increases with the increase in pressure up to 10 atms. This is due to the decrease in CO and CO 2 in the final product due to the improved thermodynamics of the combined carbonation and water gas shift reaction at high pressures. Above a pressure of 10 atms, the thermodynamics of the methanation reaction is very favorable and hence there is a loss in H 2 and an increase in the CH 4 in the final product. It can be seen that according to thermodynamic evaluation, pressure has a small effect on the purity of H 2 produced and with the increase in pressure from 1 to 40 atms the H 2 purity changes only from 96.1 to 97.3% Effect of S:C Ratio The effect of S:C ratio was first investigated at 600 ºC. It can be seen from Figure 6.7 that there is an increase in the purity of H 2 produced with the increase in S:C ratio. This is because excess steam favors the equilibrium of the water gas shift reaction in the forward direction. It was also found that with an increase in steam addition, the CH 4 composition in the product stream decreases. From Figure 6.7 it can be seen that there is a substantial increase in the H 2 purity when the S:C ratio is increased from the 1 to 2. Beyond this the increase in H 2 yield is very small Sensitivity Analysis for the Extent of Contaminant Removal from the Product H 2 As shown in Figure 6.8, it can be seen that the H 2 S in the outlet stream increases with the increase in the S:C ratio due to the inhibiting effect of steam on the 215

247 sulfidation reaction of CaO. It was also found that with the increase in temperature, the H 2 S in the outlet stream increases and maximum removal is achieved in the temperature range of around 600 ºC. The effect of steam concentration and temperature on the removal of COS by CaO was also studied and it was found that the COS in the outlet increases with the increase in temperature. This is due to the increase in the CO 2 concentration which inhibits the removal of COS by CaO. As shown in Figure 6.9, almost all the COS in the syngas stream is removed by the CaO sorbent at temperatures lower than 800 ºC. The concentration of COS in the outlet stream was also found to increase with the increase in steam addition at temperatures above 800 ºC. This is also due to the increase in the CO 2 flow rate with the increase in steam addition. The concentration of CO was found to increase with the increase in temperature at high temperatures of above 700 ºC due to the equilibrium limitation of the water gas shift reaction and due to the decrease in the CO 2 removal by the CaO sorbent at high temperatures as illustrated in Figure At temperatures below 700 ºC it was found that the S:C ratio does not have an effect on the CO concentration at the outlet of the reactor and very low concentration of CO is obtained even at low S:C ratios. This is due to the removal of CO 2 by the CaO which enhances the equilibrium of the water gas shift reaction. With the increase in temperature of above 700 ºC it was found that the CO concentration increases with the decrease in steam addition. This is due to the low CO 2 removals at temperatures of above 700 ºC. 216

248 Figure 6.11 illustrates the change in the outlet flow rate of CO 2 with the increase in temperature and S:C ratio. Almost all the CO 2 in the outlet H 2 product stream is removed by the CaO sorbent at a temperature of 600 ºC and the CO 2 concentration increases with the increase temperature. Figure 6.12 depicts the change in the outlet flow rate of CH 4 with the change in temperature and S:C ratio. The flow rate of CH 4 in the outlet H 2 product is found to decrease with the increase in steam addition. It was also found that equilibrium favors the formation of CH 4 at low temperatures of 600 ºC Sensitivity Analysis for the Cold Gas Efficiency and Overall Process Efficiency Effect of Pressure Figure 6.13 depicts the effect of pressure and S:C ratio on H 2 purity, cold gas efficiency and the process efficiency. The cold gas efficiency is the efficiency with which the energy in coal is converted to H 2 and is defined as the ratio of the HHV of the product H 2 stream to the HHV of coal. The process efficiency is the total efficiency of H 2 and electricity production which is obtained from a detailed heat integration within the process and includes the parasitic energy required for the sorbent regeneration, the gasifier, ASU, PSA etc and the energy obtained from the exothermic H 2 production reaction, cooling of hot streams, etc. At a S:C ratio of 2, the purity of H 2 increases with the increase in pressure to 5 atms. With a further increase in pressure to 20 atms, the H 2 purity falls by <1%. In contrast, the H 2 purity at a S:C ratio of 1 217

249 decreases by 3.5% with the increase in pressure from 1 to 20 atms. This decrease in H 2 purity with the increase in pressure is due to the increase in the formation of CH 4 in the H 2 production reaction from the CO and H 2 in the syngas. The cold gas efficiency as well as the process efficiency increase with the increase in pressure from 1 to 10 atms and then decrease by a small amount with the further increase in pressure to 20 atms Effect of S:C Ratio With the decrease in S:C ratio, although a small decrease in the H 2 purity is observed in Figure 6.14 due to the increased production of CH 4, there is almost no change in the cold gas efficiency or the process efficiency. This is due to the heat integration within the process which utilizes all the CH 4 in the tail gas of the PSA to provide a part of the parasitic energy requirement of the process Effect of Temperature The effect of the increase in temperature of the H 2 production reactor on the purity of H 2 produced, and efficiency of the process was also investigated and as illustrated in Figure 6.15 it was found that the purity of H 2 decreased with the increase in the temperature of the combined water gas shift, carbonation and sulfidation reaction. With the increase in temperature the removal efficiency of CO 2 by the CaO sorbent is reduced and hence the purity of H 2 is also decreased. However, the efficiency of the process remained constant with the increase in temperature of the H 2 production reactor. 218

250 Effect of Ca:C Ratio The performance of the CLP depends to a large extent on the reactivity and recyclability of the calcium sorbent. For highly reactive sorbents the Ca:C ratio required is lower than that required for naturally occurring limestone and hence the amount of solid circulation will also be low. As illustrated in Figure 6.16, the increase in Ca:C ratio results in a decrease in the cold gas efficiency as well as process efficiency. Hence by using a highly reactive sorbent, the amount of solids circulation can be reduced and the efficiency can be improved. 6.6 EFFECT OF ADDITION OF SORBENT HYDRATION TO THE CLP PROCESS Figure 6.17 illustrates the ASPEN Plus model for the CLP with sorbent hydration as a part of the carbonation - calcination cycle. The sorbent at the exit of the calciner is hydrated at a high temperature of 600 ºC and a pressure of 21 atms. The hydrated sorbent is then fed to the carbonation or H 2 production reactor. In the carbonation reactor the Ca(OH) 2 sorbent is converted to CaCO 3 and the steam produced from the Ca(OH) 2 is consumed in the water gas shift reaction. The hydration of CaO is exothermic and hence heat is extracted in the hydration reactor. A part of the exothermic energy released in the carbonation reactor due to the exothermic of carbonation and the water gas shift reaction is consumed by the endothermic decomposition of the Ca(OH) 2. The addition of sorbent hydration aids in reducing the 219

251 Ca:C ratio and the solids circulation in the CLP process. The reduction in Ca:C ratio aids in improving the overall efficiency of the process as shown in Figure TECHNO-ECONOMIC ANALYSIS OF H 2 PRODUCTION FROM COAL A comparison of the economics of the conventional coal to H 2 process with the CLP process has been discussed in this section. Figure 6.18 is a flow diagram of the conventional coal to H 2 process that was used as the base case in this study. The design, assumptions and economic analysis for the base case is based on a draft US Department of Energy (DOE) study (DOE, 2009). In the conventional plant which is the base case in this study, H 2 is produced from a GE gasifier followed by a water quench, syngas scrubber, water gas shift reactors, syngas coolers, mercury removal system, dual stage selexol system for the removal of CO 2 and H 2 S and a PSA. The PSA produces 99.9% pure hydrogen and the tail gas is combusted in a boiler to generate steam for electricity production. The H 2 S is sent to a Claus plant for the production of elemental sulfur and the CO 2 is dried and compressed for transportation and sequestration. The assumptions used for the conventional process are listed below. Key parameters for the conventional process: 1) 249 tons/hr of Illinois #6 coal is fed to the gasifier along with 243 tons/hr of O 2 from an ASU for hydrogen production. 2) 26 tons/hr of solid waste is generated from the process. 3) 26 tons/hr of 99.9% pure H 2 at 21 atms and 31 MWe is produced. 220

252 4) The process results in a net CO 2 emission of 60 tons/hr and 517 tons/hr of CO 2 is sequestered. Figure 6 illustrates a flow diagram of the CLP integrated with a GE gasifier. The syngas from the gasifier is cooled in a radiant cooler and its pressure is reduced in an expander. Since the syngas sent to the expander should be free from all particulate matter, a metallic filter is used to remove the flyash from the syngas. The syngas is then sent to the carbonation reactor along with Ca(OH) 2 sorbent from the hydrator. The Ca(OH) 2 dehydrates in the carbonation reactor producing steam which is consumed in the water gas shift reaction. The H 2 rich stream with the spent sorbent is then sent to a cyclone and metallic filter assemble to separate the sorbent from the H 2 stream. The H 2 rich stream is then further purified to 99.9% in a PSA. The spent sorbent is sent to the calciner. The energy for the calciner is supplied by the direct oxy combustion of coal and the PSA tail gas. The calcined sorbent is separated from the CO 2 stream in a cyclone and 95% of the sorbent is recovered. The remaining 5% of the sorbent is cooled with the CO 2 stream and is finally separated from the CO 2 stream in a fabric filter at low temperature. The CO 2 is then dried and compressed for transportation and sequestration. The calcined sorbent is hydrated with steam at 500C and sent to the carbonation reactor. The assumptions for the CLP are listed below. Key parameters for the CLP process: 1) The CLP process is a co-generation facility which results in the production of 23 tons/hr of 99.9% pure H 2 at 21 atms and 324 MWe of electricity. 221

253 2) A total of 367 tons/hr of Illinois #6 coal is fed to the over all process. 249 tons/hr of coal is fed to the gasifier while the rest is fed to the calciner to provide the energy for calcination. 3) A total of 592 tons/hr of O 2 is fed to the process out of which 243 tons/hr is fed to the gasifier and the rest is fed to the calciner to combust the coal and PSA tailgas. 4) A Ca:C ratio of 1.3 is used in the carbonation reactor and 5 wt% of the solids at the exit of the calciner is purged and an equivalent moles of limestone is added to the calciner as makeup. 5) 139 tons/hr of solid waste is generated from the process. 6) The process results in almost zero CO 2 emissions and 865 tons/hr of CO 2 is sequestered. Based on the process simulation and the economic assumptions used for the base case, the economic analysis was conducted for the CLP. All costs are overnight costs in 2008 dollars and the cost estimates were prepared for an Nth-of-a-kind plant. The plants are assumed to be located in Midwest US. The economic comparison of the conventional plant and the CLP is expressed on the basis of the levelized cost of H 2 in dollars per Kg of H 2. The capital charge factor used to levelize the capital costs and the levelization factors for coal, electricity and O& M costs are provided in Tables 6.12 and The annual levelized costs for the conventional and the CLP plants are provided in Tables 6.12 and 6.13 respectively. As shown in the tables, the CLP plant has higher capital costs, fixed O&M and variable O&M costs but it also produces more 222

254 than 10 times the amount of electricity produced by the conventional plant. In addition, the CLP plant has almost no CO 2 emission while the conventional plant emits 10% of the CO 2. The increased costs of capital and O&M are offset by the large amount of electricity produced and the credit obtained for the amount of CO 2 avoided. Hence the CLP plant has a levelized cost of H 2 of $1.81/Kg of H 2 while the conventional plant has a levelized cost of $2.03/Kg of H 2. A more detailed description of the two processes, assumptions for the technical and economic analysis, and the results is provided elsewhere. (Connell et al, 2010) 6.7 CONCLUSIONS The CLP integrated in a gasification system was simulated for the production of H 2 from coal using ASPEN Plus software. Two cases were explored for the production of only H 2 from the entire process and for the cogeneration of H 2 and electricity in the process. The effect of the addition of a PSA at the end of the process for the production of high purity fuel cell grade H 2 was also evaluated. Different types of gasification systems in conjunction with the CLP were evaluated and it was found that the efficiency of the coal to H 2 process depends on the composition of syngas obtained from the gasifier as well as the efficiency of the gasifier. The syngas obtained from the Shell gasifier was found to form the least amount of CH 4 in the calcium looping H 2 reactor when compared to the syngas obtained from the other gasifiers. In addition to the reduction in CH 4 production, the Shell gasifier also has a higher efficiency for the conversion of coal to syngas and hence the highest efficiency for the conversion of coal 223

255 to H 2 was obtained for the integration of the CLP with the Shell gasifier. The effect of S:C ratio, temperature and pressure were also investigated. The purity of H 2 produced from the H 2 production reactor was found to decrease by a small amount with the decrease in S:C ratio and the increase in temperature. The decrease in H 2 purity with the decrease in S:C ratio especially at high pressures was due to the increase in the formation of CH 4 in the H 2 production reactor. However, the decrease in S:C ratio and temperature did not result in a significant change in the process efficiency and the cold gas efficiency. The increase in pressure from 1 to 10 atms resulted in an increase in process efficiency. A further increase in pressure did not result in a change in the process efficiency. A decrease in the Ca:C ratio resulted in an increase in the efficiency of the process and hence a sorbent with a higher reactivity and recyclability is beneficial for the process. A techno-economic comparison of the CLP with the conventional coal to H 2 process shows that the CLP has the potential to reduce the cost of H 2 production from coal. 224

256 Proximate Analysis Wt% (As- Received) Wt%, dry Ultimate Wt% (As- Received) Moisture Moisture Wt%, dry Fixed Carbon Ash Volatiles Carbon Ash Hydrogen Total Nitrogen Chlorine HHV (MJ/kg) Sulfur Oxygen Table 6.1: Properties of Illinois # 6 coal 225

257 Syngas Composition Mole % H 2 O 2.5 N O 2 0 H CO 61.4 CO Ar 0.8 COS (ppm) 884 H 2 S 1.2 CH Temperature ( C) 1538 Pressure (atm) 36 Mass Flow Rate (Kg/hr) Table 6.2: Composition of the syngas exiting from the Shell gasifier 226

258 Stage Discharge Pressure (Mpa) Table 6.3: Intermediated pressures for compression of the CO 2 for sequestration 227

259 Component ID Type Component name Formula CH 4 CONV METHANE CH 4 CO 2 CONV CARBON-DIOXIDE CO 2 CO CONV CARBON-MONOXIDE CO H 2 CONV HYDROGEN H 2 C 2 H 6 CONV ETHANE C 2 H 6 C 2 H 4 CONV ETHYLENE C 2 H 4 H 2 0 CONV WATER H 2 O CACO 3 SOLID CALCIUM-CARBONATE-CALCITE CACO 3 CAO SOLID CALCIUM-OXIDE CAO C SOLID CARBON-GRAPHITE C H 2 S CONV HYDROGEN-SULFIDE H 2 S CAS CONV CALCIUM-SULFIDE CAS COS CONV CARBONYL-SULFIDE COS N 2 CONV NITROGEN N 2 O 2 CONV OXYGEN O 2 AR CONV ARGON AR NH 3 CONV AMMONIA H 3 N COAL NC S CONV SULFUR S CL 2 CONV CHLORINE CL 2 HCL CONV HYDROGEN-CHLORIDE HCL ASH NC O 2 S CONV SULFUR-DIOXIDE O 2 S Table 6.4: Components list for the ASPEN Plus simulation 228

260 Unit Operation Aspen Plus Model Comments / Specifications CLP Hydrogen Reactor RGibbs 1.32:1 Calcium:Carbon molar ratio based on active calcium sorbent and total carbon content in syngas, 3:1 S:C ratio, thermodynamic modeling of the water gas shift, carbonation and sulfidation reaction of CaO, isothermal operation with heat extraction Purge for solids FSplit Splits and purges 6.2% of the solids based on molar fraction Solids Make-up Mixer Combines recycle stream and fresh feed stream in terms of material and heat Calciner RGibbs Thermodynamic modeling of limestone calcination with syngas and PSA tail gas combustion operates isothermally at 850 C with 100% conversion of CaCO 3 to CaO and complete combustion of syngas and PSA tailgas in oxygen Gas-Solid Separation SSplit Operates with 100% separation efficiency HRSG MHeatX Modeling of heat exchange among multiple streams PSA Sep 90% yield of H 2 obtained from the PSA, remaining 10% H 2 and other gas components removed in the PSA tailgas stream CO 2 Compression 105 kwh electricity/tonne CO 2 to compress to 150 atm Table 6.5: ASPEN Plus models used for the simulation of the CLP 229

261 Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S Table 6.6: Material and energy balance for the CLP Continued 230

262 Table 6.6 Continued Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S Continued 231

263 Table 6.6 continued Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S Continued 232

264 Table 6.6 continued Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S Continued 233

265 Table 6.6 continued Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S Continued 234

266 Table 6.6 continued Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S Continued 235

267 Table 6.6 continued Temperature C Pressure bar Mass Flow kg/hr Volume Flow cum/hr Mole Flow kmol/hr CH CO CO H H 2 O H 2 S COS N O AR O 2 S CaCO CAO C CAS S

268 MWe Electricity from Steam Turbine Electricity from Syngas Expander Electricity Output Coal handling, milling and coal slurry pumps ASU air compressor and oxygen compressor 30 CO 2 compressor Feed water pumps H 2 Compression 11.6 Electricity Used in the Plant Net Electricity Produced Table 6.7: Power balance in the CLP process 237

269 Hydrogen and electricity Coal feed (tonnes/day) 2405 Carbon Capture(%) 100 Hydrogen(tonnes/day) 280 Hydrogen (MW) 457 Net Power(MW) 81 Overall Efficiency(%HHV) 66.5 Table 6.8: Process simulation results for the CLP process 238

270 Hydrogen Hydrogen and Electricity Coal Feed (t/day) 2,180 2,405 Carbon Capture (%) Hydrogen (t/day) Hydrogen (MW, HHV) Net Power (MW) 0 81 Overall Efficiency (%HHV) Table 6.9 Summary of the schemes investigated for the production of H 2 alone and for the coproduction of H 2 and electricity with a PSA 239

271 Hydrogen Hydrogen and Electricity Coal Feed (t/day) 2,155 2,350 Carbon Capture(%) Hydrogen (t/day) Hydrogen (MW) Net Power (MW) Overall Efficiency (%HHV) Table 6.10 Summary of the schemes investigated for the production of H 2 alone and for the coproduction of H 2 and electricity without a PSA 240

272 Hydrogen Hydrogen and Electricity Shell 62.3% 66%(81 MW) Lurgi (BGL) 55% 56(32 MW) GE 60% 63.6(104.2 MW) Table 6.11 Comparison of the efficiency of the H 2 production process for different gasifiers 241

273 Cost ($ or $/y) Capital Charge Factor and Levelization factors Levelized Annual Cost Capital (TPC) Fixed O&M Coal Electricity CO2 Emission Allowances Other Variable O&M TOTAL Levelized cost of H 2 = / = $2.03/kg H 2 Table 6.12: Levelized annual costs and levelized cost of H 2 for the conventional coal to H 2 plant (adapted from DOE, 2010) 242

274 Cost ($ or $/y) Capital Charge Factor and Levelization factors Levelized Annual Cost Capital (TPC) Fixed O&M Coal Electricity CO2 Emission Allowances Other Variable O&M TOTAL Levelized cost of H 2 = / = $1.81/kg H 2 Table 6.13: Levelized annual costs and levelized cost of H 2 for the CLP plant 243

275 Calciner CO 2 Compression CO 2 to Sequestration Radiant Cooler H 2 Production Reactor PCD Hg Removal PSA Coal GE Gasifier Gasifier Slag Steam Calcium Looping Process Pure H 2 Slurry Water ASU N 2 Rich Stream Air Figure 6.1: The CLP for coproduction of fuel cell grade H 2 and electricity from coal 244

276 MI XE R FSPLI T Coal Shell Gasifier Q-DECOMP Steam Generation DRY - CO AL DECOMP INBU RNER GASIFIER STEAM3 RYIELD RGI BBS B12 23 B10 HOTSEQCO B7 CACO3SOL 2 32 CA O CO2 H2S -O Calciner CA LC CAO,CO2 9 Integrated Reactor CARB W ASU ASU COO L P1 COMP AIRO AIR O2C N B11 B15 B8 34 COMPC O2 B13 SOLMAKUP SOLPURGE 19 1 B14 6 B B6 B2 CALCSYNG 5 B4 29 PUREH2 18 CACO3 53 Q CYCCO2 Q 11 CYCH2 B20 20 B1 8 H OTOFFGA PSA 33 PSA B16 10 Q TAILGA S H 2 at 20 bar SYNGA S COM BUS TO 30 B9 CO 2 B Figure 6.2: ASPEN simulation flow diagram for the CLP process with a PSA 245

277 FLASH2 MIXER FSPLIT Coal Shell Gasifier Steam Generation STEAM3 Q-DECOMP 40 B12 HOTSEQCO B7 32 CAO,CO2 9 DECOMP RYIELD INBURNER GASIFIER 39 RGIBBS CAO B10 B11 B3 B6 B18 B SOLPURGE B5 4 B8 COMPCO2 CO2H2S-O CACO3SOL CALCSYNG B22 17 CALC CARB 2 W 38 Q CYCCO2 B19 B20 11 B1 B2 B4 5 B21 Calcine r CACO3 18 PUREH2 CO 2 at atm psi Integrated Q Reactor H 2 at atm bar B16 ASU 35 1 Q CYCH2 SYNGAS ASU 2 P1 COOL COMP O2C AIR HYDROGEN Figure 6.3 Aspen simulation for the production of H 2 using the CLP without a PSA. 246

278 MIXER FSPLIT STEAM3 B8 27 B B12 B10 HOTSEQCO B7 CACO3SOL 2 32 CAO CO2H2S-O CALC CAO,CO2 W= Hydrogen production reactor Q Duty (Gcal/hr) W Power(kW) CARB 15 W B11 B13 SOLMAKUP Q=7 B5 SOLPURGE Q= B2 5 B4 Q=132 Q CYCCO2 Q 11 CYCH2 Q=-315 B20 B1 W=-8891 Q=-90 HOTOFFGA 33 B15 SYNGAS CACO3 PSA B9 W=311 B17 W= B21 B18 B B6 18 PUREH B16 Q=-0 Q=-2 Q TAILGAS 30 COMBUSTO Q=-37 Q=-1 B3 Q= Figure 6.4: Aspen model used for sensitivity analysis of the combined reactions occurring in the H 2 production reactor of the CLP. 247

279 248 Purity of Hydrogen Effect of Temperature on Hydrogen Purity PURE Temperature (C) Figure 6.5: Effect of temperature on the H 2 purity produced at the outlet of the carbonation reactor (S:C ratio = 3, Pressure = 10 atms) 248

280 249 Purity of Hydrogen Effect of Pressure on Hydrogen Purity Pressure (bar) (atm) Figure 6.6: Effect of pressure on the H 2 purity produced at the outlet of the carbonation reactor( S:C ratio = 3, Temperature = 600 ºC) 249

281 250 Hydrogen Purity (%) Effect of Steam to Carbon Ratio on Hydrogen Purity PURE Steam to Carbon Ratio Figure 6.7: Effect of S:C ratio on the H 2 purity produced at the outlet of the carbonation reactor ( Pressure = 10 atms, Temperature = 600 ºC) 250

282 Steam Ratio vs. H2S Output 90 Outlet H2S Flowrate (Kmoles/hr) Steam Ratio Figure 6.8: Effect of temperature and S:C ratio on the extent of H 2 S removal. 251

283 Steam Ratio vs. COS Output 1.2 Outlet COS Flowrate (Kmoles/hr) Steam Ratio Figure 6.9: Effect of temperature and S:C ratio on the extent of COS removal. 252

284 Steam Ratio vs. CO Output Outlet CO Flowrate (Kmoles/hr) Steam Ratio Figure 6.10: Effect of temperature and S:C ratio on the amount of CO impurity present in the H 2 stream. 253

285 Steam Ratio vs. CO2 Output Outlet CO2 Flowrate (Kmoles/hr) Steam Ratio Figure 6.11: Effect of temperature and S:C ratio on the extent of CO 2 removal. 254

286 Steam Ratio vs. CH4 Output Outlet CH4 Flowrate (Kmoles/hr) Steam Ratio Figure 6.12: Effect of temperature and S:C ratio on the amount of CH 4 impurity present in the H 2 product stream. 255

287 H2 Purity (%) Efficiency (%) Pressure (bar) atm S/C = 1 - H2 Purity S/C = 1 - H2 Cold gas Efficiency S/C = 1 -Process Efficiency S/C = 2 - H2 Purity S/C = 2 - H2 Cold gas Efficiency S/C = 2 -Process Efficiency Figure 6.13: Effect of pressure on the cold gas efficiency, process efficiency and H 2 purity obtained from the H 2 production reactor at various S:C ratios. 256

288 H2 Purity Efficiency (%) S/C Ratio H2 Purity H2 Cold gas Efficiency Process Efficiency Figure 6.14: Effect of S:C ratio on H 2 purity, cold gas efficiency and process efficiency 257

289 H2 Purity (%) Efficiency (%) Temperature (C) H2 Purity H2 Cold gas Efficiency Process Efficiency Figure 6.15: Effect of temperature on H 2 purity, cold gas efficiency and process efficiency (1:1, 10 atms) 258

290 H2 Purity(%) Efficiency(%) Ca/C Ratio H2 Purity H2 Cold gas Efficiency Process Efficiency 45 Figure 6.16: Effect of Ca:C ratio on H 2 purity, cold gas efficiency and process efficiency (600 ºC, 1:1, 10 atms) 259

291 MIXER F S P L I T W=57291 B12 B11 Q=-35 Steam generation 45 B29 B26 42 Q=5 Q= B19 B7 7 Q=-8 Q=38 Q= B10 32 B21 47 B B14 Q=8 B6 Q=22 33 Hydration 30 CAO B17 B B8 Q=-161 Q= B25 31 Calcination 2 38 CO2H2S-O Integrated Reactor Q Duty (Gcal/hr) W Power(kW) 27 B22 W= W=105 B Q=54 18 SOLMAKUP Q=6 B B9 Q=22 24 B2 CALC Q=124 CAO,CO2 17 Q CYCCO2 21 Q CARB W 20 B20 8 W=-8891 Q=-173 TAILGA S B1 Q=-90 B15 SYNGAS 260 SOLPURGE B5 Q=-6 4 B4 CACO3 H2 CYCH2 PSA 5 B28 PUREH2 10 Q=14 23 B16 Q=-3 Q=-1 Q HYDROGEN 53 Figure 6.17: Effect of the addition of sorbent hydration to the CLP 260

292 CO 2 Sequestration CO 2 Compression Steam Sulfuric Acid Plant H 2 S Removal 261 Air Radiant Cooler Scrubber Raw Syngas Shift Reactors Syngas Cooling Final Syngas Scrubber Mercury Removal Dual Stage Acid Gas Removal ASU PSA H 2 Coal Feed Gasifier Air Boiler Flue Gas Coal Water Slag Steam Turbine Figure 6.18: Process flow diagram of the conventional coal to H 2 plant used for the economical analysis ( DOE, 2010) 261

293 Air Radiant Cooler Metallic Filter Expander Carbonator H2 Cooling Metallic Filter PSA H 2 ASU ASU 262 Coal Feed Coal Water Gasifier Slag Sorbent Makeup Filter Coal Cal ci ner Filter Solids Purge CO2 Condenser Hydrator Figure 6.19: Process flow diagram of the CLP plant used for the economical analysis 262

294 CHAPTER 7 ENHANCED REFORMING OF HYDROCARBONS 7.1 INTRODUCTION The CLP for the reforming of hydrocarbons is similar to the CLP for the conversion of syngas to H 2. For a hydrocarbon feed, the steam reforming of the hydrocarbon is integrated with the water gas shift and carbonation reaction in a single reactor. In addition to improving the conversion of the hydrocarbon to H 2, the CLP also provides an efficient mode of internal heat integration where the endothermic energy for the reforming reaction is obtained from the exothermic energy released by the combined water gas shift and carbonation reaction. A schematic of the CLP for the reforming of hydrocarbons is shown in Figure 7.1. The CLP comprises of three reactors; the carbonation reactor where the thermodynamic constraint of the reforming and water gas shift reaction is overcome by the in-situ removal of the CO 2 product by a calcium based sorbent, the calciner where the spent calcium sorbent is regenerated and a sequestration-ready CO 2 stream is produced and the hydrator where the calcined sorbent is reactivated to improve its recyclability. 263

295 7.2 PROCESS CONFIGURATION AND THERMODYNAMICS The Carbonation Reactor System The carbonation reactor comprises either a fluidized bed, fixed fluidized bed or an entrained flow reactor that operates at pressures ranging from 1 to 30 atm and temperatures of o C. In the carbonation reactor, the reforming reaction, water gas shift reaction and CO 2 removal occur in a single reactor in the presence of a reforming catalyst and CaO sorbent. The steam reforming of the hydrocarbon occurs in the presence of the reforming catalyst, and the CO 2 produced by the combined reforming and water gas shift reaction is removed by the CaO sorbent. The concomitant carbonation of the CaO leading to the formation of CaCO 3 incessantly drives the equilibrium-limited water gas shift and reforming reaction forward by removing the CO 2 product from the gas mixture. Various reactions occurring in the carbonator are as follows: Hydrocarbon reforming: C x H y + xh 2 O xco + ( y / 2 +x) H 2 (7.1) CH 4 + H 2 O CO + 3H 2 (7.2) Water Gas Shift Reaction: CO + H 2 O H 2 + CO 2 (7.3) Carbonation Reaction: CaO + CO 2 CaCO 3 (7.4) The CLP offers several advantages. By improving the equilibrium conversion of the reforming and water gas shift reaction, steam addition can be greatly reduced. In 264

296 addition, since the combined reforming, water gas shift and carbonation reaction occurs at a high temperature of 500 to 750 ºC, the water gas shift catalyst can also be eliminated. A major advantage of the CLP is the internal heat integration that it provides to the reforming of hydrocarbons. The exothermic carbonation and water gas shift reactions convert the highly endothermic reforming of hydrocarbons into a heat neutral process thus simplifying the reforming process and reducing the temperature of reforming from >900 ºC to 650 ºC. The heat of reaction of the combined steam methane reforming, water gas shift and carbonation reaction occurring in the carbonator is shown below: Steam Methane Reforming and Water Gas Shift: CH 4 + 2H 2 O = CO 2 + 4H 2 H = +165 KJ/mole (7.5) Carbonation Reaction: CaO + CO 2 = CaCO 3 H = -178 KJ/mole (7.4) Combined Reaction: CH 4 + 2H 2 O+CaO = CaCO 3 + 4H 2 H = -13 KJ/mole (7.6) Thermodynamic analysis of reactions occurring in the carbonation reactor The equilibrium constants for the steam methane reforming (equation 7.2), steam methane reforming and water gas shift reaction (equation 7.5), and the combined reforming and carbonation reaction (equation 7.6) for various temperatures are shown in Figure 7.2. The equilibrium constants are obtained using HSC Chemistry v 5.0 (Outokumpu Research Oy, Finland). The equilibrium constant for the steam methane reforming reaction can be defined as shown below: 265

297 P C O 2 P 3 H 2 P CO K eq 1 = P CH C 4 P H O 2 H O 2 (7.7) where P CH4, P H2, P CO, P H2O are the partial pressures of CH 4, H 2, CO and H 2 O at equilibrium. The equilibrium constant of the combined reforming and water gas shift reaction is as follows: P C O 2 P 4 H 2 P K = CO2 eq 2 P CH4 C P 2 H O 2 H 2 O (7.8) The equilibrium constant for the combined reforming, water gas shift and carbonation reaction is defined as shown below: K eq 3 = P P H C P H4 H 2 O (7.9) Where K eq3 = K eq2 * K carb and K carb is the equilibrium constant of the CaO carbonation reaction. From Figure 7.2 it can be seen that the steam methane reforming reaction does not occur at temperatures below 600 ºC. The equilibrium constants of the steam methane reforming and the combined reforming and water gas shift reactions increase 266

298 with the increase in temperature beyond 600 ºC. The equilibrium constant of the combined CaO carbonation, steam methane reforming and water gas shift reaction is higher than that of the steam methane reforming reaction alone at temperatures of 0 to 800 ºC. Beyond 800 ºC, the carbonation of CaO does not occur. Hence, the CLP is capable of reducing the temperature of the steam methane reforming reaction with the aid of the conventional steam methane reforming catalyst Calciner or Sorbent Regeneration Reactor The spent sorbent at the exit of the carbonation reactor is a mixture consisting of CaCO 3 and CaO. The CaCO 3 in the spent sorbent mixture is regenerated back to CaO in the calciner. The calciner is operated at atmospheric pressure in a rotary or a fluidized bed system. The heat can be supplied directly or indirectly using a mixture of fuel and oxidant. For a directly fired calciner, the heat of calcination can be provided by the combustion of natural gas in oxygen. From the thermodynamic curve for CaO and CO 2, calcination is found to occur at temperatures above 890 o C in the presence of 1 atm of CO 2. Dilution of CO 2 in an indirectly fired calciner with steam or oxycombustion of natural gas in a direct fired calciner will permit the calcination reaction to be conducted at temperatures lower than 890 o C. The reaction occurring in the calciner is: Calcination : CaCO 3 CaO + CO 2 (7.10) 267

299 Hydrator or Sorbent Reactivation Reactor The hydrator is similar to that described in Chapter 4 for H 2 production from syngas and the reaction occurring in the hydrator is shown below: Hydration : CaO + H 2 O = Ca(OH) 2 (7.11) The Ca(OH) 2 from the hydrator is conveyed to the carbonation reactor where it dehydrates to produce high reactivity CaO and steam. The steam obtained from the dehydration reaction is consumed in the combined reforming and water gas shift reaction EXPERIMENTAL METHODS Chemicals, Sorbents, and Gases The reforming catalyst was procured from Süd-Chemie Inc., Louisville, KY and consists of a nickel oxide catalyst supported on calcium aluminate. The CaO sorbent is obtained from a PCC precursor which is synthesized from Ca(OH) 2 obtained from Fisher Scientific as described in Chapter Bench Scale Experiment Setup Figure 7.3 shows the integrated experimental setup, used for the bench scale studies of the CLP. The bench scale reactor is coupled with a set of continuous gas analyzers which detect concentrations of CO, CO 2, H 2 S, CH 4 and H 2 in the product 268

300 stream. The reactor setup is capable of handling high pressures and temperatures of up to 21 atms and 900 o C respectively, which are representative of the conditions in a commercial syngas to H 2 system. A description of the bench scale setup is provided in Chapter Steam Methane Reforming in the Presence of a Ni-based Catalyst The extent of the steam methane reforming reaction was determined in the presence of the reforming catalyst obtained from Süd-Chemie. Catalyst particles were used in a fixed bed reactor setup for all the experiments. 3 g of the catalyst was loaded into the reactor and the pressure, temperature and gas flow rates were adjusted for each run. Pure CH 4 was used as the feed gas for all the tests and it was metered into the reactor by a mass flow controller. From the mass flow controller, the CH 4 flows to the steam generating unit which also serves to preheat the CH 4 entering the reactor. The product gas mixture exiting the back pressure regulator is then cooled in a heat exchanger using chilled ethylene glycol-water mixture to condense the unconverted steam. The product gas at the exit of the heat exchanger is dried in a desiccant bed. The dry gas compositions are monitored continuously using the CO, CO 2, H 2 S, CH 4 and H 2 gas analyzers. 269

301 7.3.4 Simultaneous Steam Methane Reforming, Water Gas Shift and Carbonation The simultaneous reforming, water gas shift and carbonation reaction was conducted in the presence of the CaO sorbent and nickel based reforming catalyst. 2.5 grams of PCC was mixed with 2.5 grams of reforming catalyst and loaded into the reactor. The mixture was calcined by heating the reactor to 700 o C in a stream of N 2 until the CO 2 analyzer confirmed the absence of CO 2 in the outlet stream. At the end of calcination, the temperature of the reactor was reduced to 650 ºC and H 2 was made to flow through the catalyst and sorbent bed for an hour to reduce the catalyst to the active form for the steam methane reforming reaction. At the end of catalyst pretreatment, the temperature of the reactor was set for the combined reforming water gas shift and carbonation reaction. Pure CH 4 was used as the feed gas for all the tests and it was metered into the steam generator by a mass flow controller. The CH 4 and steam mixture was introduced into the reactor and dry gas composition of the product gas was monitored continuously using the CO, CO 2, H 2 S, CH 4 and H 2 gas analyzers Multicyclic Steam Methane Reforming and Spent Sorbent Calcination Effect of Sorbent Calcination Conditions on the Extent of Steam Reforming In these tests the effect of two sorbent calcination conditions was investigated on the extent of CH 4 reforming obtained in the presence of CaO sorbent and Ni based catalyst. In the first set of tests, the PCC sorbent and Ni-based catalyst mixture was loaded into the reactor and the calcination was conducted at 950 ºC in pure N 2. The 270

302 mixture was then exposed to H 2 to activate the catalyst at 650 ºC. The combined reforming, water gas shift and carbonation reaction was conducted at 650 ºC and 1 atm with a 100% CH 4 feed stream and a S:C ratio of 3:1. At the end of the H 2 production stage, the sorbent was calcined again at 950 ºC in pure N 2 and 3 cycles were repeated. For the 4 th cycle the spent sorbent at the end of the 3 rd cycle was calcined in a 50:50 CO 2 /H 2 O mixture at 950 ºC. The sorbent catalyst mixture was then reduced in H 2 at 650 ºC and the combined reforming, water gas shift and carbonation reaction was conducted at 650 ºC Calcination in N 2 with Sorbent Hydration 2.5 grams of PCC sorbent was loaded into the reactor and the calcination was conducted at 950 ºC in pure N 2. At the end of calcination, the sorbent was hydrated in a 80:20 H 2 O/N 2 gas mixture at 600 ºC and 11 atms. The hydrated sorbent was then mixed with the 2.5 g of reforming catalyst and loaded into the reactor. The mixture was then exposed to H 2 to activate the catalyst at 650 ºC. The combined reforming, water gas shift and carbonation reaction was conducted at 650 ºC and 1 atm with a 100% CH 4 feed stream and a S:C ratio of 3:1. At the end of the H 2 production stage, the sorbent was separated from the catalyst. The sorbent was calcined again at 950 ºC in pure N 2 and then hydrated and 4 cycles were repeated in a similar manner. 271

303 Realistic Sorbent Calcination in a Steam/CO 2 Atmosphere With Hydration 2.5 grams of PCC sorbent was loaded into the reactor and the calcination was conducted at 950 ºC in a 50:50 CO 2 /H 2 O mixture. At the end of calcination, the sorbent was hydrated in a 80:20 H 2 O/N 2 gas mixture at 600 ºC and 11 atms. The hydrated sorbent was then mixed with 2.5 grams of the reforming catalyst and loaded into the reactor. The mixture was then exposed to H 2 to activate the catalyst at 650 ºC. The combined reforming, water gas shift and carbonation reaction was conducted at 650 ºC and 1 atm with a 100% CH 4 feed stream and a S:C ratio of 3:1. At the end of the H 2 production stage, the sorbent was separated from the catalyst. The sorbent was calcined again at 950 ºC in a 50:50 CO 2 /H 2 O mixture and 4 cycles were repeated in a similar manner. 7.4 RESULTS AND DISCUSSION Base-line Steam Methane Reforming Testing Base line tests were conducted in the bench scale reactor for the steam reforming of CH 4 in the presence of the nickel based catalyst procured from Sud Chemie. The reactor was filled with a mixture 5 gms of catalyst and quartz chips. Pure CH 4 from the mass flow controller was mixed with steam in the steam generation section and sent to the reactor. Figure 7.4(a) illustrates the composition of H 2 in the product gas at the outlet of the reactor on a steam and nitrogen free basis. It can be seen in Figure 7.4(a) that for both S:C ratios of 3:1 and 5:1, the composition of H 2 increases 272

304 with the increase in temperature till the temperature reaches 700 ºC. At temperatures above 700 ºC, the purity of H 2 remains constant. With the increase in S:C ratio from 3:1 to 5:1, it was found that the purity of H 2 increases at all the temperatures investigated. Figure 7.4(b) depicts the composition of CO, CO 2 and CH 4 in the product gas at the outlet of the reactor on a steam and nitrogen free basis. The amount of CH 4 in the stream decreases with temperature and is in the ppm range at temperatures above 850 ºC. It can also be seen that with the increase in the S:C ratio the conversion of CH 4 to CO, CO 2 and H 2 increases. The CO content of the product stream increases with the increase in temperature as the conversion of CH 4 to CO increases. In addition, the conversion of CO to CO 2 decreases with an increase in temperature due to the equilibrium limitation of the water gas shift reaction. The amount of CO 2 decreases with an increase in temperature due to the thermodynamic constraint of the water gas shift reaction. The amount of CO 2 increases with the increase in S:C ratio due to the higher conversion of CO to CO 2 by the water gas shift reaction at higher steam addition rates Simultaneous Reforming with In-situ CO 2 Removal (Catalyst with CaO Sorbent) In order to study the improvement in CH 4 conversion and H 2 purity, the steam methane reforming reaction was conducted in the presence of the reforming catalyst and CaO sorbent. The mixture of sorbent and catalyst was fed into the reactor and the sorbent was calcined at 700 ºC in pure nitrogen. Following this, the catalyst and 273

305 sorbent mixture was exposed to pure H 2 at 650 ºC in order to reduce the catalyst to its active form. Pure CH 4 and steam were then fed into the reactor at 650 ºC and atmospheric pressure. Figure 7.5 depicts the concentration of H 2, CO, CO 2 and CH 4 in the product gas for the steam methane reforming reaction conducted in the presence of CaO sorbent and Ni based catalyst. It was found that >99% pure H 2 can be obtained in the pre-breakthrough region of the curve when the CaO sorbent is active. It was also found that the CH 4 is almost completely converted and the concentration of CH 4, CO and CO 2 in the product stream is only a few ppm. The removal of CO 2 by the CaO sorbent enhances the water gas shift reaction and the reforming reaction resulting in the production of a pure H 2 product stream. As the sorbent gets consumed in the fixed bed, the concentrations of CO 2, CH 4 and CO begin to increase in the product H 2 stream. This region of the curve is the breakthrough period. At the end of the breakthrough period, the CaO sorbent is completely converted to CaCO 3 and no further CO 2 capture is obtained. The conversion of CH 4 to H 2 occurs in the presence of the reforming catalyst in the post-breakthrough period. As illustrated in Figure 7.6, it can be seen that >99% conversion of CH 4 can be obtained during the pre-breakthrough period of the combined reforming, water gas shift and reforming reaction. As the sorbent gets consumed, the conversion of CH 4 decreases forming the breakthrough region of the curve. During the post-breakthrough period, the sorbent is in the form of CaCO 3 and the reforming reaction takes place in the presence of the catalyst alone. 274

306 Effect of Temperature and S:C Ratio The effect of S:C ratio and temperature was investigated on the reforming of CH 4 in the presence of the reforming catalyst and the CaO sorbent. The purity of H 2 produced is greatly enhanced by the presence of the sorbent as shown in Figure 7.7(a). Purity of H 2 increases from <80% in the presence of the catalyst alone to >90% in the presence of the catalyst and sorbent. Higher H 2 purity is obtained at 650 ºC than at 700 ºC due to the favorable thermodynamics of CO 2 removal by the CaO sorbent and the water gas shift reaction at lower temperatures. In the presence of the catalyst and CO 2 sorbent, the H 2 purity increases with the increase in S:C ratio from 2:1 to 3:1. A further increase in the S:C ratio does not produce an appreciable increase in H 2 purity. Figure 7.7(b) illustrates the effect of temperature and S:C ratio on the conversion of CH 4. The presence of the sorbent enhances the conversion of CH 4 to a large extent especially at 650 ºC when the presence of the sorbent increases the conversion of CH 4 from 83% to 94% at a S:C ratio of 3:1. At a S:C ratio of 5:1, the enhancement in CH 4 conversion due to addition of sorbent is not significant. The conversion of CH 4 with/without the sorbent almost reaches 100%. The conversion of CH 4 in the presence of the sorbent and catalyst mixture at 650 ºC is similar to that at 700 ºC. From Figure 7.8(a), it can be observed that at a particular temperature, the CO composition in the product stream for the sorbent-enhanced reaction is lower than the 275

307 case with the catalyst alone, which can be attributed to the fact that the presence of CO 2 sorbent enhances the water gas shift reaction that converts CO to CO 2. The CO composition is lower at 650 ºC than at 700 ºC as the extent of CO 2 removal and the water gas shift reaction are both thermodynamically more favorable 650 ºC than at 700 ºC. At 650 ºC, the CO composition does not change with S:C ratio as the removal of CO 2 as well as the water gas shift reaction almost reach completion even at a low S:C ratio. But at a higher temperature of 700 o C, the composition of CO in the product stream is more sensitive to the S:C ratio. On increasing the S:C ratio from 2:1 to 3:1 at 700 ºC, the CO concentration decreases from 12% to 3% and remains constant with the further increase in steam addition. Figure 7.8(b) depicts the composition of CO 2 in the product gas mixture for the reforming and the sorbent-enhanced reforming reactions. The CO 2 concentration decreases from >12% to <3% in the presence of the CO 2 sorbent. In the presence of the sorbent, the CO 2 composition increases with the increase in S:C ratio at 700 ºC. This results from a decrease in the partial pressure of CO 2 in the reactor due to the presence of excess steam which reduces the extent of CO 2 removal by the sorbent. However, at 650 ºC the CO 2 is completely removed from the gas mixture for S:C ratios ranging from 2:1 to 5:1 as CO 2 is removed to very low partial pressures at low temperatures. In the absence of the sorbent, the CO 2 composition in the product stream increases both at 650 ºC and 700 ºC with the increase in S:C ratio due to the production of a larger amount of CO 2 by the water gas shift reaction. 276

308 Effect of Pressure The effect of pressure was studied on the combined reforming, water gas shift and carbonation reaction in the presence of the catalyst and sorbent. Thermodynamics predicts a decrease in the purity of H 2 in the presence of a catalyst alone with the increase in pressure according to the Le Chatlier principle. Figure 7.9(a) shows that the H 2 purity remains almost a constant in the presence of the sorbent due to the simultaneous removal of CO 2. In the presence of the sorbent and catalyst, a high H 2 purity of >95% is obtained at pressures ranging from 1 to 11 atms. Figure 7.9(b) shows the effect of pressure on the concentration of CH 4 in the product stream. With an increase in pressure there is a small increase in CH 4 concentration from 3% at 1 atm to 5% at 11 atms in the pre-breakthrough region of the curves. This is due to the thermodynamics of the combined reforming, water gas shift and carbonation reactions governed by the Le Chatlier s principle. In the postbreakthrough region of the curve when the sorbent no longer captures CO 2, the increase in pressure of 1 to 11 atms results in a large increase from 5 to 17% of CH 4 in the product stream. Hence the presence of the sorbent results in a large increase in CH 4 conversion even at high pressures. CO 2 in the product steam was reduced to undetectable levels in the presence of the CaO sorbent in the pre-breakthrough curve at all pressures as shown in Figure 7.10(a). Figure 7.10(b) depicts the effect of pressure on the concentration of CO in the 277

309 product stream. Due to insitu CO 2 removal by the sorbent, the increase in pressure thermodynamically and kinetically improves the conversion of CO by the water gas shift reaction. Hence with the increase in pressure from atmospheric to 11 atms, the CO in the product gas decreases from 1.5% to ppm levels in the pre-breakthrough regions of the curves. Figure 7.11 illustrates the change in product gas composition with the increase in pressure in the pre-breakthrough and post-breakthrough regions. The prebreakthrough compositions are characteristic of the combined reforming, water gas shift and carbonation reaction as they are produced in the presence of the reforming catalyst and active CaO sorbent. The post-breakthough compositions are characteristic of only the reforming and water gas shift reaction as they are produced in the presence of the reforming catalyst and the spent CaCO 3 sorbent. There is almost no change in the pre-breakthrough CH 4, CO 2 and CO gas compositions with the change in pressure. An increase in the post-breakthrough CH 4 concentration is observed with the increase in pressure and this change is predictable from the Le Chatelier s principle while there is a decrease in CO concentration. The pre-breakthrough concentrations are substantially lower than the post-breakthrough concentrations due to the removal of CO 2 by the CaO sorbent. The advantage of the CLP is even more pronounced at high pressures where the CH 4 concentration is reduced by 12-15%, CO concentration is reduced by 5-10% while CO 2 concentration is reduced by 3% From the single cycle results shown above it can be inferred that the addition of 278

310 CaO sorbent significantly increases the conversion of CH 4 and the purity of H 2 with the simultaneous removal of CO 2 at a temperature of 650 ºC Effect of Sorbent Calcination Conditions on the Extent of Steam Reforming: The effect of two calcination conditions was tested on the extent of CH 4 conversion. Figures 7.12 (a) and (b) show the effect of the two calcination conditions on the purity of H 2 produced and the conversion of CH 4. The CaO sorbent used in cycles 1, 2 and 3 was obtained by calcination conducted in a pure nitrogen atmosphere at 950 ºC. It is observed that, the pre-breakthrough region of the curves during which the sorbent is active, reduces with successive cycles. From Figure 7.12(a), it can be observed that for cycles 1,2 and 3, high purity H 2 was produced for 464, 432 and 234 seconds, respectively. The earlier onset of breakthrough with increasing cycles can be attributed to the deactivation of the sorbent due to sintering. This decay in sorbent activity reduces the overall capacity of the sorbent for CO 2 capture, thus limiting highpurity H 2 production by the reforming and water gas shift reactions. However, in the post-breakthrough region (where the sorbent is completely exhausted), for the first three cycles, H 2 purity is almost constant. From 6(b) it can be seen that the concentration of CH 4 in the product stream follows the same trend as the breakthrough curves for H 2 purity. The time for which almost complete CH 4 conversion is achieved reduces with increase in the number of cycles for the first three cycles. 279

311 The CaO sorbent for cycle 4 was obtained by calcining the spent sorbent from the previous cycle (cycle 3) in a 50:50 mixture of CO 2 and steam at 950 ºC which is more representative of the realistic conditions used in commercial calciners. From Figure 7.12(a), it can be seen that for cycle 4, high-purity H 2 is not produced and there is no pre-breakthrough region observed. A similar observation is made for CH 4 concentration in the product H 2 stream from Figure 7.12(b). This is due to extensive sintering of the sorbent during calcination in the presence of CO 2 and steam. While the H 2 purity in the post-breakthrough region of cycle 4 was almost the same as H 2 purity in the post-breakthrough periods in the first three cycles the same is not true for the concentration of CH 4 in the product stream Calcination in N 2 with Sorbent Hydration The effect of sorbent reactivation by hydration was investigated on the cyclic carbonation and calcination of CaO sorbent during the production of H 2 from steam methane reforming. In Figures 7.13 (a) and (b) the CaO sorbent for all the 4 cycles was obtained by calcination in the presence of pure nitrogen at 950 ºC. No hydration was conducted before the first three cycles while sorbent hydration at 600 ºC and 11 atms was conducted before the 4 th reforming cycle. For the first three cycles, it is observed that the pre-breakthrough region of the curves during which the sorbent is active, reduces with successive cycles. From Figure 7.13(a), it is observed that for cycles 1,2 and 3, high purity H 2 is produced for 500, 285 and 130 seconds, respectively. The earlier onset of breakthrough with increasing cycles can be attributed to the 280

312 deactivation of the sorbent due to sintering. However, in the post-breakthrough region (where the sorbent is completely exhausted), for the first three cycles, H 2 purity is almost constant. This is because the reforming and water gas shift reactions occur in the presence of the nickel catalyst alone and the H 2 production is not enhanced by insitu CO 2 capture. Figure 7.13 (b) illustrates the effect of cycling on CH 4 concentration in the product gas stream. During the first three cycles, the time of the prebreakthrough region for CH 4 also decreases with the increase in cycle number. However, in the 4 th cycle, it can be observed that the conversion of CH 4 and the purity of H 2 are higher than cycles 2 and 3, and high purity H 2 is produced for about 470 seconds. This longer duration of the pre-breakthrough region can be attributed to the reactivation of the sorbent due to hydration, which improves the CO 2 capture capacity of the sorbent. Figure 7.14 illustrates the purity of H 2 produced in 3 cycles of steam CH 4 reforming in the presence of CaO sorbent which is obtained by calcination of the spent sorbent from the previous cycle followed by hydration. The calcination of the sorbent was conducted in nitrogen at 950 ºC and the hydration was conducted for every cycle at 600 ºC and 11 atms. Although there is still a decrease in the purity of H 2 produced during the three cycles, the decrease is small and lower than that observed in the first three cycles of both figures 7.12(a) and 7.13(a). High purity H 2 was produced for 620, 570 and

313 seconds in cycles 1, 2 and 3 respectively in Figure Thus, hydration helps to reduce the extent of sintering and arrest the rapid decline in sorbent activity Realistic Sorbent Calcination in a Steam/CO 2 Atmosphere with Sorbent Hydration It is observed in Figure 7.12 (a) that the pre-breakthrough region decreases from 234 sec to 0 sec due to calcination in the presence of CO 2 and steam at 950 ºC during cycle 4. The effect of hydration on the purity of H 2 produced and the extent of CH 4 converted during 4 cycles in which the calcination was conducted in the presence of a 50:50 H 2 O/ CO 2 atmosphere is shown in Figures 7.15 (a) and (b). The sorbent for all 4 cycles was obtained by calcination in a steam and CO 2 mixture at 950 ºC followed by hydration at 600 ºC and 11 atms. In Figure 7.15(a), although the sorbent was calcined in CO 2 and steam, a complete loss in sorbent reactivity is not observed as hydration was conducted every cycle. Although an initial decrease in H 2 purity is observed between the first and the second cycles, the purity is maintained at almost a constant value in the subsequent cycles. A similar observation is made for the CH 4 concentration in the H 2 product stream in Figure 7.15(b). From the multicyclic investigation conducted for different calcination conditions, it is evident that the reactivation of the sorbent by hydration aids in reducing the extent of sorbent sintering. 282

314 7.5 APPLICATIONS OF CLP IN HYDROCARBON REFORMING The CLP enhanced steam reforming of hydrocarbons can be applied to the production of H 2 from natural gas and other hydrocarbon feedstock. It can also be used for the production of electricity in a carbon constrained scenario from hydrocarbons. Another important area for its application is in the production of synfuels with carbon capture through the indirect coal conversion process comprising coal gasification and the Fisher Tropsch (F-T) Process. Further description of the integration of the CLP enhance reforming process in natural gas conversion and liquid fuel production is described in the following sections Steam Reforming of Natural Gas and Other Hydrocarbons for H 2 and Electricity Generation: Figure 7.16 shows the integration of the CLP in a natural gas reforming process in which the unit operations namely, reforming, water gas shift, CO 2 capture and sulfur removal are integrated in a single reactor system. Within the H 2 production reactor, the natural gas is reformed with steam in the presence of the reforming catalyst and CaO sorbent. The removal of CO 2 removes the thermodynamic limitation of the water gas shift and the reforming reaction and results in a high conversion of the CH 4 to H 2. The H 2 production reactor is almost heat neutral due to the exothermic energy from the water gas shift and carbonation reactions being equal to the endothermic reforming reaction heat duty. Hence the temperature of operation for the reforming reaction can 283

315 be reduced from over 900 ºC to 650 ºC resulting in cost savings for the reactor material. The spent sorbent containing CaCO 3 and CaO is separated from the H 2 and regenerated in a calciner at 900 ºC to produce a sequestration ready CO 2 stream. The CaO sorbent is then recycled back to the integrated H 2 production reactor. To improve the reactivity of the sorbent, a sorbent hydration reactor may be added downstream of the calciner and a part or all of the calcined sorbent may be hydrated before it is fed back into the H 2 production reactor Technical analysis of the natural gas to H 2 process using the CLP A preliminary process analysis has been conducted for the production of H 2 and electricity from natural gas by the CLP as shown in Figure Water, pressurized to 15 atm, is converted to high temperature steam at 650 ºC and fed to the H 2 reactor along with preheated natural gas. The CaO sorbent at 900 ºC from the calciner is also sent to the H 2 production reactor. Although the combination of the reforming, water gas shift and carbonation reactions is almost heat neutral, the heat given out by the solids that cool from 900 ºC to 650 ºC in the H 2 production reactor makes it slightly exothermic. The product gases are separated from the spent sorbent in a cyclone and sent to a Pressure Swing Absorber (PSA) for the production of high purity H 2. The final high pressure H 2 product at the exit of the PSA is cooled from 650 ºC to ambient temperature. The tail gas from the PSA is preheated from 650 to 900 ºC and fed into the calciner for combustion, to provide the energy required for the calcination reaction. The tail gas is combusted in oxygen, preheated to 900 ºC, to produce a concentrated 284

316 CO 2 stream for sequestration. Spent sorbent is completely calcined in the calciner. The hot solids at 900 ºC are conveyed to the H 2 production reactor while the CO 2 is cooled from 900 ºC to 25 ºC and compressed to 150 atms for transportation and sequestration. The results from the mass balance are shown in Table Basis for process analysis: 1) Sulfur free natural gas from the pipeline at 41 atms containing 90% CH 4, 5% ethane and 5% nitrogen with a Higher Heating Value of MJ/Kg was used for the analysis. ( DOE, 2002) 2) All the reactions were assumed to proceed to thermodynamic equilibrium. In the H 2 production reactor, the extent of various reactions used for conducting the mass and energy balance are shown in Table 7.1. The calcination reaction proceeds to completion at a temperature of 900 ºC and hence all the CaCO 3 is converted to CaO in the calciner. 3) A S:C ratio of 3 and a Ca:C ratio of 1.5 is used for the analysis. 4) For this preliminary study, heat loses in all the equipment were assumed to be minimal. 5) The turbine isentropic efficiency is 72%. 285

317 6) A five stage compression was assumed for CO 2 compression with an isentropic efficiency of 85%. The energy required for CO 2 compression was found to be 105 KWh/tonne of CO 2. 7) MWth is used to quantify heat energy and MWe to quantify electrical energy. An efficiency of 40% was applied for the conversion of heat energy to electrical energy. 8) The reforming catalyst is mostly retained in the H 2 production reactor which is a fixed fluidized bed. Hence only the calcium sorbent is transported between the carbonator and the calciner in the process Results from the process analysis The results for the energy balance are shown in Table 7.3. The cold gas efficiency of the CPL, defined as the ratio of the Higher Heating Value (HHV) of H 2 produced to the Higher Heating Value (HHV) of natural gas, is 84%. In addition to the production of H 2, the process also produces 47.7 MWe of power after accounting for all the parasitic energy required within the plant. The detailed explanation for the energy required/produced from each of the unit operations in the process, listed in Table 7.3 is provided below. Table 7.4 illustrates the energy required for the production of steam for the reforming and water gas shift reaction in the H 2 production reactor. Since a total of kmoles/hr of carbon is fed to the H 2 production reactor and a S:C ratio of 3 is 286

318 used, kmoles of steam is produced in the steam generator at 15 atms. The Cp and latent heat values were determined based on the fact that water boils at 199C at 15 atms. Table 7.5 illustrates the heat required for preheating the natural gas from -27C to 650 ºC. The CaO sorbent at the exit of the calciner is at 900 ºC and hence it releases heat in the H 2 production reactor which is operated at 650 ºC. The average Cp of the calcium sorbent over the temperature range of 650 to 900 ºC was determined and used to calculate the heat released as shown in Table 7.6. Table 7.7 shows the heat balance within the H 2 production reactor. The heat required for the endothermic reforming reaction is provided by the heat released from the exothermic water gas shift and carbonation reaction and the hot solids from the calciner. The amount of heat released by the solids is calculated in Table 7.6. The H 2 produced from the CLP is at a temperature of 650 ºC and is cooled to ambient temperatures for transportation. The heat released is calculated from Table 7.8 to be MWth. The tail gas from the PSA which is at 650 ºC is combusted in the calciner to provide heat for the calcination reaction. Since the calciner operates at 900 ºC, the tail gas is preheated to 900 ºC before being fed to the calciner and the details for the heat required is shown in Table

319 The oxygen for the combustion of the tail gas in the calciner also needs to be preheated to 900 ºC and the energy required for preheating is shown to be MWth from Table The spent calcium sorbent consisting of CaO and CaCO 3 at the exit of the H 2 production reactor is at a temperature of 650 ºC and hence absorbs heat from the combustion of tail gas in the calciner to heat up to 900 ºC. This energy has been calculated in Table 7.11 to be MWth. The tail gas is combusted in the calciner with oxygen to produce heat which is partly used for the endothermic calcination reaction. The remaining heat is used to produce electricity as shown in Table The hot CO 2 at 900 ºC from the calciner is cooled down to 25 ºC and produces 280 MWth as shown in Table The steam in the CO 2 stream is condensed out at 100 ºC and the dry CO 2 is compressed for sequestration Implementation of Carbon Capture in Liquid Fuels Production From Coal: Crude oil satisfies majority of the transportation-based energy needs of the United States and 60% of the crude oil requirement is achieved through imports. However, with the increasing fluctuation in the price of crude and the desire to achieve energy independence, there is a renewed focus on alternative technologies to satisfy the rising demand of energy. Coal-to-Liquids (CTL) through the Fischer-Tropsch (FT) process is one such promising technology which enables the production of high quality 288

320 and cleaner liquid fuels from the abundantly present fossil fuel coal. The estimated carbon footprint of a CTL plant is % higher than a petroleum-based plant. Implementation of CCS in a CTL plant can help in achieving 20% lower life cycle CO 2 emissions compared to petroleum based fuel. The CLP is capable of producing a sequestration ready CO 2 stream by capturing all the CO 2 emitted during the CTL process. In addition to achieving carbon capture, the CLP improves the efficiency of the CTL process by conversion of the Fischer Tropsch reactor s off gases to H 2. This H 2 is used to adjust the H 2 :CO ratio, making it suitable for the FT reaction as well as for the product upgrader. The CTL process can be broken down into two main blocks which consist of the gasifier block and the Fischer Trospch synthesis block. The gasifier block consists of the gasifier and other unit operations like particulate and heavy metal removal. Similarly the FT synthesis block consists of the FT reactor followed by the product separation unit, hydrocracking and hydrotreating unit, etc. Extensive studies have been conducted on the gasifier block leading to several demonstration and pilot plant studies and improvements are being made to their design and operation currently. A few of the coal gasification projects for electricity, H 2, or liquid fuels production are the 313 MW Tampa Electric s IGCC Plant in Florida, USA; the 292 MW Wabash River Gasification Repowering IGCC Project in Indiana, USA; the 253 MW Nuon Buggenum IGCC Power Plant in Buggenum, the Netherlands; the Shenhua Group Corp s 70,000+ barrel liquid fuel/day CTL project which is under construction, China. 289

321 Similarly the Fischer Tropsch synthesis block has been commercially operated since World War 2. 9 CTL plants were set up in the Germany at the end of world war II which produced 4 MMT/year of liquid fuels. In South Africa, Sasol is operating 150,000 BPD CTL plants and currently China is working with Sasol in building 2 plants which are estimated to produced 30 MMtons if liquid fuel. Although liquid fuels are being produced commercially, no large scale demonstration exists with advanced technology. The route for the conversion of coal to the FT feed consists of various stages and consumes excessive parasitic energy. The CLP could reduce the parasitic energy consumption by achieving contaminant capture at high temperature and by integrating various operations like reforming, water gas shift, CO 2 and sulfur removal in a single stage reactor system Description of the Processes Currently, the production of coal derived liquid fuels is though the coal gasification and the Fisher Tropsch (F-T) process illustrated in Figure A conventional CTL plant consists of a gasifier which produces the syngas. The H 2 to carbon monoxide (H 2 /CO) ratio of the syngas is around 0.63, which is much lower than the optimal ratio of ~2, required for liquid fuel production. Hence, in order to modify the amount of H 2 in the syngas, part of the syngas is introduced to a water gas shift reactor to be shifted to H 2. Since the gas stream contains sulfur impurities, a sulfur tolerant water gas shift catalyst is used. The rest of the syngas stream passes through a hydrolysis unit where the COS is converted into H 2 S. 290

322 The gas streams from the water gas shift reactor and the hydrolysis reactor are mixed together and passed through several gas cleanup units that consist of a mercury removal bed, bulk sulfur removal units, sulfur polishing unit, and CO 2 removal units. After the pollutants are removed, a portion of the syngas is sent to a PSA to separate H 2 for use in the product upgrader. The bulk of the clean syngas stream with a H 2 /CO ratio of around 2 is sent to the F-T reactor for the production of liquid fuel. The F-T reactor is capable of converting more than 70% syngas into a wide range of hydrocarbons ranging from CH 4 to wax. The products from the F-T reactor are sent to a product upgrader where the high molecular weight hydrocarbons are refined into liquid fuel or naphtha while the low molecular weight offgas stream is sent to a power generation block to generate electricity for the ASU and other parasitic energy consumption ( Mayer, 2005, Choi et al, 1997). The addition of CO 2 capture to the CTL process will add units to reform the C1-C4 hydrocarbons present in the offgases from the F-T reactor, water gas shift reactors to shift the CO to H 2, and CO 2 capture units, making the over all process very energy intensive. The addition of CO 2 capture to a CTL plant can be simplified by the CLP. The CLP can be integrated in a CTL plant in two configurations as shown in Figure In configuration 1, the CLP is placed downstream of the F-T reactor and converts the offgases from the F-T reactor to H 2. The F-T reactor offgas contains a mixture of C1-C4 hydrocarbons and unconverted syngas. The CLP integrates the reforming of hydrocarbons and the conversion of unconverted syngas to H 2 with the 291

323 capture of CO 2 in a single reactor leading to the production of a pure H 2 stream. The H 2 is separated from the spent sorbent. A part of the H 2 can be added to the syngas feed entering the F-T reactor to improve the H 2 /CO ratio of the F-T reactor feed. A part of the H 2 can also be fed to the product upgrader to refine the liquid fuel product. Figure 7.21 illustrates a detailed schematic of this proposed process. In configuration 2, the CLP is placed upstream of the F-T reactor and the feed to the carbonation reactor consists of the syngas from the gasifier and the offgas from the F-T reactor. The CLP achieves the following objectives: a) Converts the C1-C4 hydrocarbons and unconverted syngas from the FT process, and syngas from the gasifier, into a 2:1 H 2 :CO stream by shifting the equilibrium of the water gas shift and reforming reaction in the forward direction by removing the CO 2 product insitu, b) Achieves simultaneous CO 2 and H 2 S capture at high temperatures, c) Produces a sequestration ready CO 2 stream in the calcination stage removal d) Reduces the excess steam requirement which aids in higher levels of H 2 S As shown in Figure 7.21 the unreacted syngas and light hydrocarbons from the FT reactor are mixed with the syngas from the gasifier and sent into the single reactor system which adjusts the ratio of the H 2 :CO in the syngas stream by reforming the 292

324 hydrocarbons and shifting the syngas in the presence of CaO. The concomitant carbonation of the metal oxide leading to the formation of the metal carbonate incessantly drives the equilibrium-limited water gas shift and the reforming reaction forward by removing the CO 2 product from the gas mixture. The metal carbonate can then be regenerated by heating, to give back the metal oxide and a pure CO 2 stream. By improving the equilibrium conversion of the reforming and water gas shift reaction, steam addition can be greatly reduced. The reduction in steam consumption not only reduces energy consumption but also aids in the removal of H 2 S to ppb levels by the CaO)as steam poses an equilibrium constrain to the removal of H 2 S. Various reactions occurring in this system are as follows Reforming: C x H y + xh 2 O xco + ( y / 2 +x) H 2 (7.1) Water gas shift: CO + H 2 O H 2 + CO 2 (7.3) Carbonation: CaO + CO 2 CaCO 3 (7.4) Sulfidation: CaO + H 2 S CaS + H 2 O (7.12) Calcination: CaCO 3 CaO + CO 2 (7.13) 7.6 CONCLUSIONS Single cycle tests have shown that the conversion of CH 4 is improved to a large extent by the addition of CaO sorbent at 650 ºC. High purity H 2 is obtained at low S:C ratios of 3:1 for various pressures ranging from 1 11 atms. The purity of H 2 was found to be higher at 650 ºC than at 700 ºC due to the favorable thermodynamics of the 293

325 carbonation of CaO. Although the conversion of CH 4 in the conventional steam methane reforming process decreases with the increase in pressure, the removal of CaO during steam methane reforming reduces this effect and results in almost a constant amount of CH 4 conversion with the increase in pressure. The effect of calcination conditions on the extent of CH 4 reforming was determined. The reactivity of the sorbent is found to decrease over multiple cycles due to calcination in both pure nitrogen and in a mixture of steam and CO 2. This reduction in reactivity of the sorbent results in a decrease in both CH 4 conversion and H 2 purity. In order to improve the recyclability of the sorbent over multiple cycles, sorbent reactivation step by hydration is found to be effective. By the addition of the hydration step during every cycle, the extent of sorbent sintering is reduced. This reactivation will aid in the production of a constant purity of H 2 during the steam reforming of CH 4. The CLP for reforming of hydrocarbons can be directly applied to the production of H 2 and electricity from natural gas. This integration would benefit the steam reforming process by providing a method of internal heat integration. In the CLP, although the endothermic calciner is operated at a high temperatures( ºC) similar to the steam methane reforming it is at atmospheric pressure while the heat neutral reformer is at a high pressure and a relatively low temperature of 650 ºC. The energy required for the reformer in the conventional process is supplied to the calciner in the CLP and an additional benefit of producing a sequestration ready CO 2 stream is obtained by integrating steam reforming with the CLP. The CLP can also be integrated in a CLT plant for the conversion of F-T offgases to H 2 with CO 2 capture. 294

326 Reaction Extent of 650 ºC and 15 atms Methane Reforming CH 4 + H 2 O = CO + 3H % Ethane Reforming C 2 H 6 + 2H 2 O = 2CO + 5H % Water Gas Shift Reaction CO +H 2 O = CO 2 + H % Carbonation Reaction CaO + CO 2 = CaCO % Table 7.1: Thermodynamic extent of the various reactions occurring in the carbonator 295

327 296 Stream T ºC P bar Mass Flow kg/hr CH C 2 H N H 2 O CO H CO O CaCO CaO Mole Flow kmol/hr CH C 2 H N H 2 O CO H CO O CaCO CaO Continued Table 7.2: Stream data for the integration of the CLP in a steam methane reforming process 296

328 297 Table 7.2 continued Stream T ºC P bar Mass Flow kg/hr CH C 2 H N H 2 O CO H CO O CaCO CaO Mole Flow kmol/hr CH C 2 H N H 2 O CO H CO O CaCO CaO

329 Mass Flow Heat Power Kg/hr MWth MWe Inputs Natural Gas feed rate Steam boiler duty NG heater duty O 2 preheater duty Tail gas preheater duty Pumps 0.17 Net energy Input Outputs H 2 production rate H 2 Reactor heat duty Calciner heat duty Hot H 2 cooler duty Hot CO 2 cooler duty NG turbine 2.45 Net energy out Cold gas efficiency (HHV) 0.84 Total Net thermal heat available Net Power available 2.28 CO 2 Compression Export power from plant Table 7.3: Energy balance for the production of H 2 and electricity from natural gas using the CLP. 298

330 Steam Boiler Duty Mole flow T IN T out Cp (l) Cp (v) Latent heat Q (MWth) kmol/hr (C) (C) kj/kg-k kj/kg- K kj/kg For H 2 O Cp 199- Values C 650 ºC at 199C Table 7.4: Heat required for the production of steam at 650 ºC from water at 15 atms 299

331 Mass flow Mole flow T IN T out Energy Reqd. Q (MWth) kg/hr kmol/hr (C) (C) kj/kmole Natural Gas Preheating CH C 2 H N Total Table 7.5: Heat required for preheating the natural gas at 15 atms to 650 ºC 300

332 Mass flow Mole flow T IN T out Cp Q (MWth) kg/hr kmol/hr (C) (C) kj/kg-k Solid from calciner CaO Table 7.6: Heat released by the solids from the calciner in the H 2 production reactor. 301

333 Heat of Moles Heat Reaction Reacted Duty KJ/Mole kmoles/hr MWth Heat duty Hydrogen Production Reactor Reforming of CH Endothermic Reforming of C 2 H Endothermic Water gas shift reaction Exothermic Carbonation Exothermic Sensible heat from hot CaO solids From Table Exothermic Net Reactor heat duty Exothermic Table 7.7: Heat generated from the H 2 production reactor 302

334 Mass flow Mole flow T IN T out Energy Reqd. Q (MWth) kg/hr kmol/hr (C) (C) kj/kmole H 2 outlet cooler Table 7.8: Heat released on cooled the H 2 from 650 ºC to ambient temperature 303

335 Mass flow Mole flow T IN T out Energy Reqd. Q (MWth) kg/hr kmol/hr (C) (C) KJ/kmole Tail Gas Preheating CH C 2 H N H 2 O CO H CO Total Table 7.9: Heat required for preheating the PSA tail gas from 650 ºC to 900 ºC 304

336 Mass flow Mole flow T IN T out Energy Reqd. Q (MWth) kg/hr kmol/hr (C) (C) kj/kmole Oxygen Preheating Table 7.10: Heat required for preheating the oxygen from ambient temperature to 900 ºC 305

337 Mass flow Mole flow T IN T out Cp Q (MWth) kg/hr kmol/hr (C) (C) kj/kg-k Solids to calciner CaCO CaO Total Table 7.11: Heat absorbed by the solids from the H 2 production reactor in the calciner 306

338 Heat of Moles Heat Reaction Reacted Duty KJ/Mole kmoles/hr MWth Heat duty Calcination Reactor Calcination of CaCO Endothermic Heat needed to heat the solids From Table Endothermic Tail gas combustion Combustion of CH Exothermic Combustion of C 2 H ~0.00 Exothermic Combustion of H Exothermic Combustion of CO Exothermic Net Reactor heat duty Exothermic Table 7.12: Heat released from the calciner 307

339 Mass flow Mole flow T IN T out Cp (l) Cp (v) Latent ht Q (MWth) kg/hr kmol/hr (C) (C) kj/kg-k kj/kg-k kj/kg CO 2 cooler For H2O Cp Values ºC ºC at 100 ºC N H 2 O CO O Total Table 7.13: Heat released from cooling the CO 2 from 900 ºC to ambient temperature 308

340 Reaction Sorbent Purge Sorbent Makeup Regeneration H 2 Integrated reactor Net Heat Output Hydrocarbon Feed Dehydration : Ca(OH) 2 CaO + H 2 O Reforming : C x H y +H 2 O CO + H 2 WGSR : CO + H 2 O CO 2 + H 2 CO 2 removal : CaO + CO2 CaCO3 Sulfur : CaO + H2S CaS + H 2O Halide : CaO + 2HX CaX 2 + H 2 O Pure CO 2 gas Heat Calciner Input Calcination: CaCO3 CaO + CO2 Reactivation Hydrator Heat Output H 2 O Hydration : CaO + H 2 O Ca(OH) 2 Figure 7.1: Schematic of the CLP for the conversion of hydrocarbons to H 2 309

341 K, Equilibrium Constant Reforming Reforming + WGS Reforming + WGS + Carbonation Temperature (C) Figure 7.2: Thermodynamic data illustrating the equilibrium constants of the steam reforming of CH 4, water gas shift and carbonation reaction 310

342 Thermocouple And Pressure Guage Steam & Gas Mixture Steam Generator Water In Gas Gas Mixture Mixture Catalyst Powder + Sorbent MFC MFC MFC MFC Hydrocarbon Analyzer Back Pressure Regulator Heated Steel Tube Reactor Water Syringe Pump N 2 CH 4 C 2 H 6 C 3 H 8 Analyzers (CO, CO 2, H 2, H 2 S) Heat Exchanger Water Trap Figure 7.3: Simplified schematic of the bench scale experimental setup 311

343 90 S:C = 3 S:C = 5 85 H 2 Purity (%) Temperature (C) (a) Gas Composition (%) CH4 (3:1) CH4 (5:1) CO (3:1) CO (5:1) CO2 (3:1) CO2 (5:1) Temperature (C) (b) Figure 7.4: Effect of temperature and S:C ratio on (a)h 2 purity and (b) the amount of CO, CO 2 and CH 4 remaining in the product gas for the steam methane reforming reaction in the presence of Ni-based catalyst ( P = 1 atm) 312

344 100 H 2 Gas composition (%) CH 4 CO CO Time (sec) Figure 7.5: Breakthrough curve in the composition of the product gases obtained during the simultaneous reforming, water gas shift and carbonation reaction. (T = 650 ºC, P = 1 atm) 313

345 Methane Conversion Time (sec) Figure 7.6: CH 4 conversion obtained during the simultaneous reforming, water gas shift and carbonation reaction. (T = 650 ºC, P = 1 atm) 314

346 (a) Figure 7.7: Effect of temperature and S:C ratio on (a) H 2 purity (b) conversion of CH 4 (P = 1atm) (b) 315

347 (a) (b) Figure 7.8: Effect of temperature and S:C ratio on the amount of (a) CO and (b) CO 2 remaining in the product gas for H 2 production from methane with/without sorbent. ( P = 1 atm) 316

348 H 2 Purity (%) atm 3 atm 4.5 atm 11 atm Time (sec) (a) CH 4 in the Product Stream (%) atm 3 atm 4.5 atm 11 atm Time (sec) (b) Figure 7.9: Effect of pressure on (a) H 2 purity and (b) CH 4 concentration in the product stream. (T = 650 ºC, S:C ratio = 3) 317

349 CO 2 in the Product Stream (%) atm 3 atm 4.5 atm Time (sec) (a) CO in the Product Stream (%) atm 3 atm 4.5 atm 11 atm Time (sec) (b) Figure 7.10: Effect of pressure on (a) CO 2 and (b) CO concentration in the product stream. (T = 650 ºC, S:C ratio = 3) 318

350 Gas Composition in the Product Stream (%) CH4-Post CH4 CO-Post CO CH4-Pre CO2 CO2-Post CO-Pre CO2-Pre Pressure (atms) Figure 7.11: Effect of pressure on the pre-breakthrough and post-breakthrough concentration of CH 4, CO and CO 2 in the product stream. (T = 650 ºC, S:C ratio = 3) 319

351 Cycle 1 Cycle 2 Cycle 3 Cycle 4 H 2 Purity (%) Time (sec) (a) CH 4 in the Product Stream (%) Cycle 1 Cycle 2 Cycle 3 Cycle Time (sec) (b) Figure 7.12: Effect of calcination conditions on (a) H 2 purity and (b) CH 4 composition in the product gas for cycles 1,2,3 and 4. [(Reforming reaction conditions :T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2 and 3 are calcined in pure N 2 at 950 and the sorbent for cycle 4 is calcined in a 50:50 CO 2 /H 2 O atmosphere at 950 ºC.)] 320

352 H 2 Purity (%) Cycle 1 Cycle 2 Cycle 3 Cycle Time (sec) (a) Continued Figure 7.13: Effect of hydration on (a) H 2 purity and (b) CH 4 composition in the product gas for cycles 1, 2, 3 and 4. [(Reforming reaction conditions :T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in pure N 2, T = 950, P = 1 atm)(hydration conditions: hydration of calcined sorbent from the 3 rd cycle in a 80:20 H 2 O/N 2 atmosphere, T = 600, P = 11 atm)] 321

353 Table 7.13 continued 8 CH 4 in the Product Stream (%) Time (sec) (b) Cycle 1 Cycle 2 Cycle 3 Cycle 4 322

354 Cycle 1 Cycle 2 Cycle 3 H 2 Purity (%) Time (sec) Figure 7.14: Effect of hydration on H 2 purity for cycles 1,2,3 and 4. [(Reforming reaction Conditions: T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in pure N 2, T = 950, P = 1 atm)(hydration conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20 H 2 O/N 2 atmosphere, T = 600, P = 11 atm)] 323

355 H 2 Purity (%) Cycle 1 Cycle 2 Cycle 3 Cycle Time (sec) (a) Continued Figure 7.15: Effect of hydration on (a) H 2 purity and (b) CH 4 content in the product gas for cycles 1,2,3 and 4. [(Reforming reaction Conditions: T= 650 ºC, P = 1 atm, S:C ratio = 3) (Calcination conditions: sorbent for cycles 1, 2, 3 and 4 are calcined in a 50:50 CO 2 /H 2 O atmosphere, T = 950, P = 1 atm) (Hydration conditions: hydration for cycles 1, 2, 3 and 4 in a 80:20 H 2 O/N 2 atmosphere, T = 600, P = 11 atm)] 324

356 Figure 7.15 continued CH 4 in the Product Stream (%) 14 Cycle 1 Cycle 2 12 Cycle 3 Cycle Time (sec) (b) 325

357 Natural Gas Steam Integrated Reforming, WGSR, CO 2 and Sulfur Capture CaO Calciner Electricity Gas Solid Hydrogen Separation Chemicals and Liquid Fuels Spent Sorbent CO 2 CO 2 Compression and Sequestration Figure 7.16: Integration of the CLP in a natural gas reforming system 326

358 0.17 MWe MWth Water Pump MWth Steam Generator MWth Pure MWth PSA 2.45 MWe Hydrogen Hydrogen MWth Natural Gas Production Cyclone H 2 Cooling from Pipeline Reactor MWth NG Preheating MWth 12 Tailgas Preheating 28.4 MWth MWth Cyclone Calciner Oxygen MWe O 2 Preheating MWth CO 2 Compression CO 2 Cooling CO 2 to Sequestration Figure 7.17: Detailed schematic for H 2 production from natural gas 327

359 328 Figure 7.18: Conventional CTL plant 328

360 Conventional Flue Gas COS Hydrolysis Sulfur and CO 2 Reformer Gas Turbine Gases Selexol FT Reactor Separator /Upgrader Fuel WGSR H 2 Recovery H 2 CLP -1 H 2 CLP H 2 Gasifier Sulfur Removal FT Reactor Gases Separator /Upgrader Fuel CLP -2 CLP FT Reactor Gases Separator /Upgrader Fuel H 2 Recovery H 2 Figure 7.19: Integration of the CLP in a CTL plant in two configurations 329

361 Steam Coal Pretreatment Raw Syngas H2/CO = 0.5 Mercury Removal Hg Fly Sulfur Ash Byproduct Calciner HC Reforming + WGSR +CO 2 removal Reactivator Steam BFW High Pressure O 2 CO Gasifier C1-C4 Generator Slag Air Steam Unconverted Syngas N 2 N 2 Product Separation C5 C14 Separator II Gasoline JP-8 Fuel Water Warm Clean Syngas H 2 /CO = 2:1 F-T Reactor Hydrocraking Hydrotreatment C15 and above H 2 Figure 7.20: Integration of the CLP in a CTL plant configuration 1 330

362 Steam Raw Syngas H2/CO = 0.5 Calciner High Pressure CO 2 To Sequestration One Step Process WGSR HC Reforming CO 2 Sulfur Coal Pretreatment BFW WGSR+ HC Reforming + sulfur CaCO 3 to Calciner Candle Filter Fly Ash Steam O Gasifier Slag Air Steam C1-C4 Unconverted Syngas Generator N 2 N 2 Product Separation C5 C14 Separator II Gasoline JP-8 Fuel Water Warm Clean Syngas H 2 /CO = 2:1 F-T Reactor Hydrocraking Hydrotreatment C15 and above H 2 Figure 7.21: Integration of the CLP in a CTL plant configuration 2 331

363 CHAPTER 8 SUBPILOT SCALE TESTING AND RECOMMENDATIONS FOR FUTURE WORK 8.1 INTRODUCTION The CLP has been studied in the lab and bench scale for H 2 production from syngas and hydrocarbons. The process is being scaled up to a 25 KWth subpilot unit demonstration at the Ohio State University. The design for the subpilot scale unit is based on the thermodynamic, kinetic and sorbent reactivity studies detailed in the previous chapters. Based on the design, cold flow models were constructed and sorbent flow testing was conducted. The subpilot scale unit consists of the gas delivery system and the gas regulating panel, steam generation system, sorbent feeding system, H 2 production or carbonation reactor, gas cooling and particle capture system and gas analysis. The unit is designed to operate at a temperature of up to 900 ºC and pressure of 4.5 atms. In this chapter, the results from cold flow testing and the design of the subpilot scale unit, currently under construction are discussed. 332

364 8.2 COLD FLOW TESTING In the CLP, the carbonation reactor, calciner and hydrator are fluidized or entrained bed reactors and hence an understanding of the flow characteristics of the sorbent through these reactors is very critical to the process. Since CaO and Ca(OH) 2 are cohesive particles that are difficult to fluidize, cold flow testing was conducted before construction of the subpilot scale unit. As discussed in Chapter 3, hydration and the subsequent dehydration of the sorbent has been found to produce CaO with a particle size of 2-20 micron(d50). Hence the active sorbent that captures CO 2, sulfur and halides in the carbonation reactor has a D50 of 2-20 microns. A cold model with a diameter of 4 inches was constructed for the carbonation reactor with a sorbent feeding system, flow controllers for air and a particulate filter at the exit of the reactor. Air was used as the fluidization gas. Ca(OH) 2 sorbent with a D50 of 2-20 micron was used as the active sorbent. Since Ca(OH) 2 is difficult to fluidize, it was mixed with a fluidizing aid, ground lime which has a D50 of 600 microns. Other powders like sand and other metal oxides could also be used as fluidizing aids. The gas velocity in the carbonation reactor was designed such that the active Ca(OH) 2 sorbent is entrained while the fluidizing aid is in the turbulent regime. Hence the gas velocity in the carbonation reaction should be above the entrainment velocity of the Ca(OH) 2 particles and between the critical and entrainment velocities of the fluidizing aid. 333

365 Cold model tests were conducted to determine the critical velocity for the fluidizing aid particles in the 4 inch cold model of the carbonation reactor. Critical velocity is the minimum velocity for turbulent fluidization and is determined from the standard deviation of pressure change across the bed. The standard deviation of pressure change is a maximum at the critical velocity of the particles. Pressure taps were placed along with 2 micron inline filter along the height of the column and were connected to a U-tube manometer to measure the pressure change. The manometer was connected to a transducer and the voltage signal generated by the transducer was recorded using a data logging system. The collapsed bed height of the sorbent was ~10 cms. The air flowrate was varied between scfh and the change in pressure drop across the bed was determined at each flowrate. Figure 8.1 shows the standard deviation of pressure change(std(pa)) for different gas velocities(ug). The standard deviation is a maximum at velocities greater than 1.3 m/s and hence turbulent fluidization of the fluidizing aid would occur above a velocity of 1.3 m/s. Since the active Ca(OH) 2 particles are entrained at velocities above 1.3 m/s, the carbonation reactor can be operated at a velocity of 1.3m/s and greater. After determining the suitable gas velocity for the carbonation reactor, a continuous fluidization test was conducted with a mixture of 10% fluidizing aid (ground lime power with a D50 of 500 to 600 microns) and 90% active sorbent (Ca(OH) 2 with a D50 of 2-20 microns). The mixture of sorbent was fed to the carbonation reactor from a sorbent storage hopper by a screw feeder. The air flow rate 334

366 was set at 1400 SCFH. When the sorbent was conveyed at a flow rate of 15 g/min, ~ 80% of the total sorbent that was conveyed into the reactor was entrained at the end of 2 hours. At a higher sorbent flowrate of 60g/min, ~70% of the sorbent was entrained. From these test it can be seen that not all the active Ca(OH) 2 sorbent was entrained by the gas. At the end of the tests, agglomerates were observed in the bed material which might have been formed by the cohesive Ca(OH) 2 sorbent at room temperature. The moisture in the fluidizing air might also cause the aggregates to form in the sorbent mixture. Hence fluidization test need to be conducted at the reaction temperature of 600 ºC in the subpilot scale with air prior to testing the system with the simulated syngas stream. A cold model of the entire CLP with the carbonation reactor, calciner and hydrator was also constructed as shown in Figure 8.2(a). Figure 8.2(b) is a picture of the cold model unit. The carbonation reactor and hydrator were fluidized beds while the calciner was a rotary bed. As shown in Figure 8.2(a), the cold flow model consists of a riser (H 2 production or carbonation reactor), cyclones, a rotary calciner, a U-valve and a hydrator. Air is introduced into the riser to carry the sorbent through the riser reactor. At the top of the riser, two cyclones separate the solids from the air. The recovered solids enter the rotary calciner via a U-valve. At the outlet of the calciner, the sorbent enters the hydrator. Cold flow tests with air and a mixture of active Ca(OH) 2 sorbent and fluidization aid have shown that the sorbent can be fluidized and made to flow smoothly in the continuous reactor system. Figure 8.3 is a picture of the 335

367 cold model for the hydrator. Air was used instead of steam to fluidize a mixture of CaO and Ca(OH) 2. The sorbent mixture fluidizes well with minimum gas channeling and no slugging was observed. 8.3 DESIGN OF THE SUBPILOT SCALE UNIT The 25 KWth sub-pilot scale reactor system being constructed at OSU is shown in Figure 8.4. The calcium sorbent will be continuously fed using a sorbent hopper and a motor-driven screw feeder. This sorbent is entrained using the reactant gas mixture entering from the bottom. The product gas is analyzed using a micro GC. The subpilot scale unit consists of the gas delivery system and the gas regulating panel, steam generation system, sorbent feeding system, H 2 production or carbonation reactor, gas cooling and particle capture system and gas analysis. The gas delivered to the reactor is a simulated syngas mixture. Gas cylinders are used as the source of H 2, CO, CO 2 and N 2 (carrier gas). To ensure a continuous supply of gas to the reactor during the test a change over system is installed for each gas. 2 cylinders are connected to the changeover system and when one cylinder becomes empty, the changeover system automatically switches over to the other cylinder. An audible alarm is used to alert the operator to replace the empty cylinder while the other cylinder is in use. In this manner cylinders can be changed even when the system is in operation. The switch over valves and the cylinders are placed in gas cabinets that are provided with adequate ventilation. Gas leak alarms are also placed in 336

368 the gas delivery area for H 2 and CO 2. Safety features like flash arrestors, check valves to prevent back flow of the gases, automatic shutoff valves to stop gas flow in the event of a leak, pressure relief valve to prevent an increase in pressure and flow limit shut off valves to prevent an increase in the gas flow rates beyond the preset maximum value are included in the gas delivery system. The gases from the gas delivery room are sent to a gas mixing panel consisting of digital mass flow controllers. Safety features like flash arrestors, check valves, maximum flow limit shut-off valves and pneumatically controlled diaphragm valves have been incorporated into the gas delivery system. The mass flow controllers and the diaphragm valves are controlled using a computer interface. A nitrogen line is provided to flush the system and a line is provided from the gas panel to the vent for safety. In addition to these safety features, the gas panel is connected to safety features in the other sections of the subpilot scale unit. The flow of gas through the gas panel will be automatically shut off if the temperature exceeds a preset value in the reactor, the pressure increases beyond 4.5 atms or the pressure blower in the vent malfunctions. The gas mixture from the gas panel is passed through a preheater section consisting of a helical coil surrounded by a ceramic heater before being sent to the carbonation reactor. At the outlet of the heater, the gas composition and temperature of the gas mixture is monitored. H 2 and CO gas alarms are also connected to the gas panel to shut off the gas flow in the case of a gas leak. 337

369 Steam is generated separately using a high precision water pump which conveys water to a heated section similar to the gas preheater. The pressure and temperature of the steam is constantly monitored. The steam and preheated gas mixture are mixed just before they enter the reactor. The sorbent is stored in a hopper and is metered by a screw feeder into the reactor as shown in Figure 8.5. An airlock and two valves are provided in the sorbent transport line to prevent gas flow from the reactor to the sorbent hopper. The reactor is divided into 4 flanged parts and is surrounded by ceramic heaters to maintain the desired temperature during operation. Provision has been made for 8 temperature ports, 4 gas sampling ports and 4 pressure ports throughout the length of the reactor. The gas sampling ports and pressure ports are protected by in-line filters to avoid the solids from entering the lines. The gas sampling ports are connected to the GC for continuous gas analysis. Pressure transducers are connected to the pressure ports and thermocouples are inserted into the temperature measurement ports. The temperature ports of the reactor are integrated with the reactor heaters through a feedback loop process control system. Gas leak detectors have been placed near the reactor to alert the operators. The H 2 -rich product gas containing the calcium sorbent, from the outlet of the reactor, is then passed through a high temperature particle capture device followed by a water cooled heat exchanger shown in Figure 8.6. Before the gas enters the baghouse, 338

370 it is further diluted to reduce the temperature and prevent the formation of an explosive mixture. From the baghouse, the gas is vented out of the facility. The continuous production of H 2 by the CLP will be tested in this subpilot scale facility. Figure 8.7 is a schematic of the reactor system along with the support structure. The support structure has two platform levels and the sorbent is fed at the top level. 8.4 CONCLUSIONS The CLP has been shown to enhance H 2 yield and purity from syngas and hydrocarbons in lab and bench scale tests. The process is being scaled up to a 25 KWth subpilot unit demonstration at the Ohio State University. Cold flow test were conducted to determine the flow characteristics of the sorbent and parameters like fluidization velocity. The subpilot scale unit design is based on the thermodynamic, kinetic and sorbent reactivity studies and cold flow tests. This unit will be used to conduct continuous testing for the production of H 2 from a simulated syngas stream and a mixture of hydrocarbons. 8.5 RECOMMENDATIONS FOR FUTURE WORK The CLP has been successfully demonstrated at the lab and bench scale for carbon capture during the production of H 2 and electricity. A subpilot scale unit is being constructed for continuous testing of the concept at a higher scale. System 339

371 analysis and economic analysis have shown that the CLP has good potential to improve the efficiency and economics of H 2 production. After construction of the subpilot scale unit, shake down and start up testing will be conducted. Following this, leak testing and fluidization tests with hot air will be conducted to ensure safety and to determine the temperature and pressure profile and the flow characteristics of the sorbent at 600 ºC in the system. H 2 production tests will then be conducted with a simulated syngas stream to determine and optimize important process parameters like the residence time and Ca:C ratios. The effect of process parameters including temperature, pressure and S:C ratio will also be determined and a comparison will be made with the data obtained at the bench scale. On successful testing of the carbonation reactor for H 2 production, a fluidized bed calciner and hydrator should be integrated in the subpilot scale unit to test continuous sorbent flow through the process during H 2 production. The purity of H 2 produced from the carbonator and that of the CO 2 from the calciner should be monitored over long range tests lasting from 1 to 5 days. The performance of particle capture devices like high temperature cyclones and metallic filter for the micron sized calcium sorbent should be evaluated. At the end of this testing the subpilot scale unit should be moved to a location where a slip stream testing can be conducted from a real gasifier. The effect of other impurities in the syngas and flyash will be determined from the slip stream testing. On successful testing of the CLP concept in the 25 KW subpilot scale facility, 340

372 the process should be scale up to 250 to 500 KW to move the process to commercialization. System analysis studies and economic analysis should be updated during every scaleup testing to confirm the feasibility of the process with new information obtained from the testing. Sensitivity analysis on different coal and coal surfur contents, limestones from different locations, and process operation parameters should be conducted to determine the advantages that certain parameters or location of the plant could offer. Since the CLP produces a lot of high quality heat in the carbonation reactor and hydrator, reactor design and effective methods of heat extraction should be evaluated so that this heat can be used to produce additional electricity. Complete life cycle analysis should be conducted to determine the impact of this process on the environment and ecology and these lifecycle and economic studies should be used to guide the development of the process. Scientific studies to understand the mechanism of sintering and deactivation of the sorbent and reactivation by hydration should be conducted. These studies will help in further improving the reactivity of the sorbent. With complete understanding of the sorbent performance, scale up characteristics, system analysis, economics and life cycle analysis the CLP can be brought closer to commercialization for production of H 2, electricity, chemicals and liquid fuels. 341

373 40 30 STD (Pa) Ug (m/s) Figure 8.1: Standard deviation of pressure in the fluidized bed 342

374 Figure 8.2 (a): Schematic diagram of the cold flow model for the CLP 343

375 Figure 8.2 (b): Snapshot of the cold flow model for the CLP 344

376 Figure 8.3: Cold flow model for the hydrator 345