OPTIMIZATION OF THE SHIFT CONVERSION UNIT IN A GASIFICATION PLANT

OPTIMIZATION OF THE SHIFT CONVERSION UNIT IN A GASIFICATION PLANT Gasification Technologies Council Annual Meeting 2006 Washington, DC October 1-4, 2006 Ashok Rao Akshay Verma Advanced Power and Energy Program University of California Irvine, California 92697-3550 Douglas H. Cortez Hensley Energy Consulting LLC 412 North Coast Highway Suite B346 Laguna Beach, CA 92651

ABSTRACT With added emphasis being given to hydrogen coproduction and / or carbon dioxide capture and sequestration in a gasification plant, optimization of the shift conversion unit as an integrated feature of the gasification plant is critical in order to lower the capital and operating costs and increase the overall plant efficiency. When the gasification process produces a syngas that has a high CO to H 2 ratio a large quantity of steam addition is required to limit the temperature rise across the shift catalyst bed since shifting the CO to H 2 involves an exothermic reaction. This large steam demand tends to increase equipment sizes in the shift unit and decrease the plant efficiency. This paper presents the results of a preliminary plant design and economic study to illustrate the methods used to identify an optimum configuration for the shift unit for the coproduction of H 2 in an IGCC facility. INTRODUCTION Some of the leading entrained flow, high temperature gasification processes produce a raw syngas that have a high CO to H 2 ratio. Steam addition to the syngas entering the shift unit is required in quantities far in excess of the amount required by stoichiometry or to avoid carbon deposition on the shift catalyst. Generation of this large quantity of steam tends to decrease the plant efficiency and adds to equipment costs in the shift unit as well as the waste water treating unit. An alternate process configuration consists of a two-stage shift unit with a substantial fraction of the syngas bypassed around the first shift reactor to combine with the effluent from the first reactor and then feeding the combined stream to the second reactor. The effluent from the first reactor functions as a thermal diluent in the second reactor. The overall steam consumption is significantly lowered using this configuration as compared to a conventional two-stage shift system, i.e., without the by-pass. On the other hand, the CO concentration in the shift unit effluent gas is increased which in turn increases the amount of syngas to be processed in the shift unit and the feed gas to the Pressure Swing Adsorption (PSA). The tail gas generated by the PSA is also increased and hence the size of the tail gas compressor. This paper illustrates these above trade-offs and presents a near optimum configuration for the coproduction of H 2 using a typical gasification configuration consisting of shifting the syngas downstream of the acid gas removal unit. For this study, we have selected a configuration that feeds petroleum coke to an IGCC facility utilizing a high pressure entrained bed slurry fed twostage gasifier. Similar principles would apply to configurations using the other high temperature entrained flow gasification technology. This scheme may be employed in both sweet shift as well as sour shift applications where addition of steam to the syngas is required. International Standards Organization (ISO) ambient conditions are utilized for developing the plant performance while the feedstock consists of a typical, high sulfur delayed petroleum coke. Its characteristics are summarized in Table 1. Plant heat rejection is accomplished utilizing mechanical draft cooling towers. It is assumed that fresh makeup water is available. Two trains of F class technology gas turbines are utilized while 80 MM scfd of refinery grade 99% purity H 2 is exported from the plant. The performance of the gas turbines on syngas is estimated Rao, Verma and Cortez Page 2 of 12

utilizing performance published for the General Electric (GE) Frame 7FB gas turbine on natural gas by accounting for the reduction in firing temperature and increase in power output associated with syngas operation. The performance estimates for the gasification unit were developed utilizing information available in the public domain for the E-Gas gasifier i ii and GE gas turbine systems. The plant heat and material balances were generated using ASPEN models with proprietary improvements developed by the authors. The gasifier was modeled on a thermodynamic basis and approaches to equilibrium were utilized to account of the incompleteness of the various reactions. The gas turbine performances were developed by first calibrating a proprietary model utilizing the natural gas performance and then operating it in off-design mode to estimate its performance on syngas. Sud-Chemie provided the performance and catalyst requirements for two of the cases and a proprietary kinetic model was utilized to develop the data for the remaining cases. UOP provided the performance and cost of the PSA unit also for two of the cases while the performance and cost estimates for the remaining cases were developed from this UOP supplied data. Plant cost estimates were developed utilizing ICARUS (an Aspen suite product) and information available in the public domain but adjusted and scaled by the authors to apply to the specific cases examined in this paper. PROCESS DESCRIPTION Figure 1 depicts the overall block flow diagram for this IGCC plant coproducing H 2 from delayed petroleum coke. The nominal federate to the plant is 6,000 ST/D (as received basis) of the petroleum coke. The Air Separation Unit (ASU) consisting of two trains supplies O 2 to three 50% gasifier trains. The petroleum coke along with a fluxant is wet ground in two 50% trains of rod mills to form a slurry and introduced into the gasifiers. The gasifiers partially oxidize the coke with the O 2 to generate a hot raw gas, slag and char. The raw gas leaving the gasification zone is cooled with colder recycle gas to convert any entrained molten slag to a hardened solid prior to entering the syngas cooler, which generates high pressure (HP) steam. The char present in the gas exiting the syngas coolers is removed by a hot gas filter (metal candle filter) followed by a wet scrubber which removes any remaining solids. The char separated by the hot gas filter is recycled to the gasifier. The other contaminants such as soluble alkali salts, hydrogen halides and a portion of the ammonia are also removed. The contaminated water is sent to a sour water stripper along with a bleed from the slag bath. The sour gases stripped out of the water are routed to the Sulfur Recovery Unit (SRU). The scrubbed gas enters the Low Temperature Gas Cooling Unit (LTGCU) at its dew point and it is first preheated to about 40 F (22 C) above its dew point and then treated in a COS hydrolysis reactor to convert most of the COS to H 2 S which is easier to remove in the downstream Acid Gas i NETL DOE Report, Destec Gasifier IGCC Base Cases, PED-IGCC-98-003, September 1998, Latest Revision June 2000. ii Phil Amick, et. al., An Optimized Petroleum Coke IGCC Coproduction Plant, presented at the Gasification Technologies Council Conference, San Francisco, California, October 7-10, 2001. Rao, Verma and Cortez Page 3 of 12

Removal Unit (AGRU). The effluent from the COS hydrolysis reactor is then cooled in a series of heat exchangers while recovering heat for the syngas humidifier and vacuum condensate heating. Finally the syngas is cooled in a trim cooler against cooling water before it is supplied to the AGRU. In the case of coal, the LTGCU would also include a sulfided activated carbon bed for removal of Hg and As (substantial amounts of Cd and Se that may be present in the syngas are also expected to be removed by this activated carbon bed) from the syngas. In such a case, the cooled syngas after superheating by about 20 F (11 C) to avoid pore condensation is fed to the sulfided activated carbon bed. Within the AGRU, H 2 S, a small fraction of the COS present in the feed gas and some (about 15%) of the CO 2 are captured by the amine (MDEA) circulating through the absorber. The desulfurized syngas contains typically less than 10 ppmv H 2 S. The treated syngas is then supplied to a humidifier while the acid gas removed in the stripper of the AGRU is fed to the SRU where sulfur is recovered in elemental form. The tail gas leaving the SRU, which contains mostly CO 2 with some residual sulfur compounds as well as elemental sulfur vapor, is hydrogenated to form H 2 S from the sulfur species. Note that it is important to hydrogenate the elemental sulfur species present in the SRU tail gas to avoid condensation within the intercoolers of the recycle compressor. The hydrogenated tail gas is compressed and recycled to the gasifiers. The desulfurized syngas is humidified in a counter-current column by direct contact with hot water and then supplied to the shift unit. The heat required by the humidifier is recovered from the LTGCU. The humid syngas (Stream 1 depicted in Figure 2) is first preheated to 550 F or 288 C (Stream 2) against high temperature boiler feed water from the heat recovery steam generator (HRSG). The characteristics of Stream 2 are presented in Table 2. The amount required for H 2 coproduction is diverted to the shift unit (Stream 4) while the remainder is supplied to the gas turbines (Stream 3). Within the Shift Unit, a portion of the gas (Stream 6) is preheated against the second shift reactor effluent (Stream 12) in the Feed Preheat Exchanger, combined with steam and then supplied to the Shift Reactor 1 (Stream 8), while the remainder of the gas (Stream 5) constitutes the bypass. Effluent from the Shift Reactor 1 (Stream 9) is cooled while recovering heat for high pressure (HP) steam generation 2080 psig (143.4 barg). Next, this cooled syngas (Stream 10) is combined with the bypass syngas (Stream 5) and fed to the Shift Reactor 2 (Stream 11). Effluent from this second reactor (Stream 12) is cooled in the Feed Preheat Exchanger and then supplied to a series for heat exchangers (Stream 13) before it is fed to the PSA unit. The first of these exchangers cools the gas while generating intermediate pressure (IP) steam at a pressure of 430 psig (29.6 barg). The gas (Stream 13) is then further cooled while generating low pressure (LP) steam at a pressure of 70 psig (4.8 barg) and then cooled in a trim cooler against cooling water before it (Stream 15) is supplied to the PSA unit. The process condensate generated in these exchangers is collected and supplied as makeup to the syngas humidifier while the excess is treated in the sour water stripper of the waste water treatment unit. The amount of condensate collected and treated in the sour water stripper decreases as the amount of Shift Reactor 1 bypass is increased (steam addition in the shift unit being reduced). The sulfur content in the feed gas to these high temperature sweet shift reactors is maintained by the AGRU below the typical 100 ppmv limit for this shift catalyst. Low temperature sweet shift Rao, Verma and Cortez Page 4 of 12

reactors are not utilized in this design due to their susceptibility to sulfur poisoning. A ZnO bed would be required to desulfurize the syngas upstream of the shift reactors while utilizing the amine AGRU if a low temperature sweet shift reactor is utilized, which would increase both the capital and operating costs of the plant. The PSA unit is designed to produce 80 MMSCFD of 99% purity H2 (Stream 16) with CO content limited to 10 ppmv. The tail gas from the PSA unit (Stream 17) after compression (Stream 18) to a pressure as required by the gas turbine is preheated to 550 F or 288 C (Stream 19) against high temperature boiler feed water from the HRSG, combined with the portion of the syngas which is in excess of that required for H 2 generation (Stream 3) and is supplied to the gas turbines (Stream 20). HP N 2 from the ASU after preheating against high temperature boiler feed water is also supplied to the gas turbine combustors. The exhaust gas from the gas turbines is supplied to triple pressure Heat Recovery Steam Generators (HRSGs) which provide the superheated steam at 1920 psig / 1050 F (132.4 barg / 566 C) and reheated steam at 360 psig / 1050 F (24.8 barg / 566 C) to a single condensing steam turbine. The plant includes the necessary general facilities such as cooling water system, instrument air, flare, etc. RESULTS Tables 3 and 4 summarize the shift unit characteristics and impacts on the PSA unit for the baseline case designed with no bypass around the Shift Reactor 1 and three cases with various amounts of bypass: 25%, 50% and 75% of total gas supplied to the shift unit (i.e., percentage of Stream 4 bypasses as Stream 5). The upper limit for the bypass is set by constraints of the shift catalyst in second reactor which include its operating temperature and the minimum steam to carbon ratio required to avoid carbon deposition and reduction of the FeO to a carbide form (as the fraction of bypassed gas is increased, the Shift Reactor 2 effluent temperature increases while the steam to carbon ratio decreases). As seen in the Table 3, the steam produced downstream of the shift unit and the catalyst requirements are significantly reduced as the amount of bypass is increased. The size of the heat exchangers is also reduced. The size of the PSA unit on the other hand increases along with the tail gas compressor as seen in Table 4 as the bypass around the first shift reactor is increased since the amount of gas to be treated in the PSA unit is increased while its H 2 concentration is lower. There are also minor changes in the amount of air extracted from the gas turbine and the amount of nitrogen supplied by the ASU (the gas turbine output and the coke feed rate to the plant are held constant for each of the cases by adjusting the amount of air extracted). Table 5 summarizes the incremental plant performance and costs for various amounts of syngas bypass around the first shift reactor over the No Bypass Case. The petroleum coke feed rate and the H 2 export rate are held constant for each of the four cases. The incremental plant cost per incremental kw power available for export over the No Bypass Case steadily increases as the percentage bypass is increased. Rao, Verma and Cortez Page 5 of 12

Table 6 presents the results of the economic analysis performed on these cases. The following lists the basis for this analysis: Plant Installed Costs based on Installed Costs Overnight $2006 Costs Excludes Contractor Profit, Risk Fees & Contingency Simplified Economics based on Utility Cost of Service WACC of 8% (Typical Investor Owned Utility) Incremental Maintenance Costs of 1.5% of Capex Incremental Property Taxes & Insurance at 1% of Capex No Added Fuel Costs or H2 Revenues (Fixed by Design) Cost of Electricity (COE) Methodology General Inflation of 2.5% / year, Excluding Property Taxes (no Escalation) Simplified DCF Model Used to Compute Nominal COE Deflated to $2006 From the results of the economic analysis presented in Table 6, it may be seen that the COE reduces significantly as the percentage bypass around the Shift Reactor 1 is increased. DISCUSSION Based on the results of this analysis, it may be concluded that the economics favor increasing Shift Reactor 1 bypass for the syngas consisting of a CO to H 2 ratio of 1.91. The upper limit for the fraction of gas bypassed around the reactor is constrained by the minimum steam required to avoid carbon deposition, to avoid reduction of the catalyst to FeC, and limit the Shift Reactor 2 exit temperature. It should be noted that these conclusions are specific to sweet shift units operating in non-carbon capture IGCC plants with co-production of hydrogen, the type of gasifier and feedstock and the feed gas CO to H 2 ratio. Each application of shift technology should be evaluated on a case by case basis. The configuration consisting of bypassing the 1 st shift reactor in order to reduce the steam consumption may also be applicable to Zero Emission plants consisting of a sour shift unit followed by selective acid gas removal. In such applications, a tradeoff between the CO slip and the degree of CO 2 capture in the downstream AGRU needs to be studied. Maximum benefit of this approach may be found in applications where the PSA tail gas can be utilized as boiler fuel. In such applications the penalty both in terms of efficiency and capital cost associated with tail gas compression is minimized. Rao, Verma and Cortez Page 6 of 12

ACKNOWLEDGEMENTS The authors wish to specifically thank the University of California, Irvine, and UOP for providing the performance and cost for the PSA unit and Sud-Chemie for providing the shift reactor performance and catalyst costs. AUTHORS Ashok Rao Dr. Ashok Rao is the Chief scientist for Power Systems at the Advanced Power and Energy Program (APEP) of the University of California, Irvine. Prior to accepting the position at APEP in 2004, Dr Rao was a director in Process Engineering at Fluor Inc involved in gasification and combined cycles, and a Senior Fellow. Prior to joining Fluor in 1979, he worked for a gasification technology developer, Allis Chalmers and a gasification licensor, McDowell Wellman. With more than 30 years of experience working in industry, Dr. Rao has been involved in all phases of project development starting from process conceptualization to R&D to techno-economic feasibility studies to detailed design. He is the recipient of several patent awards in the area of gasification technology. Dr. Rao can be reached at adr@apep.uci.edu or 949-824-7302 x345. Akshay Verma Akshay Verma is a graduate student at the University of California, Irvine working towards his PhD in Mechanical Engineering. His research involves advanced membrane technologies for the separation of H 2 from syngas including the development of a mechanistic model for a H 2 separation membrane shift reactor using high temperature inorganic membranes and the necessary experiments to gather fundamental data and validate the model. He obtained his MS degree from the Materials Science department of the University of California, Irvine and his undergraduate degree in Metallurgical Engineering and Materials Sciences from IIT-Mumbai, India. He has extensive experience in performing IGCC plant simulations in Aspen. Douglas H. Cortez Dr. Cortez is Managing Director, Hensley Energy Consulting LLC. He is a consultant to leading companies in the energy industry, including utilities, oil and gas companies, financial institutions and government agencies. He has over 35 years of experience in energy project design, development and project financing with a focus on clean coal, power and gasification technologies. Prior to forming Hensley Energy Consulting in 2006, Dr. Cortez served in various management capacities with Fluor Corporation and Tosco Corporation (now ConocoPhillips). Dr. Cortez can be reached at HensleyEnergy@pobox.com or 949-697-7536. Rao, Verma and Cortez Page 7 of 12

Table 1: Petroleum Coke Feedstock Characteristics PROXIMATE ANALYSIS iii As Received Dry Moisture 4.83 Ash 0.13 Volatile Matter N/A Fixed Carbon N/A HHV kj/kg 32,871 34,537 Btu/lb 14,132 14,848 ULTIMATE ANALYSIS Carbon 88.755 Hydrogen 3.20 Nitrogen 0.90 Chlorine 0.005 Sulfur 7.00 Ash 0.14 Oxygen <0.01 iii Coke composition is not expressed in terms of Proximate Analysis. The Proximate Analysis shown was arbitrarily picked to provide the complete input required by Aspen. Rao, Verma and Cortez Page 8 of 12

Figure 1: Overall Plant Schematic for the IGCC H 2 Coproduction Facility

Figure 2: Shift and PSA Units Rao, Verma and Cortez Page 10 of 12

Table 2: Humid Syngas (Stream 2) Characteristics CO/H 2 Ratio, mole/mole 1.91 H 2 O Content, mole % 16.72 H 2 S + COS Content, ppmv < 110 Flow Rate, Moles/h 60,760 Temperature, F 550 Pressure, psia 443 Table 3: Shift Unit Characteristics Shift Reactor 1 Bypass, % 0 25 50 75 Feed Gas Flow Rate (Stream 4), moles/h 15,820 16,286 17,393 20,307 Steam Required, lb/h 428,810 332,730 235,410 137,400 HP Boiler Duty, MM Btu/h 80.04 64.16 46.24 22.79 IP Boiler Duty, MM Btu/h 64.73 73.04 81.89 91.52 LP Boiler Duty, MM Btu/h 212.69 121.45 36.77 32.95 Relative Reactor 1 Volume 1.0 0.78 0.56 0.33 Relative Reactor 2 Volume 1.0 0.92 0.86 0.71 Table 4: Impact on PSA Unit Shift Reactor 1 Bypass, % 0 25 50 75 Feed Gas Flow Rate (Stream 15), moles/h 19,690 20,000 20,720 22,590 CO Conc. in Feed Gas, mole % 2.1 3.5 6.7 14.2 H 2 Conc. in Feed Gas, mole % 51.4 50.7 49.1 45.6 Tail Gas Compression, kw 17,188 17,706 18,922 22,091 Rao, Verma and Cortez Page 11 of 12

Table 5: Overall Plant Performance & Costs Summaries Increase / Decrease over No Bypass Case Shift Reactor 1 Bypass, % 25 50 75 Increase in Steam Turbine Power, kw 2,393 5,405 12,115 Increase in In-plant Power Consumption, kw 1,121 2,524 6,984 Increase in Net Plant Power, kw 1,272 2,882 5,131 Increase in Installed Plant Cost, 1000$ 4,235 7,237 8,953 Decrease in Operating Cost (Catalyst), 1000$/yr 90 174 280 Incremental Plant Cost, ( $)/( kw) 3,330 2,511 1,745 Table 6: Cost of Incremental Power - Typical Utility COE, $2006 Shift Reactor 1 Bypass, % 25% 50% 75% $/MWh $/MWh $/MWh Weighted Ave Cost of Capital at 42.2 31.9 22.1 Incremental Fuel Cost and Hydrogen Revenues 0.0 0.0 0.0 Savings in Annual Catalyst Cost (9.0) (7.7) (6.9) Incremental Property Taxes and Insurance 4.2 3.2 2.2 Incremental Annual Fixed Maintenance 6.3 4.8 3.3 Cost of Incremental Electricity, $/MWh 43.8 32.2 20.7 Rao, Verma and Cortez Page 12 of 12