Chemical Engineering Science 59 (2004) 931–942www.elsevier.com/locate/ces

Modeling and simulation for the methane steam reforming enhanced byin situ CO2 removal utilizing the CaO carbonation for H2 production

Deuk Ki Leea ;∗, Il Hyun Baekb, Wang Lai Yoonc

aDivision of Civil and Environmental Engineering, Gwangju University, Gwangju 503-703, Republic of KoreabDivision of Energy and Environmental Research, Korea Institute of Energy Research, Daejon 305-343, Republic of Korea

cEnergy Conversion Process Research Center, Korea Institute of Energy Research, Daejon 305-343, Republic of Korea

Received 30 July 2003; received in revised form 8 December 2003; accepted 13 December 2003

Abstract

The transient behavior of catalytic methane steam reforming (MSR) coupled with simultaneous carbon dioxide removal by carbonationof CaO pellets in a packed bed reactor for hydrogen production has been analyzed through a mathematical model with reaction experimentsfor model veri8cation. A dynamic model has been developed to describe both the MSR reaction and the CaO carbonation-enhanced MSRreaction at non-isothermal, non-adiabatic, and non-isobaric operating conditions assuming that the rate of the CaO carbonation in a localzone of the packed bed is governed by kinetic limitation or by mass transfer limitation of the reactant CO2. Apparent carbonation kineticsof the CaO pellet prepared has been determined using the TGA carbonation experiments at various temperatures, and incorporated into themodel. The resulting model is shown to successfully depict the transient behavior of the in situ CaO carbonation-enhanced MSR reaction.The e<ects of major operating parameters on the transient behavior of the CaO carbonation-enhanced MSR have been investigated usingthe model. The bed temperature is the most important parameter for determining the amount of CO2 removed by carbonation of CaO,and at temperatures of 600◦C, 650◦C, 700◦C and 750◦C, the CO2 uptake is 1.43, 2.29, 3.5 and 5:09 mol-CO2=kg-CaO, respectively.Simultaneously with the increase in CO2 uptake with increasing temperature, the corresponding amounts of hydrogen produced are 1.56,2.54, 3.91 and 5:63 mol-H2/kg-CaO, at the same temperatures as above. Operation at high pressure, high steam to methane feed ratio,and the decreased feed rate at a given temperature are favorable for increasing the degree of the overall utilization of CaO pellets in thereactor bed, and for lowering the CO concentration in the product.? 2004 Elsevier Ltd. All rights reserved.

Keywords: Methane steam reforming; CaO carbonation; CO2 removal; H2 production; Modeling; Packed bed

1. Introduction

Methane steam reforming (MSR) is a major route for theindustrial production of H2. The three main reactions in aMSR reactor are represented by following equations (Xuand Froment, 1989):

CH4 + H2O ⇔ CO + 3H2;

FH298 = 206:2× 103 kJ=kmol; (1)

CH4 + 2H2O ⇔ CO2 + 4H2;

FH298 = 164:9× 103 kJ=kmol: (2)

∗ Corresponding author. Tel.: +82-62-670-2394;fax: +82-62-670-2192.

E-mail address: dklee@gwangju.ac.kr (D.K. Lee).

CO + H2O ⇔ CO2 + H2;

FH298 =−41:1× 103 kJ=kmol: (3)

Reforming reactions (1) and (2) are highly endothermic, andthermodynamically favored by high temperature and lowpressure. On the other hand, the water–gas shift (WGS) re-action given by Eq. (3) is favored at low temperature, butit has no pressure dependence. MSR is generally operatedat a temperature of 750–900◦C due to the overall endother-mic nature of the reactions (Hufton et al., 1999). Althoughhigh-temperature operation is indispensable for a substantialconversion of CH4, it facilitates the reverse WGS reaction,giving the product gas containing 8–10% CO on a dry ba-sis. For the purpose of obtaining the product gas with lessCO and more H2, it is conventional that the MSR productgas is fed to another reactor where the temperature is keptas low as 300–400◦C for the WGS reaction to take place

0009-2509/$ - see front matter ? 2004 Elsevier Ltd. All rights reserved.doi:10.1016/j.ces.2003.12.011

932 D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942

prior (Hufton et al., 1999). To obtain the H2 product stream,the eKuent is then cooled and fed to a multicolumn pressureswing adsorption (PSA) process.During industrial operation of the MSR process, it is cus-

tomary to release the product CO2, which is a greenhousegas with potential to contribute to global warming, to the en-vironment. To be ready for possible obligation to decreasethe amount of CO2 released, it is important for industries todevelop means to capture/sequestrate CO2 from their pro-cesses. As far as the MSR for the production of H2 is con-cerned, in situ capture of CO2 from the MSR reaction systemhas advantages of not only providing a chance to seques-trate the greenhouse gas instead of its release to the atmo-sphere, but also bringing the following separation-enhancedreaction merits: higher CH4 conversion and H2 yield thanthose equilibrium-limited at a given condition. Provided thata CO2-acceptor is available in the reaction zone, the con-version of CH4 to CO2 through Eq. (2) is enhanced, as isthat produced via CO intermediate.Several studies concerning MSR or WGS enhanced by in

situ separation of CO2 employing adsorbents or chemicalacceptors are available in the literature. Using dolomite asCO2 acceptor in a Muidized bed reactor containing Ni-basedcatalyst, Brun-Tsekhovoi et al. (1986) reported a high en-hancement of CH4 conversion to H2. Han and Harrison(1994) investigated H2 production via the WGS reactionusing dolomite as a CO2 acceptor in the temperature rangeof 500–600◦C. At the most favorable conditions, the totalconcentration of carbon oxides in the product gas was lowas 50 ppm. A concept called sorption-enhanced reactionprocess (SERP) was suggested by Circar and co-workers(Carvill et al., 1996; Hufton et al., 1999; Waldron et al.,2001) for H2 production by MSR using a packed columncontaining an admixture of MSR catalyst and adsorbentto remove CO2 from the reaction zone. In the SERP, theadsorbent, a potassium carbonate promoted hydrotalciteover which CO2 could reversibly adsorb in the tempera-ture range of 300–500◦C in the presence of excess steam,was periodically regenerated by PSA. They reported thatthe SERP concept allowed direct production of high-purityH2 (¿ 95 mol%) at high methane to hydrogen conversion(¿ 80%) with dilute CH4 and trace carbon oxide impuritiesat 450◦C and 4:8 bar. Once adsorbents have been saturatedwith CO2, the reaction-enhancement e<ect does not exist anymore. Therefore, such methods as pressure-, concentration-,and temperature-swing operations, or reactive regenera-tion are developed to regenerate the spent adsorbents (Xiuet al., 2002). MSR employing hydrotalcite-based CO2 ad-sorbents was studied experimentally and theoretically indetail recently (Ding and Alpay, 2000; Xiu et al., 2002).Ding and Alpay (2000) analyzed the transient behavior ofa tubular reactor when a Ni-based catalyst was admixedwith the adsorbent. Considerable enhancement of the CH4

conversion was experimentally demonstrated, and predictedby a mathematical model. Xiu et al. (2002) reported thathydrogen-enriched stream with traces of CO2 and CO could

be produced by a sorption-enhanced MSR process withreactive regeneration of adsorbent where the purge stepwas performed at a low temperature of 400◦C (comparedwith the reaction temperature of 450◦C) with 10% H2 innitrogen at atmospheric pressure. It is noticeable abovethat the CO2 removal-enhanced MSR processes mostly em-ploy hydrotalcite-based sorbents. Reversible CO2 uptakecapacity of the hydrotalcite-based sorbent was reported tobe about 0:45 mol=kg at 400◦C after 10 cycles of adorp-tion/desorption (Hufton et al., 1999). Such a low uptakecapacity of the adsorbent means a very short period of ser-vice, requiring frequent regeneration of the spent sorbentseven in a large reactor.The application of CaO as a CO2-acceptor to MSR in a

8xed bed reactor was reported by Balasubramanian et al.(1999), demonstrating that hydrogen with a purity of morethan 95% could be produced in a single step MSR processusing dolomite particles as a CaO-based CO2-acceptor inthe reactor at 650◦C and above. Later, Ortiz and Harrison(2001) reported the e<ect of regeneration conditions of thespent dolomite as a function of temperature and regenera-tion gas composition on the performance in the subsequentH2 production cycle, showing only moderate activity lossunder most of the regeneration conditions. The carbonationreaction of CaO with CO2 is represented by

CaO + CO2 ⇔ CaCO3;

FH298 =−178:8× 103 kJ=kmol: (4)

CaO carbonation as a kind of non-catalytic gas–solid reac-tion was reported to not proceed to the complete conversionof CaO, with ultimate conversions in the range of 70–80%for fresh powder samples (Bhatia and Perlmutter, 1983).However, the ultimate carbonation conversion of CaO after14 cycles of CaO-carbonation/CaCO3-calcination was re-ported to decrease to 20% (Abanades, 2002). Even with thislowered carbonation conversion, the CO2-uptake capacityof CaO is calculated to 3:57 mol=kg, much higher than thatof the hydrotalcite-based sorbent. This large amount of CO2

uptake is of great advantage to the MSR process utilizingthe CaO carbonation for in situ CO2 removal as comparedwith the process utilizing the hydrotalcite-based sorbent.Whereas the CO2 removal-enhanced MSR reaction utilizinghydrotalcite-based sorbents was well analyzed using math-ematical models (Ding and Alpay, 2000; Xiu et al., 2002),no studies concerning transient modeling of the MSR reac-tion combined with the CaO carbonation as a non-catalyticgas–solid reaction have been reported.In this work, the transient behavior of the catalytic MSR

over a Ni-catalyst coupled with the simultaneous carbona-tion of CaO pellets in a packed bed reactor has been ana-lyzed for H2 production through a mathematical model andreaction experiments for the model veri8cation. Modelingsuch a process requires information on the CaO carbona-tion kinetics as well as reaction kinetic models for reactions(1)–(3). CaO carbonation kinetics needed for modeling has

D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942 933

been obtained using TGA experiments at various tempera-tures for CaO pellets prepared with inorganic binders. Thereaction kinetic model proposed by Xu and Froment (1989)has been employed for reactions (1)–(3). Throughout mod-eling, e<ects of major operating parameters on the transientbehavior of the CaO carbonation-enhanced MSR have beeninvestigated and discussed. Because the CaO carbonationreaction generates a signi8cant amount of heat, as repre-sented by reaction (4), the thermal e<ect in the reactor onthe reaction performance has been also determined.

2. Experimental

2.1. Carbonation of CaO pellets in TGA

For the in situ CO2 removal, spherical pellets (3 mm indiameter, surface area = 5 m2/g, porosity = 0:045) of CaOcontaining inorganic binders, feldspar (5%) and bentonite(5%), were prepared. Carbonation conversion of the CaOpellet was obtained using a TGA (TA Instrument). For aCaO pellet (about 22:8 mg) placed on the TGA sample pan,the weight change by the carbonation was measured as func-tion of time by Mowing pure CO2 (100 ml=min) at a con-trolled temperature. For accurate measurement and controlof the temperature of the sample pellet during the TGA ex-periment, the sensing tip of a thermocouple was placed at aposition as close to the sample as possible. The durabilityof the CaO pellet was tested in a separate experiment us-ing TGA by repeating 10 cycles of carbonation (at 750◦C)followed by complete calcination (at 900◦C) under Mow of100 ml=min 10% CO2=N2.

2.2. MSR

A commercial Ni-based reforming catalyst (ICI 57-7, 1/8”cylindrical pellets) was used in combination with the CaOpellets prepared. Catalyst (16:4 g) and CaO (83:6 g) pelletswere admixed, and packed into a stainless steel tube reac-tor of internal diameter 24 mm, and the length of the reac-tive packing bed was 290 mm. Ahead of the reactive bed,spherical pellets of �-alumina were packed for preheatingand even distribution of the feed. Reactor bed temperatureswere monitored with two thermocouples, one located at thepreheating zone and the other at the middle section of the re-active packing bed, and controlled with two electric heaters.The reactor entrance was connected to the lines for feedgases (CH4, H2O), reducing gas (50% H2/He) and purgingor process gas (N2, H2). The feed Mow of CH4 was regu-lated using a massMow controller, and steam was suppliedto the reactor via water-vaporizer using a HPLC pump. Thereactor eKuent gas for analysis was introduced to water con-denser followed by a moisture trap to remove water. A GC(Donam, DS6200) equipped with a TCD detector and a Po-rapak Q column was used for analysis of the dried gas sam-ple employing a 6-port auto-sampling valve. The reaction

experiment was carried out as follows: (i) heat the reactor ata rate of 5◦C=min at atmospheric pressure under the Mow of50% H2/He up to 800◦C, at which the reactor was kept foradditional 12 h for catalyst reduction; (ii) adjust the reactortemperature to a desired reaction temperature; (iii) supplywater and H2 to the reactor so that the molar ratio of H2O toH2 is the same as the desired H2O=CH4 ratio in the reaction,(iv) pressurize the reactor using a backpressure regulator,(v) replace the H2 Mow with CH4 to start the MSR reaction.

3. Mathematical modeling

3.1. CaO carbonation kinetics

TGA-determined carbonation conversions of the CaOpellet with time at temperatures 650–750◦C are shown inFig. 1. The CaO pellet exhibits a relatively low degree ofconversion at each temperature. Even at 750◦C, carbona-tion was limited to a low 8nal conversion of about 30%.This may be attributed to the large pellet size with thelow surface area and porosity, by which pore plugging inthe outer shell of the pellet is strongly e<ected, preventinga large amount of intra-pellet CaO from being carbon-ated. A smaller pellet size would have been preferable forhigher ultimate conversion. However, this sample did notshow any sign of deterioration in the ultimate carbonationconversion at the durability test repeating 10 cycles of car-bonation/calcination in the TGA. Lee (2004) suggested anapparent kinetic expression for the carbonation of CaO asfollows:

dXdt

= kc

(1− X

Xu

)2; (5)

where X denotes the fractional conversion of CaO to CaCO3,kc is an apparent kinetic rate constant which is dependent

0 20 40 60 80Time, min

0

10

20

30

40

50

Car

bona

tion

conv

ersi

on, %

Temperature, °C650700750predicted

Fig. 1. Carbonation conversions of the prepared CaO pellets under pureCO2.

934 D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942

on temperature, and Xu is the ultimate conversion of CaO.Integration of Eq. (5) leads to

X =kcbtb+ t

; (6)

where b is de8ned as the time taken to reach half the ultimateconversion in the carbonation reaction at a given tempera-ture. By substituting the relation X =Xu=2 at t=b to Eq. (6),the ultimate conversion is given by Xu = kcb. Least-squareregression for the conversion data in Fig. 1 using Eq. (5)gives values of kc and b at each temperature. The temper-ature dependencies of the two parameters have been deter-mined as follows:

kc = 96:34 exp(−12171=T ); (7)

b= 4:49 exp(4790:6=T ): (8)

The prediction using Eq. (6) by substituting kc and b inEqs. (7) and (8) is in a good agreement with the conversiondata shown in Fig. 1, supporting that apparent carbonationreaction of the CaO pellets in the TGA takes place on a rategiven by Eq. (5). Therefore, the molar rate of CO2 removalper unit mass of CaO can be represented as

rcbn =1

MCaO

(dXdt

); (9)

where MCaO is molecular weight of CaO. Eq. (9) is givenindependent of the concentration of CO2. Bhatia and Perl-mutter (1983) reported that the rate of CaO carbonation wasindependent of CO2 partial pressures, except for a slight ef-fect at a very early stage of the carbonation. Dedman andOwen (1962) also reported that the reaction was zero orderwith respect to CO2 pressures.

3.2. Reforming and shift reaction kinetics

The reaction kinetic model of Xu and Froment (1989)can be summarized as

R1 =k1P2:5H2

(PCH4PH2O − P3

H2PCO

K1

)/(DEN)2; (10a)

R2 =k2P3:5H2

(PCH4P

2H2O − P4

H2PCO2

K2

)/(DEN)2; (10b)

R3 =k3PH2

(PCOPH2O − PH2PCO2

K3

)/(DEN)2; (10c)

DEN = 1 + KCOPCO + KH2PH2 + KCH4PCH4

+KH2OPH2O=PH2 ; (10d)

where Rj (j = 1; 2; 3) denotes the reaction rate of the MSRreactions (1) and (2), and the WGS reaction (3). Pi (i =CH4, H2O, H2, CO, CO2) corresponds to the partial pres-sure of species i. Expressions for the rate constants kj,the equilibrium constants Kj (j = 1; 2; 3), and adsorption

constants Ki (i=CO; H2; CH4; H2O) as functions of tem-perature have been given by Xu and Froment (1989), andwere well summarized by Xiu et al. (2002). The rate of for-mation or consumption of species i is then calculated usingEqs. (10a)–(10c):

rCH4 =−(R1 + R2); (11a)

rH2O =−(R1 + 2R2 + R3); (11b)

rH2 = 3R1 + 4R2 + R3; (11c)

rCO = R1 − R3; (11d)

rCO2 = R2 + R3: (11e)

3.3. Governing equations

A dynamic model is developed to describe both the MSRreaction and the CaO carbonation-enhanced MSR reactionat non-isothermal, non-adiabatic, and non-isobaric operatingconditions. The model assumptions are summarized as fol-lows: axial plug Mow, ideal gas behavior, Mat radial pro8leof temperature or concentration, constant bed void fractionand uniform particle size of the reactor-packed materials,and constant wall temperature. Based on the above assump-tions, the pseudo-homogeneous model (Feyo De Azevedoet al., 1990) of the di<erential mass balance for a packedbed can be written as

�@Ci

@t=−@(uCi)

@z+ �(1− �)�catri − (1− �)�CaOrcbn ; (12)

where Ci is the molar concentration of species i, � is the bedvoid fraction, u is the super8cial velocity, � is the catalyste<ectiveness factor, �cat and �CaO are the apparent density ofthe catalyst and CaO pellets, respectively. The last term onthe right-hand side of Eq. (12) represents the molar rate ofCO2 removed by CaO carbonation per unit bed-volume ofthe reactor. The molar rate of CO2 removal per kg CaO, rcbn,can be represented as a function of the fractional conversionof CaO, X , using Eqs. (5) and (9):

rcbn =�

MCaO

(1− X

Xu

)2; (13)

where � is a parameter introduced to account for that thelocal rate of CaO carbonation along the packed bed may begoverned by kinetic limitation or by mass transfer limitationof the reactant CO2. For the former, the rate of the CaOcarbonation at a given bed temperature is determined bykc given in Eq. (7). For the latter case, it is determinedby the availability of CO2, which depends on the amountentrained with upcoming convection Mow and the generationby catalytic reactions within a local bed zone. Therefore,

�= kc under kinetic limitation; (14a)

�=MCaO

(1− �)�CaO[ CO2 + �(1− �)�catrCO2]

under mass transfer limitation: (14b)

D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942 935

For a local bed zone where the CaO carbonation rate isunder mass transfer limitation, the quantity of CO2 re-movable by the carbonation of CaO corresponds to theamount of CO2 available within the zone minus that ofCO2 escaping from the zone with a convective Mow inthe equilibrium concentration (Ceq;CO2) of CO2 in theCaO-carbonation/CaCO3-calcination reaction (4). In Eq.(14b), CO2 is introduced to take into account part of theconvection-delivered CO2, which is supposed to be re-moved by carbonation of CaO within a bed zone of a 8nitebed length Fz:

CO2 =uCCO2 − uCeq;CO2

Fz: (15)

Ceq;CO2 at temperature T in Kelvin (Barin, 1989) is providedas

Ceq;CO2= 4:137× 1012 exp(−20474=T )

(1× 10−5

0:082T

):

(16)

For a time increment Ft, the change in the carbonationconversion of a local bed CaO at time t can be obtained as

Xt+Ft = Xt +MCaO

∫ t+Ft

trcbn dt: (17)

The pseudo-homogeneous energy balance for a packedbed in a plug Mow can be written as

[(1− �)�sCps + ��gCpg]@T@t

=− @(u�gCpgT )@z

− (1− �)�cat

∑j

�RjHRj

− (1− �)�CaOrcbnHcbn + hW (TW − T )4DR

; (18)

where �s, the average apparent density of the two solidsin the reactor, is given by �s = (Wcat +WCaO)=(Wcat=�cat +WCaO=�CaO), withWcat andWCaO denoting the weights of thecatalyst and CaO pellets packed in the reactor, respectively.�g is the gas phase density, and Cps and Cpg are the solidand gas heat capacities, respectively. HRj denotes the heatof reaction j in reactions (1)–(3), and Hcbn denotes the heatof CaO carbonation reaction in reaction (4). The last term inthe right-hand side of Eq. (18) accounts for the heat transferthrough the reactor tube wall at a constant temperature TW .Wall-bed heat transfer coeScient, hW , is given for a packedtube reactor of inside diameter DR as Eq. (19a) by Li andFinlayson (1977), and as Eq. (19b) by DeWash and Froment(1972) at very low-feed Mow rates.

hW = 2:03(

kgDR

)Re0:8p exp

(−6dp

DR

)

(Rep = 20− 7600; dp=DR = 0:05− 0:3); (19a)

hW = 6:15(

k0zDR

)(as Rep → 0); (19b)

where Rep = u��gdp=%, and k0z = kg[�+ (1− �)={0:139�−0:0339 + 2=3(kg=ks)}], in which % is the viscosity of gas,

Table 1Values of parameters used

Parameters Values

Cpg 8:45 kJ=(kg K)Cps 0:98 kJ=(kg K)dp 3× 10−3 mDR 2:4× 10−2 mkg 2:59× 10−4 kJ=(m sK)ks 1× 10−3 kJ=(m sK)MCaO 56 kg=kmolWCaO 83:6× 10−3 kgWcat 16:4× 10−3 kg� 0:5% 2:8× 10−5 Pa s�CaO 1257 kg=m3

�cat 246 kg=m3

�s 1503 kg=m3

� 0.3

and kg and ks denote the thermal conductivity of gas and ofthe packed solid, respectively.Pressure distribution in the packed bed is described by

the Ergun equation (Ergun, 1952)

dPdz

=−�gu2

dp

1− ��

[150(1− �)%

dp�gu+ 1:75

]× 10−5: (20)

3.4. Numerical method

The initial and boundary conditions in Eqs. (12), (18) and(20) are set as follows:

Ci = Ci; init ; T = Tinit at t = 0;

Ci = Ci;f; T = Tf; P = Pf at z = 0;

where i denotes the gas species CH4, H2O, H2, CO, CO2.The initial bed concentrations of H2 and H2O were set equalto those of CH4 and H2O in the feed, respectively, and theinitial concentrations of other gas species were set to zero.The feed concentration of H2 was in fact zero, however, itwas assumed as 10−6 of the feed concentration of CH4 sothat the denominator in Eq. (10d) would not be zero.The model equations were solved with MATLAB pro-

gramming. Using the method of lines, the partial di<erentialequations (12) and (18) were converted to a set of ordinarydi<erential equations with initial conditions by discretizingthe spatial derivative in z-direction using backward 8nitedi<erences. For 8nite di<erences, the reactor of bed lengthL=0:29 m was divided into 50 sections with 51 nodes. Theinitial value ordinary di<erential equations comprised of 306equations and other explicit algebraic equations at a time twere simultaneously solved using ’ode15s.m’, a MATLABbuilt-in solver for initial value problems for sti< ordinarydi<erential equations. A summary of the values of the pa-rameters used in the simulation is given in Table 1.

936 D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942

4. Results and discussion

4.1. Veri:cation of the model

Simulated pro8les of the product gas composition on adry basis are shown in Fig. 2 in comparison with experimen-tal data obtained from the in situ CaO carbonation-enhancedMSR reaction under the operating conditions: 700◦C, 3 bar,CH4 feed rate= 11:2 NL=h, molar H2O=CH4 feed ratio= 3.Total feed Mow rate through the reactor bed correspondsto 341:5 h−1 in gas hourly space velocity (GHSV). Dottedlines represent the simulated results for the reaction in the ab-sence of the CaO pellets under the same operating conditionsas above, indicating that the reaction reaches at a steady stateshortly after the start of run. In the simulation, feed temper-ature, initial bed temperature, and reactor wall temperaturewere all set to 700◦C, and the feed pressure was 3 bar. Ex-cept for some scattered experimental data in CH4 composi-tion, excellent agreement exists between the experiment and

70

80

90

100simulated with CaOsimulated without CaOexperimental

0

3

6

9

12

0

2

4

6

0 40 80 120 160 200Time on stream, min

0

3

6

9

12

H2

CO2

CH4

CO

Pro

duct

gas

com

posi

tion,

%

Fig. 2. Product gas compositions on a dry basis with the reactiontime on stream: comparison of the simulation with the experiment(temperature = 700◦C, pressure = 3 bar, feed CH4 = 11:2 NL=h, feedH2O=CH4 = 3, GHSV = 341:5 h−1).

the simulation for the CaO carbonation-enhanced MSR. Theconcentration of CO2 in the product gas is kept low for the8rst 28 min of reaction due to its in situ removal by the car-bonation of CaO pellets in the reactor bed. During such a pre-breakthrough period, reaction enhancements are observed inthe H2 production, the CH4 conversion and the CO reduc-tion. As the degree of carbonation of all the CaO pellets inthe bed approaches the ultimate conversion, which dependson the local bed temperature, the eKuent concentration ofCO2 rises to a steady value, with the enhancement e<ectsbeing diminished. For both the prebreakthrough and post-breakthrough periods, the product gas composition is foundclose to that of equilibrium at each period. Equilibrium com-position for a steady-state period of prebreakthrough nearlycoincides with the simulation results represented by solidlines while that of postbreakthrough by dotted lines.Fig. 3 shows the reaction time-dependent axial distribu-

tions of CH4 conversion, reactor-bed temperature, and CaOcarbonation conversion. As shown in Fig. 3(A), CH4 con-version towards equilibrium has been accomplished mostlyin the 8rst quarter of the reactor bed. The reactor bed tem-perature at this MSR reaction-active region rapidly drops toabout 610◦C from the initial 700◦C, as shown in Fig. 3(B),due to the large endothermic heat of the MSR reaction. En-dothermic MSR in this region is fast enough to reach atthe equilibrium conversion of CH4, whereas the exothermicCaO carbonation reaction is quite slow, resulting in a largetemperature decrease. A more reactive CaO (larger poros-ity) would result in these rates being more equal, wouldlessen the large temperature decrease near the reactor en-trance. But, the heat e<ect by the exothermic carbonation ofCaO would not last for a long period of time because thecarbonation of CaO at near the reactor entrance would reachat the state of breakthrough, and after that, no more heatwould be generated. In this cool bed region, the rate of CO2

removal by the carbonation of CaO before its exhaustion ismostly kinetically limited because of the lowered temper-ature regardless of an abundance of CO2 available for thecarbonation. This also leads to a gradual increase in the CaOcarbonation conversion up to a certain position along thebed length, as shown in the pro8les of Fig. 3(C). In the sub-sequent high-temperature bed after this cool bed region, theendothermic MSR does not occur to a degree of practicalsigni8cance, but only the carbonation reaction of the packedCaO proceeds as long as it has not been exhausted. The rateof CaO carbonation in this bed region changes to be limitedby the small amount of CO2 available for the reaction. Lo-cal temperature of a bed region where only the carbonationreaction is occurring increases due to the exothermic heatof the carbonation reaction. A maximum increase to 25◦Cabove the initial temperature is observed at the reactor posi-tion 0:05 m from the entrance after 1 min of the reaction. Itcan be observed that the bed zone with such a temperatureincrease where the CaO carbonation is occurring moves to-ward the reactor exit with the reaction time on stream. Thetemperature rise at the reactor exit position is 15◦C after

D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942 937

Fig. 3. Time dependent spatial distribution of CH4 conversion (A), thereactor-bed temperature (B) and the carbonation conversion of CaO (C)in the simulated run shown in Fig. 2.

about 28 min of the reaction time on stream, which can beregarded as a breakthrough time of CaO in the packed bed.Like the present simulation, Han and Harrison (1994) re-ported a bed temperature increase of 10–15◦C in the WGSreaction using dolomite as a CO2 acceptor in the tempera-ture range of 500–600◦C.The carbonation conversion of CaO in the bed after

180 min of run ranges from 9.7% in the vicinity of the reac-tor entrance to 20.7% at the exit. Such simulated behaviorof the CaO carbonation can be reasonably acceptable. Itcan be therefore stated that the kinetics of the MSR reac-tion and that of the CaO carbonation, and the assumptionsmade in the current modeling are adequate to describe thetransient behavior of the in situ CaO carbonation-enhancedMSR reaction in a packed bed reactor.

60

70

80

90

100Temperature, °C

600650700750

0

5

10

15

0

5

10

15

0 40 80 120 160 200Time on stream, min

0

5

10

15

H2

CO2

CH4

CO

Pro

duct

gas

com

posi

tion,

%

Fig. 4. Simulated pro8les of the product gas composition on a dry basiswith the change in temperature (pressure=3 bar, feed CH4 =11:2 NL=h,feed H2O=CH4 = 3, GHSV = 341:5 h−1).

4.2. E;ect of temperature

For the in situ CaO carbonation-enhanced MSR reaction,temperature is a critical variable by which the extent of theultimate conversion of CaO as well as the rate of the car-bonation is determined. Kinetic rate and equilibrium gascomposition of the MSR reaction are also a<ected by tem-perature. Fig. 4 shows the simulated pro8les of the prod-uct gas composition with the change of temperatures in arange of 600–750◦C. The other operating conditions em-ployed here are the same as those in Fig. 2. At the lowtemperature of 600◦C, the maximum carbonation conver-sion of the CaO pellets is as low as 8.4% at the end of180 min run, resulting in the absence of any discrete pre-breakthrough period. At 650◦C, the prebreakthrough periodis extended to about 10 min by virtue of the increase inthe maximum conversion of CaO as 13.5%. At 750◦C, al-though there exists a long prebreakthrough period due tothe much enhanced conversion of CaO ranging from 12.8 atthe vicinity of the reactor entrance to 30% at the exit, CO2

938 D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942

Table 2Amounts of CO2 uptake and those of additional H2 production per kg ofCaO in the reactor bed by the in situ carbonation removal of CO2

Simulated operating conditions CO2 uptake H2 productivity(mol/kg-CaO) (mol/kg-CaO)

Temperature, (◦C) 600 1.43 1.56at 3 bar, 3 H2O=CH4, 650 2.29 2.54341:5 h−1 700 3.50 3.91

750 5.09 5.63

Pressure (bar) 3 3.50 3.91at 700◦C, 3 H2O=CH4, 7 3.52 3.98341:5 h−1 15 3.55 4.08

Feeding ratio of H2O=CH4 3 3.52 3.98at 700◦C, 7 bar, 341:5 h−1 5 3.53 4.00

7 3.54 4.01

GHSV (h−1) 341.5 3.54 4.01at 700◦C, 7 bar, 683.0 3.50 3.923 H2O=CH4 1366.0 3.37 3.74

concentration in the product gas during that period is ob-served as high as 3:1 mol% on a dry basis. This is becausethe equilibrium pressure of CO2 in the reversible reaction(4) of CaO-carbonation/CaCO3-calcination increases withincreasing temperature, as given by Eq. (16).Higher reaction temperature favors thermodynamically

the formation of CO by reactions (1) and (3) as well as theconversion of CH4, as resulted in Fig. 4. Resultantly, purityof H2 in the product gas during the prebreakthrough perioddecreases with higher concentration of CO as the tempera-ture increases. However, the total amount of H2 producedadditionally per kg of the bed-packed CaO pellets is in-creased with increasing temperature, as summarized in Table2. It can be found in Table 2 that corresponding amounts ofCO2 removed per kg CaO in the reactor by the carbonationincrease with temperature. Therefore, if the process has tobe operated for the direct production of high purity H2 whilekeeping CO as low as possible, low temperature operationis preferable at the sacri8ce of the decrease in the amount ofCO2 removed. Operation at higher temperature is preferredfor maximal carbonation–sequestration of CO2 and forthe maximal production of H2 regardless of the com-position of CO.

4.3. E;ect of pressure

Fig. 5 shows the simulated pro8les of the product gascomposition on a dry basis with the change in pressures be-tween 3 and 15 bar for the in situ CaO carbonation-enhancedMSR reaction (700◦C, CH4 feed rate = 11:2 NL=h, feedH2O=CH4=3, GHSV=341:5 h−1). The pressure drop alongthe total reactor bed in the simulation is small enough to beneglected. Pressure increase in the MSR reaction exerts toshift the equilibrium toward suppressing CH4 conversion, as

60

70

80

90

100Pressure, bar

3715

0

5

10

15

0

5

10

15

20

0 40 80 120 160 200Time on stream, min

0

3

6

9

12

H2

CO2

CH4

COP

rodu

ct g

as c

ompo

sitio

n, %

Fig. 5. Simulated pro8les of the product gas composition on a dry basiswith the change in pressure (temperature=700◦C, feed CH4=11:2 NL=h,feed H2O=CH4 = 3, GHSV = 341:5 h−1).

suggested by reaction (1). The MSR at higher pressure afterthe breakthrough of CaO generally gives a product gas con-taining less CO and more CO2, as shown in Fig. 5, becausereaction (1) forming CO is suppressed by pressure increase.Pro8les of CO2 composition in the product gases indicatethat the in situ CaO carbonation-enhanced MSR reaction isa<ected by the change of pressure. However, no direct e<ectof the total pressure on the CaO carbonation is consideredbecause the carbonation kinetics has been modeled assum-ing temperature dependence only. Instantaneous axial pro-8les at 10 min of reaction of CH4 conversion, temperature,and CaO conversion are displayed in Fig. 6. At that instant,the bed fronts where the CaO carbonation is taking place arelocated in the middle section of the reactor with only a smalle<ect of pressure. The most important thing noteworthy isthe degree of bed temperature rise, as shown in Fig. 6(B).The maximum temperature rises at that instant are 16, 12and 8◦C at the pressures 3, 7 and 15 bar, respectively. Oper-ation at higher pressure leads to a lower rate of CO2 produc-tion by suppressing the conversion of CH4, and resultantly,

D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942 939

0

20

40

60

80

100C

H4

conv

ersi

on, %

Pressure, bar3715

600

620

640

660

680

700

720

Tem

pera

ture

,o C

0 0.05 0.1 0.15 0.2 0.25 0.30

5

10

15

Ca

O c

on

vers

ion

,%

(A)

(B)

(C)

Axial distance,m

Fig. 6. Axial instantaneous pro8les of CH4 conversions (A), temperatures(B), and CaO conversions (C) at 10 min of the reaction time on streamin the simulated run shown in Fig. 5.

the rate of CaO carbonation is slow. Accordingly, the bedtemperature rise is small. As mentioned earlier, higher bedtemperature rise gives a product stream with higher equilib-rium concentration of CO2 while CaO is still available forfurther carbonation in the bed, as shown in Figs. 5 and 6.Operation at higher pressure also contributes to maintain-

ing the reactor temperature close to the set value, as shownin Fig. 6(B). The degree of reactor bed cooling appearssmaller during operation at higher pressure by suppressingthe endothermic CH4 conversion. Corresponding to this, theobserved degree of CaO carbonation conversion is higher,as shown in Fig. 6(C). The total amount of H2 producedadditionally per kg of the bed-packed CaO pellets is some-what increased with pressure, as listed in Table 2. The cor-responding amounts of CO2 removed in the reactor by theCaO carbonation are increased as well.

4.4. E;ects of the feed H2O=CH4 and the feed spacevelocity

Fig. 7 shows the simulated composition pro8les of theproduct gas on a dry basis as function of the molar ratio offeed H2O=CH4 for the in situ CaO carbonation-enhancedMSR reaction (700◦C, 7 bar). In this simulation, the feed

60

70

80

90

100H2O/CH4 ratio

357

0

5

10

15

0

5

10

15

0 40 80 120 160 200

Time on stream, min

0

3

6

9

12

H2

CO2

CH4

CO

Pro

duct

gas

com

posi

tion,

%

Fig. 7. Simulated pro8les of the product gas composition on a dry basiswith the change in the molar ratio of feed H2O=CH4 (temperature=700◦C,pressure = 7 bar, feed CH4 = 11:2, 7.47, 5:6 NL=h for the feedH2O=CH4 = 3, 5, 7, respectively, GHSV = 341:5 h−1).

Mow rates of CH4 are varied to 11.2, 7.47, and 5:6 NL=h cor-responding to a change of the feed H2O=CH4 ratio to 3, 5,and 7, respectively, to get the total feed rate be the same as44:8 NL=h, equivalent to 341:5 h−1 in GHSV. After the CaObreakthrough, the MSR at higher feed ratio of H2O=CH4

gives a product stream containing less CO and more CO2,with higher conversion of CH4, as shown in Fig. 7. Due tothe decreased feed rate of CH4, the prebreakthrough periodis extended with increasing feed ratio of H2O=CH4. At thefeed ratios of 5 H2O=CH4 and above, it can be noticed thatthe production of H2 with a purity of about 98% is possibleduring the prebreakthrough period. The total amounts of H2

produced additionally per kg of the bed-packed CaO pelletsare not largely di<erent by the feed ratios of H2O=CH4,being given by about 4:0 mol-H2/kg-CaO, as shown inTable 2.Fig. 8 shows the simulated composition pro8les of the

product gas on a dry basis at di<erent space velocities

940 D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942

70

80

90

100GHSV, hr-1

341.56831366

0

5

10

15

0

2

4

0 40 80 120 160 200Time on stream, min

0

2

4

6

H2

CO2

CH4

CO

Pro

duct

gas

com

posi

tion,

%

Fig. 8. Simulated pro8les of the product gas composition on a dry basiswith the change in the GHSV (temperature=700◦C, pressure=7 bar, feedH2O=CH4 = 7, feed CH4 = 5:6, 11.2, 22:4 NL=h for the GHSV = 341:5,683, 1366 h−1, respectively).

(341.5, 683, and 1366 h−1 in GHSV), 700◦C, 7 bar, feedH2O=CH4 = 7. As can be expected, the increase of theGHSV results in a decrease of the prebreakthrough period.The temperature distribution of the reactor bed along the ax-ial distance is a<ected by the GHSV because higher GHSVmakes the cool region of the bed wider in the axial direc-tion, where endothermic MSR takes place. As a result, thecarbonation-conversion pro8les are di<erent due to the dif-ferent temperature distribution of the bed. Fig. 9 shows theaxial pro8les of the 8nal carbonation-conversions of theCaO at the end of the simulated run shown in Fig. 8. It canbe seen that lower conversion of CaO is attained locallyalong the reactor bed with higher GHSV. In this simulatedrun, the amounts of CO2 removed and those of H2 producedadditionally by the carbonation of CaO throughout the bedare decreased with increasing GHSV, as shown in Table2. From the standpoint of carbonation-utilization of CaOpacked in the bed, it is better to operate the process at lowerGHSV.

0 0.05 0.1 0.15 0.2 0.25 0.3Axial distance, m

0

5

10

15

20

25

Ca

O c

on

vers

ion

, %

GHSV, hr -1

341.56831366

Fig. 9. Axial bed pro8les of the 8nal carbonation-conversions of the CaOat the end of the simulated run shown in Fig. 8.

4.5. Operating options of the in situ CaOcarbonation-enhanced MSR

As a summary for some 8ndings above about the tran-sient in situ CaO carbonation-enhanced MSR reaction, itshould be 8rst mentioned that the reaction at lower temper-atures than 650◦C fails to give a signi8cant conversion ofthe CaO pellets prepared in this study. As shown in Table2, the amount of CO2 uptake and the productivity of H2

per kg-CaO by CaO carbonation are increased with increas-ing temperature up to 750◦C. However, operation at highertemperatures su<ers not only from higher outlet CO2 con-centration due to the increased partial pressure in the equilib-rium of CaO-carbonation/CaCO3-calcination, but also fromthe production of CO in higher concentration. If the pro-cess is operated in pursuit of as large CO2 sequestration aspossible and large production of H2 regardless of CO con-centration, higher temperature up to a maximum 750◦C willbe better. The amount of CO2 uptake by CaO carbonation,as shown in Table 2, appears much larger than those by thehydrotalcite-based chemisorbents with about 0:45 mol=kg(Hufton et al., 1999). For a process option of the direct pro-duction of H2 with as low carbon oxides concentrations aspossible, operation at 650◦C is the best choice with the dis-advantage of very limited carbonation conversion at lowertemperature than that.As shown in Table 2 and discussed above, operation at

higher pressure is bene8cial for increasing the conversionof CaO in the reactor-bed at a given condition, and also forlowering the concentration of CO in the product. Raisingthe feed ratio of H2O=CH4 is helpful for increasing theconversion of CH4 as well as lowering the product CO

D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942 941

Table 3Analysis for the reaction performance and the eKuent product gases during the prebreakthrough period in the simulated run with the changes in pressure

Pressurea (bar) 7 15 30 50 80

CH4 conversion (%) 98.7 97.4 95.3 93.0 90.0H2 production rate (mol/h) 1.10 1.08 1.05 1.01 0.94H2 purityb (mol%) 99.3 99.1 98.7 98.1 97.1CO compositionb (ppm) 1230 571 282 170 112CO2 compositionb (ppm) 3020 1430 727 456 315CH4 compositionb (mol%) 0.32 0.66 1.22 1.87 2.84

aThe other operating conditions: temperature = 650◦C, feed CH4 = 5:6 NL=h, and feed H2O=CH4 = 7 (GHSV = 341:5 h−1).bPurity and compositions on a dry basis.

concentration. The degree of overall utilization of CaO pel-lets in the reactor-bed could be increased by lowering theGHSV. Increased pressure, high ratio of H2O=CH4, and de-creased GHSV are favorable for both of the two processoptions at the temperatures investigated here.For the case of a process option to produce a pure H2

stream, there is a need to investigate the e<ect of pressure onthe composition of CO in the product stream at the operat-ing conditions as follows: temperature=650◦C, feeding rateof CH4 =5:6 NL=h, and feed H2O=CH4 =7 (GHSV of totalfeed gas = 341:5 h−1). Table 3 provides the analysis of thegas products obtained during the prebreakthrough period inthe simulated run at such conditions with changing pressureup to 80 bar. The prebreakthrough periods are maintainedfor 30, 33, 35, 37, and 40 min at the pressures of 7, 15, 30,50, and 80 bar, respectively. At these simulation conditions,the data show that a very high pressure is required for reduc-ing the product concentration of CO down to the level of afew hundreds ppm. Hufton et al. (1999) demonstrated thatit was possible to produce 95 mol% H2 stream containing5 mol% CH4 and less than 50 ppm carbon oxide impuritiesin their SERP using a CO2 chemisorbent at the operationconditions: temperature=450◦C, pressure=4:8 bar, and thefeed H2O=CH4 = 6. A key to the production of a H2 streamwith such a low carbon oxide concentration in their experi-ment is that the chemisorbent was working at reaction tem-peratures as low as 450◦C. At 650◦C of the simulated reac-tion temperature, the pressure required for the product withabout 50 ppm of carbon oxides is expected above 100 bar.Consequently, it is thought that targeting the direct produc-tion of pure H2 with carbon oxides less than 50 ppm is notappropriate in the in situ CaO carbonation-enhanced MSRreaction as long as any CaO-based chemisorbent with a con-siderable working capacity at temperatures much lower than650◦C has not been developed.

5. Conclusions

A dynamic model derived for the CaO carbonation-en-hanced MSR reaction at non-isothermal, non-adiabatic,and non-isobaric operating condition was shown to ac-curately depict the transient behavior of the in situ CaO

carbonation-enhanced MSR reaction. The reaction at lowertemperatures than 650◦C failed to give a practical conver-sion of the CaO pellets used in this study. The amount ofCO2 uptake and the productivity of H2 per kg-CaO by theCaO carbonation were increased with increasing tempera-ture up to 750◦C. High CO2 uptake by the CaO carbonationwould be very advantageous in the process application forthe maximal sequestration of CO2 and production of H2

compared to the hydrotalcite-based chemisorbent. Oper-ation at higher temperatures, however, resulted not onlyin the release of CO2 in higher concentration but also theproduction of CO in higher concentration during the pre-breakthrough period. Operations at high pressure, high ratioof feed H2O=CH4, and decreased feed rate at a given tem-perature were shown favorable for increasing the degree ofoverall carbonation-utilization of CaO pellets in the reactorbed, and for lowering the product concentration of CO.In particular, the operation at higher pressure was helpfulfor maintaining the reactor temperature close to the initialset value by adequately suppressing the endothermic MSRreactions and the exothermic CaO carbonation reaction,by which a front cool zone and a rear hot zone along thereactor bed were brought about, respectively. For the appli-cation of the in situ CaO carbonation-enhanced MSR to thedirect production of pure H2 with carbon oxides less than50 ppm, a CaO-based chemisorbent with a considerableworking capacity at temperatures much lower than 650◦Cshould be developed.

Notation

b time taken for half the ultimate conversionof CaO, s

Ceq;CO2equilibrium concentration of CO2 in reac-tion (4), kmol=m3

Ci concentration of component i, kmol=m3

Ci;f concentration of component i in the feed,kmol=m3

Ci; init initial concentration of component i,kmol=m3

Cpg heat capacity of gas, kJ/kmolKCps heat capacity of solid, kJ/kmolK

942 D.K. Lee et al. / Chemical Engineering Science 59 (2004) 931–942

dp particle diameter, mDR inside diameter of reactor, mhW heat transfer coeScient through reactor

wall, kJ=m2 sKHcbn reaction heat of CaO carbonation, kJ/kmolHRj reaction heat of reaction j (j=1–3),

kJ/kmolkc apparent rate constant of CaO carbona-

tion, s−1

kg thermal conductivity of gas, kJ/m sKkj rate constant of reactions j, j = 1; 2:

kmol bar0:5=kg-cat s; j=3:kmol/kg-cat s bar

ks thermal conductivity of solid, kJ/m sKk0z static e<ective thermal conductivity,

kJ/m sKKi adsorption constant for component i,

i=CO, H2, CH4: bar−1; i=H2O: dimen-

sionlessKj equilibrium constant of reactions j,

j = 1, 2: bar2; j = 3: dimensionlessMCaO molecular weight of CaO, kg/kmolP pressure, barPf pressure at reactor entrance, barPi initial reactor bed pressure, barrcbn rate of CO2 removal by CaO carbonation,

kmol/kg-CaO sri rate of formation of component i,

kmol/kg-cat sRj rate of reaction j (j=1–3), kmol/kg-cat sRep particle Reynolds number, dimensionlesst time, sT temperature, KTf feed gas temperature, KTinit initial temperature of the reactor bed, KTW reactor wall temperature, Ku super8cial velocity of gas, m/sWCaO weight of the CaO pellets packed in the

reactor, kgWcat weight of catalyst pellets packed in the

reactor, kgX fractional carbonation conversion of CaO,

dimensionlessXu ultimate fractional carbonation conversion

of CaO, dimensionlessz axial distance of reactor from the entrance,

dimensionless

Greek letters

� parameter de8ned in Eq. (14), s−1

� bed void fraction, dimensionless� catalyst e<ectiveness factor, dimension-

less% viscosity of gas, Pa s�CaO apparent density of CaO pellets, kg=m3

�cat apparent density of catalyst pellets, kg=m3

�s average apparent density of the two mixedsolids in reactor, kg=m3

�g density of gas, kg=m3

CO2 parameter de8ned in Eq. (15), kmol=m3 s

Acknowledgements

This work was carried out through a Strategic NationalR&D Program with 8nancial support from the Ministry ofScience and Technology (MOST), Republic of Korea.

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