Kinetic simulation of a compact reactor system for hydrogen production by steam reforming of higher hydrocarbons.

Rakib, M.A. ; Grace, J.R. ; Elnashaie, S.S.E.H. 等

INTRODUCTION

Hydrogen is frequently discussed as a future energy carrier. Key applications are as a carbon-free fuel, and as a fuel for hydrogen fuel cells for automotive and other applications. Hydrogen has been used effectively in a number of internal combustion engine vehicles mixed with natural gas (hythane) (Johnston et al., 2005). Hydrogen can also be combined electrochemically with oxygen without combustion to produce direct-current electricity in fuel cells, and is used in a growing number of fuel cell vehicles.

As a feedstock in chemical processes, the demand for hydrogen is increasing, both for the petrochemical industries and for petroleum refining processes. Synthesis gas, a mixture of hydrogen, carbon monoxide and carbon dioxide in various proportions, is used by Fisher-Tropsch catalytic technology to produce a wide range of chemicals from methanol to diesel. Steam-reforming-based hydrogen plants are installed in refineries to meet the fast-rising demand-supply gap in their daily operations (Rostrup-Nielsen and Rostrup-Nielsen, 2002).

Hydrogen is used in the metallurgical industry to create a reducing atmosphere in metal extraction (Eliezer et al., 2000), and in annealing of steel. It is also used in the electronics industry to manufacture semiconductor devices, and in the food industries for hydrogenation of fats and oils (Ramachandran and Menon, 1998; Eliezer et al., 2000).

Thus the demand of hydrogen is projected to increase, and this has motivated research into improving methods of hydrogen production, separation, purification, storage, and transportation. Many of the hydrogen uses put special demand on the purity of the hydrogen from these reformers.

Steam reforming remains the leading pathway of hydrogen from hydrocarbon sources, especially natural gas (Rostrup-Nielsen and Rostrup-Nielsen, 2002; Rostrup-Nielsen et al., 2002). The greatest advantage of the steam-reforming pathway is that hydrogen is extracted not only from a hydrocarbon, but from steam as well, thereby enhancing [H.sub.2] production, giving the maximum [H.sub.2] production per mole of hydrocarbon. The presence of excess steam in the reaction mixture suppresses coking reactions, the extent of which depends largely on the reaction temperature and the type of hydrocarbon.

Currently, methane is the major feedstock for production of synthesis gas, as well as pure hydrogen. However, compared to liquid hydrocarbons, the volumetric hydrogen density remains low, even after natural gas is compressed to liquid for transportation, although the H/C ratio of methane is high (Kaila and Krause, 2004). Therefore, an easily deliverable and safely storable hydrogen source, such as gasoline and diesel, is preferred for mobile applications (Melo and Morlane, 2005). On-board hydrogen generation systems prefer liquid hydrocarbon feedstocks, such as gasoline, kerosene, and diesel oil, which have a higher energy density and a wider distribution network, compared to methanol (Zhu et al., 2005). In addition, many refineries benefit from flexibility in feedstocks, taking advantage of the surplus of various hydrocarbons in the refinery.

Traditional steam reforming plants have a fixed bed steam reformer. For naphtha steam reforming, the desulfurized hydrocarbon is fed to a pre-reformer, which is operated adiabatically, where the higher hydrocarbons are directly converted to methane, giving a methane-rich gas feed for the reformer (Christensen, 1996). In the primary reformer there are hundreds of externally fired catalyst-packed tubes, in which steam reforming of methane takes place. The fixed bed reformer is followed by the shift high temperature and low temperature (HTS and LTS) reactors section for further reaction of carbon monoxide with steam to enhance hydrogen yield. The gas purification system consists of a C[O.sub.2] removal unit, a Methanator, and finally a Pressure Swing Adsorption unit to produce pure hydrogen.

Steam reforming is limited by diffusional resistances inside the catalyst pellet, resulting in very low effectiveness factors, of the order of [10.sup.-2] to [10.sup.-3] (Soliman et al., 1988; Elnashaie and Adris, 1989; Elnashaie and Elshishini, 1993). In addition, with external firing needed for the highly endothermic reactions, formation of hot spots can lead to problems related to temperature control. Pressure drop limitations block attempts to improve the effectiveness factor by using smaller diameter particles. Adris (1989) and Elnashaie and Adris (1989) proposed a novel Fluidized Bed Steam Reformer, with the heat supplied through immersed heat transfer tubes. Heat transfer limitations of the fixed bed reactor are also minimized in the fluidized bed because of better mixing characteristics.

The other major limitation for the steam reforming reactions is thermodynamic equilibrium. Removal of the main products can drive the reaction towards completion, following Le Chatelier's principle. Permselective membranes of Pd or Pd-based alloys can remove [H.sub.2], thus serving dual objectives: enhancing the hydrocarbon conversion by favourably shifting the equilibrium conversion, and producing a stream of pure [H.sub.2] as permeate (Adris et al., 1991, 1997; Grace et al., 2005).

This study deals with modelling a fluidized bed membrane reactor (FBMR) for steam reforming of higher hydrocarbons, carried out to size an experimental reformer set-up. Typically, naphtha consists predominantly of saturated hydrocarbons (>90 percent by volume), the balance being composed mainly of aromatics, and some unsaturated hydrocarbons (Chen, 2004). n-Heptane is treated in the current simulations as a model compound for steam reforming of naphtha, as also earlier assumed by Chen (Chen et al., 2003; Chen, 2004), Tottrup (1982), Christensen (1996), and Darwish et al. (2004). Others (Kaila and Krause, 2004; Puolakka and Krause, 2004; Zhu et al., 2005) have assumed it to be a model component for gasoline. A hydrocarbon feed mixture composed of n-heptane and n-hexane (in a weight ratio of [C.sub.7]/[C.sub.6] = 2) was taken as a synthetic feed for steam reforming of naphtha by Melo et al. (2005).

IRREVERSIBILITY OF STEAM REFORMING OF HIGHER HYDROCARBONS

Equilibrium calculations, in Figure 1 show that the steam reforming of heptane is practically irreversible, indicated by its complete consumption at the representative conditions of reaction. The temperature was varied from 400 to 800[degrees]C at four different pressures from 1 to 20 bar, and equilibrium compositions were predicted using a Gibbs Reactor in HYSYS simulation software. The feed composition, consisting of n-heptane, steam, and H2, used for the equilibrium predictions are listed in Table 1, and is the same as employed for the base simulation conditions in the kinetic model.

For higher hydrocarbons, the reaction can be written (Rostrup-Nielsen and Rostrup-Nielsen, 2002; Chen et al., 2003; Chen, 2004; Chin et al., 2006) as:

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.] (1)

Once [H.sub.2] and CO are available by steam reforming of higher hydrocarbons, a reverse steam reforming reaction (reverse of Equation 2) produces C[H.sub.4] (methanation reaction), and thereafter the process proceeds as simple steam reforming of methane (Chen et al., 2003; Chen, 2004; Chin et al., 2006):

C[H.sub.4] + [H.sub.2]O [left and right arrow] CO + 3[H.sub.2] [r.sub.2], [DELTA][H.sup.o.sub.298] = 206.1 kJ/mol (2)

CO + [H.sub.2]O [left and right arrow] C[O.sub.2] + [H.sub.2] [r.sub.3], [DELTA][H.sup.o.sub.298] = -41.1 kJ/mol (3)

C[H.sub.4] + 2[H.sub.2]O [left and right arrow] C[O.sub.2] + 4[H.sub.2] [r.sub.4], [DELTA][H.sup.o.sub.298] = -165 kJ/mol (4)

Although methane is not present in the feed, it immediately starts to appear in the system due to the methanation reactions (reverse of reactions (2) and (4)), once [H.sub.2], CO, and C[O.sub.2] appear in the system by reactions (1) and (3). The methane yield decreases with increasing temperature due to the endothermicity of the steam reforming reaction of methane. As a result, [H.sub.2] yield continues to increase. If this [H.sub.2] is selectively removed from the system, C[H.sub.4] yield will decrease further, due to forward equilibrium shift of reaction (2).

[FIGURE 1 OMITTED]

The irreversibility for steam reforming applies to all higher hydrocarbons with different degrees of reactivity. The higher hydrocarbons are generally more reactive than methane, with aromatics showing the lowest reactivity, approaching that of methane (Rostrup-Nielsen, 1984). Industrially, with proper desulfurization, it has been possible to convert light gas oils and diesel fuel into syngas with no trace of higher hydrocarbons in the product gas (Rostrup-Nielsen and Rostrup-Nielsen, 2002). Pilot scale experiments on adiabatic pre-reforming of natural gas, which also contained higher hydrocarbons in the range [C.sub.2]-[C.sub.7], showed that the concentration of all higher hydrocarbons decreased continuously through the bed and that no intermediate compounds were observed (Christensen et al., 1996).

KINETIC MODELLING OF A FLUIDIZED BED MEMBRANE REACTOR

A two-phase model of an FBMR was prepared to assist with the sizing of an experimental reactor. The bubbling bed regime of operation has been adopted for the simulations for this paper since the experimental reformer will be focused mainly on this regime. Pd-based membrane panels supplied by Membrane Reactor Technologies Limited, a Vancouver based company, will be used in the reactor immersed in the fluidized bed of the catalyst. A distributor design has been adopted in the experimental design which minimizes any effect of jetting. The geometry and reaction base conditions are tabulated in Table 1. Simulations were performed for a 1 m membrane length. Figure 2 shows a schematic of the model developed.

[FIGURE 2 OMITTED]

Double-sided membrane panels are inserted through vertical slits along the height of the reformer shell. The membrane panels pass through the centreline of the reformer, dividing the cross-section into two communicating sections. Thus, the membranes will be in contact with the bubble and dense phases nearly proportionally to the fractions they occupy in the fluidized bed.

[FIGURE 3 OMITTED]

MODEL ASSUMPTIONS

1. Steady-state reactor conditions.

2. Isothermal bed. The experimental reactor set-up will be externally heated to overcome the high endothermicity of the reaction in addition to allowing isothermal operation.

3. Only the lower dense catalyst bed is simulated; the lean freeboard regime is not treated in this paper.

4. The lower dense catalyst bed is treated as two parallel phases made up of a dense phase and a bubble phase.

5. Plug flow behaviour is assumed for the dense phase as well as the bubble phase. The high aspect ratio of the FBMR simulated justifies this assumption.

6. Catalyst diffusion resistance is taken to be negligible. Very fine catalyst particles with a mean particle size of 100 wm will be used for the experiments.

7. Catalyst deactivation is neglected in this paper.

8. Any jetting above the distributor is neglected.

Model Equations for Reactor Side

Mole balance for ith species in the bubble phase

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.] (5)

[FIGURE 4 OMITTED]

Mole balance for ith species in the dense phase

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.] (6)

Subscripts b and d refer to the bubble and dense phases, respectively; [y.sub.ij] is the stoichiometric coefficient of component i in the jth reaction (negative for species consumed and positive for products); [Q.sub.ib] and [Q.sub.id] are the permeation rates per unit length from the reactor side to permeation side for the bubble phase and the dense phase, respectively, for species i.

MODEL EQUATIONS FOR SEPARATION SIDE

The differential mole balance equation for the permeate hydrogen is written as:

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.] (7)

The hydrogen permeation rate from each phase is calculated from Sievert's law:

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.] (8)

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.] (9)

The membranes are assumed to be impermeable to all other species; where [P.sub.M0] is the pre-exponential factor for permeation = 0.00207 mol/(m min [atm.sup.0.5]), and E[H.sub.2] is the activation energy for permeation = 9180 J/mol.

INTERPHASE MASS EXCHANGE COEFFICIENT

The interphase mass exchange coefficient is calculated based on the correlation by Sit and Grace (1981). For the ith component:

[k.sub.ig] = [U.sub.mf]/3 + [4[D.sub.ie][[epsilon].sub.mf][U.sub.b]/[pi][d.sub.b]] (10)

where [D.sub.ie] is the effective diffusivity of component i in the gas mixture and is calculated based on the average composition of the bubble and the dense phases, using the correlation of Wilke and Lee (1955):

1 - [x.sub.i]/[D.sub.ie = [n.summation over (i=1)]([x.sub.i]/[D.sub.ij]), i [not equal to] j (11)

where [D.sub.ij] is the binary diffusivity of components i and j.

[FIGURE 5 OMITTED]

[FIGURE 6 OMITTED]

RESULTS AND DISCUSSION

Figure 3 shows the predicted species concentrations for the two phases for operation at 650[degrees]C (close to the current maximum temperature of palladium membrane) and 10 bar absolute pressure. As can be seen, although the reaction occurs predominantly in the dense phase, and there is almost no reaction in the bubble phase, the species concentrations in the two phases are almost identical. This is attributable to the relatively fast mass transfer between the two phases at the temperature of the reformer.

Figure 4 shows that as hydrogen is withdrawn from the reaction mixture, the methane yield decreases, enhancing the hydrogen production. Thus, while on the one hand pure hydrogen is produced due to membrane separation, on the other hand, overall hydrogen yield increases, which is a measure of the reactor performance in this case. Retentate hydrogen yield, which represents the hydrogen left inside the reactor, goes on decreasing as more and more hydrogen permeates through the membranes.

Figure 5 shows that heptane conversion is completed within a few centimetres after the entrance, especially for higher steam-to-carbon ratios. The rest of the reactor then proceeds as in steam reforming of methane.

[FIGURE 7 OMITTED]

As seen from Figure 6, with increasing steam-to-carbon ratio, the hydrogen permeate yield is predicted to be enhanced, correspondingly increasing the overall hydrogen yield.

Based on these observations, as shown in Figure 7, the FBMR, can be considered to be composed of two overlapping zones: Zone 1, a short zone, where steam reforming of heptane is completed, and Zone 2, for steam reforming of methane. Thus, in this bi-functional reaction and separation set-up, a separate pre-reformer is not needed, since with hydrogen permeation, the reaction can proceed towards completion in the same unit. In view of the pure hydrogen permeation, PSA units are also not required.

The main challenge for the competitiveness of this technology lies with membrane issues, in particular in assuring pin-hole-free high-flux perm-selective membranes. Figure 8 shows the effects of decreasing the membrane thickness for a reformer operating at 650[degrees]C and 10 bar. Thinner membranes could minimize the residual methane and hydrogen in the reformer, and maximize the pure (permeate) hydrogen yield.

Figure 9 shows the increase of hydrogen permeate yield with increasing specific membrane surface area for a reformer operating at 650[degrees]C and 10 bar. Steam reforming reactions being very rapid, and hydrogen permeation being slow, an important parameter is the membrane packing factor, a, defined as the membrane surface area per unit volume of reactor. As this factor is increased, the reformer performance as measured in terms of pure hydrogen yield, is significantly enhanced, and a significantly smaller reformer can be used.

Thus this multifunctional reactor is predicted to be able to combine the units from a pre-reformer, reformer and hydrogen purifier into a single unit. The sequence of events can be considered to be:

i. Steam reforming of higher hydrocarbon, depicted in Figure 5.

ii. Methanation, indicated in Figure 4a when the peak is attained for the methane yield.

iii. Steam reforming of methane, depicted in Figure 4a, when the methane conversion becomes zero, thus completing the full conversion of the hydrocarbons.

iv. Hydrogen permeation until the hydrogen partial pressure in the retentate equalizes with that in the permeate stream, evident from Figure 4b.

v. In parallel with step (iv), net interphase mass transfer between the bubble and dense phases is also completed.

[FIGURE 8 OMITTED]

When this sequence of events is complete, the species concentrations in the two phases do not change any further, and the concentration profiles remain flat thereafter, as in Figure 3. The reformer heights corresponding to this sequence of events depend on the operating parameters including reformer pressure, membrane permeate side pressure, reformer temperature, steam-to-carbon ratio in the feed, and superficial velocity.

[FIGURE 9 OMITTED]

CONCLUSIONS

n-Heptane was used as a model component for higher hydrocarbons, close to the naphtha cut. In situ permselective membranes should be able to produce ultra-pure hydrogen as required by some sectors like the fuel cell industry. Higher conversion of methane (produced by the methanation reaction) allows the reformer to be operated at much lower temperature to achieve the same hydrogen yield as for much higher temperatures without membranes. An FBMR system for higher hydrocarbons can result in a compact reformer system combining the units from a pre-reformer, reformer and hydrogen purification into a single unit.

However, for the system to be economically viable and competitive, major challenges remain for the membranes. Desirable membrane features are:

* High flux.

* High selectivity to hydrogen.

* Low cost.

* Longevity.

* High membrane packing, while maintaining a minimum separation requirement in a fluidized space to prevent solids bridging and gas bypassing.

Challenges specific to higher hydrocarbons include catalyst deactivation and possible membrane fouling. These have not been considered in this paper, but will be key factors to be examined in the experimental work.

The model considers the bubbling bed mode of operation, as this will be the main operating regime in the forthcoming experiments. However, many industrial fluidized bed reactors are operated in the turbulent regime in view of the higher throughput and advantageous features (Bi et al., 2000). The transition from bubble to turbulent flow happens earlier for powders with smaller mean particle diameter and wider particle size distributions (Sun and Grace, 1992). The experimental work will include determination of this transition at the temperature and pressure of the reformer, and investigate how it affects the hydrogen yield.

ACKNOWLEDGEMENTS

The authors gratefully acknowledge financial support provided by the Natural Sciences and Engineering Research Council of Canada (NSERC).

NOMENCLATURE

a membrane packing factor ([m.sup.2]
 membrane area per [m.sup.3] of reactor)
[a.sub.b] specific surface area of gas bubbles
 ([m.sup.2]/[m.sup.3])
A cross-sectional area of bed ([m.sup.2])
[A.sub.P] membrane permeation area per unit length
 of membrane ([m.sup.2]/m)
Ar Archimedes number
[C.sub.i] concentration of species on reaction side
 (mol/[m.sup.3])
[d.sub.b] bubble diameter (m)
[d.sub.bm] maximum bubble diameter (m)
[d.sub.bo] bubble diameter just above distributor
 (m)
[d.sub.p] mean particle diameter (m)
D reactor diameter (m)
[D.sub.ie] effective molecular diffusivity of
 component i ([m.sup.2]/s)
[D.sub.ij] binary diffusivity of components i and j,
 with i not equal to j
[MATHEMATICAL EXPRESSION activation energy for permeation (J/mol)
NOT REPRODUCIBLE IN ASCII.]
[F.sub.i] molar flow rate of species i on the
 reaction side (mol/s)
[F.sub.P] flow rate of permeate [H.sub.2] (mol/s)
[F.sub.T] total molar flow rate including all
 species along bed crosssection (mol/s)
h vertical position above the distributor
 (m)
[DELTA]H heat of reaction (kJ/mol)
HTS high temperature shift
[MATHEMATICAL EXPRESSION permeation flux of hydrogen through Pd
NOT REPRODUCIBLE IN ASCII.] membranes (mol/[m.sup.2] s)
L vertical distance measured from feed
 inlet (m)
LTS low temperature shift
[P.sub.i] partial pressure of species i (bar)
[MATHEMATICAL EXPRESSION partial pressure of hydrogen in reactor
NOT REPRODUCIBLE IN ASCII.] side in bubble phase
[MATHEMATICAL EXPRESSION partial pressure of hydrogen in reactor
NOT REPRODUCIBLE IN ASCII.] side in dense phase (bar)
[MATHEMATICAL EXPRESSION partial pressure of hydrogen in permeate
NOT REPRODUCIBLE IN ASCII.] side (bar)
[P.sub.M0] pre-exponential factor for permeation
 (mol/(mmin [atm.sup.0.5]))
[Q.sub.i] permeation rates per unit length from
 reactor side to sweep gas side for
 species i (mol/m s)
[r.sub.j] rate of jth reaction (mol/kg-cat s)
R universal gas constant (J/mol K)
[Re.sub.mf] Reynolds number at minimum fluidization
 velocity
SCR steam-to-carbon molar ratio
[X.sub.i] mole fraction of component in gas mixture
[U.sub.b] rising velocity of bubble (m/s)
[U.sub.o] superficial gas velocity (m/s)
[U.sub.mf] superficial gas velocity at minimum
 fluidization (m/s)
[MATHEMATICAL EXPRESSION thickness of hydrogen selective membranes
NOT REPRODUCIBLE IN ASCII.] (m)
[[epsilon].sub.mf] bed voidage at minimum fluidization
[[epsilon].sub.b], volume fraction of catalyst bed occupied
[[epsilon].sub.d] by bubble and dense phase, respectively
[[phi].sub.b], volume fraction of catalyst bed occupied
[[phi].sub.d] by solid particles in bubble and dense
 phase, respectively

Appendix A

A.1 Kinetic Expressions for Reactions in Reformer

* Steam reforming of higher hydrocarbons: (Tottrup, 1982)

[C.sub.n][H.sub.m] + n[H.sub.2]O [right arrow] nCO + (n + m/2) [H.sub.2][r.sub.1]

For Heptane, n = 7, [DELTA][H.sup.o.sub.298] = 1108 kJ/mol

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.]

[k.sub.1] = 8 x [10.sup.-5] exp(- 67 800/RT) (mol/[g.sub.catalyst] h bar)

[K.sub.a] = 25.2 [bar.sup.-1]

[K.sub.b] = 0.077

* Steam reforming of methane: (Xu and Froment, 1989)

C[H.sub.4] + [H.sub.2]O [left and right arrow] CO + 3[H.sub.2] [r.sub.2], [DELTA][H.sup.o.sub.298] = 206.1 kJ/mol

CO + [H.sub.2]O [left and right arrow] C[O.sub.2] + [H.sub.2] [r.sub.3], [DELTA][H.sup.o.sub.298] = -41.1 kJ/mol

C[H.sub.4] + 2[H.sub.2]O [left and right arrow] C[O.sub.2] + 4[H.sub.2] [r.sub.4], [DELTA][H.sup.o.sub.298] = 165.0 kJ/mol

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.]

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.]

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.]

[MATHEMATICAL EXPRESSION NOT REPRODUCIBLE IN ASCII.]

The equation parameters are available in Xu and Froment (1989).

A.2 Hydrodynamic Equations for the Two-Phase Model

Bubble size distribution (Mori and Wen, 1975):

[d.sub.b] = ([d.sub.bm] - [d.sub.bo]) [e.sup.-0.3h/D]

[d.sub.bm] = 1.64[A([[U.sub.o] - [U.sub.mf])].sup.0.4]

[d.sub.b0] = 1.38/[g.sup.0.2] [A([U.sub.o] - [U.sub.mf])/[N.sub.or]]

Fraction of bed occupied by bubbles:

[[epsilon].sub.b] = [U.sub.o] - [U.sub.mf]/[U.sub.b]

Bubble rise velocity (Davidson and Harrison, 1963):

[U.sub.b] = [U.sub.O] - [U.sub.mf] + 0.711 [([gd.sub.b]).sup.1/2]

Minimum fluidization velocity (Wen and Yu, 1966a,b):

[Re.sub.mf] = [[[(33.7).sup.2] + 0.0408 Ar].sup.1/2] -33.7

Volume fraction of solids:

[[phi].sub.d] = (1 - [[epsilon].sub.b]) (1 - [[epsilon].sub.mf]), [[phi].sub.b] = 0.001 [[epsilon].sub.b]

Manuscript received January 4, 2008; revised manuscript received February 15, 2008; accepted for publication February 21, 2008.

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M. A. Rakib, (1) * J. R. Grace, (1) S. S. E. H. Elnashaie, (2) C. J. Lim (1) and Y. G. Bolkan (3)

(1.) Department of Chemical and Biological Engineering, University of British Columbia, 2360 East Mall, Vancouver, BC, Canada V6T I Z3

(2.) Environmental and Sustainable Engineering, Pennsylvania State University at Harrisburg, Capital College, 777 W. Harrisburg Pike, Middletown, PA 17057-4898, U.S.A.

(3.) Department of Chemical and Petroleum Engineering, University of Calgary, 2500 University Drive NW, Calgary, AB, Canada T2N 1N4

* Author to whom correspondence may be addressed. E-mail address: mrakibaa chmLabc.ca Can. J. Chem. Eng. 86:403-412, 2008

Table 1. Reactor geometry and base simulation parameters

Reformer empty cross-sectional area 2.0x[10.sup.-3] Reformer
 [m.sup.2]
Specific membrane area 64 [m.sup.2]/
 [m.sup.3] of
 reactor volume
Total membrane length (along 1 m
height of reformer)

Catalyst type Ni-[Al.sub.2] Catalyst
 [O.sub.3]
Catalyst particle mean diameter 100 [micro]m
Catalyst particle density 2270 kg/[m.sup.3]

Steam:carbon ratio in feed 3 Process
n-Heptane mole fraction in feed 0.0454 operating
Steam mole fraction in feed 0.9538 conditions
[H.sub.2] mole fraction in feed 0.0008
Feed temperature 650[degrees]C
Feed pressure 10 bar abs
Membrane permeate side pressure 0.3 bar abs
Reactor inlet gas superficial velocity 0.23 m/s