paper 2 (pdf)

Chemical Engineering Science 55 (2000) 3929}3940

Adsorption-enhanced steam}methane reforming

Y. Ding, E. Alpay*Department of Chemical Engineering and Chemical Technology, Imperial College of Science, Technology and Medicine, Prince Consort Road,

London SW7 2BY, UK

Received 8 September 1999; received in revised form 6 December 1999; accepted 13 December 1999

Abstract

Experimental and theoretical studies of steam}methane reforming in the presence of a hydrotalcite-based CO2

adsorbent arepresented. Attention is given to the analysis of the transient behaviour of a tubular (integral) reactor when an Ni-based catalyst isadmixed with the adsorbent. Considerable enhancement of the methane conversion is experimentally demonstrated. Enhancementarises from the favourable shifts in the reaction equilibria of the reforming and water}gas shift reactions towards further CO

2production. As predicted, the potential for conversion enhancement is shown to increase under the conditions of a high reactor spacetime, high operating pressure, or a low steam-to-methane feed ratio, i.e. when reaction equilibrium limitations are important.A mathematical model, accounting for mass transfer limited adsorption kinetics, non-linear (Langmuirian) adsorption equilibria anda general reaction kinetic model, is shown to accurately predict the observed elution pro"les from the reactor, and thus the degree ofconversion enhancement. ( 2000 Elsevier Science Ltd. All rights reserved.

Keywords: SMR; Sorption enhancement; Transient kinetics; Ni-catalyst; Hydrotalcite; CO2

adsorbent; Mathematical modelling

1. Introduction

The advantages of coupling reaction systems withsome form of in situ separation have been widely re-ported in the literature. Such hybrid con"gurations maysubstantially improve reactant conversion or productselectivity and, for reversible reactions, establish a morefavourable reaction equilibrium than that which could beachieved under conventional reactor operation. Reactionenhancement may enable a lower temperature of opera-tion, which in turn may alleviate the problems associatedwith catalyst fouling, high process energy requirementsand poor energy integration within the plant environ-ment. For gas-phase catalytic reactions, the separationcan be based on adsorption, selective permeationthrough a membrane, or through simultaneous reactionof the targeted molecule (e.g. the reaction inhibitor) witha chemical acceptor.

A comprehensive review on membrane-based reactionsystems has been given by Armor (1995). Advances havebeen made in the use of metallic membranes (oftenGroup VIII metals which only small molecules like

*Corresponding author. Tel.: 0044-171-594-5625; fax: 0044-171-594-5604.

E-mail address: [email protected] (E. Alpay).

hydrogen can permeate) and, more recently, polymeric,ceramic and zeolitic membranes. The membranes mayact as permselective barriers, or as an integral part of thecatalytically active surface. Practical issues such as mem-brane pore blockage, thermal and mechanical stability,and the dilution caused by the need for sweep (i.e. per-meate purge) gases, have limited the usefulness of themembrane reactor systems. Nevertheless, the bene"ts ofthe membrane systems have been demonstrated thougha wide number of experimental reaction studies, exam-ples of which include the dehydrogenation of ethane(Tsotsis, Champagnie, Vasileiadis, Zraka & Minet, 1992),cyclohexane (Sun & Khang, 1988), ethylbenzene (Wu,Gerdes, Pszczolowski, Bhave & Liu, 1990), and acetylene(Itoh, Xu & Sathe, 1993), CO production via thewater}gas shift reaction (Uemiya, Sato, Ando & Kikuchi,1991), and steam}methane reforming (Adris, Lim& Grace, 1994, 1997).

In comparison to the membrane reactors, a relativelysmall amount of work has been carried out on systemscombining reaction with adsorption or chemical accep-tor-based separation processes. Even so, such processeso!er distinct advantages to the membrane-based systemsin terms of the material tolerance to high temperaturesand pressures, and the wide choice and availability ofadsorbents for achieving the desired separations under

0009-2509/00/$ - see front matter ( 2000 Elsevier Science Ltd. All rights reserved.PII: S 0 0 0 9 - 2 5 0 9 ( 9 9 ) 0 0 5 9 7 - 7

reaction conditions. Furthermore, even through use ofa purge gas for regeneration, the e!ective separation ofthe primary adsorbate from other non- or weakly adsor-bing species can be achieved. Some examples of recentworks employing chemical acceptors or adsorbents forreaction enhancement are now summarised.

Han and Harrison (1994) studied hydrogen productionvia the water}gas shift reaction using CaO as a CO

2acceptor in a tubular reactor. CO conversions werereported which exceeded that of the thermodynamicequilibrium conversion under the speci"ed operatingconditions. Brun-Tsekhovoi, Zadorin, Katsobashvili andKourdyumov (1986) showed a very signi"cant enhance-ment of CH

4conversion to H

2in a #uidised bed reactor

containing Ni-based catalyst balls, and a specially treatedform of dolomite as adsorbent. Typical industrial operat-ing conditions were considered in this work, i.e. pressurelevels of 103}104 kPa, and an operating temperature of6273C. Goto, Tagawa and Oomiya (1993) studied thedehydrogenation of cyclohexane over a Pt}alumina cata-lyst and CaNi

5alloy as a hydrogen acceptor. The

workers showed that at 150}1903C and ambient pres-sure, the overall conversion of cyclohexane to benzenecould be exceeded by three-fold when compared to thecatalyst-only case. Most recently, Carvill, Hufton andSircar (1996) and Hufton, Mayorga and Sircar (1999)describe the general concept of the Sorption EnhancedReaction Process (SERP), which utilises pressure andconcentration swing adsorption principles for reactionenhancement; see also Vaporciyan and Kadlec (1989) andAlpay, Chatsiriwech and Kershenbaum (1995). CO pro-duction via the reverse water}gas shift reaction was spe-ci"cally considered by Carvill et al. (1996), in which NaXzeolite was used as an adsorbent for water. The authorsshowed that at 2503C and 480 kPa, a CO

2conversion of

36% could be achieved; a conventional plug #ow reactorrequired an operating temperature of 5653C for the sameconversion. Furthermore, the process generated a highpurity (#99% (v/v)) CO stream as product. Hufton et al.(1999) applied the SERP concept to H

2production via

the steam}methane reforming (SMR) reactions. In speci-"c, using a hydrotalcite-based CO

2adsorbent, and

a commercial Ni-based catalyst, the authors showed thatat 4503C and 480 kPa, #95% (v/v) H

2could be produc-

ed directly from reactor. The CH4

to H2

conversion was82%, which could only be achieved at approximately6503C with a conventional SMR reactor.

The present work also considers the sorption-en-hanced steam}methane reforming (SE-SMR) process.The key reactions of the SMR process are given as:

CH4#H

2O Q CO#3H

2, *H

298"206 kJ/mol, (1)

CH4#2H

2O Q CO

2#4H

2, *H

298"164.9 kJ/mol,

(2)

CO#H2O Q CO

2#H

2, *H

298"!41 kJ/mol. (3)

Reforming reactions (1) and (2) are strongly endothermic,so the forward reaction is favoured by high temperatures,while the water}gas shift reaction (3) is moderatelyexothermic and is therefore favoured by low temper-atures. The reforming reactions will also be favoured atlow pressures, whereas the water}gas shift reaction islargely una!ected by changes in pressure. In the presenceof a selective CO

2adsorbent, the conversion of CH

4to

CO2

though reaction (2) is favoured, as is the productionof CO

2through CO intermediate. For a reaction which

is not kinetically limited, the use of an adsorbent will thusenable a lower operating temperature for a desired con-version. However, on equilibration of the adsorbent, theseparation e!ect is, of course, lost. This necessitates theperiodic regeneration of the adsorbent and thus, forexample, the pressure and concentration swing type op-erations mentioned above. In other words, sorption-en-hanced reaction processes are inherently dynamic inoperation. Adequate design and scale-up of such pro-cesses will thus require information on the kinetics ofadsorption and desorption, as well as reaction kineticmodels under transient conditions in the presence of anadsorbent.

Research work on the kinetics of the SMR processdates back to the 18th century (see Sabatier, 1922; Marek& Hahn, 1932), with the "rst extensive study by Akersand Camp in 1955. A good review of the work up until1970 is given by Van Hook (1980), which covers kineticstudies over porous nickel catalysts and nickel foilover large temperature (260}10003C) and pressure(100}5000 kPa) ranges. A considerable amount of workon the kinetic aspects of the SMR process has beencarried out since 1970; see, for example, Schnell (1970),Ross and Steel (1973), Allen (1975), Phung Quach andRouleau (1975), Munsted and Grabke (1981), De Deken,Deves and Froment (1982), and Xu and Froment (1989a,b). To date, the rate models proposed by Xu and Fro-ment (1989a) are considered to be the most general inform, and have been extensively tested under typicalindustrial operating conditions (see, also, Elnashaie,Adris, Al-Ubaid & Soliman, 1990). However, like mostprevious work, the models are applicable to steady-statekinetics, and untested for forced-dynamic operation. It isinteresting to note, however, that where some attentionhas been given to the transient kinetics, ideal surfaceconditions were maintained though vacuum operations;see, for example, Ross and Steel (1973). This, of course,limits the applicability of such models for process designapplications.

In this work, attention is given to the analysis of SMRreaction kinetics under transient conditions depictingSERP-type operation, both in the presence and absenceof a selective absorbent for CO

2. Particular attention is

given to the transient analysis of the Xu and Froment(1989a) kinetic model. The work then considers the in#u-ence of operating parameters on the degree of separation

3930 Y. Ding, E. Alpay / Chemical Engineering Science 55 (2000) 3929}3940

Fig. 1. Schematic representation of the experimental system, MFC* mass#ow controller, PC * personal computer.

enhancement. In doing so, mathematical models of theprocess are veri"ed, which can ultimately be used in thedesign, analysis or scale-up of pressure or concentrationswing based adsorptive reactors.

2. Experimental

A commercial Ni-based catalyst (United Catalyst Inc.)containing 25}35% Ni, 25}35% NiO, 5}15% MgO and15}25% sodium silicate, was used in this work. Theoriginal catalyst (1/8A cylindrical pellets; BET area of137.6 m2/g) was crushed and sieved into two particle sizegroups: 0.11}0.25 and 0.25}0.5 mm. The CO

2adsorbent

consisted of an industrially supplied potassium promotedhydrotalcite, which was previously measured for its capa-city and stability under wet gas conditions (see Ding& Alpay, 2000). The adsorbent was also crushed andsieved into the particle size groups mentioned above. Asa catalyst diluent, e.g. in the absence of adsorbent, siliconcarbide particles were employed. High-purity methane(99.95% (v/v)) and hydrogen (99.995% (v/v)) gases weresupplied from gas cylinders; steam supply to the SMRreactor was generated from distilled water.

A schematic diagram of the experimental apparatus isgiven in Fig. 1. The reactor consisted of a stainless-steeltubular column of internal diameter 12.4 mm and length220 mm, packed with a mixture of catalyst and adsorbent(or silicon carbide) particles. The reactor was "tted withinlet and outlet lines for introducing feed gas (CH

4and

H2O), purge as (H

2/He and H

2O) and reducing gas (H

2).

The inlet CH4#ow rate was controlled by a Brooks

5850E mass#ow controller. An HPLC pump was used tosupply water to the reactor via a vaporiser (i.e. a heatedtubular column packed with silicon carbide particles);both the reactor and water vaporiser were mounted in

a convection oven. The outlet #ow rate from the reactorwas monitored with an Aalborg GFM-17 mass #owmeter and a soap-bubble #owmeter. The catalyst bedtemperature was detected with a type K thermocouple(positioned at the centre of the reactor along the centralaxis), and a back pressure regulator used to maintaina constant reactor pressure. The reactor e%uent #ow wassplit into sample and vent lines, each equipped witha water condenser. For the former, time delays in thesample analysis were minimised by use of 1/16A transferlines, and a low-volume condenser unit. The sample linewas connected to a Valco 16-loop valve, the operationalschedule of which was computer controlled. The sampleline and the sample valve were heat traced, and thetemperature controlled at 1103C by a PID controller.A Shimadzu gas chromatograph (GC-14B), equippedwith a TCD detector and a Porapak-Q column, was usedfor sample analysis. In addition, two on-line TeleganCO

2analysers (0}30 and 0}5%FSD) were used to moni-

tor the reactor e%uent CO2

concentration.For the reaction studies in the absence of adsorbent,

approximately 7.2 g of catalyst was admixed with densesilicon carbide particles (&1 : 3 mass ratio), and packedinto the reactor. For the sorption-enhanced reactionstudies, approximately 7.2 g of catalyst was admixed with14.8 g of CO

2adsorbent. Reactor operating conditions

for both the adsorbent and adsorbent-free systems aresummarised in Table 1.

As mentioned above, conversion enhancement arisesprior to the equilibration of the adsorbent. In this work,transient operation was imposed by means of step cha-nges in feed concentration. Due to the negligible resi-dence time of methane in the reactor, even in the presenceof adsorbent (i.e. typically less than 0.2 s), the conversionenhancement factor at any given time (E(t)) can be quan-ti"ed by the normalised conversion of methane (X

CH4) in

the presence (ad) and absence (nad) of adsorbent, i.e.

E(t)"(X

CH4)!$!(X

CH4)/!$

(XCH4

)/!$

100. (4)

Thus, values of E(t)'0 indicate conversion enhance-ment in the presence of adsorbent. At steady state, andfor similarly packed reactors (i.e. catalyst mass loadingand distribution), E(t) should approach 0. A typical ex-perimental cycle involved the following steps: (i) heat-upof the reactor at atmospheric pressure under a hydrogenenvironment to the preset temperature, (ii) water supplyto the system so that the molar ratio of H

2O to H

2is

approximately the same as the desired H2O-to-CH

4ratio in the reaction step, (iii) pressurisation of the systemto the preset pressure, (iv) switch from H

2to CH

4to

initiate the reaction step, (v) depressurisation of the unit,(vi) low-pressure purge of the unit with H

2and steam,

and (vii) reduction of the catalyst with H2

at 4803C for3 h. Steps (vi) and (vii) were carried out as precautions

Y. Ding, E. Alpay / Chemical Engineering Science 55 (2000) 3929}3940 3931

Table 1List of experimental!

Run type" Space time(g-cat h/mol-CH

4)

H2O/CH

4Pressure(kPa)

Temperature(3C)

Particle size(mm)

WA/NA 5.37}17.9 3 445.72 455 0.25}0.5WA/NA 10.7 2}6 445.72 455 0.25}0.5WA/NA 10.7 3 308}721.5 457.5 0.25}0.5WA/NA 10.7 3 445.72 428}467.5 0.25}0.5WA 10.7 3 445.72 459 0.11}0.25

!Initial reactor environment: H2O : H

2(He)"H

2O : CH

4"WA * with adsorbent, NA * no adsorbent.

towards any carbonaceous deposits on the catalyst, andthus to ensure a reproducible catalyst activity from oneexperiment to the next. Reproducibility was subsequentlycon"rmed through repetition of the experiments. Themeasurement of the reactor e%uent concentration pro-"les in the reaction step (step (iv)) enabled both transientand steady-state reaction kinetics to be tested. Note thatdue to non-isothermal nature of the SMR process, themeasured wall temperature at the middle point of thereactor is taken as the characteristic temperature in thiswork.

3. Mathematical modelling

3.1. Governing equations

A dynamic model accounting for non-isothermal,non-adiabatic, and non-isobaric operation, was de-veloped to describe both the SMR and SE-SMRprocesses. For the SE-SMR process, the reactions andadsorption were assumed to take place on the surfaces ofthe catalyst and adsorbent, respectively. A Langmuirmodel was used to describe the adsorption equilibria ofCO

2, and a linear driving force (LDF) model for the

intraparticle mass transfer of the adsorbent; further de-tails of equilibria and kinetic measurements, and math-ematical model development under non-reactingconditions, are given by Ding and Alpay (2000). As men-tioned above, the general reaction kinetic model pro-posed by Xu and Froment (1989a) was considered in thiswork. Other principal model assumptions can be sum-marised as: axially dispersed plug #ow, perfect gas behav-iour, no radial concentration or temperature gradients,and a catalyst/adsorbent packing of uniform voidage andparticle size.

Based on the above assumptions, component massbalances for the packed-bed reactor can be written as

LLt

(etC

i#o

b(!$)qi)"

LLzADz

LCi

Lz B!L(uC

i)

Lz#g

iob(#!5)

ri,

(5)

where i denotes CH4, H

2O, H

2, CO

2CO, and He. the

semi-empirical correlation proposed by Edwards andRichardson (1968) was used to estimate the axial disper-sion coe$cient D

z. The catalyst e!ectiveness factor, g

i,

was set as unity for the crushed catalyst used in this work;see Section 4 for details. The reaction kinetic model of Xuand Froment (1989a) can be summarised as

RI"

k1

P2.5H2APCH4

PH2O

!

P3H2

PCO

KIBN(DEN)2, (6a)

RII"

k2

P3.5H2APCH4

P2H2O

!

P4H2

PCO2

KII

BN(DEN)2, (6b)

RIII"

k3

PH2APCO

PH2O

!

PH2

PCO2

KIII

BN(DEN)2, (6c)

DEN"1#KCO

PCO

#KH2

PH2

#KCH4

PCH4

#KH2O

PH2O

/PH2

, (6d)

where Rj( j"I!III) denotes the reaction rate of the

SMR reactions (1) and (2) and the water}gas shift reac-tion (3). The formation rate of component i, r

i, was

then calculated by using Eqs. (1)}(3) and (6a)}(6d);for example, r

CH4"!(R

I#R

II), r

H2"3(R

I#R

III)#

(RII#R

III). The rate constants k

i(i"1}3) and the ad-

sorption constants Kj( j"I}III) are function of temper-

ature, details of which are given by Xu and Froment(1989a).

Pressure distribution in the packed-bed reactor wasdescribed by the Ergun equation (Ergun, 1952)

LP

Lz"!K

Du!K

Vu2, (7)

where KD

and KV

are parameters corresponding to theviscous and kinetic pressure loss terms, respectively.Semi-empirical relationships for K

Dand K

Vhave been

derived by Ergun (1952) (see also MacDonald, El-Sayed,Mow & Dullien, 1979).

For compressible #ow, the energy balance for thereactor is given by (see, for example, Bird, Stewart &


Table 2Parameters (constants) used in the simulations

Parameters (constants) Value Unit

bCO2

SE-SMR: 2.36]10~4; SMR:0 Pa~1

dp

(average) 1.8]10~4, 3.57]10~4 mD

r1.27]10~2 m

Cpg

42 J/mol KC

ps850 J/kg K

e 0.35 *

kg

0.09 J/m Kks

0.3 J/m KH

!$ CO2SE-SMR: !17000, SMR: 0 J/mol

¸ (reactor length) 0.22 mm

CO2SE-SMR: 0.65, SMR: 0 mol/kg

b1

0.95 *

eb

0.47 *

et

0.64 *

U 0.2 *

c 0.667 *

js

1.0 *

kg

2.87]10~5 Pa sob(!$)

SE-SMR: 609.3, SMR:0 kg/m3

ob(#!5)

139.0 kg/m3

Lightfoot, 1960)

LLt

(ogC

pget¹#o

bC

ps¹)

"et

LP

Lt#

LLzAKz

L¹Lz B!

LLz

(oguC

pg¹)

!ob(!$)

+iAH!$i

Lqi

Lt B#+i

(Riob(#!5)

giH

Ri)

#

4;0

Dr(¹

w!¹). (8)

The bed e!ective conductivity, Kz, can be expressed as

(see Yagi & Kunii, 1957; Kunii & Smith, 1960; De Wasch& Froment, 1972; Li & Finlayson, 1977; Wakao& Kaguei, 1982)

Kz

kg

"

K0z

kg

#a(Pr) (Rep), (9)

where K0z

is the static e!ective conductivity accountingfor the e!ects of conduction and radiation, see Kunii andSmith (1960) and Ding and Alpay (2000).

For a bed packed with spherical particles, the wall}bedheat transfer coe$cient, ;

0, in Eq. (8) is given by Li and

Finlayson (1977) as

;0D

rkg

"2.03Re0.8p

expA!6d

pD

rB

(Rep"20}7600, d

p/D

r"0.05}0.3). (10)

Eq. (10) is not applicable to very low feed #ow rates, as;

0should approach a "xed value where Re

pP0. How-

ever, De Wasch and Froment (1972) give the followingcorrelation for ;

0at very low Re

p:

;0DRep?0

"6.15K0

zD

r

(11)

which can be linearly combined with Eq. (10) to approx-imate ;

0over the entire range of Re

prelevant to this

work.The Langmuir model for CO

2adsorption can be

written as

qHCO2

"

mCO2

bCO2

PCO2

1#bCO2

PCO2

. (12)

Other reaction components were considered to be non-adsorbing. The LDF model is given by

Lqi

Lt"k

i(qH

i!q

i), (13)

where k is the e!ective mass transfer coe$cient. Note thatLq

i/Lt"0 for non-adsorbing species in the SE-SMR pro-

cess, and Lqi/Lt"0 for all species in the SMR process. As

mentioned above, Langmuir and LDF parameters havebeen previously measured for the CO

2}H

2O-hydro-

talcite system; see Ding and Alpay (2000).

3.2. Boundary and initial conditions

For the simulation of the reaction step, initial condi-tions (t"0, z3(0, ¸)) were set to depict a clean andisothermal bed, but with a pre-imposed pressure pro"le:

qi"0, (14a)

LP

Lz"!K

Du0!K

Vu20, (14b)

LP/Lt"0, (14c)

¹"¹0, (14d)

Ci"P

i/ (R¹

0) (14e)

At the onset of the reaction step (i.e. the point at whichCH

4is supplied to the reactor), the following boundary

conditions were used in the simulations:(i) Reactor entrance (z"0)

!Dz(LC

i/Lz)"u(C

fi!C

i), (15a)

!Kz(L¹/Lz)"uo

gC

pg(¹

f!¹), (15b)

u"Qf/A. (15c)

(ii) Reactor outlet (z"¸)

LCi/Lz"0, (15d)

L¹/Lz,0, (15e)

P"PL. (15f )

Qf

in Eq. (15c) is the feed volumetric #ow rate measuredunder the local conditions.


Fig. 2. E%uent concentration (mole fraction) pro"les on a water-freebasis: comparison between experiments and simulations (reactorpressure"445.7 kPa, temperature"450$23C, feed #ow rate ofCH

4"250 ml/min (STP), molar ratio of H

2O to CH

4"6, cata-

lyst/adsorbent (or carbide) particle size"0.25}0.5 mm).

Fig. 3. Comparison of methane conversion pro"les and CO2

break-through curve (reactor pressure"445.7 kPa, temperature"450$23C, feed #ow rate of CH

4"250 ml/min (STP), molar ratio

of H2O to CH

4"6, catalyst/adsorbent (or carbide) particle

size"0.25}0.5 mm).

Eqs. (5)}(15) were simultaneously solved in thegPROMS modelling environment (Process Systems En-terprises Ltd.). The spatial discretisation method of or-thogonal collocation on "nite elements was employed.Second-order collocation on 100 elements was found togive a converged solution in which component balanceerrors (associated with the numerical integration) did notexceed 1%. A summary of parameters (constants) usedfor the simulations is given in Table 2.

4. Results and discussion

Typical reactor e%uent concentration pro"les mea-sured in the SMR and SE-SMR experiments undersimilar operating conditions (4503C, 445.7 kPa, H

2O/

CH4"6) are shown in Fig. 2(a) and (b), respectively. It

can be seen that in the presence of CO2

adsorbent, thetransient period (&5 min) is much longer than that of

the SMR run (&1.5 min). However, since the experi-mental conditions are nearly identical, the steady-stateconcentrations of CH

4, H

2, and CO

2are approximately

equal, indicating no strong direct catalyst/adsorbent in-teraction. The measured e%uent CO concentrations werefound to be very low ((0.2% (v/v)) for all these runs.The longer transient time of the SE-SMR run is mainlydue to the interaction between the adsorption and reac-tion processes, i.e. the retention of CO

2and its sub-

sequent in#uence on local reaction kinetics andequilibria. Adsorption of CO

2not only decreases the

gas-phase CO2

partial pressure, but also releases someheat, both of which are favourable for the reformingreactions (1) and (2). On the other hand, an increase in theSMR reaction rate will result in a decrease in the reactortemperature due to the endothermic nature of the re-forming reactions; see below for further discussions onthe reactor temperature pro"les in the presence and ab-sence of adsorbent. Also shown in Fig. 2(a) and (b) are thesimulated e%uent concentration pro"les. The excellentagreement between the experiments and simulations forboth the SMR and SE-SMR processes indicates that therate expressions proposed by Xu and Froment (1989a)are suitable for both the transient and steady-state peri-ods of operation, even in the presence of adsorbent. Thissuggests that the microkinetic dynamics of reaction arerelatively fast, and that the physically admixed nature ofcatalyst and adsorbent precludes any local e!ect of ad-sorption on reaction intermediates, and hence on mo-lecular kinetic steps.

Fig. 3 shows typical measured temporal pro"les ofmethane conversion and CO

2exit concentration. Corre-

sponding measured and calculated enhancement factorsare shown in Fig. 4. It can be seen that there exists a peakin conversion enhancement, which disappears on the


Fig. 4. Conversion enhancement: comparison between experimentsand simulations (reactor pressure"445.7 kPa, temperature"450$23C, feed #ow rate of CH

4"250 ml/min (STP), molar ratio

of H2O to CH

4"6, catalyst/adsorbent (or carbide) particle

size"0.25}0.5 mm, measured steady-state conversion"23.98%).

Fig. 5. Temperature distribution: Comparison between SMR and SE-SMR (4503C, 445.7 kPa, feed #ow rate of CH4"250 ml/min (STP) H

2O to

CH4"6, catalyst/adsorbent (or carbide) particle size"0.25}0.5 mm).

onset of the breakthrough of CO2. A slight negative

enhancement (of 1}2%) after the sorption-enhancementpeak is also observed in both experiments and simula-tions. This negative enhancement may be attributed tothe propagation of a weak thermal (cold) wave prior theapproach to the steady state. Thus, whilst CO

2adsorp-

tion enhances overall reaction rates through favourableequilibria shifts, as the adsorbent becomes equilibrated,a slightly lower temperature pro"le in the reactor resultsin a small (and temporary) reduction in the conversionrelative to the catalyst-only reactor. This e!ect is demon-

strated in Fig. 5, in which calculated axial temperaturepro"les for the SMR and SE-SMR processes are shownat di!erent times from the onset of the reaction step.

The complex interaction of adsorption and reactionprecludes any intuitive design of forced-periodic adsor-ptive reactors. However, some insight into favourableoperating regimes can be gained by considering the speci-"c e!ects of key operating and design parameters on theenhancement factor and the overall yield of product. Inthis work, the following parameters have been consideredin turn: space time (g-cat h/mol-CH

4), H

2O-to-CH

4feed

molar ratio, temperature, pressure and adsorbent/cata-lyst particle size. Fig. 6 shows the e!ect of space time onthese enhancement factor; other operating conditions areset as 445.7 kPa, 4553C, H

2O : CH

4"3 mol/mol. Under

these conditions, higher conversions are achieved, andapproach the equilibrium conversion of approximately20% at 4553C. The bene"cial e!ects of local CO

2adsorp-

tion are realised as the reaction equilibrium conversion isapproached. For a relatively low space time, the reactionkinetics dictate a low production rate of CO

2, and adsor-

bent equilibration (i.e. CO2

breakthrough) occurs rela-tively rapidly. Where conversions are so low that thebackward reaction rates are small, the advantages ofCO

2removal from the reaction zone will be minimised.

Thus, a low space time will generally lead to poor conver-sions, for which the advantages of in situ separation willnot be realised due to the rapid propagation velocity ofthe CO

2concentration front, and the diminished inhibi-

tion of CO2

on the net reaction rates. It is important tonote however, that in this case, reactor space time (i.e.feed #ow rate) has a signi"cant in#uence on heat transferrates and subsequently on the reactor temperature pro-"le. In speci"c, a high space time will lead to poor heat


Fig. 6. E!ect of space time on the conversion enhancement (reactorpressure"445.7 kPa, reactor temperature"455$53C, molar raito ofH

2O to CH

4"3, catalyst/adsorbent size"0.25}0.5 mm, measured

steady-state conversions"14.50%, 17.0%, 14.90%, 15.2%, 16.50%,15.75% and 18.10% for space times"17.89, 13.50, 10.7, 8.96, 7.68, 6.72and 5.37 g-cat h/mol-CH

4correspondingly).

Fig. 7. E!ect of feed molar ratio of H2O to CH

4on the conversion

enhancement (reactor pressure"445.7 kPa, reactor temperature"455$43C, space time"10.7 g-cat h/mol-CH

4, catalyst/adsorbent

size"0.25}0.5 mm, measured steady-state conversions"13.0%,14.87%, 16.65%, 20.50% and 23.70% for H

2O : CH

4"2, 3, 4, 5 and

6 correspondingly).

transfer, and thus a cooler reactor. Hence, the decrease inthe steady-state conversion with increasing space time(see "gure caption of Fig. 6) is to expected, althoughequilibrium constrains dominate for relatively high spacetimes.

The e!ect of the feed molar ratio of H2O to CH

4on

the enhancement factor is shown in Fig. 7; other operat-ing conditions are set as 445.7 kPa, 4553C, 10.7 g-cath/mol-CH

4. Methane conversion enhancement is fa-

voured by low H2O : CH

4ratios, i.e. at relatively low

partial pressures of feed H2O. For low H

2O partial

pressures, the equilibrium conversion is relatively low, i.e.the backward reactions for methane and water produc-

tion are signi"cant, and thus greater potential for separ-ation-enhanced conversion exists. Conversely, at highpartial pressures of water, a relatively high equilibriumconversion exists, and thus the signi"cance of CO

2ad-

sorption on conversion enhancement decreases. Accord-ing to reactions (1) and (2), the ideal ratio of H

2O to CH

4is between 1 and 2, depending on the desired "nal reac-tion products. For the production of H

2, this ratio

should be 2. However, a H2O : CH

4ratio greater than

this stoichiometric value is often required to avoid cata-lyst deactivation due to carbon deposition. Carbon de-position principally occurs at high temperatures throughthe decomposition of methane (CH

4"C#2H

2,

*H298

"75 kJ/mol), and to a smaller extent by theBoudouard reaction (2CO"C#CO

2, *H

298"

!173 kJ/mol) which is thermodynamically favoured atlow temperatures (see Rostrup-Nielsen, 1984). Althoughsorption-enhanced steam}methane reforming takes placeat a considerably lower temperature than that used inthe conventional reactor, some carbon formation via theBoudouard reaction cannot be avoided. Furthermore,the adsorption of CO

2is likely to shift the decomposition

of CO to further carbon formation. As a consequence,whilst it is feasible to operate the SE-SMR at lowerH

2O/CH

4ratios, this again will be dictated by the level

of catalyst coking. However, in this work, no catalystdeactivation was apparent, possibly due to relativelyshort times-on-stream of catalyst (i.e. &200 h), and thevery low CO partial pressures measured in all experi-ments.

The e!ect of operating temperature on the enhance-ment factor is shown in Fig. 8; other operating conditionsare set as 445.7 kPa, H

2O : CH

4"3 mol/mol, 10.7 g-cat

h/mol-CH4. As temperature decreases, the steady-stage

yield of CO2

decreases, i.e. CH4

conversions decreasefrom 16.3 to 12.7% as the temperature is reduced from467.5 to 4283C. As for the case of low reactor space times,the bene"ts of in situ separation are negligible, eventhough the adsorption a$nity of CO

2is relatively high.

However, at relatively high temperatures of operation,a high yield of CO

2is achieved, but further enhancement

of conversion is di$cult due to the relatively low adsorp-tion a$nity. Thus, as indicated in Fig. 8, an optimaltemperature exists for maximum enhancement. In orderto achieve an equivalent product yield at a temperaturelower than that of an adsorbent-free (conventional) reac-tor, careful selection of the adsorbent and catalyst isneeded such that the kinetics of reaction and the capacityof the adsorbent are in accord. When the adsorbent needsto be operated at a lower temperature, it may be possibleto have a two-stage reactor consisting of a high-temper-ature catalyst-only stage, and a lower temperature cata-lyst#adsorbent stage. Further discussion on analyticalcriteria for favourable catalyst and adsorbent combina-tions is given by Sheikh, Kershenbaum and Alpay(1998).


Fig. 8. E!ect of reactor temperature on the conversion enhancement(reactor pressure"445.7 kPa, molar ratio of H

2O to CH

4"3, space

time"10.7 g-cat h/mol-CH4, catalyst/adsorbent size"0.25}0.5 mm,

measured steady-state conversions"12.70%, 14.15%, 14.87% and16.25% for reactor temperatures"428, 445, 458 and 467.53C corre-spondingly).

Fig. 9. E!ect of reactor pressure on the conversion enhancement (reac-tor temperature"445.5$13C, molar ratio of H

2O to CH

4"3, space

time"10.7 g-cat h/mol-CH4, catalyst/adsorbent size"0.25}0.5 mm,

measured steady-state conversions"17.8%, 14.87%, 12.13% and10.5% for reactor pressures"308, 445.72, 583.68 and 721.5 kPa corre-spondingly).

Fig. 10. E!ect of particle sizes of catalyst and adsorbent on the transi-ent kinetics (reaction temperature"459$13C, reactor pressure"445.7 kPa, space time"10.7 g-cat h/mol-CH

4, H

2O : CH

4"3).

Fig. 9 shows the e!ect of reactor pressure on theenhancement factor; other operating conditions areset as 457.53C, 10.7 g-cat h/mol-CH

4, H

2O : CH

4"

3 mol/mol. It can be seen that the enhancement factorincreases with increasing reactor pressure. Higher pres-sures will favour greater CO

2adsorption, and will result

in a greater potential for reaction enhancement due toa reduction in the equilibrium conversion of CH

4. Thus,

although SMR reactions (1) and (2) are favoured by lowpressures, the above result suggests that the sorption-enhanced SMR reactor can be operated at higher pres-sures to achieve the same product yield. This, of course, isparticularly bene"cial where the direct high-pressure de-livery of H

2is of economic advantage.

Fig. 10 shows the e!ect of particle size of catalyst andadsorbent on the methane conversion; other operatingconditions are set as 4593C, 445.7 kPa, 10.7 g-cat h/mol-CH

4, H

2O : CH

4"3 mol/mol. It can be seen that the

methane conversion for experiments with 0.11}0.25 mmparticles can be up to 20% higher than that with0.25}0.5 mm particles in the transient period. However,this di!erence diminishes when the reactor approachesthe steady stage. Given that the e!ect of particle size isonly apparent during the transient stage, it is concludedthat the above results are likely to be associated with themass transfer resistances inside the adsorbent particles,i.e. the e!ectiveness factor of the catalyst approachesunity. As mentioned above, a reduction in the #ow rate ofreactant, or an increase in reactor length, will reduce thesigni"cance of the adsorption and desorption kinetics.

As mentioned in Section 2, the transient experiments inthis work consisted of a step change from a hydro-gen}steam feed to a methane}steam feed. Experimentshave also been performed in which H

2is replaced with

He prior to the onset of the reaction step (H2O : He

&3 mol/mol), i.e. a reactor environment free of H2

priorto the introduction of methane and steam. A comparisonof the CH

4conversion for the two cases is given in

Fig. 11; operating conditions for both cases are set as4453C, 445.7 kPa, 6.27 g-cat h/mol-CH

4, H

2O : CH

4"

3 mol/mol. It can be seen that at early times, the methaneconversion is higher under an initial He}H

2O environ-

ment. This inhibitive e!ect of H2

on the reaction kineticsis well documented; see, for example. Bodrov, Apel'baumand Tempkin (1967, 1968), Phung Quach and Rouleau(1975) and Ross and Steel (1973). For example, PhungQuach and Rouleau (1975) studied the kinetics of SMRover Ni/a-Alumina catalyst in a continuous stirred tankreactor at 350}4503C; a Langmuir}Hinshelwood}Hougen}Watson type rate equation was found to "t theirkinetic data satisfactorily, and the equation obviously


Fig. 11. E!ect of reactor initial environment on the transient kinetics(reaction temperature"4453C, reactor pressure"445.7 kPa, spacetime"6.27 g-cat h/mol-CH

4, H

2O : CH

4"3, catalyst/adsorbent size

"0.25}0.5 mm).

implied the inhibitive e!ect of the presence of H2

in thefeed. Nevertheless, the in#uence of the initial presence ofH

2on the dynamics of conversions is short-lived, i.e. less

than 1 min. This suggests that under pressure andconcentration swing type operations, and for thetemperatures considered in this work, adsorbent regen-eration can be achieved under a hydrogen-free environ-ment without adverse e!ect on the catalyst activity, andthus the dynamics of the in situ separation and reactionstep.

Having validated transient reaction kinetics in thepresence of adsorbent, future work will involve model-based optimisation methods for the selection and designof adsorptive reactor con"gurations; see, for example,Yongsunthon and Alpay (1999). In particular, considera-tion will be given to process design based on distinctobjectives, such as: (i) the lowest energy utilisation fora speci"ed product yield, and (ii) the highest concentra-tion (bulk separation) of hydrogen for a speci"ed conver-sion. Ultimately, the stability of the adsorbent undercyclic reactor conditions, i.e. under constant temporaland spatial variations of gas composition and temper-ature, will govern the commercial feasibility. Thus, fur-ther experimental work will be required, at least ata semi-technical scale, to address such issues under iden-ti"ed process con"gurations.

5. Conclusions

The steady-state kinetic model of Xu and Froment(1989a) for steam}methane reforming has been shown tobe applicable to transient reactor operation, both in thepresence and absence of adsorbent. In particular, a reac-tor model accounting for non-isothermal, non-adiabaticand non-isobaric operation, and mass transfer limited

adsorption, was found to accurately predict the elutionpro"les of the reaction species over an admixture of Nicatalyst and hydrotalcite adsorbent. Kinetic and equilib-rium parameters for adsorption and reaction were em-ployed from previous (independent) measurements; nomodel "tting parameters were needed. Comparison of theelution pro"les of methane for the adsorbent-free andadsorbent-present cases led to a measure of the reaction(conversion) enhancement. Thus, direct evidence for theincrease in net reaction rates due to in situ separationwas attained. Furthermore, investigations on the in#u-ence of some key operating parameters on the degree ofenhancement led to the following conclusions: (i) a highreactor space time is favourable for minimising the e!ectsof adsorbent intraparticle mass transfer resistances and,of course, for overcoming kinetic limitation to reaction,(ii) for a given conversion of CH

4or yield of CO

2, the

SE-SMR process enables operation with lower steam-to-methane ratios, and/or higher operating pressures,and (iii) under the conditions of SE-SMR operation,negligible deactivation of the catalyst occurs. Currently,separation-enhanced reaction studies in a semi-technicalscale reactor are under way, in which the typical operat-ing steps of a pressure swing based adsorption unit arebeing imposed. The work will enable the further testing ofkinetic and process models under cyclic operation.

Notation

a constant in the empirical correlation for the ef-fective thermal conductivity, dimensionless

A cross-sectional area of the reactor, m2

bi

Langmuir model constant for component, i,Pa~1

Cfi

gas-phase concentration of component i in thefeed, mol/m3

Ci

molar concentration of gas-phase component i,mol/m3

Cpg

gas-phase heat capacity, J/mol KC

pssolid-phase heat capacity, J/kg K

dp

particle diameter, mD

rinner diameter of the reactor, m

Dz

axial dispersion coe$cient, m2/sE(t) conversion enhancement, %H

!$iadsorption heat of component i (on adsorbentsurface), J/mol

HRi

reaction heat of reaction i, J/molk LDF mass transfer coe$cient, s~1

kg

gas-phase thermal conductivity, J/m Kki

rate constant of reaction i, i"1, 2; mol Pa0.5/kg-cat s, i"3: mol/kg-cat s Pa

ks

solid-phase thermal conductivity, J/m KK

DErgun equation coe$cient, N s/m4

Ki

equilibrium constant of reactions (1)}(3), i"I,II: Pa2, i"III: dimensionless


Kj

adsorption constant for component j (on cata-lyst surface), j"CO, H

2, CH

4: Pa~1, j"H

2O:

dimensionlessK

VErgun equation coe$cient, N s2/m5

K0z

static e!ective thermal conductivity, J/m KK

ze!ective thermal conductivity, J/m K

¸ reactor length, mm

iLangmuir model constant for component i,mol/kg

P pressure, PaPi

partial pressure of gas-phase component i, PaPL

outlet pressure, PaPr Prandtl number, dimensionlessqi

solid-phase concentration (average over an ad-sorbent particle), mol/kg

qHi

equilibrium solid-phase concentration,mol/kg

Qf

volumetric #ow rate of the feed gas (inlet),m3/s

ri

formation rate of component i, mol/kg-cat sR universal gas constant, J/mol KRe

pparticle Reynolds number, dimensionless

Rj

rate of reaction j ( j"1}3), mol/kg-cat st time, s¹ temperature, K¹

ffeed gas temperature (inlet), K

¹0

initial bed temperature, reference temperature,K

¹w

wall temperature, Ku super"cial velocity, m/su0

initial super"cial velocity, m/s;

0overall bed}wall transfer coe$cient, J/m2 K

XCH4

conversion of CH4, %

Greek letters

et

total voidage of the adsorbent bed,dimensionless

gi

catalyst e!ectiveness factor, dimensionlesskg

gas-phase viscosity, Pa sob(!$)

bulk density of the adsorbent, kg/m3

ob(#!5)

bulk density of the catalyst, kg/m3

og

gas-phase density, mol/m3

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