Adsorption-enhanced steam}methane reforming - CiteSeerX

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 CO adsorbent are presented. Attention is given to the analysis of the transient behaviour of a tubular (integral) reactor when an Ni-based catalyst is admixed with the adsorbent. Considerable enhancement of the methane conversion is experimentally demonstrated. Enhancement arises from the favourable shifts in the reaction equilibria of the reforming and water}gas shift reactions towards further CO production. As predicted, the potential for conversion enhancement is shown to increase under the conditions of a high reactor space time, 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 and a general reaction kinetic model, is shown to accurately predict the observed elution pro"les from the reactor, and thus the degree of conversion enhancement. 2000 Elsevier Science Ltd. All rights reserved. Keywords: SMR; Sorption enhancement; Transient kinetics; Ni-catalyst; Hydrotalcite; CO adsorbent; Mathematical modelling

1. Introduction The advantages of coupling reaction systems with some form of in situ separation have been widely reported in the literature. Such hybrid con"gurations may substantially improve reactant conversion or product selectivity and, for reversible reactions, establish a more favourable reaction equilibrium than that which could be achieved under conventional reactor operation. Reaction enhancement may enable a lower temperature of operation, which in turn may alleviate the problems associated with catalyst fouling, high process energy requirements and poor energy integration within the plant environment. For gas-phase catalytic reactions, the separation can be based on adsorption, selective permeation through a membrane, or through simultaneous reaction of the targeted molecule (e.g. the reaction inhibitor) with a chemical acceptor. A comprehensive review on membrane-based reaction systems has been given by Armor (1995). Advances have been made in the use of metallic membranes (often Group VIII metals which only small molecules like * Corresponding author. Tel.: 0044-171-594-5625; fax: 0044-171-5945604. E-mail address: [email protected] (E. Alpay).

hydrogen can permeate) and, more recently, polymeric, ceramic and zeolitic membranes. The membranes may act as permselective barriers, or as an integral part of the catalytically active surface. Practical issues such as membrane pore blockage, thermal and mechanical stability, and the dilution caused by the need for sweep (i.e. permeate purge) gases, have limited the usefulness of the membrane reactor systems. Nevertheless, the bene"ts of the membrane systems have been demonstrated though a wide number of experimental reaction studies, examples 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 the water}gas shift reaction (Uemiya, Sato, Ando & Kikuchi, 1991), and steam}methane reforming (Adris, Lim & Grace, 1994, 1997). In comparison to the membrane reactors, a relatively small amount of work has been carried out on systems combining reaction with adsorption or chemical acceptor-based separation processes. Even so, such processes o!er distinct advantages to the membrane-based systems in terms of the material tolerance to high temperatures and pressures, and the wide choice and availability of adsorbents 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

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

reaction conditions. Furthermore, even through use of a purge gas for regeneration, the e!ective separation of the primary adsorbate from other non- or weakly adsorbing species can be achieved. Some examples of recent works employing chemical acceptors or adsorbents for reaction enhancement are now summarised. Han and Harrison (1994) studied hydrogen production via the water}gas shift reaction using CaO as a CO acceptor in a tubular reactor. CO conversions were reported which exceeded that of the thermodynamic equilibrium conversion under the speci"ed operating conditions. Brun-Tsekhovoi, Zadorin, Katsobashvili and Kourdyumov (1986) showed a very signi"cant enhancement of CH conversion to H in a #uidised bed reactor containing Ni-based catalyst balls, and a specially treated form of dolomite as adsorbent. Typical industrial operating conditions were considered in this work, i.e. pressure levels of 10}10 kPa, and an operating temperature of 6273C. Goto, Tagawa and Oomiya (1993) studied the dehydrogenation of cyclohexane over a Pt}alumina catalyst and CaNi alloy as a hydrogen acceptor. The workers showed that at 150}1903C and ambient pressure, the overall conversion of cyclohexane to benzene could be exceeded by three-fold when compared to the catalyst-only case. Most recently, Carvill, Hufton and Sircar (1996) and Hufton, Mayorga and Sircar (1999) describe the general concept of the Sorption Enhanced Reaction Process (SERP), which utilises pressure and concentration swing adsorption principles for reaction enhancement; see also Vaporciyan and Kadlec (1989) and Alpay, Chatsiriwech and Kershenbaum (1995). CO production via the reverse water}gas shift reaction was speci"cally considered by Carvill et al. (1996), in which NaX zeolite was used as an adsorbent for water. The authors showed that at 2503C and 480 kPa, a CO conversion of 36% could be achieved; a conventional plug #ow reactor required an operating temperature of 5653C for the same conversion. Furthermore, the process generated a high purity (#99% (v/v)) CO stream as product. Hufton et al. (1999) applied the SERP concept to H production via the steam}methane reforming (SMR) reactions. In speci"c, using a hydrotalcite-based CO adsorbent, and a commercial Ni-based catalyst, the authors showed that at 4503C and 480 kPa, #95% (v/v) H could be produc ed directly from reactor. The CH to H conversion was 82%, which could only be achieved at approximately 6503C with a conventional SMR reactor. The present work also considers the sorption-enhanced steam}methane reforming (SE-SMR) process. The key reactions of the SMR process are given as: CH #H O 0 CO#3H , *H "206 kJ/mol, (1) CH #2H O 0 CO #4H , *H "164.9 kJ/mol, (2) CO#H O 0 CO #H , *H "!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 moderately exothermic and is therefore favoured by low temperatures. The reforming reactions will also be favoured at low pressures, whereas the water}gas shift reaction is largely una!ected by changes in pressure. In the presence of a selective CO adsorbent, the conversion of CH to CO though reaction (2) is favoured, as is the production of CO through CO intermediate. For a reaction which is not kinetically limited, the use of an adsorbent will thus enable a lower operating temperature for a desired conversion. However, on equilibration of the adsorbent, the separation e!ect is, of course, lost. This necessitates the periodic regeneration of the adsorbent and thus, for example, the pressure and concentration swing type operations mentioned above. In other words, sorption-enhanced reaction processes are inherently dynamic in operation. Adequate design and scale-up of such processes will thus require information on the kinetics of adsorption and desorption, as well as reaction kinetic models under transient conditions in the presence of an adsorbent. Research work on the kinetics of the SMR process dates back to the 18th century (see Sabatier, 1922; Marek & Hahn, 1932), with the "rst extensive study by Akers and Camp in 1955. A good review of the work up until 1970 is given by Van Hook (1980), which covers kinetic studies over porous nickel catalysts and nickel foil over large temperature (260}10003C) and pressure (100}5000 kPa) ranges. A considerable amount of work on the kinetic aspects of the SMR process has been carried out since 1970; see, for example, Schnell (1970), Ross and Steel (1973), Allen (1975), Phung Quach and Rouleau (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 Froment (1989a) are considered to be the most general in form, and have been extensively tested under typical industrial operating conditions (see, also, Elnashaie, Adris, Al-Ubaid & Soliman, 1990). However, like most previous work, the models are applicable to steady-state kinetics, and untested for forced-dynamic operation. It is interesting to note, however, that where some attention has been given to the transient kinetics, ideal surface conditions were maintained though vacuum operations; see, for example, Ross and Steel (1973). This, of course, limits the applicability of such models for process design applications. In this work, attention is given to the analysis of SMR reaction kinetics under transient conditions depicting SERP-type operation, both in the presence and absence of a selective absorbent for CO . Particular attention is given to the transient analysis of the Xu and Froment (1989a) kinetic model. The work then considers the in#uence of operating parameters on the degree of separation


enhancement. In doing so, mathematical models of the process are veri"ed, which can ultimately be used in the design, analysis or scale-up of pressure or concentration swing based adsorptive reactors.

2. Experimental A commercial Ni-based catalyst (United Catalyst Inc.) containing 25}35% Ni, 25}35% NiO, 5}15% MgO and 15}25% sodium silicate, was used in this work. The original catalyst (1/8 cylindrical pellets; BET area of 137.6 m/g) was crushed and sieved into two particle size groups: 0.11}0.25 and 0.25}0.5 mm. The CO adsorbent consisted of an industrially supplied potassium promoted hydrotalcite, which was previously measured for its capacity and stability under wet gas conditions (see Ding & Alpay, 2000). The adsorbent was also crushed and sieved into the particle size groups mentioned above. As a catalyst diluent, e.g. in the absence of adsorbent, silicon carbide particles were employed. High-purity methane (99.95% (v/v)) and hydrogen (99.995% (v/v)) gases were supplied from gas cylinders; steam supply to the SMR reactor was generated from distilled water. A schematic diagram of the experimental apparatus is given in Fig. 1. The reactor consisted of a stainless-steel tubular column of internal diameter 12.4 mm and length 220 mm, packed with a mixture of catalyst and adsorbent (or silicon carbide) particles. The reactor was "tted with inlet and outlet lines for introducing feed gas (CH and H O), purge as (H /He and H O) and reducing gas (H ). The inlet CH #ow rate was controlled by a Brooks 5850E mass#ow controller. An HPLC pump was used to supply water to the reactor via a vaporiser (i.e. a heated tubular column packed with silicon carbide particles); both the reactor and water vaporiser were mounted in

a convection oven. The outlet #ow rate from the reactor was monitored with an Aalborg GFM-17 mass #ow meter and a soap-bubble #owmeter. The catalyst bed temperature was detected with a type K thermocouple (positioned at the centre of the reactor along the central axis), and a back pressure regulator used to maintain a constant reactor pressure. The reactor e%uent #ow was split into sample and vent lines, each equipped with a water condenser. For the former, time delays in the sample analysis were minimised by use of 1/16 transfer lines, and a low-volume condenser unit. The sample line was connected to a Valco 16-loop valve, the operational schedule of which was computer controlled. The sample line and the sample valve were heat traced, and the temperature controlled at 1103C by a PID controller. A Shimadzu gas chromatograph (GC-14B), equipped with a TCD detector and a Porapak-Q column, was used for sample analysis. In addition, two on-line Telegan CO analysers (0}30 and 0}5%FSD) were used to moni tor the reactor e%uent CO concentration. For the reaction studies in the absence of adsorbent, approximately 7.2 g of catalyst was admixed with dense silicon carbide particles (&1 : 3 mass ratio), and packed into the reactor. For the sorption-enhanced reaction studies, approximately 7.2 g of catalyst was admixed with 14.8 g of CO adsorbent. Reactor operating conditions for both the adsorbent and adsorbent-free systems are summarised in Table 1. As mentioned above, conversion enhancement arises prior to the equilibration of the adsorbent. In this work, transient operation was imposed by means of step changes in feed concentration. Due to the negligible residence time of methane in the reactor, even in the presence of adsorbent (i.e. typically less than 0.2 s), the conversion enhancement factor at any given time (E(t)) can be quanti"ed by the normalised conversion of methane (X ) in !& the presence (ad) and absence (nad) of adsorbent, i.e. (X ) !(X ) !& 100. E(t)" !& (X ) !&

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

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(4)

Thus, values of E(t)'0 indicate conversion enhancement in the presence of adsorbent. At steady state, and for similarly packed reactors (i.e. catalyst mass loading and distribution), E(t) should approach 0. A typical experimental cycle involved the following steps: (i) heat-up of the reactor at atmospheric pressure under a hydrogen environment to the preset temperature, (ii) water supply to the system so that the molar ratio of H O to H is approximately the same as the desired H O-to-CH ratio in the reaction step, (iii) pressurisation of the system to the preset pressure, (iv) switch from H to CH to initiate the reaction step, (v) depressurisation of the unit, (vi) low-pressure purge of the unit with H and steam, and (vii) reduction of the catalyst with H at 4803C for 3 h. Steps (vi) and (vii) were carried out as precautions

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Table 1 List of experimental Run type

Space time (g-cat h/mol-CH )

H O/CH

Pressure (kPa)

Temperature (3C)

Particle size (mm)

WA/NA WA/NA WA/NA WA/NA WA

5.37}17.9 10.7 10.7 10.7 10.7

3 2}6 3 3 3

445.72 445.72 308}721.5 445.72 445.72

455 455 457.5 428}467.5 459

0.25}0.5 0.25}0.5 0.25}0.5 0.25}0.5 0.11}0.25

Initial reactor environment: H O : H (He)"H O : CH WA * with adsorbent, NA * no adsorbent.

towards any carbonaceous deposits on the catalyst, and thus to ensure a reproducible catalyst activity from one experiment to the next. Reproducibility was subsequently con"rmed through repetition of the experiments. The measurement of the reactor e%uent concentration pro"les in the reaction step (step (iv)) enabled both transient and steady-state reaction kinetics to be tested. Note that due to non-isothermal nature of the SMR process, the measured wall temperature at the middle point of the reactor is taken as the characteristic temperature in this work.

3. Mathematical modelling 3.1. Governing equations A dynamic model accounting for non-isothermal, non-adiabatic, and non-isobaric operation, was developed to describe both the SMR and SE-SMR processes. For the SE-SMR process, the reactions and adsorption were assumed to take place on the surfaces of the catalyst and adsorbent, respectively. A Langmuir model was used to describe the adsorption equilibria of CO , and a linear driving force (LDF) model for the intraparticle mass transfer of the adsorbent; further details of equilibria and kinetic measurements, and mathematical model development under non-reacting conditions, are given by Ding and Alpay (2000). As mentioned above, the general reaction kinetic model proposed by Xu and Froment (1989a) was considered in this work. Other principal model assumptions can be summarised as: axially dispersed plug #ow, perfect gas behaviour, no radial concentration or temperature gradients, and a catalyst/adsorbent packing of uniform voidage and particle size. Based on the above assumptions, component mass balances for the packed-bed reactor can be written as

* * *C *(uC ) G ! G #g o (e C #o q )" D r, @ G G @ G *t R G *z X *z *z (5)

where i denotes CH , H O, H , CO CO, and He. the semi-empirical correlation proposed by Edwards and Richardson (1968) was used to estimate the axial dispersion coe$cient D . The catalyst e!ectiveness factor, g , X G was set as unity for the crushed catalyst used in this work; see Section 4 for details. The reaction kinetic model of Xu and Froment (1989a) can be summarised as

k P P R " P P ! & !' P !& & K & '

(DEN),

(6a)

k P P R " P P ! & !- '' P !& & K & ''

(DEN),

(6b)

P P k R " P P ! & !- ''' P !- & K ''' &

(DEN),

(6c)

DEN"1#K P #K P #K P !- !& & !& !& #K

P /P , & - & - &

(6d)

where R ( j"I!III) denotes the reaction rate of the H SMR reactions (1) and (2) and the water}gas shift reaction (3). The formation rate of component i, r , was G then calculated by using Eqs. (1)}(3) and (6a)}(6d); for example, r "!(R #R ), r "3(R #R )# !& ' '' & ' ''' (R #R ). The rate constants k (i"1}3) and the ad'' ''' G sorption constants K ( j"I}III) are function of temperH ature, details of which are given by Xu and Froment (1989a). Pressure distribution in the packed-bed reactor was described by the Ergun equation (Ergun, 1952) *P "!K u!K u, " 4 *z

(7)

where K and K are parameters corresponding to the " 4 viscous and kinetic pressure loss terms, respectively. Semi-empirical relationships for K and K have been " 4 derived by Ergun (1952) (see also MacDonald, El-Sayed, Mow & Dullien, 1979). For compressible #ow, the energy balance for the reactor is given by (see, for example, Bird, Stewart &


Lightfoot, 1960)

3.2. Boundary and initial conditions

* (o C e ¹#o C ¹) @ NQ *t E NE R

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

*P * *¹ * "e # K ! (o uC ¹) R *t *z X *z *z E NE *q G # (R o gH ) !o H G *t G @ G 0G @ G G 4; # (¹ !¹). Dr U

q "0, G *P "!K u !K u , " 4 *z (8)

The bed e!ective conductivity, K , can be expressed as X (see Yagi & Kunii, 1957; Kunii & Smith, 1960; De Wasch & Froment, 1972; Li & Finlayson, 1977; Wakao & Kaguei, 1982) K K X " X #a(Pr) (Re ), (9) N k k E E where K is the static e!ective conductivity accounting X for the e!ects of conduction and radiation, see Kunii and Smith (1960) and Ding and Alpay (2000). For a bed packed with spherical particles, the wall}bed heat transfer coe$cient, ; , in Eq. (8) is given by Li and Finlayson (1977) as

; D 6d P "2.03Re exp ! N N D k E P (Re "20}7600, d /D "0.05}0.3). (10) N N P Eq. (10) is not applicable to very low feed #ow rates, as ; should approach a "xed value where Re P0. How N ever, De Wasch and Froment (1972) give the following correlation for ; at very low Re : N K (11) ; " N "6.15 X 0C D P which can be linearly combined with Eq. (10) to approximate ; over the entire range of Re relevant to this N work. The Langmuir model for CO adsorption can be written as m b P qH " !- !- !- . (12) !1#b P !- !Other reaction components were considered to be nonadsorbing. The LDF model is given by *q G "k (qH!q ), G G G *t

3933

(13)

where k is the e!ective mass transfer coe$cient. Note that *q /*t"0 for non-adsorbing species in the SE-SMR proG cess, and *q /*t"0 for all species in the SMR process. As G mentioned above, Langmuir and LDF parameters have been previously measured for the CO }H O-hydro talcite system; see Ding and Alpay (2000).

*P/*t"0,

(14a) (14b) (14c)

¹"¹ , (14d) C "P /(R¹ ) (14e) G G At the onset of the reaction step (i.e. the point at which CH is supplied to the reactor), the following boundary conditions were used in the simulations: (i) Reactor entrance (z"0) !D (*C /*z)"u(C !C ), X G DG G !K (*¹/*z)"uo C (¹ !¹), X E NE D u"Q /A. D (ii) Reactor outlet (z"¸) *C /*z"0, G *¹/*z,0,

(15a) (15b) (15c)

(15d) (15e)

P"P . (15f ) * Q in Eq. (15c) is the feed volumetric #ow rate measured D under the local conditions.

Table 2 Parameters (constants) used in the simulations Parameters (constants)

Value

Unit

b !d (average) N D P C NE C NQ e k E k Q H !- ¸ (reactor length) m !b e @ e R U c j Q k E o @ o @

SE-SMR: 2.36;10\; SMR:0 1.8;10\, 3.57;10\ 1.27;10\ 42 850 0.35 0.09 0.3 SE-SMR: !17000, SMR: 0 0.22 SE-SMR: 0.65, SMR: 0 0.95 0.47 0.64 0.2 0.667 1.0 2.87;10\ SE-SMR: 609.3, SMR:0 139.0

Pa\ m m J/mol K J/kg K * J/m K J/m K J/mol m mol/kg * * * * * * Pa s kg/m kg/m

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Eqs. (5)}(15) were simultaneously solved in the gPROMS modelling environment (Process Systems Enterprises Ltd.). The spatial discretisation method of orthogonal collocation on "nite elements was employed. Second-order collocation on 100 elements was found to give a converged solution in which component balance errors (associated with the numerical integration) did not exceed 1%. A summary of parameters (constants) used for the simulations is given in Table 2.

4. Results and discussion Typical reactor e%uent concentration pro"les measured in the SMR and SE-SMR experiments under similar operating conditions (4503C, 445.7 kPa, H O/ CH "6) are shown in Fig. 2(a) and (b), respectively. It can be seen that in the presence of CO adsorbent, the transient period (&5 min) is much longer than that of

Fig. 2. E%uent concentration (mole fraction) pro"les on a water-free basis: comparison between experiments and simulations (reactor pressure"445.7 kPa, temperature"450$23C, feed #ow rate of CH "250 ml/min (STP), molar ratio of H O to CH "6, cata lyst/adsorbent (or carbide) particle size"0.25}0.5 mm).

the SMR run (&1.5 min). However, since the experimental conditions are nearly identical, the steady-state concentrations of CH , H , and CO are approximately equal, indicating no strong direct catalyst/adsorbent interaction. The measured e%uent CO concentrations were found to be very low ((0.2% (v/v)) for all these runs. The longer transient time of the SE-SMR run is mainly due to the interaction between the adsorption and reaction processes, i.e. the retention of CO and its sub sequent in#uence on local reaction kinetics and equilibria. Adsorption of CO not only decreases the gas-phase CO partial pressure, but also releases some heat, both of which are favourable for the reforming reactions (1) and (2). On the other hand, an increase in the SMR reaction rate will result in a decrease in the reactor temperature due to the endothermic nature of the reforming reactions; see below for further discussions on the reactor temperature pro"les in the presence and absence of adsorbent. Also shown in Fig. 2(a) and (b) are the simulated e%uent concentration pro"les. The excellent agreement between the experiments and simulations for both the SMR and SE-SMR processes indicates that the rate expressions proposed by Xu and Froment (1989a) are suitable for both the transient and steady-state periods of operation, even in the presence of adsorbent. This suggests that the microkinetic dynamics of reaction are relatively fast, and that the physically admixed nature of catalyst and adsorbent precludes any local e!ect of adsorption on reaction intermediates, and hence on molecular kinetic steps. Fig. 3 shows typical measured temporal pro"les of methane conversion and CO exit concentration. Corre sponding measured and calculated enhancement factors are shown in Fig. 4. It can be seen that there exists a peak in conversion enhancement, which disappears on the

Fig. 3. Comparison of methane conversion pro"les and CO break through curve (reactor pressure"445.7 kPa, temperature" 450$23C, feed #ow rate of CH "250 ml/min (STP), molar ratio of H O to CH "6, catalyst/adsorbent (or carbide) particle size"0.25}0.5 mm).


Fig. 4. Conversion enhancement: comparison between experiments and simulations (reactor pressure"445.7 kPa, temperature" 450$23C, feed #ow rate of CH "250 ml/min (STP), molar ratio of H O to CH "6, catalyst/adsorbent (or carbide) particle size"0.25}0.5 mm, measured steady-state conversion"23.98%).

onset of the breakthrough of CO . A slight negative enhancement (of 1}2%) after the sorption-enhancement peak is also observed in both experiments and simulations. This negative enhancement may be attributed to the propagation of a weak thermal (cold) wave prior the approach to the steady state. Thus, whilst CO adsorp tion enhances overall reaction rates through favourable equilibria shifts, as the adsorbent becomes equilibrated, a slightly lower temperature pro"le in the reactor results in a small (and temporary) reduction in the conversion relative to the catalyst-only reactor. This e!ect is demon-

3935

strated in Fig. 5, in which calculated axial temperature pro"les for the SMR and SE-SMR processes are shown at di!erent times from the onset of the reaction step. The complex interaction of adsorption and reaction precludes any intuitive design of forced-periodic adsorptive reactors. However, some insight into favourable operating regimes can be gained by considering the speci"c e!ects of key operating and design parameters on the enhancement factor and the overall yield of product. In this work, the following parameters have been considered in turn: space time (g-cat h/mol-CH ), H O-to-CH feed molar ratio, temperature, pressure and adsorbent/catalyst particle size. Fig. 6 shows the e!ect of space time on these enhancement factor; other operating conditions are set as 445.7 kPa, 4553C, H O : CH "3 mol/mol. Under these conditions, higher conversions are achieved, and approach the equilibrium conversion of approximately 20% at 4553C. The bene"cial e!ects of local CO adsorp tion are realised as the reaction equilibrium conversion is approached. For a relatively low space time, the reaction kinetics dictate a low production rate of CO , and adsor bent equilibration (i.e. CO breakthrough) occurs rela tively rapidly. Where conversions are so low that the backward reaction rates are small, the advantages of CO removal from the reaction zone will be minimised. Thus, a low space time will generally lead to poor conversions, for which the advantages of in situ separation will not be realised due to the rapid propagation velocity of the CO concentration front, and the diminished inhibi tion of CO on the net reaction rates. It is important to note however, that in this case, reactor space time (i.e. feed #ow rate) has a signi"cant in#uence on heat transfer rates and subsequently on the reactor temperature pro"le. In speci"c, a high space time will lead to poor heat

Fig. 5. Temperature distribution: Comparison between SMR and SE-SMR (4503C, 445.7 kPa, feed #ow rate of CH "250 ml/min (STP) H O to CH "6, catalyst/adsorbent (or carbide) particle size"0.25}0.5 mm).

3936


Fig. 6. E!ect of space time on the conversion enhancement (reactor pressure"445.7 kPa, reactor temperature"455$53C, molar raito of H O to CH "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, 6.72 and 5.37 g-cat h/mol-CH correspondingly).

Fig. 7. E!ect of feed molar ratio of H O to CH on the conversion enhancement (reactor pressure"445.7 kPa, reactor temperature" 455$43C, space time"10.7 g-cat h/mol-CH , 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 O : CH "2, 3, 4, 5 and 6 correspondingly).

transfer, and thus a cooler reactor. Hence, the decrease in the steady-state conversion with increasing space time (see "gure caption of Fig. 6) is to expected, although equilibrium constrains dominate for relatively high space times. The e!ect of the feed molar ratio of H O to CH on the enhancement factor is shown in Fig. 7; other operating conditions are set as 445.7 kPa, 4553C, 10.7 g-cat h/mol-CH . Methane conversion enhancement is fa voured by low H O : CH ratios, i.e. at relatively low partial pressures of feed H O. For low H O 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 separation-enhanced conversion exists. Conversely, at high partial pressures of water, a relatively high equilibrium conversion exists, and thus the signi"cance of CO ad sorption on conversion enhancement decreases. According to reactions (1) and (2), the ideal ratio of H O to CH is between 1 and 2, depending on the desired "nal reaction products. For the production of H , this ratio should be 2. However, a H O : CH ratio greater than this stoichiometric value is often required to avoid catalyst deactivation due to carbon deposition. Carbon deposition principally occurs at high temperatures through the decomposition of methane (CH "C#2H , *H "75 kJ/mol), and to a smaller extent by the Boudouard reaction (2CO"C#CO , *H " !173 kJ/mol) which is thermodynamically favoured at low temperatures (see Rostrup-Nielsen, 1984). Although sorption-enhanced steam}methane reforming takes place at a considerably lower temperature than that used in the conventional reactor, some carbon formation via the Boudouard reaction cannot be avoided. Furthermore, the adsorption of CO is likely to shift the decomposition of CO to further carbon formation. As a consequence, whilst it is feasible to operate the SE-SMR at lower H O/CH ratios, this again will be dictated by the level of catalyst coking. However, in this work, no catalyst deactivation was apparent, possibly due to relatively short times-on-stream of catalyst (i.e. &200 h), and the very low CO partial pressures measured in all experiments. The e!ect of operating temperature on the enhancement factor is shown in Fig. 8; other operating conditions are set as 445.7 kPa, H O : CH "3 mol/mol, 10.7 g-cat h/mol-CH . As temperature decreases, the steady-stage yield of CO decreases, i.e. CH conversions decrease from 16.3 to 12.7% as the temperature is reduced from 467.5 to 4283C. As for the case of low reactor space times, the bene"ts of in situ separation are negligible, even though the adsorption a$nity of CO is relatively high. However, at relatively high temperatures of operation, a high yield of CO is achieved, but further enhancement of conversion is di$cult due to the relatively low adsorption a$nity. Thus, as indicated in Fig. 8, an optimal temperature exists for maximum enhancement. In order to achieve an equivalent product yield at a temperature lower than that of an adsorbent-free (conventional) reactor, careful selection of the adsorbent and catalyst is needed such that the kinetics of reaction and the capacity of the adsorbent are in accord. When the adsorbent needs to be operated at a lower temperature, it may be possible to have a two-stage reactor consisting of a high-temperature catalyst-only stage, and a lower temperature catalyst#adsorbent stage. Further discussion on analytical criteria for favourable catalyst and adsorbent combinations 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 O to CH "3, space time"10.7 g-cat h/mol-CH , catalyst/adsorbent size"0.25}0.5 mm, measured steady-state conversions"12.70%, 14.15%, 14.87% and 16.25% for reactor temperatures"428, 445, 458 and 467.53C correspondingly).

Fig. 9. E!ect of reactor pressure on the conversion enhancement (reactor temperature"445.5$13C, molar ratio of H O to CH "3, space time"10.7 g-cat h/mol-CH , catalyst/adsorbent size"0.25}0.5 mm, measured steady-state conversions"17.8%, 14.87%, 12.13% and 10.5% for reactor pressures"308, 445.72, 583.68 and 721.5 kPa correspondingly).

Fig. 9 shows the e!ect of reactor pressure on the enhancement factor; other operating conditions are set as 457.53C, 10.7 g-cat h/mol-CH , H O : CH " 3 mol/mol. It can be seen that the enhancement factor increases with increasing reactor pressure. Higher pressures will favour greater CO adsorption, and will result in a greater potential for reaction enhancement due to a reduction in the equilibrium conversion of CH . Thus, although SMR reactions (1) and (2) are favoured by low pressures, the above result suggests that the sorptionenhanced SMR reactor can be operated at higher pressures to achieve the same product yield. This, of course, is particularly bene"cial where the direct high-pressure delivery of H is of economic advantage.

3937

Fig. 10. E!ect of particle sizes of catalyst and adsorbent on the transient kinetics (reaction temperature"459$13C, reactor pressure" 445.7 kPa, space time"10.7 g-cat h/mol-CH , H O : CH "3).

Fig. 10 shows the e!ect of particle size of catalyst and adsorbent on the methane conversion; other operating conditions are set as 4593C, 445.7 kPa, 10.7 g-cat h/molCH , H O : CH "3 mol/mol. It can be seen that the methane conversion for experiments with 0.11}0.25 mm particles can be up to 20% higher than that with 0.25}0.5 mm particles in the transient period. However, this di!erence diminishes when the reactor approaches the steady stage. Given that the e!ect of particle size is only apparent during the transient stage, it is concluded that the above results are likely to be associated with the mass transfer resistances inside the adsorbent particles, i.e. the e!ectiveness factor of the catalyst approaches unity. As mentioned above, a reduction in the #ow rate of reactant, or an increase in reactor length, will reduce the signi"cance of the adsorption and desorption kinetics. As mentioned in Section 2, the transient experiments in this work consisted of a step change from a hydrogen}steam feed to a methane}steam feed. Experiments have also been performed in which H is replaced with He prior to the onset of the reaction step (H O : He &3 mol/mol), i.e. a reactor environment free of H prior to the introduction of methane and steam. A comparison of the CH conversion for the two cases is given in Fig. 11; operating conditions for both cases are set as 4453C, 445.7 kPa, 6.27 g-cat h/mol-CH , H O : CH " 3 mol/mol. It can be seen that at early times, the methane conversion is higher under an initial He}H O environ ment. This inhibitive e!ect of H on the reaction kinetics is well documented; see, for example. Bodrov, Apel'baum and Tempkin (1967, 1968), Phung Quach and Rouleau (1975) and Ross and Steel (1973). For example, Phung Quach and Rouleau (1975) studied the kinetics of SMR over Ni/a-Alumina catalyst in a continuous stirred tank reactor at 350}4503C; a Langmuir}Hinshelwood} Hougen}Watson type rate equation was found to "t their kinetic data satisfactorily, and the equation obviously

3938


Fig. 11. E!ect of reactor initial environment on the transient kinetics (reaction temperature"4453C, reactor pressure"445.7 kPa, space time"6.27 g-cat h/mol-CH , H O : CH "3, catalyst/adsorbent size "0.25}0.5 mm).

implied the inhibitive e!ect of the presence of H in the feed. Nevertheless, the in#uence of the initial presence of H on the dynamics of conversions is short-lived, i.e. less than 1 min. This suggests that under pressure and concentration swing type operations, and for the temperatures considered in this work, adsorbent regeneration can be achieved under a hydrogen-free environment without adverse e!ect on the catalyst activity, and thus the dynamics of the in situ separation and reaction step. Having validated transient reaction kinetics in the presence of adsorbent, future work will involve modelbased optimisation methods for the selection and design of adsorptive reactor con"gurations; see, for example, Yongsunthon and Alpay (1999). In particular, consideration will be given to process design based on distinct objectives, such as: (i) the lowest energy utilisation for a speci"ed product yield, and (ii) the highest concentration (bulk separation) of hydrogen for a speci"ed conversion. Ultimately, the stability of the adsorbent under cyclic reactor conditions, i.e. under constant temporal and spatial variations of gas composition and temperature, will govern the commercial feasibility. Thus, further experimental work will be required, at least at a semi-technical scale, to address such issues under identi"ed process con"gurations.

5. Conclusions The steady-state kinetic model of Xu and Froment (1989a) for steam}methane reforming has been shown to be applicable to transient reactor operation, both in the presence and absence of adsorbent. In particular, a reactor model accounting for non-isothermal, non-adiabatic and non-isobaric operation, and mass transfer limited

adsorption, was found to accurately predict the elution pro"les of the reaction species over an admixture of Ni catalyst and hydrotalcite adsorbent. Kinetic and equilibrium parameters for adsorption and reaction were employed from previous (independent) measurements; no model "tting parameters were needed. Comparison of the elution pro"les of methane for the adsorbent-free and adsorbent-present cases led to a measure of the reaction (conversion) enhancement. Thus, direct evidence for the increase in net reaction rates due to in situ separation was attained. Furthermore, investigations on the in#uence of some key operating parameters on the degree of enhancement led to the following conclusions: (i) a high reactor space time is favourable for minimising the e!ects of adsorbent intraparticle mass transfer resistances and, of course, for overcoming kinetic limitation to reaction, (ii) for a given conversion of CH or yield of CO , the SE-SMR process enables operation with lower steamto-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-technical scale reactor are under way, in which the typical operating steps of a pressure swing based adsorption unit are being imposed. The work will enable the further testing of kinetic and process models under cyclic operation.

Notation a A b G C DG C G C NE C NQ d N D P D X E(t) H G H 0G k k E k G k Q K " K G

constant in the empirical correlation for the effective thermal conductivity, dimensionless cross-sectional area of the reactor, m Langmuir model constant for component, i, Pa\ gas-phase concentration of component i in the feed, mol/m molar concentration of gas-phase component i, mol/m gas-phase heat capacity, J/mol K solid-phase heat capacity, J/kg K particle diameter, m inner diameter of the reactor, m axial dispersion coe$cient, m/s conversion enhancement, % adsorption heat of component i (on adsorbent surface), J/mol reaction heat of reaction i, J/mol LDF mass transfer coe$cient, s\ gas-phase thermal conductivity, J/m K rate constant of reaction i, i"1, 2; mol Pa / kg-cat s, i"3: mol/kg-cat s Pa solid-phase thermal conductivity, J/m K Ergun equation coe$cient, N s/m equilibrium constant of reactions (1)}(3), i"I, II: Pa, i"III: dimensionless


K H K 4 K X K X ¸ m G P P G P * Pr q G qH G Q D r G R Re N R H t ¹ ¹ D ¹ ¹ U u u ; X !&

adsorption constant for component j (on catalyst surface), j"CO, H , CH : Pa\, j"H O: dimensionless Ergun equation coe$cient, N s/m static e!ective thermal conductivity, J/m K e!ective thermal conductivity, J/m K reactor length, m Langmuir model constant for component i, mol/kg pressure, Pa partial pressure of gas-phase component i, Pa outlet pressure, Pa Prandtl number, dimensionless solid-phase concentration (average over an adsorbent particle), mol/kg equilibrium solid-phase concentration, mol/kg volumetric #ow rate of the feed gas (inlet), m/s formation rate of component i, mol/kg-cat s universal gas constant, J/mol K particle Reynolds number, dimensionless rate of reaction j ( j"1}3), mol/kg-cat s time, s temperature, K feed gas temperature (inlet), K initial bed temperature, reference temperature, K wall temperature, K super"cial velocity, m/s initial super"cial velocity, m/s overall bed}wall transfer coe$cient, J/m K conversion of CH , %

Greek letters e R g G k E o @ o @ o E

total voidage of the adsorbent bed, dimensionless catalyst e!ectiveness factor, dimensionless gas-phase viscosity, Pa s bulk density of the adsorbent, kg/m bulk density of the catalyst, kg/m gas-phase density, mol/m

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