Ind. Eng. Chem. Res. 2008, 47, 6759–6764
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Effect of Reaction Temperature on the Performance of Thermal Swing Sorption-Enhanced Reaction Process for Simultaneous Production of Fuel-Cell-Grade H2 and Compressed CO2 from Synthesis Gas Ki Bong Lee, Michael G. Beaver, Hugo. S. Caram, and Shivaji Sircar* Department of Chemical Engineering, Lehigh UniVersity, Bethlehem, PennsylVania 18015
A novel cyclic thermal swing sorption-enhanced reaction (TSSER) process concept was recently proposed for the simultaneous production of fuel-cell-grade H2 and compressed CO2 from synthesis gas containing CO and H2O. The process carried out the catalytic water-gas shift (WGS) reaction (CO + H2O T CO2 + H2) with simultaneous removal of CO2 from the reaction zone by a reversible, water-tolerant, CO2-selective chemisorbent in order to circumvent the thermodynamic limitation of the WGS reaction and enhance the rate of the forward reaction. The chemisorbent was periodically regenerated using the principles of thermal swing adsorption by purging the sorber-reactor with superheated steam at different pressures and temperatures. Several intermediate process steps were employed to produce a pure and compressed CO2 byproduct during the thermal desorption process. The present work reports (a) new experimental data demonstrating the concept of the sorption-enhanced WGS reaction at different temperatures using a commercial WGS catalyst and Na2Opromoted alumina as the CO2 chemisorbent and (b) the effect of the sorption-reaction temperature on the TSSER process performance estimated by model simulation. Relatively slower kinetics of the sorption-enhanced WGS reaction imposes a lower bound (∼200 °C), whereas the thermal stability of the chemisorbent and the use of carbon steel sorber-reactors set the upper bound (∼550 °C) of temperatures for practical operation of the TSSER process. Simulated process performances (sorption-reaction at 200 and 400 °C and regeneration at 550 °C) show that the operation of the sorption-reaction step at 200 °C increases the H2 and CO2 productivities of the process by ∼38% and 35%, respectively, without changing (a) the number of moles of H2 produced per mole of CO in the feed gas or (b) the net CO2 recovery as a compressed byproduct gas. The total steam duty for the sorbent regeneration increases by ∼14% for operation at the lower sorption-reaction temperature. Another major benefit of operation at the lower reaction temperature is a very large increase in the pressure of the CO2 byproduct (e.g., 40 and 21 atm at 200 and 400 °C, respectively) when the reactor feed gas contained 20% CO + 80% H2O at a total pressure of 15 atm. Introduction The recently proposed thermal swing sorption-enhanced reaction (TSSER) process concept is designed to produce fuelcell-grade H2 directly and compressed CO2 as a byproduct gas by reacting CO and H2O from synthesis gas produced by the gasification of coal (after removal of trace sulfur impurities).1,2 A single sorber-reactor packed with an admixture of a water-gas shift (WGS) reaction catalyst and a CO2-selective chemisorbent is used for this purpose. The byproduct CO2 of the exothermic WGS reaction (CO + H2O T CO2 + H2, ∆HR ) -41 kJ/mol), which is thermodynamically controlled, is removed from the reaction zone by chemisorption to directly produce fuel-cell-grade H2 from the sorber-reactor. The chemisorbed CO2 is then periodically removed by a superheated steam purge using the principles of thermal swing adsorption so that the chemisorbent can be reused. Part of the desorbed CO2 is withdrawn as a compressed byproduct gas by employing several complementary intermediate steps such as a high-pressure CO2 rinse step and a high-pressure steam purge step. Figure 1 provides a schematic description of the five steps of the TSSER process consisting of (a) sorption-enhanced reaction; (b) highpressure CO2 rinse; (c) batch heating step; (d) high-pressure steam purge; and (e) multitask regeneration consisting of substeps such as depressurization, cooling, low-pressure steam purge, and pressurization. The sorption-reaction step (step a) * To whom correspondence should be addressed. E-mail: sircar @aol.com. Tel.: (610) 758-4469. Fax: (610) 758-5057.
and the CO2 rinse step (step b) are carried out isobarically at pressure PR and temperature TR. The sorber-reactor is indirectly heated batchwise from a temperature of TR to TH in step c, when the gas-phase pressure increases from PR to PH. The isobaric high-pressure steam purge (step d) is carried out at pressure PH and temperature TH. The sorber-reactor is depressurized from PH to a near-ambient pressure level as it undergoes indirect cooling from TH to TR, is purged with steam at that near-ambient pressure and temperature TR, and is finally repressurized from ambient pressure to PR using steam at TR in step e. A more detailed description of the process, the key functions of the individual process steps, and the key advantages of the concept can be found elsewhere.1,2 However, it should be emphasized here that the process is designed to satisfy one of the major goals of the hydrogen economy vision by the U.S. National Academy of Engineering and the U.S. National Research Council that “hydrogen can be produced from domestic energy sources in a manner that is affordable and environmentally benign”.3 Two recently developed, reversible, water-tolerant, and highly CO2-selective chemisorbents have been extensively characterized and employed to simulate the TSSER process performance. They are K2CO3-promoted hydrotalcite1,4 and Na2O-promoted alumina.2,5 Both materials offer the unique ability to reversibly and selectively chemisorb CO2 in the presence of steam, CO, H2, and CH4 with decent CO2 working capacity, low isosteric heats of chemisorption for CO2, and relatively fast CO2 chemisorption kinetics. Both materials are thermally stable
10.1021/ie071372k CCC: $40.75 2008 American Chemical Society Published on Web 07/29/2008
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Figure 1. Schematic representation of the five-step TSSER process concept.
below 600 °C. Thus, these materials offer the unique opportunity of using superheated steam as the desorbent gas for the removal of CO2 from these materials using a pressure or thermal swing mode of operation in the sorption-enhanced reaction (PSSER or TSSER) concepts.1,2,6,7 Process simulation results showed that both materials can be used to simultaneously produce fuel-cell-grade H2 and pure compressed CO2 from synthesis gas by the TSSER concept.1,2 These studies also showed that operation of the process using promoted alumina as the CO2 chemisorbent in conjunction with a synthesis gas reactor feed containing 1:4 CO/H2O at 15 atm pressure (PR) and reaction (TR) and regeneration (TH) temperatures of 400 and 550 °C, respectively, yielded (a) ∼18% lower H2 productivity at a slightly reduced CO to H2 conversion, (b) comparable productivity of compressed CO2 byproduct gas, (c) higher recovery of CO2 byproduct, and (d) lower CO2 product pressure (22.6 vs 33.1 atm). However, the steam duty for regeneration of the chemisorbent was reduced by ∼50%.2 It is obvious that the CO2 working capacity of the chemisorbent in the TSSER process and, consequently, its overall performance can be improved by lowering the temperature of the sorption-reaction step (TR) and increasing the temperature swing (TH - TR) between the sorption and desorption steps, provided that the kinetics of the WGS reaction and the kinetics of CO2 chemisorption are acceptable at the chosen value of TR. Lowering the reaction temperature also favors the thermodynamic conversion of CO to H2 by the exothermic WGS reaction. The reaction temperature (TR) can not be reduced significantly below 400 °C when the promoted hydrotalcite is used as the CO2 chemisorbent because of slow kinetics of chemisorption. On the other hand, CO2 chemisorption on the promoted alumina is relatively fast even at 250 °C.2 Consequently, we experimentally investigated the dynamics of the SER concept and simulated the performance of the TSSER process at different temperatures using the promoted alumina as the chemisorbent. The results are discussed below. Characteristics of CO2 Chemisorption on Na2O-Promoted Alumina We briefly review the key thermodynamic and kinetic characteristics of chemisorption of CO2 on promoted alumina
Table 1. Parameters of Chemisorption-Surface Reaction Model for the Sorption of CO2 on Na2O-Promoted Alumina T (°C)
m (mol/kg)
a
KC (atm-1)
KR (atm-a)
KH ) mKC [mol/(kg atm)]
250 350 450
0.295 0.295 0.295
2.0 1.7 1.5
536 48.3 8.47
8 2 0.73
158 14.2 2.45
that were reported earlier,5 because they are used in the data analysis and process simulations presented in this work. CO2 Chemisorption Isotherms. The experimentally measured isotherms for the chemisorption of pure CO2 on the promoted alumina at 250, 350, and 450 °C in the pressure range of 0-3 atm can be described very well by a novel, analytical, chemisorption-surface complexation reaction model4,5 n*(P,T) )
mKCP[1 + (a + 1)KRPa] [1 + KCP + KCKRP(a+1)]
(1)
where n* (mol/kg) is the specific amount of CO2 chemisorbed at a gas-phase CO2 partial pressure of P (atm) at temperature T (K). The parameters of model eq 1 are the saturation CO2 chemisorption capacity of the chemisorbent surface, m (mol/ kg); the equilibrium constant for surface chemisorption of CO2, KC (atm-1); the equilibrium constant for the additional surface complexation reaction between the gaseous and chemisorbed molecules of CO2, KR (atm-a); and the stoichiometric coefficient for the complexation reaction, a. The thermodynamic constants (KC and KR) are exponential functions of temperature qC d ln KC )- 2 dT RT KC ) KC° exp(qC/RT)
∆HR d ln KR )- 2 dT RT KR ) KR° exp(∆HR/RT)
(2) (3)
where qC and ∆HR (kJ/mol) are the molar isosteric heat of chemisorption and the heat of additional surface reaction, respectively. KC° (atm-1) and KR° (atm-a) are constants. The values of the model parameters are given in Table 1. The isosteric heat of chemisorption (qC) and the heat of surface reaction (∆HR) of CO2 were estimated to be 64.9 and 37.5 kJ/ mol, respectively. The corresponding pre-exponential constants (eq 3) were found to be 0.000 164 atm-1 and 0.001 417 atm-a, respectively. The parameter a is an exponential function of temperature
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[ 4.39 (kJ/mol) ] RT
a ) 0.72644 exp
(4)
The relatively low heats of chemisorption and surface complexation are unique properties of this chemisorbent. They permit reversible chemisorption and relatively easy desorption of CO2. CO2 Chemisorption Kinetics. The kinetics of chemisorption of CO2 on the promoted alumina can be described by the conventional linear driving force (LDF) model8 dn(t) ) k[n*(t) - n(t)] (5) dt Equation 5 gives the local rate of adsorption/desorption of CO2 from an inert gas on the chemisorbent. The variable n(t) is the specific amount (mol/kg) of CO2 chemisorbed at time t, and n*(t) is the specific equilibrium adsorption capacity of CO2 in the prevailing gas phase at pressure Po, temperature T(t), and CO2 mole fraction of y(t). n* is given by eq 1 for a given set of values of Po, T, and y. The parameter k (time-1) is the LDF mass-transfer coefficient for CO2 adsorption/desorption. Table 2 lists the experimentally measured values of the parameter k at different temperatures.2,5 They were estimated by the dynamic model analysis of breakthrough data for sorption and desorption of CO2 from a promoted alumina column. A continuous stirred tank reactors- (CSTRs-) in-series model, described in the next section, was used for this purpose. The CO2 mass-transfer coefficients on the promoted alumina were found to be independent of CO2 concentrations in the range of test data (0.2 < y° < 0.6) for a given temperature, and they were identical for both the sorption and desorption of CO2.2,5 The coefficients were moderate function of temperature as shown in Table 2. CSTRs-in-Series Model. The well-known CSTRs-in-series model was adapted for dynamic simulation of (a) the sorption and desorption of CO2 from a packed bed of promoted alumina and (b) the sorption-reaction and thermal regeneration steps of the TSSER process using a packed bed of an admixture of promoted alumina and WGS catalyst. The model simultaneously solves the algebraic and ordinary differential equations describing the component and overall mass balances and the energy balance for each tank in series. The Matlab function ODE15s, which is a variable-order solver based on the numerical differential formulas, was used for the mathematical integration process. It was found that 250 tanks in series were adequate to simulate the sharp sorption-column breakthrough data for sorption of CO2 on promoted alumina.5 The local CO2 chemisorption isotherm on the promoted alumina (eq 1), the local rate of CO2 chemisorption (eq 5), and an analytical model describing the WGS reaction kinetics were used in the CSTRsin-series model.9 The model is able to simulate isothermal or adiabatic operation, and it can also account for heat exchange between the packed bed and the surroundings using an overall wall heat-transfer coefficient. More details about the model can be found elsewhere.10 Experimental Demonstration of WGS Sorption-Enhanced Reaction (SER) Concept A single-column sorption apparatus was used to experimentally demonstrate the concept of sorption-enhanced WGS reaction. The key components of the test unit included a sorption-reaction column (inner diameter, D ) 1.73 cm; length, L ) 86.4 cm) that was surrounded by three different heating tapes with a feedback temperature controller, gas heating and cooling exchangers, flow measuring devices (F), and switch
Figure 2. Transient effluent gas concentration (dry and argon-free basis) from the sorber-reactor at 150 °C.
Figure 3. Transient effluent gas concentration (dry and argon-free basis) from the sorber-reactor at 200 °C.
valves. A layer of insulation was wrapped around the column over the heating tapes. Several thermocouples were used to monitor the column temperature at three different heights (midpoint and gas entrance and exit ends). A more detailed description of the apparatus can be found elsewhere.4,5 The tubular reactor was packed with an admixture of a commercial WGS catalyst (Cu/ZnO/Al2O3 low-temperature-shift catalyst produced by Sud Chemie of Basel, Switzerland) and promoted alumina as the CO2 chemisorbent. The ratio of chemisorbent to catalyst was 70:30 (wt %). The reactor was heated to and maintained at a constant temperature (150-450 °C) using the heating tapes. It was initially filled with argon at 1 atm. A gaseous mixture consisting of 10.9% CO, 56.4% H2O, and 32.7% Ar at ambient pressure was preheated to the reaction temperature and passed through the reactor. The total gas flow rate was 458.6 cm3/min. The composition of the effluent gas was continuously analyzed using a quadrupole mass spectrometer. Figures 2–4 show the transient reactor effluent gas compositions of H2, CO2, and CO on a dry and argon-free basis from these runs at reactor temperatures of 150, 200, and 300 °C, respectively. Most of these experiments were repeated several times to check their reproducibility. It can be seen from Figure 2 that very little pure H2 was produced as the reactor effluent gas at the reaction temperature of 150 °C. The H2 mole fraction in the effluent gas was initially unity, but it continuously decreased as the experiment progressed and finally leveled off at a value of ∼0.43. The effluent gas was initially free of CO, but the CO concentration progressively
6762 Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008
Figure 4. Transient effluent gas concentration (dry and argon-free basis) from the sorber-reactor at 300 °C. Table 2. LDF Mass-Transfer Coefficients for the Adsorption/ Desorption of CO2 on Promoted Alumina temperature (°C)
k (min-1)
250 350 450
4.0 5.0 5.8
Table 3. Inlet Gas Flow Rates [mmol/(cm2 min)] for Different Steps of the TSSER Process step
case a: 400 °C
case b: 200 °C
a b d e
223.5 258.0 351.0 37.0
309.2 318.0 550.0 63.0
increased before it leveled off at ∼0.15. The effluent gas was free of CO2 for a period of time (∼30 min), and then the CO2 mole fraction rose fairly sharply to be equal to the H2 mole fraction at longer times. The effluent gas H2 and CO profiles from the reactor at 150 °C indicate that the rate of the WGS reaction on the catalyst used in the test was low, which caused nearly immediate and diffused breakthrough of CO from the reactor. On the other hand, the CO2 produced by the WGS reaction was rapidly chemisorbed by the promoted alumina at 150 °C. This explains the absence of CO2 from the effluent gas for a substantial period of time. The CO2 broke through the column when its capacity on the promoted alumina at the prevailing conditions (thermodynamically governed product composition of the WGS reaction at the feed conditions) was exhausted. The effluent gas consisted of only CO and H2 (dry and Ar-free basis) until the CO2 broke through. Consequently, the decrease of the H2 mole fraction in the effluent gas during this period corresponded to the breakthrough of unreacted CO from the reactor due to the slow kinetics of the WGS reaction. The dashed lines in Figure 2 show the simulated reactor effluent gas profiles at 150 °C generated by using the CSTRs-in-series model. They describe the experimental data only qualitatively. The key reason for this discrepancy is that the model used to describe the kinetics of the WGS reaction9 fails to trace the slow kinetics of the reaction at 150 °C. Figures 3 and 4 show the experimental reactor effluent gas composition profiles (circles) and the corresponding model simulations (dashed lines) for operation of the sorber-reactor at 200 and 300 °C, respectively. The other conditions of these tests were identical. These figures show that COx-free H2 was produced from the reactor by reaction between CO and H2O (WGS reaction) for a
period of time. This was caused by the very high rate of the WGS reaction and the simultaneous removal of CO2 from the reaction zone by chemisorption on the promoted alumina. The thermodynamic limitation of the WGS reaction was circumvented, and a pure H2 product gas (dry and Ar-free basis) was produced. This sorption-enhanced reaction process continued until the CO2 chemisorption capacity of the promoted alumina was exhausted. Thereafter, the H2 concentration in the effluent gas rapidly fell, and the CO and CO2 concentrations rapidly rose to a composition (∼49.85% H2 + 49.85% CO2 + 0.3% CO on a dry and argon-free basis) that was dictated by the thermodynamics of the WGS reaction at the conditions of the test (feed gas composition and reactor temperature). Very high rates of the catalytic WGS reaction and subsequent CO2 chemisorption on the promoted alumina at these temperatures caused CO and CO2 breakthrough to occur simultaneously, thereby enabling the SER concept to produce pure H2 as the reactor effluent gas. Figures 3 and 4, therefore, provide an experimental demonstration of the SER concept for the production of pure H2 by the WGS reaction carried out in the presence of Na2O-promoted alumina as a CO2 chemisorbent. The dashed lines in Figures 3 and 4 represent the reactor effluent gas composition profiles for the conditions of the experiment simulated by the CSTRsin-series model. It can be seen that the simulated reactor effluent gas profiles trace the experimental profiles very well, indicating the validity of the model for this application at reaction temperatures above 150 °C and justifying the use of the independently measured CO2 chemisorption equilibria and kinetics on the promoted alumina as inputs to the dynamic simulation model. It can also be seen from Figures 2–4 that the breakthrough times for CO2 from the sorber-reactor progressively increased as the reaction temperature decreased from 300 to 150 °C because the CO2 chemisorption capacity of the promoted alumina increased as the temperature decreased. These times were 30, 26, and 19 min at reactor temperatures of 150, 200, and 300 °C, respectively. These data clearly demonstrate that the sorption-reaction step of the TSSER process can be operated efficiently without kinetic limitations for sorption and reaction when the reactor temperature is g200 °C. Simulated Performance of the TSSER Process at Different Reaction Temperatures We simulated the performance of the TSSER process for simultaneous production of fuel-cell-grade H2 and compressed CO2 from synthesis gas using the CSTRs-in-series model. The simulation used a shell-and-tube heat exchanger as the sorber-reactor. The tube side (1.73 cm in diameter and 500 cm in length) was packed with an admixture of the WGS catalyst and the promoted alumina (90 wt %). The relative ratio of catalyst to chemisorbent admixture inside the sorber-reactor was an independent variable in the model. The packed tube was indirectly heated or cooled using a cross-flow of superheated steam at ambient pressure in the shell side. The reactor feed gas was a mixture of CO (20 mol %) and H2O at a pressure of 15 atm (PR) and a temperature (TR) of 400 °C (case a) or 200 °C (case b). It was assumed that the LDF mass-transfer coefficients for CO2 chemisorption were functions of temperature only and were not affected by changes in the CO2 concentration in the gas phase. The highest regeneration temperature (step d) was 550 °C (TH) for both cases. The design cycle times for process steps a-e were 5.0, 0.5, 5.0, 5.0, and
Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008 6763
Figure 5. Simulated CO2 loading profiles inside the sorber-reactor.
Figure 6. Simulated CO2 desorption profiles from the sorber-reactor.
4.5 min, respectively. Table 3 lists the inlet gas mass flow rates (based on the empty cross section of the reactor) for the steps. Simulation Results. The process was operated to directly produce a fuel-cell-grade H2 product containing ∼10 ppm CO at 15 atm pressure during step a and a pure and compressed CO2 byproduct at pressure PH (>PR) during step d irrespective of the reaction temperature. Two key simulation results are described below. Figure 5 shows the simulated CO2 loading profiles inside the sorber-reactor at the end of steps a and e in terms of the specific CO2 loading amount (nCO2, mol/kg) on the promoted alumina as a function of the dimensionless distance (L/Lc) inside the reactor. The variable L is the actual distance inside the reactor from the feed end, and Lc is the total length of the packed bed in the reactor. The areas under these profiles represent the net cyclic CO2 working capacities of the TSSER process. The figure shows that the local CO2 loadings inside the sorber-reactor during step a of the process are much higher when the reaction temperature is lower (200 °C), as expected. However, the CO2 loading profiles inside the reactor at the end of step e of the process are nearly identical because they are governed by the final regeneration temperature, which is the same for both cases. Consequently, the net cyclic CO2 working capacity of the process and the net amount of hydrogen produced during step
a of the process are much higher for the lower-temperature sorption-reaction case (case b). The residual amount of CO2 in the sorber-reactor at the end of step e was only ∼2.0% of the total CO2 present in the sorber-reactor at the end of step b. Such stringent cleanup of the sorber-reactor during steps d and e is necessary to produce high-purity H2 product during step a without sacrificing the CO2 working capacity. Figure 5 also shows that more than 40% of the sorber-reactor at the H2 product end is free of CO2 at the end of step e, as required to maintain the purity of the H2 product during the next cycle. Figure 6 shows the fraction of CO2 present in the sorber-reactor at the end of step b that was desorbed during steps d and e of the TSSER process as a function of the cumulative amount of steam purge during these steps. The fraction of CO2 desorbed during step d was produced at a total pressure of PH and could be recovered as nearly pure CO2 at that pressure by condensing out the steam in a constant-pressure condenser. The fraction of CO2 desorbed during step e was produced at a near-ambient total pressure. It could be wasted or recovered after condensing out the water and recompressing. It can be seen from Figure 6 that nearly 90% of the CO2 present in the sorber-reactor was desorbed during step d for both cases. This high fractional CO2 removal during step d was deliberately chosen in the design to increase the recovery of compressed byproduct CO2 by the TSSER process. More steam purge was necessary to achieve this for case b because of higher initial amount of CO2 in the sorber-reactor. However, it can be seen from Table 4 that the total amounts of steam required per unit amount of H2 product are comparable for cases a and b. The amounts of steam required to remove another 8% of CO2 initially present in the sorber-reactor (step e) were nearly same for both cases. Comparative Performance of the Two Cases. The simulated performances of the TSSER process using reaction temperatures of 200 and 400 °C are compared in Table 4. Fuelcell-grade H2 and compressed pure CO2 byproduct gas can be produced in both cases. However, the higher cyclic CO2 working capacity offered by the TSSER process and the favorable thermodynamics of the exothermic WGS reaction at the lower reaction temperature (200 °C) translate into ∼38% more H2 product and ∼35% more CO2 byproduct than obtained in the higher reaction temperature (400 °C). Conducting the reaction step at a lower temperature also produces the CO2 byproduct at a much higher pressure (∼40.5 atm) during step d than the higher-temperature case (∼21.5 atm) when the feed synthesis gas contains 20% CO2 + 80% H2O at 15 atm total pressure. This is a result of the larger thermal swing in the process for case b. The increased pressure for the CO2 byproduct of the TSSER process operated at 200 °C might be advantageous for subsequent recompression and sequestration of CO2. The increased wall thickness of the reactor tubes needed to withstand the higher pressure is only marginal.
Table 4. Comparative Performances of the TSSER Process at Different Reaction Temperatures parameter
case a: 400 °C
case b: 200 °C
moles (net) of H2 product per mole of CO in feed gas high-purity H2 produced per cycle (∼10 ppm CO, 15 atm) net amount of CO2 produced (∼100% CO2) net CO2 recovery (%) CO2 removed from bed steam duty (steps d and e)
0.913 0.585 mol/kg of solid 0.466 mol/kg of solid (PH ) 21.5 atm) 72.9 90.0% (step d), 98.0% (steps d and e) 9.86 mol/mol of H2 ) 0.230 ton/MSCF of H2 $0.57a of steam/MSCF of H2
0.910 0.807 mol/kg of solid 0.631 mol/kg of solid (PH ) 40.5 atm) 71.1 90.0% (step d), 98.0% (steps d and e) 11.24 mol/mol of H2 ) 0.262 ton/MSCF of H2 $0.65a of steam/MSCF of H2
cost of steam purge (steps d and e) a
Steam cost at $2.50/ton.
6764 Ind. Eng. Chem. Res., Vol. 47, No. 17, 2008
A very important result of the TSSER process simulations described by Table 4 is that the total amount of steam required (steps d and e) per unit amount of H2 produced for removal of ∼98% of the CO2 present in the sorber-reactor at the end of step b is fairly small (∼0.23-0.26 ton/MSCF of H2) for both cases. The steam duty increases by only ∼14% when the reaction temperature is 200 °C instead of 400 °C. The nominal cost of this steam is ∼$ 0.7/MSCF of H2 (at $2.50/ton of steam), which is less than 12% of the nominal price of H2 at $6.0/ MSCF. Consequently, the steam duty and its cost for the TSSER process are small. However, the steam cost is only part of the overall cost of the proposed TSSER process. A full economic estimate of the concept is required to optimize the process. Summary The sorption-enhanced reaction (SER) concept for producing pure H2 by reacting a mixture of CO + H2O (synthesis gas) over an admixture of a water-gas shift (WGS) reaction catalyst and a water-tolerant, reversible CO2-selective chemisorbent such as Na2O-promoted alumina was experimentally demonstrated. The results show that the kinetics of the WGS reaction becomes limiting for the SER concept when the reaction temperature is
