Stripping, WaterâGas Shift Reactors, and CO2

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Ind. Eng. Chem. Res. 2010, 49, 4766–4781

Integrated Gasification Combined Cycle Dynamic Model: H2S Absorption/ Stripping, Water-Gas Shift Reactors, and CO2 Absorption/Stripping Patrick J. Robinson and William L. Luyben* Department of Chemical Engineering, Lehigh UniVersity, Bethlehem, PennsylVania 18015

Gasification could potentially emerge as the premier unit operation in the energy and chemical industries. In the future, plants are predicted to be a hybrid between power and chemical with the ability to handle unavoidable swings in both power demand and biomass feed composition without a loss of efficiency. The coupling of a power plant with a chemical plant provides an additional control degree of freedom, which fundamentally improves the controllability of the process. The coupling of an integrated gasification combined cycle (IGCC) power plant with a methanol chemical plant handles swings in power demand by diverting hydrogen gas from a combustion turbine and syn gas from the gasifier to a methanol plant for the production of an easily stored, hydrogen-consuming liquid product. This paper presents an extension of the dynamic gasifier model, which uses a high-molecular weight hydrocarbon (with a 1:1 hydrogen to carbon ratio) as a pseudo-biomass feed stock. Using this gasifier model, the downstream units of a typical IGCC can be modeled in the widely used process simulator Aspen Dynamics. Dynamic simulations of the H2S absorption/stripping unit, water-gas shift (WGS) reactors, and CO2 absorption/stripping unit are essential for the development of stable and agile plantwide control structures of this hybrid power/chemical plant. Because of the high pressure of the system, hydrogen sulfide is removed by means of physical absorption. SELEXOL (a mixture of the dimethyl ethers of polyethylene glycol) is used to achieve a gas purity of less than 5 ppm H2S. This desulfurized synthesis gas is sent to two water-gas shift reactors that convert a total of 99% of carbon monoxide to hydrogen. Physical absorption of carbon dioxide with Selexol produces a hydrogen-rich stream (90 mol % H2) to be fed into combustion turbines or to a methanol plant. Steady-state economic designs and plantwide control structures are developed in this paper. 1. Introduction Although gasification has been used on a relatively limited scale for many years,1,2 it is likely to emerge as the premier unit operation for both the power and chemical industries. For the past several decades, steam-methane reforming has been the key method for producing synthesis gas: a gas that mainly consists of hydrogen, carbon monoxide, and carbon dioxide. However owing to the escalating prices of natural gas and the abundant supply of domestic solid hydrocarbons, a switch to oxygen-blown gasification of coal, petroleum coke, or biomass could be imminent. Traditionally, gasifiers were “air blown” meaning that air, water, and solid hydrocarbons were the main feeds. Because of the vast amount of nitrogen present, these gasifiers produce a low calorific fuel gas. Modern gasifiers use fairly pure oxygen and are “oxygen blown,” so a higher heating value synthesis gas is generated. This synthesis gas can be used in a combustion turbine or as a feedstock for the production of a wide variety of chemical products.3 1.1. Conventional IGCC Power Plants. There is a vast amount of literature dealing with IGCC power plants, with more and more papers appearing each year. This work deals almost exclusively with steady-state conditions and designs. The dynamics and control of an IGCC power plant have been hardly mentioned. The scarcity of published information makes it difficult to work on these processes and requires that engineering judgment be used for certain aspects when modeling. Controllability studies are vital because of the complex interacting nature of these plants. They feature recycle of material and energy that can lead to dynamic instability * To whom correspondence should be addressed. E-mail: WLL0@ Lehigh.edu. Tel.: 610-758-4256. Fax: 610-758-5057.

problems. A major potential dynamic disturbance is the hourto-hour changes in power demand, which can be quite large from day to night. The inherent variability of the composition of the hydrocarbon fuel source (coal) will result in additional upsets. The control systems developed for the plant must be designed to handle these disturbances. An integrated gasification combined cycle plant (IGCC) is a power plant that turns coal into synthesis gas, removes impurities, and combusts the gas in a turbine (whose waste heat is passed to a steam turbine system). Including the Department of Energy’s Clean Coal Initiative with Carbon Capture and Sequestration (CCS), an IGCC plant can be developed for nearzero emissions and high-efficiency. After purifying the synthesis gas from an oxygen-blown gasifier, more hydrogen can be produced by means of the water-gas shift (WGS) reaction: CO + H2O f CO2 + H2. The resulting gas is a mixture of hydrogen and carbon dioxide, where CO2 can be separated by physical absorption. To cut greenhouse gas emissions, the hope is that the carbon dioxide can be sequestered in underground or undersea reservoirs. A hydrogen-rich stream is then fed into a combustion turbine to generate power with nonpolluting water as the product. As electric power demand changes from hour to hour, the amount of hydrogen fed to the combustion turbine should change. Since the gasifier is the most sensitive and expensive unit, changing biomass feed rates is to be avoided. The economic feasibility for storing hydrogen gas (such as in salt-dome caverns or using liquefaction or hydrates) has yet to be established, and safety issues remain unresolved. Alternatively an easily stored, hydrogen-consuming chemical product (such as methanol or ammonia) could be produced from the excess hydrogen during periods of low power demand.

10.1021/ie901549s  2010 American Chemical Society Published on Web 04/26/2010

Ind. Eng. Chem. Res., Vol. 49, No. 10, 2010

1.2. Hybrid Power/Chemical Plant. A hybrid process that couples an electric power plant and a chemical plant, featuring an oxygen-blown gasifier as the key unit operation, has the potential of improving efficiency, rangeability, and controllability. The coupling provides an additional control degree of freedom, which fundamentally improves the controllability of the process. This hybrid process will have the ability to handle unavoidable swings in both biomass feed composition and electric power while base-loading the gasifier to avoid disturbing this sensitive unit. During periods of low power demand, hydrogen and syn gas are diverted from the combustion turbine to a chemical plant. During periods of high power demand, the gas fed to the chemical plant is minimized, as limited by the lowest possible turndown rates. Three different scenarios will be studied during this work. Scenario 1 is the base IGCC power plant with CCS, shown in Figure 1 a, with all of the syn gas generated in the two parallel gasifiers (71 340 lb-mol/h) used for generating electric power. All of the synthesis gas produced by the gasifiers is shifted to more hydrogen, the carbon dioxide is captured, and a hydrogenrich stream is burned in the combustion turbine. The total electric power is 774 MW. This is the stand-alone, power-only IGCC process. There is no chemical plant. If the gasifier is base-loaded, electric power production cannot be reduced. The flowsheet shown in Figure 1b gives Scenario 2, which is the “half-power” case. This is the case where energy demand is at its lowest (early morning period). The electric power generated is 387 MW with the same gasifier loading. A portion (23%) of the syn gas is sent to a methanol plant. A portion (35%) of the hydrogen-rich stream from the CO2 removal unit is also sent to the methanol plant in order to provide precisely the stoichiometric amount of hydrogen to react with the CO and CO2 in the syn gas. The methanol production rate is 5714 lb-mol/h, and the methanol plant is designed for this maximum capacity. The final scenario is shown in Figure 1c. We assume that the minimum turndown of the methanol process is 20% of the maximum capacity, that is, methanol production is reduced from 5714 to 1143 lb-mol/h. The portions of the syn gas and hydrogen streams diverted to the methanol plant are reduced to only 5%. It is important to note that the total electrical power is 696 MW in this case, which is less than the 774 MW that could be generated by shutting down the methanol plant completely. So the assumed minimum turndown is an important dynamic parameter that will be studied in future work. The purpose of this paper is to extend the simple dynamic gasifier model4 and present a dynamic model of the entire IGCC system (scenario 1). The design and dynamic analysis will utilize the widely used process simulators Aspen Plus and Aspen Dynamics. The steady-state predictions of the model will be compared against the results given in the National Energy Technology Laboratory study of General Electric (GE) Gasification Technology with Carbon Sequestration.5 The four major sections modeled in the IGCC are the two gasifiers in parallel, a H2S removing absorber/stripper unit, two WGS reactors, and a CO2 removing absorber/stripper unit in that corresponding sequential order. In some IGCC plants, the WGS reactors are capable of handling H2S impurities and are called “sour” water-gas shift reactors (SWR). This allows for absorption and stripping of H2S and CO2 to occur after the SWR to help on energy and capital costs. However in the hybrid flowsheet, the H2S is removed prior to the WGS reactors because the synthesis gas is used for both production of hydrogen and methanol. H2S is a very poisonous chemical to most catalysts

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in the hydrocarbon industry and should be removed from the gas before going to the methanol reactor. Some processes were not included in this study because they were out of the scope or did not apply to the project. These included an elevated pressure air separation unit that feeds oxygen to the gasifier, a Claus Plant that deals with H2S waste, carbon sequestration after it has been separated, particulate and mercury removal, and the turbine system. 2. Two Parallel Gasifiers Two parallel gasifiers are simulated using the methodology of an approximate simplified gasifier.4 This approach looks at General Electric (GE) Gasification Technology with Carbon Sequestration considered by NETL.5 The GE technology is described as an oxygen-blown entrained flow gasifier. Water is introduced into the gasifier in liquid form via a coal/water slurry mixture. For improved reliability and productivity, two gasifiers are operated in parallel. The gasifier consists of three sections: partial oxidizer (POX), radiant synthesis gas cooler (RSC), and water quench. Both the slurry and oxygen are fed into the top of the POX where the synthesis gas is produced. The gas and molten slag (consisting of unreacted carbon and ash materials) are then passed through the RSC and quenched in water. Temperatures in the POX are around 2500 °F to promote near complete conversion of the hydrocarbon to synthesis gas. The RSC generates 2000 psia saturated steam when cooling the gas to 1100 °F. A water quench drops the product gas to a temperature around 440 °F. The gasifier is operated at high pressure (800 psia) so that synthesis gas can be provided at high pressure. Two gasifiers running in parallel at steady-state conditions are shown in Figure 2. Steady-state Aspen simulation tools are capable of handling solid hydrocarbon materials. Complex features such as particle size distribution, slag formation, and detailed coal analysis (proximate and ultimate) can be handled.6 However, the ability to export the steady-state simulation to Aspen Dynamics is not available at this time, but Aspen Technology7 is working on this feature. The fundamental chemical differences among the hydrocarbon raw materials of natural gas, petroleum, and coal are the atomic ratios of hydrogen to carbon. In natural gas, the ratio is 4:1; in petroleum, it is 2:1; and in coal, it is 1:1. Therefore, the simple approximate4 approach uses a heavy component currently found in the Aspen Component Library. The complex cyclical aromatic used for the simulation as the pseudocoal is C18H20. For many plantwide dynamic studies, a rigorous high-fidelity dynamic model of the gasifier is not needed. The dynamics are very fast and the gasifier gas volume is a relatively small fraction of the total volume of the entire plant. The proposed approximate model captures the essential macro-scale thermal, flow, composition, and pressure dynamics. This dynamic model will be the basis for the feed for rest of the modeled IGCC. Two parallel dynamic models with the published control structure are used for the upstream units of this paper. 3. Aspen Plus Simulation of the IGCC 3.1. H2S Removal. A. Process Description. Sulfur components are very poisonous to most water-gas shift and methanol catalysts. Both catalysts contain copper, and sulfiding is very favored under reactor conditions. It is important to prevent even very low levels of sulfur from contacting the

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Figure 1. Hybrid IGCC/methanol plant scenarios: (a) full energy production (typical IGCC), (b) half power for maximum methanol production, and (c) 20% turndown of the methanol plant.


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Figure 2. Steady state of two parallel gasifiers using “pseudocoal” C18H20.

catalyst.8 Therefore, removing hydrogen sulfide upstream will be a necessity for the hybrid power and chemical plant. A number of technologies are available for acid gas removal, but the two most applicable in gasification facilities are physical or chemical solvents.9 Chemical solvents are favorable at low acid gas partial pressures, while physical solvents are favored at high partial pressures. This is because chemical solvents have high absorption capacity at relatively low acid gas partial pressure, but plateau at higher pressures. On the other hand, the solubility of acid gases in physical solvents increases linearly with acid gas partial pressure (Henry’s Law). The gasifier operates at a very high pressure to produce synthesis gas, so physical absorption is the most reasonable choice. Another key benefit for using physical absorption is that it also allows for the solvent to be partially regenerated by pressure reduction, lessening the energy requirement to remove the absorbed acid from the solvent. The physical absorption technology SELEXOL is currently owned and trademarked by Union Carbide Corporation (Dow Chemical Company subsidiary). This physical solvent has been used in gasification for many years and continues to be the preferred choice for acid-gas removal. This solvent consists of a mixture of dimethyl ethers of polyethylene glycol of the formula CH3O(C2H4O)xCH3 where x is between 3 and 9. In the Aspen Plus component library, the heaviest dimethyl ether of polyethylene glycol has x ) 5 and is the dimethyl ether of pentaethylene glycol (DME-PEG, Chemical Abstracts Services Registry Number 1191-87-3). Although in reality the solvent is a mixture of many different polyethylene glycols with very little water,10 a reasonable mixture of 94 wt % DME-PEG and the water as the remainder is used as the solvent for these simulations. The synthesis gas generated from the two parallel gasifiers is sent to the H2S absorption/stripping unit to obtain a treated gas stream with 5 ppm H2S. The NETL study required that the content of the synthesis gas be reduce to less than 30 ppm of total sulfur content before utilization of the synthesis gas. This conservative specification of 5 ppm will ensure no catalyst is poisoned even during very large upsets.

Before entering the absorption column, all the water has to be knocked out of the synthesis gas. The gas is sent through a heat exchanger, cooled to 103 °F, and the water is removed in a knockout drum. As shown in Figure 3, the outlet gas is then fed into the bottom of the absorption column where most of the H2S and some CO2 and CO are physically absorbed with the DME-PEG solvent. The synthesis gas exits the top of the column with only 5 ppm H2S and is ready for production of either hydrogen or methanol. The rich solvent stream coming out of the bottom of the absorber is preheated with the lean solvent in a feed effluent heat exchanger (FEHE) prior to entering the stripping column. The H2S is stripped off with a design specification of only 1 ppm H2S in the solvent recycled back to the absorber. The stripper distillate is cooled to 110 °F to ensure most of the water is not lost overhead. The partial condenser sends the vapor stream for sulfur recovery in a Claus Plant.11 Any water that was lost in the process is added to the reflux drum. The reflux consists of mainly water and is sent back to the top of the stripping column. The recycled solvent coming out of the bottom of the stripper is cooled to 110 °F by means of the FEHE and a cooling water heat exchanger. It is then sent back to the top of the absorber. Very little DME-PEG is lost in the process, but some is added to the bottom of the column to maintain composition. Plate towers for gas absorption are sometimes used, particularly when the liquid load is more than can be handled in a packed tower of about 3.5 feet.12 They are also useful when the liquid rate is sufficient to flood a packed tower. This is the situation in this process because of the high throughputs and solvent needed for this application. Figure 3 shows an optimal designed flowsheet of this process. The NETL IGCC design uses a SWR to shift the synthesis gas to hydrogen prior to acid-gas removal, so validation of results are performed downstream. B. Design and Optimization Methodology. Acid-gas removal plants were designed to minimize total annual cost (TAC) while maintaining product specifications. Considering the real SELEXOL solvent composition is more complex then the simple

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Figure 3. Optimal design of the H2S removal via physical absorption with DME-PEG.

mixture of DME-PEG and water as simulated here, results may be somewhat different than those used in the real plant design. However, the essential thermal, flow, composition, and pressure dynamics are captured. In general, a cubic equation of state property method is appropriate for gas processing. Peng-Robinson is used throughout the absorption and stripping columns for H2S removal. The Newton algorithm was used for convergence in Aspen Plus on both columns for this highly nonideal distillation. The synthesis gas from the gasifier was cooled to 110 °F by means of a HEATER block. An adiabatic FLASH2 separator was used for the knockout drum. The vapor stream is then sent to a RADFRAC block with no condenser or reboiler used to model the absorber. Sieve trays are used with a tray spacing of 2 feet. Maintaining a high pressure is very important for physical absorption, so the gas outlet was chosen to be 750 psia. A stage pressure drop of 0.3 psi was selected to ensure pressure drops in the column are larger than twice the pressure drop due to liquid head. The acid-gas-rich solvent out of the absorber bottom is fed into a FEHE modeled by a HEATX with a hot outlet-cold inlet temperature difference of 20 °F. The overall heat transfer coefficient, U, was specified as a constant 150 BTU/(h · ft2 · °F). Pressure drops for the cold and hot sides were 5 psi. The cold outlet stream is then sent to a stripping column modeled by a RADFRAC block with a reboiler only. The top tray pressure was reduced to 30 psia to partially regenerate the solvent and lessen the energy requirement. The stripping column has the same tray type, tray spacing, and pressure drop as the absorption column. The design specification of only 1 ppm of H2S in the stripping column bottoms was maintained by varying the reboiler duty. The stripping column bottoms are sent to the hot side of the HEATX and then to a HEATER. It is designed to be cooled to 110 °F before entering back in the absorption column.

Figure 4. Solvent flow and the effect of H2S removal at different absorption tray numbers.

The condenser was modeled in a separate block for ease of convergence and to recover most of water in the reflux. Distillate was cooled to 110 °F by means of a HEATER for the condenser and an adiabatic FLASH2 separator for the reflux drum. Reflux was sent back to the top of the column. Make-up water is added at the reflux drum, and makeup DME-PEG is added at the bottom stage of the stripping column. Calculator blocks were used to adjust the amounts of makeup streams needed. The first design step is to size the absorber to obtain 5 ppm of H2S in the exit gas stream. At different absorber tray numbers (Nabs), the required 94 wt % DME-PEG solvent was found. Figure 4 shows the effect of solvent flow rate on the composition of H2S in the absorber gas outlet for columns with different


numbers of trays. For 40 trays, a solvent flow of 32 700 lb · mol/h was needed to obtain 5 ppm; for 50 trays, 31 600 lb · mol/h; and for 58 trays, 30 500 lb · mol/h. As expected, increasing the number of trays would require less solvent needed to absorb H2S. To minimize the capital cost of each absorption column, the number of tray liquid passes was varied to find a minimum column diameter. For each case, any increase above two tray passes had little to no effect on minimizing the column diameter. An absorption column with 58 trays has a diameter of 22.8 feet with two tray passes. For each sized absorption column and its respective needed solvent flow, the TAC of the whole absorption/stripping process must be analyzed to find an optimal design. This includes the capital cost of the absorption column, FEHE, stripping column, and heat exchangers along with the annual energy cost for the reboiler in the stripping column. Simple economic objective functions usually are adequate to serve the purpose of economically optimizing a design. The units of capital investment are in U.S. dollars and are represented in eqs 1 and 2: heat exchanger capital ) 1557(area in ft2)0.65

(1)

column vessel capital ) 1917(diameter in ft)1.066(length in ft)0.802 (2) A payback period of 3 years is used to obtain the annual cost of the capital investment. This is then combined with the annual energy cost to obtain the TAC. The unit cost of energy ranges from $6.08 to $9.83 per GJ depending on the steam pressure needed in the reboiler.13 The heat transfer coefficients for the reboiler and condenser are 100 and 150 BTU/(h · ft2 · °F), respectively. The solvent used is assumed to cost $0.50/lb. To minimize reboiler duty, the feed tray for the stripping column is varied. For 58 stages in an absorber, 10 stages in the stripping column, and a solvent flow of 30 500 lb · mol/h, feeding on tray three requires the minimum reboiler duty of 34.5 MW. This calculation is done for each case. Similar to the absorption column, two tray passes were used for all the stripping columns to minimize diameter. Using more than two liquid passes showed little to no improvement. For a given absorber size and solvent flow, the TAC cost was found for the absorber/stripping unit by varying the number of stages in the stripping column (NStr). However, increasing past 10 trays would increase TAC of the absorption/stripping unit because of the ease of removing the H2S from DME-PEG. Columns that contain fewer than 10 stages are unreasonable and were not considered. Using this methodology for each absorber size and solvent flow, the economic optimum design is found to have 58 stages in the absorber, 10 stages in the stripping column, and a solvent flow of 30 500 lb · mol/h as shown in Figure 5. The total flowsheet is shown in Figure 3. C. Steady-State Calculations for Control Structure. Temperatures are widely used to provide inferential control of compositions in distillation columns. Temperature sensors are inexpensive, reliable, and introduce small measurement lags in a control loop. Maintaining a constant pressure in a column, temperature is uniquely related to composition in a binary system and expensive composition analyzers would not be needed. Several alternative criteria can be used to select the location of a temperature control tray in the stripper to preserve the design specification of 1 ppm of H2S in the solvent recycle flow out of the stripping column bottoms. The slope criterion consists of studying the changes in temperature of tray to tray; the

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Figure 5. Economic optimal design of the H2S absorption/stripping unit.

sensitivity criterion finds the largest changes in tray temperature while varying the manipulated variable; and the singular value decomposition (SVD) criterion decomposes the steady-state gains between all the tray temperatures and the manipulated variables.14 The stripping column distillate will be maintained at 110 °F to ensure the majority of water is knocked out of the acid gas and returned in the reflux. The amount of H2S coming out of the bottom of the stripping column will be adjusted by reboiler duty. As shown in Figure 6, the top left graph gives the temperature profile at design conditions. The difference in temperature between adjacent trays is shown in the bottom left graph. There is a large change at the feed stage 3 due to the introduction of feed. This feed provides a large increase in composition of gaseous products significantly lowering the temperature at that stage in the column. This is not a good location for temperature control. The top right graph gives the openloop steady-state gains between tray temperature and the manipulated reboiler duty. A minor change in reboiler duty (+0.1%) shows that stage 5 is very sensitive to changes in heat input. The bottom right graph shows the SVD analysis with similar results to the sensitivity results. This suggests that stage 5 should maintain product specification by manipulating reboiler duty. 3.2. Water-Gas Shift Reactors. A. Process Description. When dealing with CO2 separation and capture in an IGCC, gasifier products must be converted to a hydrogen-rich synthesis gas.5 After the majority of H2S has been removed, the CO in the synthesis gas is reacted with water to produce more H2 and CO2 in two adiabatic tubular reactors packed with catalyst. The WGS reaction is shown in eq 3. CO + H2O S CO2 + H2

(3)

The clean synthesis gas coming out of the H2S absorption column top is combined with saturated steam to give approximately a 2:1 ratio of H2O to CO. This ratio will swing the equilibrium to the desired product, promoting high conversion of the CO. The temperature of the gas entering each reactor is controlled by heat exchangers at the inlet. The WGS reaction is exothermic, so the heat exchanger between the two reactors must remove heat. The temperature is high enough to generate steam, which is used as the water added upstream of the first reactor.

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Figure 6. Steady-state criteria for control structure selection in the H2S stripping column.

Figure 7. Optimal design of the WGS reactors for 99% conversion of CO to H2.

Approximately 96-98% conversion of the CO is achieved in the NETL studies.5 The water-gas shift reactors in this paper are modeled for 99% conversion to ensure an even higher reliability when unavoidable disturbances occur. Figure 7 is the optimal design of two WGS reactors (high-temperature shift (HTS) reactor and low-temperature shift (LTS) reactor) using the clean synthesis gas coming out of the top of the H2S absorption column. B. Reaction Kinetics. As with all reactor simulations, the most vital and the most difficult aspect of the simulation is to get correct kinetic expressions. The heterogeneous kinetics reported by Choi and Stenger15 are used in this study. Their kinetic expression was adjusted from the small catalyst loading in their experiment to a more industrially realistic particle load density, Fcat, of 150 lb/ft3. The catalyst used in the study was a typical LTS catalyst that consisted mainly of Cu/ZnO/Al2O3. Because this catalyst contains copper, it is poisoned by sulfur. The free zinc oxide present in the catalyst will prevent any deactivation when the H2S exceeds concentration limits. Hightemperature shift kinetics are not used in this paper due to unreliability and scarcity in literature. The units for overall reaction rates given in the cited paper are mol g-1 h-1 and pressures are in atmospheres. These parameters must be converted into required Aspen units of kmol

sec-1 m-3 and Pascals. The converted parameter values are given below in eq 4. RF ) kFe-EF/RTPCOPH2O RR ) kRe-ER/RTPCO2PH2

(4)

where RF ) forward reaction (kmol sec-1 m-3); kF ) forward pre-exponential factor ) 2.074 × 10-7 (kmol sec-1 m-3 Pa-2); EF ) forward activation energy ) 47 400 (kJ/kmol); Pj ) partial pressure of component j (Pascals); RR ) reverse reaction (kmol sec-1 m-3); kR ) reverse pre-exponential factor ) 1.575 × 10-5 (kmol sec-1 m-3 Pa-2); ER ) reverse activation energy ) 85 460 (kJ/kmol). C. Design and Optimization Methodology. The WGS reactors are modeled as a series of adiabatic plug flow tubular reactors (PFR) with interstage cooling. The first heat exchanger heats up the synthesis gas to the desired inlet temperature, while the others cool the process gas in steam-generating interstage coolers. Boiler feedwater (BFW) is fed on the shell side of these heat exchangers, and the steam generated is used to help maintain the 2:1 H2O:CO ratio upstream. To design this unit the minimum number of reactors, NR, is found to obtain 99% conversion of CO into H2. The inlet


Figure 8. The required catalyst for one WGS shift reactor at various inlet temperatures.

temperatures of each reactor, Tn, where n usually varies between 1 and 3, are the design optimization variables. The sizes of all the reactors are assumed to be identical with the same amount of catalyst present, Wcat. The Peng-Robinson base method was used for the Aspen simulation. Steam is added to the clean synthesis gas exiting the top of the H2S absorption column by means of a calculator block to maintain the desired ratio. Each reactor is modeled with an adiabatic RPLUG block of a given volume. Catalyst is present in the reactor with a bed voidage, ε, of 0.5 and a particle density of 150 lb/ft3. The heat exchangers are modeled simply with a HEATER block to give the desired inlet temperature. Pressure drop through the units is assumed to be 5 psi each. The first step in the design methodology is picking the number of reactors. Assuming a certain reactor size, the inlet temper-

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atures can then be varied to obtain the required CO conversion. To obtain the 99% conversion in one WGS reactor is unreasonable. Figure 8 shows the amount of catalyst needed for the reaction to obtain equilibrium at various inlet temperatures. Even cooling down the synthesis gas to 75 °F, 99% convergence does not occur at a catalyst load of over 350 tons. As expected, a higher inlet temperature speeds up the reaction and requires less catalyst, but 99% conversion of CO to H2 cannot be attained because the higher temperature reduces the equilibrium constant. By increasing the number of reactors to two, the size of each reactor becomes much more reasonable. For a certain amount of catalyst in each reactor, an optimum inlet temperature can be found to maximize total conversion. In the left graph of Figure 9, reactor inlet temperatures (T1 and T2) were manipulated to find optimal conversion for two WGS reactions containing 5 tons of catalyst and a total volume of 67 ft3 (L ) 11 ft, D ) 2.75 ft). An optimal inlet temperature of T1 and T2 equal to 520 and 480 °F provide an overall conversion of 95.5% of CO to H2. As expected, the first reactor would have a higher inlet temperature than the second. At the low inlet temperatures of 480 and 490 °F into the first reactor, the kinetics are too slow to maximize the amount of hydrogen produced. The higher inlet temperature of 560 °F promotes fast kinetics, but it is equilibrium limited and produces less than an inlet of 520 °F. The right graph in Figure 9 shows the composition profile of reactor 1 given each inlet temperature. Similar plots are done for the second reactor given its respective inlet concentration. Optimal inlet temperatures for a certain sized reactor are found when the temperature gives a high enough equilibrium constant for a high conversion and at the same time gives specific reaction rates that are fast enough to reach the equilibrium constraint. By increasing the weight of catalyst and in turn the volume of the reactor, the specification of 99% conversion can be achieved. Figure 10 gives the optimal T1 and T2 with the corresponding maximum conversion for each reactor size. As

Figure 9. Varying inlet temperatures of two WGS reactors, T1 and T2, to find an optimum.

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Figure 10. The amount of catalyst needed for 99% conversion of CO to H2 in two WGS reactors given the optimal inlet temperatures.

expected, the inlet temperatures decrease for a larger catalyst weight while conversion increases. The optimal design calls for two reactors, each with 41 tons of catalyst and 1100 ft3 in volume. The inlet temperatures of these designs are 320 and 310 °F for the inlet of the HTS and LTS, respectively. The optimal flowsheet of the WGS reactors is shown in Figure 7. 3.3. CO2 Removal. A. Process Description. The process for CO2 removal is identical to that of the H2S removal, in which an absorption/stripping unit uses the physical solvent SELEXOL to absorb CO2 out of the H2/CO2 rich synthesis gas. However, the amount of solvent needed to absorb the CO2 is much larger than that needed for the H2S because of the solvent’s higher selectivity toward H2S. Another big difference is that the pressure on the top tray of the stripping column cannot be lowered to near atmospheric pressure because this CO2 rich stream will be pressurized and used for sequestration. Therefore, an optimum stripper pressure must be found that balances compression costs with stripper reboiler energy cost.

CO2 is absorbed in the solvent to create an exit gas stream with a composition of 90 mol % H2. The CO2 is then stripped from the solvent with a design specification of only 0.1 mol % CO2 in the lean solvent recycled back to the absorber. The stripper distillate is cooled to 110 °F to ensure most of the water is not lost overhead in the stripper. Figure 11 shows the optimal design for the CO2 absorption/ stripping unit. The amount of solvent has increased by nearly an order of magnitude compared to the H2S removal. Diameters of the column are very large due to this enormous throughput. Owing to the higher pressure in the stripping column, a lot more heat is needed in the reboiler to strip off the CO2. These results clearly show the difficulty of absorbing CO2 from a H2/CO2 rich stream. Even though the sizes and energy required to remove CO2 seem large, the composition, flow, pressure, and thermal results are similar to those reported in the NETL study. Results for the hydrogen rich stream are almost identical. The difference in composition of CO and CO2 is the result of higher conversion in this simulation’s WGS reactors. The CO2 stream from the SELEXOL unit is reported in the NETL study as over 93 mol % CO2 prior to sequestration. The results in this obtain a higher purity of 96%. This paper has more conservative pressure drops throughout the units so that realistic dynamics and control studies can be achieved. B. Design and Optimization Methodology. The design specification is different than the H2S removal unit, but the methodology is the same with one added step. The absorption column is sized to obtain an exit gas composition of 90 mol % H2. This specification provides a stream pure enough to be sent to a combustion turbine with little greenhouse effect. In the hybrid power/chemical process, a portion of this stream is fed to the methanol plant to provide additional hydrogen since there is not enough hydrogen in the synthesis gas to satisfy the stoichiometry of the methanol reactions. The top graph in Figure 12 shows the solvent flow needed to obtain 90% H2 composition with absorber tray numbers of 14, 30, and 50. The amount of solvent needed does not decrease much while increasing the number of stages in the absorber.

Figure 11. Optimal design of the CO2 removal via physical absorption with DME-PEG.


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Figure 12. (a) Solvent flow and its effect on mole fraction of H2 in the gas outlets at various absorption trays. (b) Optimal stripping column top pressure based on TAC and its effect on reboiler duty and compressor work.

To obtain 90% H2, the required flow rate is 108 100 lb · mol/h of 94% DME-PEG for 14 trays; 106 500 lb · mol/h for 30 trays; and 104 800 lb · mol/h for 50 trays. The diameter of each column was minimized by using twopass trays. Decreasing the number of trays in the stripper also decreased the TAC of the total unit. A minimum of 10 trays was chosen to be reasonable. The optimal feed tray is at stage 8 of the stripper. The one extra step is to find the optimum stripping column pressure. The TAC of the system must be minimized, so varying the top pressure would vary the compressor energy needed to pressurize the CO2 for sequestration, the reboiler energy needed to strip off the CO2, and their corresponding capital costs. The compressor cost is given by the simple economic expression:16 gas compressor capital )

(517.5)(bhp) ( M&S 280 )

0.82

(2.11 + 1)

(5) An M&S index of 1310 was used, and the bhp is the brake horsepower needed in the compressor. A mechanical efficiency of 80% is assumed. Compression energy costs of $16 per GJ were used. The bottom graphs in Figure 12 show that an economic minimum is found at a stripper pressure of 160 psia. As expected, when pressure goes up, compressor work goes down, and reboiler duty increases. Similar to the H2S absorption column, an economic minimum for the CO2 absorption column was found with 10 stripping trays, but 14 absorption trays. Therefore, both the absorption and stripping columns for removing CO2 have small heights and large diameters. The reason for these short columns is the number of stages has very little effect on the required solvent flow rate. Using the real SELEXOL solvent, instead of just DME-PEG, could provide higher selectivity, smaller-diameter columns, and a smaller TAC. The optimal flowsheet of this process is shown in Figure 11. C. Steady-State Calculations for Control Structure. The steady-state criteria studied for the H2S stripping column were

applied to the CO2 stripping column as shown in Figure 13. The biggest difference in temperature is on stage 11 (column bottoms). Also, the sensitivity and SVD analysis of a +0.1% change in reboiler duty give the best results for stage 11. Therefore, the temperature of column bottoms is controlled with the reboiler heat input. The big change in temperature on stage 9 was due to it being the feed tray that introduces a large change in vapor composition. 4. Aspen Dynamics Simulation of the IGCC 4.1. H2S Removal. A. Equipment Sizing. The physical dimensions of the various pieces of equipment must be provided to convert a steady-state simulation into a dynamic one. The various equipment sizes required are provided in Figure 14. The sump of the absorption column uses the same diameter as the column (22.8 feet) and has an entering liquid flow rate of 76 100 ft3/h, so a base height of 31.1 and 22.8 ft in diameter is required to provide 5 min of holdup at 50% level. The stripping column sump is 30.3 feet in height and 24.7 ft in diameter to handle an inlet liquid flow rate of 87 100 ft3/h. The knockout drum and condenser were sized to provide a 5 min holdup at 50% level and their dimensions are also shown in Figure 14. The heat exchanger dynamics are modeled by the weight of the tubes and volume of the tubes. B. Control Structure. With all equipment sized, the steadystate Aspen Plus file is exported to Aspen Dynamics as a pressure-driven simulation. Figure 14 gives the simple PI control structure developed to maintain the 5 ppm H2S in the absorber gas product. The loops are described below: (1) The flow rate of the inlet synthesis gas is manipulated to control pressure in the upstream gasifiers. (2) The temperature of the inlet gas is controlled by manipulating heat removal before entering the water knockout drum. (3) Level in the knockout drum is level controlled by the bottoms flow rate.

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Figure 13. Steady-state criteria for control structure in the CO2 stripping column.

Figure 14. H2S absorption/stripping equipment sizes and control structure.

(4) Flow rate of the solvent is ratioed to the flow rate of the absorber feed. (5) This ratio is adjusted by a composition controller that maintains the H2S impurity in the gaseous product leaving the absorber top. (6) Absorber pressure is controlled by the control valve on the gas leaving the absorber. (7) Liquid level in the base of the absorber is controlled by manipulating the bottoms flow rate. (8) Stripper reflux controls the liquid level in the reflux drum.

(9) Stripper pressure is controlled by the control valve on the gas stream leaving the reflux drum. (10) Make-up water is ratioed to the flow rate of the distillate flow rate. Most of the water lost from the system goes out in the vapor distillate from the stripping column. (11) Stripper base level is controlled by the flow rate of makeup DME-PEG. The flow rate of this makeup stream is very small compared to the solvent flow rate, so the base level of the stripping column is not controlled tightly. However, the 5 min base holdup is sufficient to handle dynamic disturbances


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Figure 15. Controller faceplates in Aspen Dynamics of the H2S absorption/stripping unit.

when the solvent flow rate changes, which eventually changes the feed to the solvent-recovery column to bring the base level back to near its steady-state level. (12) Stripper reboiler duty is manipulated to control stage 5 temperature. Figure 15 shows the controller faceplates. Notice that flow controllers for the solvent flow into the absorber and the makeup water are on cascade. A 1 min deadtime is inserted in the stage 5 temperature controller “S1-T5_TC.” The concentration controller, “H2S_CC,” has a transmitter range of 5-20 ppm of H2S and a controller output range of 0.25-1.5 (ratio of absorber feed flow to solvent flow on a molar basis). A 3 min deadtime for the loop is used. The output of the concentration controller is the ratio manipulating the set point of the solvent flow control as shown in “S1_S/F” multiplier. Input signal 1 is the gas flow into the absorption column, input signal 2 is the OP of the composition controller, and the output signal is the set point of the solvent flow controller. The tuning constants for the H2S composition controller are Kc ) 0.16 and τI ) 30.36 min determined by a relay-feedback test and the Tyreus-Luyben rule.17 Using relay-feedback testing and Tyreus-Luyben tuning gives controller constants for the temperature and concentration controllers. Table 1 provides these constants. Level controllers are proportional-only controllers with a gain of 2. C. Results. The effectiveness of the control structure is evaluated by disturbing the process with feed flow rate and feed composition disturbances. Stable regulatory control is achieved for this system. Figure 16a gives results for 10% step changes in the flow. Figure 16b gives the results for a (0.4 mol % H2S composition change from 0.56 to 0.96 and from 0.56 to 0.16 mol % (step change in H2S flow rate: 400 lb · mol/h to 685 and 115 lb · mol/h). Figure 16c gives the response of the composition disturbance if the composition controller was in open loop (constant S/F ratio). If more sulfur is present in the feed, it limits the amount of H2 produced by forming H2S. Therefore, the amount of hydrogen was adjusted to keep constant flow when performing composition changes. Solid lines are for a distur-

bance increase, while dashed lines are for a decrease. The disturbances are made at time equals 6 min on each figure. The composition controller maintains exit gas composition quite close to the 5 ppm H2S specification for the 10% flow changes as shown in Figure 16 (a). A peak of about 5.9 ppm occurs 20 min after the initial disturbance. Composition returns to steady state within two hours. The solvent to feed molar flow ratio (S/F) varies in the range of 0.63 to 0.69. The sulfur content is kept well below poisoning levels for the methanol and WGS catalysts downstream. More reboiler duty needs to be added to remove the increased amount of H2S. Considering the huge variability of sulfur content in coal or any feedstock of gasification, a large disturbance in H2S composition is needed to show system robustness. The composition was nearly doubled (0.56-0.96 mol % H2S) or cut by a factor of 4 (0.56-0.16 mol % H2S) and the responses are shown in Figure 16b. The dynamics are slower than with the flow disturbances, but sulfur content leaving the absorption column is successfully controlled near 5 ppm and returned to steady state within 4 h. A maximum concentration of H2S of about 5.2 is reached after an hour from the initial disturbance. With an increase in composition, a larger solvent-to-feed ratio is needed and vice versa for a decrease when steady state is obtained. The stripper distillate indicates that the poisonous gases are successfully absorbed and stripped when disturbances are applied. The ppm of hydrogen sulfide has an inverse response when the system is initially disturbed with composition changes. This is due to the selectivity of the solvent toward H2S over CO2. Figure 16c demonstrates the response of the composition change with the composition controller in open loop (S/F is a constant). The initial inverse response of H2S is clearly shown as is the increase in CO2 composition leaving the top of the absorber. Since the solvent-to-feed ratio is not changed (composition controller on manual), the H2S in the offgas eventually increases to 7.2 ppm when it reaches steady state within 4 h. Steady state is not shown simply because of the needed emphasis on the

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Ind. Eng. Chem. Res., Vol. 49, No. 10, 2010 Table 1. H2S Absorption/Stripping Controller Tuning Constants H2S_CC

controlled variable manipulated variable transmitter range controller output range Kc τI dead time

ppm H2S ) 5 S/F 0-20 0.25-1.5 0.16 30.36 min 3 min

S1-HX4_TC


Tsolv ) 110 °F QHX4 90-150 °F 0-59.18 MW removed 0.63 1.09 min 1 min

S1-HX3_TC


Tcond ) 110 °F QHX3 90-150 °F 0-25.4 MW removed 0.11 0.80 min 1 min

S1-TR5_TC

controlled variable manipulated variable transmitter range controller output range Kc τi dead time

TTray5 ) 197.9 °F QReb 170-230 °F 0-70.6 MW added 1.41 5.00 min 1 min

Table 2. WGS Controller Tuning Constants

Figure 16. Graphic shows (a) 10% flow rate disturbances of the H2S absorption/stripping column; (b) 0.4 mol% H2S disturbance of the H2S absorption/stripping column.

inverse response in open-loop. The S/F ratio, sump level, and makeup DME-PEG are not shown simply because they are held constant. 4.2. Water-Gas Shift Reactors. A. Equipment Sizing. Equipment sizing is shown in Figure 17 for the water-gas shift

T1_TC


T1 ) 320 °F QHX1 220-420 °F 0-9.9 MW added 18.51 27.72 min 1 min

T2_TC


T2 ) 310 °F QHX2 210-410 °F 0-156.7 MW removed 0.44 1.23 min 1 min

reactors. The reactor sizing assumes an aspect ratio of four (L/ D). Heat exchanger dynamics are modeled by the weight of the tubes and volume of the tubes needed to obtain the desired temperature. B. Control Structure. With all equipment sized, the steadystate Aspen Plus simulation of the WGS is exported to Aspen Dynamics as a pressure-driven simulation. Figure 17 gives the simple PI control structure developed. To maintain a 99% conversion of CO to H2, the following control structure is implemented: (1) Flow rate of the inlet syngas to WGS is controlled to maintain pressure in the upstream H2S removal absorber column. (2) Steam flow rate is manipulated to give a steam-to-gas flow ratio of 0.85. (3) Inlet temperature of the first WGS reactor is controlled by manipulating the heat input of the heat exchanger WHX1. (4) Inlet temperature of the second WGS reactor is controlled by heat removal of the heat exchanger WHX2. (5) Pressure of the reactors are maintained by a valve downstream. Relay-feedback testing and Tyreus-Luyben tuning were used to obtain controller constants for both temperature controls. Table 2 provides these constants.


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Figure 17. Water-gas shift reactors equipment sizes and control structure.

Figure 18. Load disturbances in the WGS reactors.

C. Results. The effectiveness of the control structure is evaluated by disturbing the process with load changes. The steam is ratioed to the amount of synthesis gas coming into the reactor at 0.828. Figure 18 gives results for 10% step changes in the set-point of the inlet flow controller. The disturbance is made at the time of 6 min. As expected, the dynamics are quite fast with peak temperature deviations occurring in less than 1 min for T1 and 3 min for T2. Peak temperatures for the inlet of the high-temperature shift are about 2 °F, while the low-temperature peaks are 10 °F. An increase in synthesis gas flow decreases the temperature of T1 because more of the cold synthesis gas is present. This lower temperature promotes a higher conversion in the first reactor and a higher exit temperature. Because of the higher exit temperature out of the first reactor, T2 increases with an increase in synthesis gas. It takes about 15 min to get back to the steady-state condition of T2 ) 310 °F. With an increase inlet temperature for reactor 2, a decrease in total conversion occurs as expected because the equilibrium constant is smaller.

When flow rate increases, the CO to H2 conversion goes down as expected. However, even with this 20% throughput change, the conversion is still quite close to 99%. 4.3. CO2 Removal. A. Equipment Sizing. The various equipment sizes required for the CO2 absorption and stripping unit are shown in Figure 19. These equipment sizes seem rather large and would differ if the real SELEXOL solvent was available in the Aspen component library instead of just dimethyl ether of pentaethylene glycol (DME-PEG). The sump in the absorption and stripping column would need a length of 15.6 and 17.6 feet for 277 400 and 278 700 ft3/h, respectively. The lengths are much smaller due to the enormous diameters of the column. The knockout drum and condenser were sized to provide a 5 min holdup at 50% level. The heat exchanger dynamics are modeled by the weight of the tubes and volume of the tubes needed to obtain the desired temperature. B. Control Structure. The control structure is identical to the H2S absorption and stripping columns. The hydrogen

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Figure 19. CO2 absorption/stripping equipment sizes and control structure. Table 3. CO2 Absorption/Stripping Controller Tuning Constants CO2_CC


mf H2 ) 0.9 S/F 0.8-1.0 0-10 0.33 43.56 min 3 min

S2-HX4_TC


Tsolv ) 110 °F QHX4 90-150 °F 0-137.8 MW removed 18.31 5.28 min 1 min

S2-HX3_TC


Tcond ) 110 °F QHX3 90-150 °F 0-144.5 MW removed 0.24 1.25 min 1 min

S2-TR11_TC


TTray11 ) 300.8 °F QReb 200-400 °F 0-270.4 MW added 10.03 11.88 min 1 min

concentration controller manipulates the feed to solvent ratio. Relay-feedback testing and Tyreus-Luyben tuning provide the controller constants. These constants are provided in Table 3. The level is controlled by proportional-only controllers with a gain of 2. C. Results. The effectiveness of the control structure is evaluated by disturbing the process with load and composition changes. Figure 20a gives results for 10% step changes in the flow at time equals 6 min. Figure 20b gives the results for a (10 mol % CO2 composition change of 29-39 and 19 mol %. When altering the composition of CO2, the balance was made up with hydrogen.

Figure 20. Graphic shows (a) 10% inlet flow disturbances in the CO2 absorption/stripping unit, (b) CO2 concentration disturbances in the CO2 absorption/stripping unit.


The regulatory control structure maintains a 90 mol % rich hydrogen stream when experiencing flow rate and concentration disturbances. A low of 0.895 mol fraction is obtained when increasing the flow rate into the system. The dynamics are fast and steady-state conditions are obtained within 2 h. The amount of solvent needed increases as the flow rate increases in the system. A higher reboiler duty is due to the increased amount of CO2 removed. As expected, an increase in CO2 molar composition increases the flow rate out of the stripping column distillate while decreasing the amount out of the absorber tops. Steady state results are obtained within 3 h. 5. Conclusion This paper has presented a method for generating a dynamic model of an entire IGCC that runs in the current version of Aspen Dynamics 2006.5. This model adequately captures the macroscale thermal, pressure, flow, and composition dynamics. This is needed to perform dynamic studies on the control of the hybrid power/chemical plant and its ability to handle unavoidable swings in both feed and power demand. Acknowledgment Financial support for this work provided by the Zisman Family Foundation is greatly appreciated. Literature Cited (1) Higman, C.; van der Burgt, M. Gasification; Elsevier: New York, 2003. (2) Mayfield, G. G.; Agreda, V. H. Eastman chemical process for acetic anhydride from coal. Energy Prog. 1986, 6, 214–218. (3) Luyben, W. L. Chemical Reactor Design and Control; John Wiley & Sons: New York, 2007.

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(4) Robinson, P. J.; Luyben, W. L. Simple dynamic gasifier model that runs in Aspen Dynamics. Ind. Eng. Chem. Res. 2008, 47 (20), 7784–7792. (5) National Energy Technology Laboratory. Cost and Performance Baseline for Fossil Energy Power Plants Study. Volume 1: Bituminous Coal and Natural Gas to Electricity, May 2007; www.netl.doe.gov. (6) Aspen Plus 2004-1, Getting Started Modeling Processes with Solids; Aspen Technology, Inc.: Cambridge, MA, 2004. (7) Talley, P. Capitalizing on the Emerging Coal Gasification Market. Aspen Technology Webinar, December 11, 2007. (8) Breckenridge, W., Holiday, A., Ong, J. O. Y., Sharp, C. Use of SELEXOL process in coke gasification to ammonia project, paper presented at the Laurance Reid Gas Conditioning Conference, University of Oklahoma, Norman, OK,Feb 27-Mar 1, 2000. (9) Twigg, M. V., Ed. The Catalyst Handbook; Wolf Scientific: London, 1989. (10) Jamiel, A.; Seymour A. F. Solvent Composition Useful in Acide Gas Removal from Gas Mixtures. U.S. Patent 3,737,392, June 5, 1973. (11) Gary, J. H. Handwerk, G. E. Petroleum Refining Technology and Economics; Marcel Dekker, Inc.: New York, 1984. (12) Coulson, J. M.; Richardson, J. F. Particle Technology and Separation Processes; Chemical Engineering, Vol. 2; Pergamon Press: Oxford, 1999. (13) Turton, R.; Bailie, R. C.; Whiting, W. B.; Shaeiwitz, J. A. Analysis, Synthesis, and Design of Chemical Processes, 2nd ed.; Prentice Hall PTR: Upper Saddle River, NJ, 2003. (14) Luyben, W. L. Distillation Design and Control Using Aspen Simulation; John Wiley & Sons: New York, 2006. (15) Choi, Y.; Stenger, H. G. Water-gas shift reaction kinetics and reactor modeling for fuel cell grade hydrogen. J. Power Sources 2003, 124, 432–439. (16) Douglas, J. M. Conceptual Design of Chemical Processes; McGrawHill, Inc.: New York, 1988. (17) Luyben, W. L. Plantwide Dynamic Simulators in Chemical Processing and Control; Marcel Dekker, Inc.: New York, 2002.

ReceiVed for reView October 2, 2009 ReVised manuscript receiVed March 30, 2010 Accepted April 6, 2010 IE901549S

Stripping, WaterâGas Shift Reactors, and CO2

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Stripping, WaterâGas Shift Reactors, and CO2