Ind. Eng. Chem. Res. 2007, 46, 5003-5014
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Novel Thermal-Swing Sorption-Enhanced Reaction Process Concept for Hydrogen Production by Low-Temperature Steam-Methane Reforming Ki Bong Lee, Michael G. Beaver, Hugo S. Caram, and Shivaji Sircar* Department of Chemical Engineering, Lehigh UniVersity, Bethlehem, PennsylVania 18015
Hydrogen production by steam reforming of natural gas is a well-established technology. The possibility of using hydrogen, a nonpolluting fuel, in fuel cells has brought new interest in developing small, efficient, fuel-cell grade hydrogen production units for residential or industrial use. A novel, step-out, low-temperature, steam-methane reforming (SMR) process concept called “thermal-swing sorption-enhanced reaction” (TSSER) is described. The concept simultaneously carries out the SMR reactions at 490-590 °C and removes the byproduct CO2 from the reaction zone in a single unit operation, thereby (a) circumventing the thermodynamic limitations of the SMR reactions and (b) directly producing a fuel-cell grade H2 product with very high CH4-to-H2 conversion. A K2CO3 promoted hydrotalcite is used as the CO2 selective chemisorbent in the reactor, which is periodically regenerated by steam purge at 590 °C. Model simulations of the TSSER process using recently measured CO2 chemisorption characteristics of the promoted hydrotalcite indicate that a very compact H2 generation unit can be designed that requires relatively low amounts of steam for regeneration. New CO2 desorption data from the chemisorbent and its thermal stability are reported. Introduction
Table 1. Performance of Batch SMR Reactor with and without SER Concept
The most common industrial process for production of hydrogen is steam-methane reforming (SMR) using Ni on alumina catalyst.1 Commercial reaction conditions include a pressure of 50-400 psig, a temperature of 700-900 °C, and a feed steam-methane ratio of 2.4-6.0. The SMR reactor effluent gas containing 70-72% H2, 6-8% CH4, 8-10% CO, and 1014% CO2 (dry basis) is cooled to 300-400 °C (steam produced) and subjected to the water gas shift (WGS) reaction, cooled again, and purified in an activated carbon-zeolite containing multicolumn pressure-swing adsorption (PSA) unit. The H2 product contains 99.999+ H2 (20 ppm CO since it interferes with the noble metal catalysts used in the fuel-cell stack.2,3 Obviously, the large and complex process described above may not be appropriate for this application, and novel compact concepts for H2 generation are needed. Sorption-Enhanced Reaction (SER) Concept The key reactions in the SMR reactor of Figure 1 include the following:1
(a) Endothermic steam-methane reforming (SMR) reaction: CH4 + H2O T CO + 3 H2
∆H ) +206 kJ/mol
(1)
(b) Exothermic water gas shift (WGS) reaction: CO + H2O T CO2 + H2
∆H ) -41 kJ/mol
(2)
Net reaction:
A particular need for reducing the capital cost and the footprint of a hydrogen generator is called for by a residential * Corresponding author. E-mail:
[email protected]. Phone: (610) 7584469. Fax: (610) 758-5057.
CH4 + 2 H2O T CO2 + 4 H2
∆H ) + 165 kJ/mol
(3)
Both SMR and WGS reactions are controlled by thermodynamics. Consequently, the composition of the reaction products from the SMR reactor and the overall conversion of CH4 to H2 are
10.1021/ie0701064 CCC: $37.00 © 2007 American Chemical Society Published on Web 06/13/2007
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Figure 1. Conventional steam-methane reforming (SMR) route for hydrogen production.
Figure 2. KSMR, KWGS, and CH4-to-H2 conversion at different temperatures.
strong functions of the reactor pressure (P), the temperature (T), and the CH4/H2O ratio in the reactor feed gas. Figure 2 shows the thermodynamic equilibrium constants and the limits of CH4to-H2 conversion (6:1 H2O/CH4 feed gas, reactor pressure ) 1.5 atm) as functions of reactor temperatures. The equilibrium constants (K) as functions of temperature (T) were estimated using the following published correlations:4
KSMR )
1 , atm2 3 exp(0.2513Z - 0.3665Z 0.58101Z2 + 27.1337Z - 3.277) 4
conversion of CH4 to H2 (∼90%) can be achieved at a temperature of ∼700 °C. The SER concept circumvents this thermodynamic limitation by carrying out the SMR-WGS reactions in the presence of a high-temperature CO2 selective chemisorbent, which removes the CO2 byproduct from the reaction zone as shown below:
(4)
KWGS ) exp(-0.29353Z3 + 0.63508Z2 + 4.1778Z + 0.31688) (5) where Z ) 1000/T - 1 (T in K). It may be seen from Figure 2 that KSMR decreases rapidly while KWGS increases moderately as T is reduced. The maximum
Thus, according to Le Chatelier’s principle, (a) the reaction is driven to the product side, bypassing the thermodynamic limit and (b) the rate of the forward reaction is enhanced. These advantages can be employed to (a) produce a H2 product which is nearly free of COx impurities, (b) achieve very high conversion of CH4 to H2, and, very importantly, (c) permit the
Ind. Eng. Chem. Res., Vol. 46, No. 14, 2007 5005 Table 2. SER Concepts for Production of H2 by SMR Using CaO or Dolomite product H2 purity authors
H2 (%)
temperatures (°C)
CO
CO2
react.
regen.
Brun-Tsekhovoi et al. (1986)6
92-96
627
Balasubramanian et al. (1999)7 Lopez-Oritz & Harrison (2001)8
92-96 95+
450-750 650
975 800-950
Yi & Harrison (2005)9 Johnsen et al. (2006)10
92+ 98+
440 600
850
Li et al. (2006)11
90+
630
850
∼11 ppm 2-5%
process plan fluidized-bed reactor, separation of catalyst and sorbent by gravity followed by external thermal regeneration of sorbent same as above fixed-bed reactor, thermal regeneration with N2, N2 + O2, or CO2 purge not disclosed fluidized-bed reactor, in situ thermal regeneration under N2 purge fixed-bed, thermal regeneration under inert purge
Table 3. Cyclic PSA Based SER Process Performance for H2 Production (ND ) Not Detectable; 1 psia ) 6.9 kPa) gas quantities (m-lbmoles/lb of total solid in sorber-reactor)
H2 product purity (dry basis)
feed P (psia)
feed
purge steam
product H2
H2 (%)
CH4 (%)
CO2 (ppm)
CO (ppm)
CH4-to-H2 conversion (%)
26.2 66.5
0.60 0.54
1.88 0.77
0.25 0.16
94.4 88.7
5.6 11.3
40 136
ND ND
73 54
operation of the reactor at a relatively lower temperature than that used in the conventional SMR reactor while meeting goals (a) and (b). Thermodynamic Model of Batch SMR Reactor The thermodynamic equilibrium constants for the SMR and WGS reactions are functions of temperature (T) only:
P2[yH2]3[yCO] [yCO2][yH2] ; KWGS(T) ) (7) KSMR(T) ) [yCH4][yH2O] [yCO][yH2O] For a constant-temperature (T) and constant-pressure (P) batch reactor containing an admixture of the SMR catalyst and a CO2 selective chemisorbent, the component mole fractions of the equilibrium gas phase are given by5
y H2 )
(3β + δ) (β - δ) ; yCO ) ; (1 + 2β - fδ) (1 + 2β - fδ) yCO2 )
yCH4 )
(1 - f)δ (8) (1 + 2β - fδ)
1 - β(1 + R) ; (1 + R)(1 + 2β - fδ) yH2O )
R - (β + δ)(1 + R) (9) (1 + R)(1 + 2β - fδ)
where R is the molar ratio of feed H2O/CH4 introduced into the batch reactor at P and T. β is the moles of CO produced by the SMR reaction per mole of feed gas. δ is the moles of CO reacted by the WGS reaction per mole of feed gas. f is the fraction of CO2 produced by the WGS reaction that is removed from the gas phase by selective chemisorption. The variables R, β, and δ are related by the following thermodynamic equations:
KSMR(T) ) P2(β - δ)(3β + δ)3(1 + R)2 [1 + 2β - fδ]2[1 - β(1 + R)][R - (1 + R)(β + δ)] KWGS(T) )
(1 - f)δ(3β + δ)(1 + R) (β - δ)[R - (β + δ)(1 + R)]
(10) (11)
The CH4-to-H2 conversion in the batch reactor (χ) is given by
χ)
(3β + δ)(1 + R) 4
(12)
Equations 7-12 can be solved simultaneously for a given set of P, T, R, and f to obtain the equilibrium gas-phase composition inside the batch reactor and the corresponding CH4-to-H2 conversion. Table 1 shows an example of these calculations for a batch SMR reactor at T ) 480 °C, P ) 1.5 atm, and R )6.0 using different values of f. The first row of Table 1 shows that the equilibrium gas phase of the batch reactor containing the catalyst alone (no CO2 removal, f ) 0) produces a very impure H2 product at 480 °C containing large amounts of CO and CO2, and the conversion of CH4 to H2 is poor (only ∼50%). On the other hand, the table shows that progressively increased fraction of CO2 removal from the gas phase of the reactor using a chemisorbent produces remarkable process intensification by the SER concept by direct production of very high purity H2 product with very high conversion (>95+%) of CH4 to H2. It should be re-emphasized here that the CH4-to-H2 conversion of an SMR reactor with catalyst alone and operating under the same conditions of feed pressure and composition cannot exceed ∼90% even at a temperature of 700 °C (Figure 2), and even then the product gas is very impure H2 [78.35% H2 + 13.96% CO2 + 7.51% CO + 0.18% CH4 (dry basis)] requiring further WGS reaction and PSA purification as shown by Figure 1. The SER concept combines the reaction and separation processes in a single unit operation. CO2 Chemisorbents for SER Concept The SER concept for H2 production by SMR by its nature will require the development of a cyclic process where the SMR-WGS reaction step, combined with in situ CO2 removal from the reaction zone, will be carried out until the CO2 chemisorption capacity of the sorbent is exhausted, followed by regeneration of the chemisorbent (desorption of CO2) so that it can be reused. Furthermore, the key requirements for the CO2 chemisorbent to be used in the SER concept are (a) selective chemisorption of CO2 at the reaction temperature in the presence of steam, CO, CH4, and H2, (b) acceptable CO2 sorption capacity at the reaction condition, (c) high rate of chemisorption, (d) reversibility and ease of desorption of CO2 from the chemisor-
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Figure 3. Schematic flow diagram of TSSER process concept.
Figure 4. Equilibrium chemisorption isotherms of CO2 on promoted hydrotalcite.
bent, and (e) thermal and cyclic stability of the chemisorbent under the operating conditions of the SER process.5 The key requirements for the design of the cyclic SER process are (a) direct production of high-purity H2 product at feed gas pressure during the sorption-reaction step of the cyclic process, (b) supply of the endothermic heat of reaction during the reaction step, (c) design of an efficient and practical regeneration scheme, (d) supply of the heat of desorption of CO2, and (e) reduction of the footprint of the overall process system. CaO or Dolomite as the CO2 Chemisorbent Several articles promoting the use of CaO or dolomite as the chemisorbent for CO2 in an SMR reactor have been published in recent years.6-11 Table 2 gives a list of the key works on
this subject. The principal chemisorption reactions are as follows:
CaO + CO2 T CaCO3; CaO‚MgO + CO2 T CaCO3‚MgO ∆H ) -178 kJ/mol (13) The temperature of the sorption-reaction step reported in these articles ranges from 440-750 °C, and the purities of the H2 products are reported to be >90%. The concentrations of CO and CO2 in the H2 product are often fairly high or not reported at all, except for the work of Yi and Harrison,9 who report production of a fuel-cell grade H2 (dry basis), containing ∼10 ppm CO, by operating the SMR reactor at 440 °C and at 3 bar pressure using a calcined sample of dolomite.
Ind. Eng. Chem. Res., Vol. 46, No. 14, 2007 5007 Table 4. Parameters of the Novel Chemisorption-Surface Reaction Model24 T (°C)
m (mol/kg)
a
KC (atm-1)
KR (atm-a)
KH ) mKC (mol/kg/atm)
400 520
0.25 0.25
2.5 2.5
37.4 21.2
2.5 0.8
9.35 5.30
Table 5. Physical Dimensions and Operating Conditions of TSSER Reactor length of a tube in the reactor ) 200 cm; diameter of a tube ) 1.73 cm bulk density of catalyst/chemisorbent admixture ) 0.82 g/cm3 weight fraction of catalyst in reactor ) 0.10 feed to the reactor: H2O/CH4 ) 4:1; pressure ) 1.5 bar; temperature ) 490 °C feed flow rate ) 8.3 m-gmol/cm2/min ) 2.4 lbmol/ft2/h time for sorption-reaction step ) 10.0 min regeneration steam to reactor: pressure ) 1.0-1.5 bar; temperature ) 590 °C regeneration steam flow rate: 38.5 m-gmol/cm2/min ) 11.1 lbmol/ft2/h total time for regeneration step ) 10.0 min shell-side temperature ) 590 °C Table 6. LDF Mass-Transfer Coefficients and Reaction Rate Constants Used in Simulation temp (°C)
k (LDF) (min-1)
kSMR (mmol atm0.5/g-cat min)
kWGS (mmol/atm g-cat min)
490 590
4.5 6.3
2.57 206.2
839.0 2858.3
A critical issue with the use of CaO as the chemisorbent, however, is high-temperature regeneration. Typically, a temperature of 850-1000 °C is necessary for this purpose. Consequently, the SER process must be designed to accommodate this step. Use of a fluidized-bed reactor with separation of solid catalyst and the chemisorbent and external thermal regeneration of the chemisorbent by transporting it to another unit may not be very practical. In situ thermal regeneration of the chemisorbent in a fixed- or fluidized-bed reactor by purging with an extraneous gas may not also be cost-effective. A key advantage of the CaO is the potentially high capacity of CO2 chemisorption.15 However, important issues like (a) formation of Ca(OH)2 during the SMR reaction,12 (b) high-temperature sintering,13,14 (c) very high heat of CO2 chemisorption (∼200 kJ/mol),15 and (d) slow desorption of CO215 must be evaluated. Furthermore, the published literature does not provide any information on the actual cyclic CO2 working capacity of CaO under any realistic process conditions. K2CO3 Promoted Hydrotalcite as the CO2 Chemisorbent Recently, it has been demonstrated that a K2CO3 promoted hydrotalcite can selectively and reversibly chemisorb CO2 in the temperature range of 400-550 °C.5,16,17 The material satisfies all of the above-described properties required for use in an SER process for production of high-purity H2 by lowtemperature SMR reaction. A pressure-swing adsorption (PSA) based SER concept (PSSER) was developed by Air Products and Chemicals, Inc., using this material that directly produced fuel-cell grade hydrogen by steam-methane reforming at a reaction temperature of ∼400-500 °C without sacrificing the CH4-to-H2 conversion.5,18 The chemisorbent was periodically regenerated by purging it with steam at the reaction temperature under a subatmospheric pressure (∼0.34-0.68 bar). A fixed-bed sorberreactor packed with an admixture of a commercial SMR catalyst and the CO2 chemisorbent was used in the process. The cyclic
process consisted of four steps: (a) sorption-reaction at a superambient pressure to produce the H2 product at feed gas pressure, (b) countercurrent depressurization to near-ambient pressure, (c) countercurrent steam purge at subatmospheric pressure, and (d) countercurrent pressurization with steam to feed pressure. This PSSER process steps were operated under a nearly isothermal condition. The process was demonstrated in a pilot-scale unit. Table 3 provides a few examples of the experimental cyclic steady-state process performance data. They were measured using a reactor feed gas consisting of 6:1 H2O/ CH4 at various pressures and a temperature of 490 °C.18 The sorber-reactor (2.54 cm diameter, 6.1 m long) was packed with an admixture of the SMR catalyst (33% by weight) and the promoted hydrotalcite. The data of Table 3 shows that the promoted hydrotalcite can be successfully used as a CO2 chemisorbent to produce fuelcell grade H2 by SMR reaction at a much lower temperature (∼490 °C) than that (∼850 °C) of the conventional process (Figure 1). The data also shows that this chemisorbent can be regenerated by a PSA scheme operated at ∼490 °C. Thus, there were two key advantages of this material compared to CaO as the CO2 chemisorbent for a SER process. They were (a) the lower regeneration temperature and (b) the use of steam as the purge gas during the CO2 desorption process. A shell-and-tube reactor design was suggested for the abovedescribed PSA-SER process.18 Two different types of indirect heat-transfer methods were also proposed for supplying the endothermic heat of SMR reaction and CO2 desorption.17,19 They consisted of (a) flowing a vaporized heat-transfer liquid (e.g., therminol VP 1) through the shell side of the reactor so that the condensing vapor would supply the heat of reaction in the reactor and maintain a constant reactor temperature during all steps of the process and (b) indirect gas heating (IGH) by flowing a hot flue gas through the shell side of the reactor with finned tubes to supply the heat of reaction. Several theoretical model analyses of the original PSSER process concept as well as other process variations have also been published by various authors.19-22 Novel Thermal-Swing Sorption-Enhanced Reaction (TSSER) Concept The purpose of this work is to simulate a new concept for direct production of fuel-cell grade H2 by a thermal-swing sorption-enhanced reaction process (TSSER) employing the K2CO3 promoted hydrotalcite as the chemisorbent. The concept uses only two cyclic steps: (a) Sorption-Reaction Step. The sorption-reaction step is one where a mixture of H2O and CH4 is fed at a pressure of ∼1.5-2.0 bar and a temperature of ∼490 °C into a fixed-bed reactor, which is packed with an admixture of the SMR catalyst and the chemisorbent and which is preheated to ∼590 °C. The effluent from the reactor is fuel-cell grade H2 at feed pressure. (b) Thermal-Regeneration Step. The thermal-regeneration step is one where the reactor is simultaneously depressurized to near-ambient pressure and countercurrently purged with superheated steam at ambient pressure and at ∼590 °C, followed by countercurrent pressurization of the reactor with steam at ∼590 °C to the feed pressure. The reactor effluent is a CO2 rich waste gas. The key advantages of the proposed concept over the original PSSER process are (a) elimination of the usually expensive, subatmospheric steam purge step for desorption of CO2 and the consequent use of a rotating machine (vacuum pump) in the process; (b) direct supply of the heat of endothermic SMR
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Figure 5. CO2 column breakthrough data at (a) 400 °C and (b) 520 °C.
Figure 6. Dynamics of CO2 desorption from promoted hydrotalcite by purge at 520 °C.
reaction from the sensible heat stored in the reactor at the start of step (a); (c) higher utilization of the specific CO2 capacity of the chemisorbent in the cycle due to more stringent regeneration; (d) higher conversion of CH4 to H2; (e) higher purity of H2 product; and (f) lower steam purge requirement per unit amount of H2 product. The proposed TSSER process will potentially provide a relatively simple and very compact alternative for production of fuel-cell grade hydrogen by low-temperature SMR without producing export steam. Figure 3 is a schematic drawing of a two-column embodiment of the concept using a shell-and-tube design of the sorber-reactors. The tubes will be packed with an admixture of the SMR catalyst and the CO2 chemisorbent. The outside walls of the tubes will be maintained at a constant temperature by cross-flowing superheated steam in the shell side. This design is one of many possibilities required for reducing the cycle time for the TSSER process.23 Figure 3 clearly exhibits the compactness of the proposed idea compared with the rather-
involved flow sheet for the conventional SMR-WGS-PSA route of Figure 1. Characteristics of K2CO3 Promoted Hydrotalcite The information required for mathematical simulation of the above-described TSSER concept includes (a) models to describe the equilibrium chemisorption isotherms of CO2 at different temperatures, (b) heats of chemisorption, (c) models to describe kinetics of chemisorption and desorption of CO2, and (d) models to describe the thermodynamics and kinetics of the SMR and WGS reactions. The CO2 chemisorption properties of the K2CO3 promoted hydrotalcite were recently measured in detail using a column apparatus and a sample of the material donated to the Lehigh University by Air Products and Chemicals, Inc.24 Models describing the SMR and WGS reactions were available in the published literature.4,25 These properties are summarized below:
Ind. Eng. Chem. Res., Vol. 46, No. 14, 2007 5009
Figure 7. Cyclic stability of promoted hydrotalcite at 600 °C.
Figure 8. Shape and location of RMTZ inside the sorber-reactor at time t.
CO2 Chemisorption Isotherms. Figure 4 shows the experimental chemisorption isotherms of CO2 (points) on the promoted hydrotalcite at 400 and 520 °C.24 The equilibrium amounts of CO2 chemisorbed (mol/kg) are plotted as functions of the gasphase CO2 partial pressure (atm) at different temperatures. These experimental data could not be described by the classical Langmuir model for chemisorption.26 A new analytical model that accounted for simultaneous chemisorption of CO2 on the surface of the promoted hydrotalcite and additional reaction between the gaseous CO2 and the chemisorbed molecules was developed to describe the data as shown by the solid and dashedlines in the figure. The model isotherm equation is given below.24
n*(P, T) )
mKCP[1 + (a + 1)KRPa] [1 + KCP + KCKRP(a+1)]
(14)
where n* is the specific equilibrium amount (mol/kg) of CO2 chemisorbed at pressure P (atm) and temperature T (K). The parameters KC (atm-1) and KR (atm-a) are, respectively, the equilibrium constants for the Langmuirian chemisorption and the additional surface reaction. The parameter a is the stoichiometric constant for the additional surface reaction. m is the saturation capacity (mol/kg) for the chemisorption. The parameters KC and KR are exponential functions of temperature only,
d ln KC qC d ln KR ∆HR ) - 2; )dT dT RT RT2
(15)
where qC and ∆HR (kcal/mol) are, respectively, the isosteric heat of chemisorption of CO2 and the heat of additional surface reaction. It can be shown that (i) eq 14 has the same form as the Langmuir model (eq 16) in the low-pressure region and (ii)
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Figure 9. Temperature profiles inside the sorber-reactor at time t.
n* asymptotically approaches the value of m(a + 1) at the limit of P f ∞. For small values of P or n*:
n*(P, T) )
mKCP [1 + KCP]
(16)
Table 4 reports the parameters of eq 14 corresponding to the isotherms of Figure 4. Equation 15 was then used to estimate the values of the heat of CO2 chemisorption (qc ) 21.0 kJ/mol) and the heat of additional surface chemical reaction (∆HR ) 42.16 kJ/mol), as well as the values of the pre-exponential coefficients [KC° ) 0.8778 atm-1 and KR° ) 1.34 × 10-3 atm-2.5]. The relatively low values of these heats indicate that the CO2 chemisorption on K2CO3 promoted hydrotalcite is energetically weak and, therefore, easy to desorb from the surface. This is a very favorable property exhibited by this chemisorbent vis a vis CaO, which has a very high heat of CO2 chemisorption (∼200 kJ/mol) and a high activation energy for CO2 desorption (∼300 kJ/mol), resulting in very slow desorption.15 The exothermic chemisorption of CO2 during the simultaneous sorption- reaction step of the SER concept reduces the net heat requirement for the endothermic reaction given by eq 3. This is an additional benefit of the SER concept. Thus, the reduction in heat duty for the reaction using the promoted hydrotalcite is at least ∼12%, if the reaction conditions are such that only the low-pressure part of the chemisorption isotherm (eq 14) is utilized during the SER process. Table 4 also reports the Henry’s Law constants (KH) for the isotherms. CO2 Chemisorption Kinetics. Figure 5 shows an example of the experimental CO2 breakthrough data from a clean column (101.6 cm long, 1.73 cm diameter) packed with the promoted hydrotalcite and filled with pure N2 at ambient pressure.24 The data are shown for two temperatures. The plots are normalized effluent gas CO2 mole fraction from the column (y/y°) as functions of normalized time (t/t*) for a feed gas containing 40% CO2 (y°) in N2 at atmospheric pressure. The variable t* (s) is the stoichiometric breakthrough time. The near vertical breakthrough behavior at t/t* ) 1 in the concentration range {0 < y/y° < 0.7} indicates that the kinetics of sorption of CO2 on the chemisorbent is fast at both temperatures. The deviation of the breakthrough plots from the vertical line at higher values of y/y° is caused by column nonisothermality.24,27 Figure 5 also shows that the CO2 breakthrough curves can be simulated well by a “continuous stirred tank reactor (CSTR) in series” column dynamic model in conjunction with the
simplified linear driving force (LDF) model for the local chemisorption kinetics inside the column at any time t,24,28
dn(t) ) k[n*(t) - n(t)] dt
(17)
where n* is the equilibrium amount of CO2 adsorbed, defined by the new chemisorption isotherm model (eq 14), at the local conditions inside the column, and k is the LDF mass transfer coefficient for CO2 chemisorption. The extracted k values at 400 and 520 °C were, respectively, 3 and 5 min-1. k was found to be independent of feed gas CO2 concentration in the range of the experimental data (0.4 e y° e 0.6).24 Dynamics of CO2 Desorption from the Chemisorbent. Figure 6 shows newly measured desorption characteristics of chemisorbed CO2 from a column packed with the promoted hydrotalcite at 520 °C by purge with N2. It plots the fraction of CO2 desorbed from the column as a function of cumulative amount of N2 (mol/kg of chemisorbent) introduced as purge gas. The column was initially equilibrated with a stream of 40% CO2 in N2 at ambient pressure and 520 °C. The figure also shows the simulated desorption characteristics of CO2 employing the same column dynamic model used for simulation of the breakthrough curves of Figure 5. The equilibrium isotherm of eq 14, the LDF kinetic model of eq 17, and the same mass transfer coefficients extracted from the chemisorption breakthrough experiments were used in the simulation of the desorption profile. The good match between the experiment and the simulation in Figure 6 demonstrates that the chemisorption of CO2 is reversible on this material, and there is no difference in the kinetic time constants between the CO2 sorption and the desorption processes. Thermal Stability of Promoted Hydrotalcite. The CO2 selectivity of K2CO3 promoted hydrotalcite and its thermal stability in the presence of steam (partial pressure ) 10 bar) at 400 °C was previously investigated by repeatedly chemisorbing CO2 (partial pressure ) 0.3 bar) from N2 and desorbing it by pure N2 purge at 400 °C using a sample of the material.5,17 The cyclic CO2 chemisorption capacity initially decreased but stabilized after several cycles of operation.5 We recently tested the thermal stability of a sample of promoted hydrotalcite by cyclically measuring its pure CO2 chemisorption capacity at ambient pressure and 400 °C followed by heating it in dry N2 to 600 °C for desorption of CO2. A commercial microbalance was used for these tests. Figure 7 shows the results. The CO2 capacity was retained at its initial
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Superheated steam was passed in the shell side (cross-flow) to maintain the outside tube walls at a constant temperature (Tw) at all times. The well-known CSTR-in-series model29 was adapted for dynamic simulation of the sorption-reaction and thermal-regeneration steps of the TSSER process. The key model assumptions included (i) ideal gas behavior, (ii) instantaneous thermal equilibrium between the gas and the solid inside the sorber-reactor tubes, and (iii) absence of axial dispersion and column pressure drop. The ordinary differential equations of the model representing (i) the transient component and the overall mass balances and (ii) the transient energy balance including heat transfer from the shell side to the tube side (overall heat-transfer coefficient ) Uo) at any time (t) are given below:
Overall molar balance in the gas phase: NoutA ) NinA - FbVstage
[∑ ∑ i
Figure 10. Specific H2 productivity vs COx impurities in the product gas.
value within the experimental scatter after nine cycles of sorption-desorption tests. This demonstrated the thermal stability of the material at 600 °C, albeit for a few cycles of sorptiondesorption operation. Much more extensive testing of the stability of the material at 600 °C will be necessary.
∑j (1 - w)Rads,j
dt
The following three empirical but analytical expressions describing the relevant SMR and WGS reaction kinetics for a commercial nickel on alumina catalyst were published by Xu and Froment.25 These kinetic models were chosen for use in the simulation of the TSSER process proposed in this work,
(1) CH4 + H2O T CO + 3H2 rate )
k1 pH22.5
[
(2) CO + H2O T CO2 + H2 rate )
pCH4 pH2O -
]
pH23pCO KSMR
[
rate )
pH2
3.5
[
/[A] (18)
pCH4 pH2O -
pH24 pCO2 KSMR KWGS
nt
{∑
NinA(yin j - yj) + FbVstage
{∑ ∑
-yjFbVstage
i
wRij +
j
dnt dt
}
wRij + (1 - w)Rads,j
}
i
∑
(1 - w)Rads,j
j
]
(22)
(23)
Energy balance: ηFbVstagecps
dT dt
) NinAcpg(Tin - T) - FbVstage
]
[∑
w∆HiRi +
i
∑j (1 - w)∆Hads,jRads,j + πdcLstageUo(Tw - T) nt )
]
2
)
1
-
(24)
Ideal gas law:
pH2pCO2 k2 pCOpH2O /[A]2 (19) pH 2 KWGS
(3) CH4 + 2H2O T CO2 + 4H2 k3
2
[
]
Component molar balance in the gas phase: dyj
Kinetics of WGS and SMR Reactions
wRij +
j
]
/[A]2 (20)
A ) {1 + βCOpCO + βH2pH2 + βCH4 pCH4 + βH2OpH2O/pH2} (21) where pi is the gas-phase partial pressure of component i in the catalytic reactor. k1, k2, and k3 are, respectively, the specific rate constants for reactions 1, 2, and 3. βi is an empirical, temperature-dependent parameter for component i. CSTR-in-Series Model for Simulation of TSSER Process A conventional shell-and-tube type heat exchanger was used as the sorber-reactor vessel in the process. The tube side (diameter ) dc, cross-sectional area ) A, and length ) Lc) of the exchanger was packed with an admixture of the SMR catalyst and the chemisorbent (bulk density ) Fb, heat capacity ) cps, weight fraction of catalyst ) w, and void fraction ) ).
VstageP RT
(25)
Molar balance in the solid phase (LDF model): dnads,j / ) kads,j(nads,j - nads,j) dt
(26)
where cpg is the heat capacity of the gas phase. ∆H and ∆Hads are, respectively, the net heats of the SMR-WGS reactions and the heat of CO2 chemisorption. kads is the LDF mass transfer coefficient. N is the total gas flux. nads, nads*, and nt are, respectively, the solid-phase loading, the equilibrium solid-phase loading, and the total moles in the gas phase. Ri and Rads are, respectively, the rate of chemical reaction and CO2 chemisorption. Lstage and Vstage are, respectively, the length and the volume of a CSTR stage (cell). yj is the mole fraction of component j in the gas phase, P is the total gas pressure, T is the temperature, and R is the gas constant. η ()1.2) is a factor to account for the heat capacity of the exchanger tubes and body. Equations 22-26 were simultaneously solved using the Matlab function ODE15S, which is a variable-order solver based on the numerical differentiation formulas. A fairly large number of tanks in series (250-500) were used to simulate the sharp mass transfer zones for chemisorption of CO2. It was found that 250 tanks in series was adequate for the present purpose. The new chemisorption isotherm model (eqs 14 and 15) and
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Ind. Eng. Chem. Res., Vol. 46, No. 14, 2007
Figure 11. CO2 loading profiles inside the sorber-reactor during the regeneration step.
Figure 12. Reactor temperature profiles at different times during regeneration.
the SMR-WGS reaction kinetic models (eqs 18-21) were used in the simulation. It should be noted from Figure 4 that the chemisorption isotherm model of eq 14 somewhat overpredicts the amount of CO2 sorbed above a CO2 partial pressure of ∼1.2 atm at higher temperatures. However, that factor is not critical for the present simulation since the CO2 partial pressures always remain below ∼1.0 atm during the process steps of the present simulation. Tables 5 and 6 summarize the key design parameters as well as the process operating conditions used in the simulation. Simulation Results Sorption-Reaction Step. The feed gas mixture reacts to form CO2, CO, and H2. The CO2 is selectively chemisorbed, which drives both SMR and WGS reactions to near completion, thus producing a stream of essentially pure H2 (dry basis) from the sorber-reactor. A reaction-mass transfer zone (RMTZ) is formed near the entrance of the reactor, where all of the chemical reactions and the subsequent chemisorption of CO2 takes place. The RMTZ moves forward to the H2 product end as more feedgas is introduced. The section of the reactor behind the RMTZ remains essentially equilibrated with a gas mixture whose composition is determined by the thermodynamics of the reactions without the chemisorption of CO2. Figure 8 shows the simulated movement of the RMTZ through the sorber-reactor at different times for the reactor
Figure 13. Cumulative amount of steam purge and degree of regeneration.
operating conditions given by Table 5. It plots the specific CO2 loadings on the chemisorbent (mol/kg) as functions of dimensionless distance in the sorber reactor (L/Lc) at different times. L is the actual distance in the column from the feed end. The sorption-reaction step is stopped at a time of 10 min when the leading edge of the RMTZ reaches the reactor-end, causing incipient breakthrough of COx. Figure 9 shows the temperature profiles inside the reactor during the sorption-reaction step of Figure 8. The feed-end temperature decreases to 450 °C as soon as the feed gas is introduced in order to supply the endothermic heat of the SMR
Ind. Eng. Chem. Res., Vol. 46, No. 14, 2007 5013 Table 7. Potential Advantages of the Proposed TSSER Process Concept
a
process
H2 product purity (dry basis)
PSSER18 6:1 H2O/CH4 feed P ) 1.78 bar; T ) 490 °C catalyst ) 33% TSSERa 4:1 H2O/CH4 feed P ) 1.5 bar; T ) 490 °C catalyst ) 10%
H2 ) 94.4% CH4 ) 5.6 % CO )
