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Jun 29, 2005 - The performance of a novel gasoline steam reformer with facing catalytic burner and ... reference catalytic SR reactor facing, as the h...
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Ind. Eng. Chem. Res. 2005, 44, 9422-9430

Combining Catalytic Combustion and Steam Reforming in a Novel Multifunctional Reactor for On-Board Hydrogen Production from Middle Distillates G. A. Petrachi, G. Negro, S. Specchia, G. Saracco, P. L. Maffettone,* and V. Specchia Dipartimento di Scienza dei Materiali ed Ingegneria Chimica, Politecnico di Torino, Corso Duca degli Abruzzi 24, 10129 Torino, Italy

The performance of a novel gasoline steam reformer with facing catalytic burner and reformer both based on the microchannel concept (system A) was modeled and then compared with a reference catalytic SR reactor facing, as the heat source, a microchannel heat exchanger fed with the flue gases from an external burner (system B). Reactions in the catalytic reformer and the catalytic burner occurred heterogeneously on Pt-CeO2 catalysts. The developed dynamic model [25 partial differential equations (PDEs) for system A; 15 PDEs for system B; reforming and combustion kinetics from the scheme proposed by Pacheco et al. (Appl. Catal. 2003, 250, 161)] was numerically solved. The simulation results showed that the integrated burnerreformer microchannel reactor (system A) was able to transfer heat from the hot side to the cold side very efficiently, thus enabling high compactness. However, the best design conceived (the catalytic burner facing toward the reformer microchannel) showed mass-transfer limitations inside the burner catalyst layer because of the low values of the effective diffusion inside the solid catalytic phase. As a consequence, there was not an “ignited” condition in the catalytic burner (confirmed by the lack of a peak in the burner temperature profile) but a diffuse heat generation along the microchannel length. This condition involves a longer reformer start-up but at the same time allows one to limit the temperature increase in the catalyst solid layer, preserving its activity and stability. From the dynamic analysis, a longer start-up time resulted for system A, which could be a limitation for automotive applications. Conversely, this system proved to guarantee a fast response to sudden load changes: hydrogen production reached almost instantaneously the new steady-state conditions, as needed for vehicle applications. 1. Introduction Population growth and the related increase in the energy demand for both residential and mobile applications require new ways to generate electric power with low emissions; hydrogen fuel cell technology could actually help in achieving this goal. Automotive applications require fuel cells with high specific power, high efficiency, low costs, long durability, and compact dimensions. Proton-exchange-membrane fuel cells (PEM-FCs) are typically implemented in cars because they are small and light, they operate at relatively low temperature, and they are operationally flexible. The required power supply can be obtained by stacking several elements. A fuel cell, which uses pure hydrogen as fuel, produces only steam with zero polluting emissions; however, for mobile applications, apart from the high cost of the on-board tanking, the lack of infrastructures for pure hydrogen refuelling is currently hampering wide application. In the short term, the only feasible way to produce hydrogen, sufficient enough to satisfy the actual and future demand, is hydrocarbon conversion (primarily gasoline and diesel oils) by autothermal reforming (ATR) and steam reforming (SR), thereby obtaining hydrogen-rich gases. ATR is a self-sustaining process because the fuel is at the same time burnt and steam-reformed within the * To whom correspondence should be addressed. Tel.: +39-011-5644649. Fax: +39-0115644699. E-mail: [email protected].

same reactor, but it produces a reformate gas with low hydrogen partial pressure, which limits the fuel cell efficiency. This problem can be overcome using a SR process that produces more hydrogen with a higher partial pressure, but this requires a heat source and a large heat-exchange surface, which limit drastically its possible applications. A comparison between ATR and SR fuel processors shows that the latter enables a slightly higher global efficiency than the former when different primary fuels (gasoline, light diesel, and biodiesel) are used,2 but commonly it is not particularly suitable for automotive applications because of the large space needed. Multifunctional reactors hosting catalytic combustion and steam reforming of fuel at the opposite sides of a heat-exchange surface appear to be very promising for achieving maximum compactness, a mandatory property for an auxiliary power unit placed on board road vehicles, yachts, and aircrafts. Microchannel reactors can satisfy this requirement because of their high specific exchange area and high heat-transfer coefficients,3 saving at the same time energy costs because of their low pressure drop. In fact, on the basis of the friction factor estimation in microchannel systems using various literature correlations,3-5 the calculated pressure drops for the microchannel SR reactor considered in this work (see the following paragraphs) are about 20-80 Pa. Neither theoretical nor experimental study exists in the technical literature related to the above compact

10.1021/ie050215n CCC: $30.25 © 2005 American Chemical Society Published on Web 06/29/2005

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external burner (system B; Figure 1B). In both cases, the SR outlet stream, being rich in CO, has to be purified because CO poisons the PEM-FC catalyst. Usually, the cleanup is performed by the water-gas shift (WGS) reaction, followed by preferential CO oxidation (PrOx); the cleaned gas is finally fed to the fuel cell. The hydrogen off-gas from the anode fuel cell is usually recycled to an afterburner so as to recover its thermal energy: in system A, it is burnt in an external catalytic afterburner to heat up the combustion air for the catalytic integrated fuel burner facing the steam reformer; in system B, the hydrogen off-gas is burnt together with the fuel in the external burner for the production of the SR heating phase. This paper deals with the SR reactor only: the novel system A is at first modeled and numerically simulated; then, its performance is compared to that of the more conventional system B, assumed as reference. The aim of this theoretical work is to develop a dynamic model to understand the behavior of the novel integrated reactor, to characterize its transient response under changing loads, and to obtain useful information for the design of a prototype integrated steam reformer. SR reactors for automotive applications are, in fact, typically operated under rapid variation of operating conditions, and thus a dynamic description of the process could be helpful for both design and control purposes. 2. Model Description

Figure 1. Schematic of a gasoline SR fuel processor for automotive applications: SR heated by a gasoline catalytic burner (A) or by the flue gases of an external noncatalytic combustor (B).

integrated reactor hosting gasoline by both the SR and combustion processes separated by a metal wall. Only a work on methanol steam reforming in a microchannel reactor was found,6 even if in this case an electric heater was used to sustain the endothermic reaction. With regards to the patents area, three patents only consider the microchannel technology with integrated heat exchange applied to both generic exothermic and endothermic reactions (combinatorial chemistry)7 and to the Fischer-Tropsch synthesis.8,9 Moreover, an international patent deals with a fuel processor10 equipped with an internally heated microchannel steam reformer, even if it does not consider explicitly catalytic combustion as the heat source. The typical configuration of the hydrogen production process (fuel processor) based on a SR reactor, starting from gasoline as the fuel, is sketched in Figure 1. Gasoline and water are evaporated and then fed to the steam reformer; a burner coupled in some way with the reformer supplies the required heat. Two technical approaches for the steam reformer are considered: a novel one, based on the microchannel technology, where a reactor with a catalytic combustor and a catalytic steam reformer faces the same steel wall from opposite sides (system A; Figure 1A); a microchannel SR reactor, where gasoline steam reforming is sustained by internal heat exchange with the flue gases produced by a traditional noncatalytic combustion carried out in an

For system A, the reactor facing the catalytic burner can be divided into elementary units, composed of two gas phases (the burner side and the reformer side), two catalyst layers (one for the reformer and the other for the burner), and a steel wall that separates the two reaction volumes. A broad section of this elementary unit is shown in Figure 2. For system B, the elementary unit of the reactor heated with flue gases from the noncatalytic external burner is similar to that of system A: obviously, the oxidizing catalyst layer is not present in the heatexchange side. For the microchannel geometry, the width/height ratio W is, in particular, the most important parameter to maximize heat exchange: its optimal value is W ) 0.1.11 In another theoretical work, validated by experimental data, it is shown that high microchannel heat-transfer coefficients are achieved only when the microchannel hydraulic diameter Dh is smaller than 1.167 mm.12,13 Therefore, a Dh value of 1 mm was chosen for the reactor configuration used for modeling, thus obtaining the elementary unit dimensions listed in Table 1. The desired hydrogen flow rate for a fixed fuel cell power can be produced with a SR reactor obtained by coupling several elementary units. For example, with a reactor constituted of about 200 burner-reformer units, by properly cleaning up the reformate stream, the hydrogen amount necessary to operate a 1 kWe PEM-FC can be produced. A model was ad hoc developed to qualitatively describe the reactor behavior, and several simplifications were considered to avoid cumbersome analysis. Both the reformer and the catalytic burner were modeled as heterogeneous reactors consisting of a gas phase in the microchannel and a catalytic layer deposited on the steel wall. A one-dimensional system was considered: only the longitudinal coordinate was taken into account. Both the burner and the reformer were operated under

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Figure 2. Elementary unit of the integrated SR: system A with a microchannel catalytic burner and system B with microchannel flue gases. Table 1. Dimensions of the Integrated SR Elementary Unit

Table 3. Kinetic Parameters Used in the SR Reactor Modeling

5.50 × 10-3 0.55 × 10-3 190 × 10-3 0.05 × 10-3 0.5 × 10-3

microchannel height (m) microchannel width (m) microchannel length (m)a catalyst layer thickness (m) steel wall thickness (m)

a For the steam reformer heated by a catalytic combustor (system A) and operated as described in condition 1, microchannels of 380 × 10-3 m long were chosen.

Table 2. Reaction Model and Kinetic Expressions Used in the SR Reactor Modeling C8H18 + 25/2O2 f 8CO2 + 9H2O r1 ) k1pC8H18pO2

C8H18 + 8H2O f 8CO + 17H2

C8H18 + 8CO2 f 16CO + 9H2

r2 )

k2 pH 2

pC8H18pH2O -

(

k5 r5 ) pH 2

)

3.5

)

(3) 4

(

pCOpH2O -

(

(2)

K3pC8H18pCO2

1 + KH2O

(

2

pCO2pH22

)

pH2 pCO2

pC8H18pH2O2 -

k4 pH 2

K2

p H 2O 1 + KH2O pH2

(

)

pH23pCO

(

r3 ) k3pC8H18pCO2 1 -

C8H18 + 16H2O f 8CO2 + 25H2 r4 )

CO + H2O f CO2 + H2

(

(1)

K4 pH2O 2 pH 2

)

)

(4)

pH2pCO2 K5

)

p H 2O 1 + KH2O pH 2

parameter

preexponential factor

activation energies and water heat of adsorption (kJ/mol)

k1 (mol/gcat‚s‚bar2) k2 (mol‚bar0.5/gcat‚s) k3 (mol/gcat‚s‚bar2) k4 (mol‚bar0.5/gcat‚s) k5 (mol/gcat‚s‚bar) KH2O

2.58 × 108 2.61 × 109 2.78 × 10-5 1.52 × 107 1.55 × 101 1.57 × 104

166 240.1 23.7 243.9 67.1 88.7

Mass- and heat-transfer coefficients were assumed constant, independent of the axial position or temperature. As for the heat-transfer coefficient, a specific correlation developed for microchannels was found in the literature,5 where the Nusselt number under laminar conditions depends only on the microchannel W characteristic value. In the heat balance of the metal wall, the heat losses to the surroundings were also considered, taking into account the radiant mechanism due to the high-temperature driving force between the metal wall and the external air, assumed at a constant temperature of 25 °C. Considering a fuel processor for 10 kWe, the global heat losses from the external surface area of the integrated microchannel burner-SR reactor were calculated; the heat losses were then uniformly shared to about 2000 burner-reformer elementary units constituting the whole integrated system, obtaining a small average heat loss value. Therefore, taking into account the small external heat-exchange area of the elementary unit, the latter was assumed not to be insulated; in this way, conservative calculations were developed during the system simulations.

2

(5)

constant pressures (fixed independently of each other) and laminar conditions. Longitudinal diffusion and gas thermal conductivity were taken into account for both the reformer and the catalytic burner gas phases; they were assumed constant independently from the temperature and composition. Longitudinal thermal conductivity was also taken into account in either the steel wall or the catalytic layers.

3. Reaction Kinetics and Operating Conditions Commercial gasoline is a mixture of a large number of hydrocarbons, so it is really difficult to develop a model taking into account all of the reacting species. Therefore, as is commonly done in studies about internal combustion engines, pure isooctane was assumed to be representative of gasoline. Kinetic expressions for isooctane combustion and steam reforming are not common in the literature, but in a recent work,1 a reaction model and the respective

Ind. Eng. Chem. Res., Vol. 44, No. 25, 2005 9425 Table 4. Summary of Operative Conditions for Both SR Systems and Heating-Side Conditions condition 1

heating side

reformer side

electric power attainablea C8H18 fed to burner air excess C8H18 fed to reformer steam-to-carbon ratio burner/reformer isooctane initial temperature inlet temperature pressure inlet velocity inlet molar flow C8H18 molar fraction O2 molar fraction CO2 molar fraction H2O molar fraction N2 molar fraction initial temperature inlet temperature pressure inlet velocity inlet molar flow C8H18 molar fraction H2O molar fraction

We mol/s % mol/s °C °C bar m/s mol/s % % % % % °C °C bar m/s mol/s % %

condition 2

system A (catalytic burner)

system B (flue gases)

system A (catalytic burner)

system B (flue gases)

5 2.08 × 10-6 20 2.43 × 10-6 3 0.856 25 580 1.5 2.91 1.89 × 10-4 1.1 16.5 10.4 9.9 62.1 25 500 2 0.64 6.07 × 10-5 4 96

5 1.68 × 10-6 94.4 2.43 × 10-6 3 0.691 25 1000 1.5 5.50 2.41 × 10-4 0 8.3 13.7 14.1 63.9 25 500 2 0.64 6.07 × 10-5 4 96

5 1.41 × 10-6 20 2.43 × 10-6 3 0.580 25 760 1.5 2.62 1.41 × 10-4 1.0 15.0 14.0 13.4 56.6 25 500 2 0.64 6.07 × 10-5 4 96

5 2.31 × 10-6 195 2.43 × 10-6 3 0.951 25 800 1.5 8.81 4.54 × 10-4 0 12.4 8.4 8.7 70.5 25 500 2 0.64 6.07 × 10-5 4 96

a Electric power attainable with H produced by the designed elementary unit in a fuel processor with WGS and PrOx reactors in the 2 cleanup section.

Langmuir-Hinshelwood-Hougen-Watson kinetic expressions for a Pt-CeO2 catalyst adapted from previous works of Xu and Froment14,15 on methane were found. Table 2 lists the related reactions and kinetic expressions, whereas Table 3 shows the kinetic parameters used. Operating conditions, the steam-to-carbon (S/C) ratio and inlet temperatures in particular, are really important in a SR process. Generally, it is advisable to operate with high S/C value not only to increase the isooctane conversion and hydrogen equilibrium concentration but also to avoid coke formation and deposition, thus preserving the catalytic activity.16 In automotive applications too, high S/C ratios are difficult to cope with because of the need to close the water balance with no external water feed and because of the small space available on the vehicle for heat exchangers and economizers. Under these circumstances, it is reasonable to operate with a S/C ratio equal to 3 and a reformer inlet temperature of 500 °C. The gas phases flow cocurrently in the microchannels; this is, in fact, the arrangement that enables maximum hydrogen production and isooctane conversion. The temperature inside the reactor should indeed be higher at the beginning, so as to promote isooctane conversion, and lower at the end to assist the exothermic WGS reaction, so as to increase the hydrogen partial pressure. This optimal condition is achievable only if reformer and burner gases are fed cocurrently. The operating conditions for systems A and B are listed in Table 4. Before start-up of the reformer, both the heating and reforming sides needed to be preheated from room temperature to a suitable temperature necessary to overcome computational problems linked with the particular kinetic equations employed. System A was preheated by feeding the catalytic burner with an airfuel mixture at the fixed burner inlet temperature, whereas the reformer was fed by overheated steam; the

system B preheating was, instead, carried out only by feeding the flue gases to the SR heating side. The preheating phase was not investigated in this work, which was at the moment dedicated only to the reformer dynamics analysis without taking into account the whole fuel processor. For both systems A and B, after the preheating procedure, the feeding of the isooctane-steam mixture to the reformer microchannel was started. Concerning the inlet temperature at the heating side of the integrated reforming reactor for both systems A and B, two different conditions were considered: 1. Condition 1: a minimum value for the catalytic burner of system A (580 °C), close to the isooctane autoignition temperature, and a maximum value for the heating side of system B (1000 °C), compatible with the use of a construction material (Fe-Cr alloy, for example) very resistant to high temperature but very expensive. 2. Condition 2: a higher temperature for system A (760 °C), as obtained by vaporizing-heating of the catalytic burner feeds (isooctane and combustion air), and a lower temperature of system B (800 °C), compatible with the thermal resistance of less costly metallic materials. In the catalytic burner of system A, isooctane combustion occurs with an air excess of 20%, so as to prevent CO formation and to ensure a complete fuel burnout. Conversely, the external noncatalytic burner works with global air excess of about 95% and 195% for conditions 1 and 2, respectively, which at the same time ensure very low NOx emissions. A conventional burner operating at high air excess, especially such as that of condition 2, could not ensure a stable isooctane combustion with the whole air stream fed to the burner; to avoid this (notwithstanding it is not relevant for the analysis and the results of this work), the whole air rate was divided into the two streams (see Figure 1) of primary (directly fed to the burner to properly operate

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Table 5. Balance Equations for System A

Table 6. Balance Equations for System B

Mass Balances

(

)

(6)

Rcb j )

(7)

∂ybj ∂yj ∂2ybj av(mcb/Fcb) T b cb yj - ybj ) -vb + Db 2 + Kov ∂t ∂x ∂x Vb T cb ∂ycb j ∂t

(

)

T cb b mcb RT cb cb + (Rj - ycb yj - ycb j j T  Pb

) Kovav

∑

(

j

∂yrj ∂yrj ∂2yrj av(mcr/Fcr) T r cr yj - yrj ) -vr + Dr 2 + Kov ∂t ∂x ∂x Vr T cr ∂ycr j ∂x

) Kovav

(

T

cr

T

r

)

cr

yrj - ycr + j

m RT 

P

r

cr cr (Rcr j - yj

Mass Balances equations (8) and (9)

)

∑R

cr j )

(8)

(9)

j

(10)

cb U1 av ∂T cb k (1 - ) ∂2T cb ) cb cb + cb cb(T b - T cb) + ∂t F cp ∂x2 cp F U3 Apm pm rI∆Hr,I (T - T cb) - cb (11) cpcb mcb cp (1 - )

U1 av(mcr/Fcr) cr ∂T r ∂T r kr ∂2T r ) -vr + r r + r r (T - T r) 2 ∂t ∂x F cp ∂x F cp Vr cr U2 av ∂T cr k (1 - ) ∂2T cr ) + cr cr(T r - T cr) + cr cr 2 ∂t F cp ∂x cp F U4 Apm pm (T - T cr) cpcr mcr rII∆Hr,II + rIII∆Hr,III + rIV∆Hr,IV + rV∆Hr,V

cpcr(1 - )

U1 Apm ∂T b ∂T b kb ∂2T b + b b b (T pm - T b) - vb + b b 2 ∂t ∂x F cp ∂x F cp V

(15)

U2 Apm ∂T pm kpm ∂2T pm + pm pm(T b - T pm) + ) pm pm 2 ∂t F cp ∂x cp m U3 av(mcr/Fcr)mpm b σr Aperd (T - T pm) + pm pm [Test4 - (Tpm)4] pm pm cp m cp m (16) Table 7. Boundary and Initial Conditions for System A

Heat Balances

U1 av(mcb/Fcb) cb ∂T b ∂T b kb ∂2T b ) -vb + b b + (T - T b) ∂t ∂x F cp ∂x2 Fbcpb Vb

Heat Balances equations (12), (13), and

(12)

(13)

U2 av(mcb/Fcb) cb ∂T pm kpm ∂2T pm ) pm pm + (T - T pm) + ∂t F cp ∂x2 cppm mpm U4 av(mcr/Fcr) cr σr Aperd (T - T pm) + pm pm [Test4 - (T pm)4] (14) pm pm cp m cp m

the latter) and secondary air (to be mixed with the flue gases from the burner to reach the fixed temperature level). 4. Balances The following chemical compounds are present in the steam reformer: C8H18, H2O, CO, CO2, and H2. The heating side contains C8H18, H2O, CO2, O2, and N2. For system A, the mass and heat balances were written for each chemical component in the four subsystems: reformer gas phase, burner gas phase, and their respective catalytic solid layers; a heat balance was also written for the steel wall. The model equations for system A are listed in Table 5. In system B, the mass balance equations were instead written only for the reformer gas phase and its respective catalytic solid phase; the gas-phase composition of the heating side, in fact, remains the same. As for the heat balance equations for system B, because there is not a catalyst layer at the heating side, no heat balance equation was considered for this phase. The mass and heat balance equations for system B are reported in Table 6.

b ybj (0,x) ) yj,0 b b yj (t,0) ) yj,in b JN,j |(t,L) ) -vb(Pb/RT b)ybj cb cb yj (0,x) ) yj,0 cb JN,j |(t,0) ) 0 cb JN,j |(t,L) ) 0 r r yj (0,x) ) yj,0 r yrj (t,0) ) yj,in r JN,j|(t,L) ) -vr(Pr/RT r)yrj cr ycr j (0,x) ) yj,0 cr JN,j|(t,0) ) 0 cr JN,j |(t,L) ) 0 Tb(0,x) ) Tb0 Tb(t,0) ) Tbin

(17) (18) (19) (20) (21) (22) (23) (24) (25) (26) (27) (28) (29) (30)

JbQ|(t,L) ) -vbFbcpbT b Tcb(0,x) ) Tcb 0 Jcb Q |(t,0) ) 0 Jcb Q |(t,L) ) 0 T r(0,x) ) Tr0 T r(t,0) ) Trin JrQ|(t,L) ) -vrFrcprT r T cr(0,x) ) Tcr 0 Jcr Q |(t,0) ) 0 cr JQ |(t,L) ) 0 T pm(0,x) ) Tpm 0 Jpm Q |(t,0) ) 0 Jpm Q |(t,L) ) 0

(31) (32) (33) (34) (35) (36) (37) (38) (39) (40) (41) (42) (43)

Because ideal conditions were assumed for the gas phases, the local volumetric flow rate depends on both the local temperature and average molar mass. As already mentioned, the pressures in both reactors were assumed to be constant (negligible pressure drop). The average mass density and heat capacity were treated as pseudo-steady-state quantities and, thus, appeared as coefficients of the temperature space derivative, whose value was continuously updated. The model was based on 25 partial differential equations (PDEs) for the steam reformer integrated with the catalytic burner (system A) and on 15 PDEs for the steam reformer operating with flue gases (system B). The 27 boundary and initial conditions for system A are listed in Table 7. Conversely, for system B, because of a lack of the heating side mass balance and the mass and heat balance for the heating side catalytic layer, only the conditions in Table 7 from (23) to (31) and from (35) to (40) were used. It should be mentioned that the specific heats of both the burner and reformer catalyst were assumed to be equal to that of the CeO2 support because the presence of Pt in the structure was negligible. Reaction enthalpies were considered to be temperature-independent. The set of PDEs was integrated for both systems with the method of lines based on a Matlab ODE solver implementing an implicit Runge-Kutta formula with a first stage that is a trapezoidal rule step and a second stage that is a backward differentiation formula of order 2. The results presented below were obtained with a spatial discretization of 97 points because this value ensures numerical convergence, good accuracy, and acceptable computation time.

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Figure 4. Steady-state isooctane molar fraction profiles for system A in the bulk gas and catalyst phases of both the burner and reformer for conditions 1 and 2.

Figure 3. Steady-state isooctane conversion profiles for systems A and B and operating conditions 1 and 2.

5. Results and Discussion The isooctane conversions at the steam reformer and heating sides of both systems A and B are compared in Figure 3 for feed conditions 1 and 2. It is important to remark again on the different microchannel lengths: to achieve a final conversion, system A at condition 1 (named system 1-A) needs a longer microchannel (380 mm) than the one obtained by the other cases examined (190 mm). System A with the catalytic burner fed at 760 °C (system 2-A in Figure 3) showed a superior performance: an isooctane conversion over 99.9% is reached at both the burner and reformer sides, even when using the lowest isooctane molar flow at the burner inlet. The smaller conversion in the steam reformer for system 1-A depends on the low inlet temperature in the steam reformer and the low masstransfer rate from the burner gas bulk to the catalyst solid phase that decreases the isooctane combustion rate. As a consequence, the low heat flow generated at low temperature by the combustion reaction limits the SR reaction rate; a greater reaction volume is needed to achieve a 98% SR isooctane conversion. The systems B operating with flue gases (named 1-B and 2-B) always showed, as reported in Figure 3, good isooctane conversion, independently from the gas inlet temperature in the integrated heat exchanger. This is due to the high temperature anyway obtained that ensures a really high SR reaction rate. For system A, at the burner side the mass-transfer rate is the controlling resistance, whereas at the reformer side, the overall rate is kinetically controlled, as confirmed by Figure 4. The molar fraction of isooctane in the burner solid phase is indeed always near zero and quite different from that in the bulk gas phase, confirming the large mass-transfer resistance; in the reformer microchannel, instead, the two concentration profiles are very near each other and really higher than zero, demonstrating the primary role of the reaction kinetic resistance on the overall rate. The mass-transferlimitation problem depends only on the low value of the isooctane effective diffusivity in the burner catalytic

Figure 5. Steady-state concentration profiles in the reformer gas phase of system A for condition 1 (1-A) and for condition 2 (2-A).

layer because the laminar convective mass-transfer coefficient is quite high as a result of the small microchannel size. This situation is really disadvantageous for catalyst layers where most of the active species are distributed within the bulk of the catalyst. The reformer bulk gas-phase concentration profiles at steady state for system A for the two inlet conditions are shown in Figure 5. Also, for the steam reformer operating with flue gases (system B), the reformer concentration profiles are similar to those reported in Figure 5 as a consequence of the similarity of the isooctane conversion profiles in both systems A and B. In this case, with the steam reformer, the hydrogen molar fraction is almost 50%, a quite high value that can even be increased in the downstream WGS processing step. Because the CO outlet concentration is, in fact, about 9%, a hydrogen concentration of approximately 58-58.5% at the WGS reactor outlet may be expected.

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Figure 6. Comparison between H2 and CO concentrations in the reformer gas phase for systems 1-A and 2-A.

Figure 8. Start-up and transient response results for systems A and B and conditions 1 and 2.

Figure 7. Steady-state temperature profile for both systems A and B and conditions 1 and 2.

The H2 and CO concentration profiles of the reformer gas phase for both systems 1-A and 2-A are shown in Figure 6 as a function of the microchannel length. The different profiles for CO are worth noticing: for system 2-A, the CO concentration reaches higher values inside the microchannel because of the higher temperature, which increases the SR reaction rate, even if at the microchannel end it assumes practically the same value of system 1-A. System 1-A showed a slightly higher H2 outlet concentration than system 2-A: probably the lower reformer outlet temperature (see Figure 7) and the larger reaction volume for system 1-A, favoring the WGS equilibrium, allow one to obtain a slightly H2-richer stream at the reactor outlet. Both systems 1-A and 2-A reach about the same outlet composition, but system 2-A has a a smaller reactor (190 mm long versus 380 mm for system 1-A) and a lower isooctane consumption in the catalytic burner (1.41 × 10-6 mol/s versus 2.08 × 10-6 mol/s for system 1-A; see Table 4). Thus, it is clear that with system 2-A a more effective exploitation of the reaction volume can be achieved. The temperature steady-state profiles for both systems A and B and both inlet conditions 1 and 2 are reported in Figure 7. The lack of a marked temperature

peak in the burner curve of both systems 1-A and 2-A is noteworthy, as shown in other works on similar macroscale systems.17 This confirms the absence of a single ignition point in the reactor and the presence of a diffuse, but limited, heat generation along the microchannel length. The good overlapping of the temperature profiles is also remarkable, which demonstrates the high heat-transfer efficiency between the hot and cold microchannels, also under different conditions, as was predicted during previous investigations.3 The temperature increase for system 2-A is smaller than that for system 1-A: the heat generated by isooctane combustion is almost fully used to sustain the SR reaction, faster than in system 1-A because of the higher reaction temperature level of system 2-A. The hydrogen production from the four examined systems during the start-up phase and transient conditions under step-changing loads is shown in Figure 8. System 1-A reaches steady-state mode after about 1000 s (a little more for system 2-A), whereas both systems 1-B and 2-B (operating with flue gases) need only about 400 s to achieve the steady-state hydrogen production rate of about 5 × 10-5 mol/s, practically identical for both of the systems, corresponding to an electric power production from the elementary unit of 5 We. During the start-up heating phase, because of the thermic inertia of the system, the temperatures are higher at the first part of the microchannel length and lower near its end; this situation promotes at the reactor beginning the SR reaction and at the end the WGS one, therefore justifying some hydrogen overproduction as shown in Figure 8. Starting from steady-state conditions, the SR inlet global molar flow rate (with the composition remaining constant) was first decreased by 20% and then, after settling, increased again up to its nominal value. Both systems showed a really good response under stepchanging loads: the new steady-state conditions were reached almost instantaneously, showing also a small undershoot in the deceleration step and a small overshoot in the acceleration step. It should be remarked,

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however, that the sharp productivity change is partly due to the relatively high gas-phase velocity, which makes the diffusion effects really negligible with respect to convective mass transfer. 6. Final Remarks In this work, the behavior of a gasoline steam reformer facing catalytic burner-reformer microchannels (system A) was analyzed by employing isooctane as a model molecule, compared with a similar reactor using flue gases from an external noncatalytic burner as the heat source (system B). Two different operating conditions, one requiring more expensive construction materials (condition 1) and another one characterized by comparatively lower operating temperatures (condition 2), were examined. The reactions in the burner and the reformer occurred heterogeneously on a Pt-CeO2 catalyst. The developed dynamic model, consisting of 25 PDEs for system A and 15 PDEs for system B, was numerically solved. The reforming and combustion kinetics were treated with the scheme proposed by Pacheco et al.1 The simulations showed that the burner-reformer microchannel reactor (system A) enables efficient heat transfer from the hot side to the cold side, thus proving to be a compact unit potentially suitable for automotive applications. The catalytic integrated burner-reformer microchannel operating under condition 2 showed to be the best design option because it leads to high reformer isooctane conversion with the lowest isooctane consumption at the burner side, even if mass-transfer limitations inside the burner catalytic layer are present, because of the low effective diffusion values inside the solid phase. For this reason, it was not possible to reach an “ignited” condition in the catalytic burner, as confirmed by the lack of a peak in the burner temperature profile. This low heat generation rate involved longer start-up periods but at the same time allowed one to preserve both catalyst activity and stability, reducing the hotspot generation probability. Thus, for system A, operating with a lower burner inlet temperature, a longer microchannel unit was needed to achieve the same high isooctane conversion obtained in the device operating with flue gases (system B) and also in the same system A working with a higher burner inlet temperature (condition 2). Anyway, this result suggests the need to develop effective burner catalysts with most of the active sites located at the external surface, so as to limit the negative effect of the internal mass transfer on the overall combustion rate. Moreover, the dynamic analysis showed a longer start-up time for system A, which could limit the applications of this integrated device for automotive applications unless tailored effective start-up procedures are conceived. Conversely, the catalytic-combustionintegrated microchannel reformer showed a really fast response to step load changes: hydrogen production reached almost instantaneously (i.e., in a few seconds) the new steady-state conditions, as was needed in automotive applications. Nomenclature Aperd ) exchange surface for heat losses Apm ) heat-exchange surface between the burner microchannel and steel wall av ) catalyst specific surface (3304 m2/m3)

cp ) specific heat (kJ/kg/K) D ) gas diffusivity (m2/s) Dh ) hydraulic diameter (m) ∆Hr ) reaction enthalpy (kJ/kmol) JN ) mass flux (kmol/m2/s) JQ ) heat flux (kJ/m2/s) k ) thermal conductivity (kW/m/s) Kov ) global mass-transfer coefficient (m/s) m ) mass (kg) P ) pressure (bar) R ) ideal gas constant (8.314 kJ/kmol/K) r ) reaction rate (kmol/s/kgcat) Rj ) generation rate of the jth component (kmol/s) t ) time (s) U ) global heat-transfer coefficient (kW/m2/K) v ) gas-phase velocity (m/s) V ) volume (m3) x ) space (m) yj ) molar fraction of the jth component (kmol/kmol) W ) microchannel width/length ratio (m/m) Greek Letters  ) catalyst void fraction (0.19) r ) metal wall emissivity (0.35) F ) density (kg/m3) σ ) Stephan-Boltzmann constant (5.67 × 10-11 kW/m2/ K4) τ ) catalyst tortuosity (2) Subscripts 0 ) initial value est ) surrounding environment I, II, III, ... ) chemical reaction number in ) inlet value j ) reactant index L ) outlet value Superscripts b ) burner gas phase cb ) burner catalyst phase cr ) reformer catalyst phase pm ) metal wall r ) reformer gas phase

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(10) Whyatt, G.; Fischer, C.; Davis, J. Rapid start fuel reforming systems and techniques. International Patent WO2004104140, 2004. (11) Ryu, J. H.; Choi, D. H.; Kim, S. J. Numerical optimization of the thermal performance of a microchannel heat sink. Int. J. Heat Mass Transfer 2002, 45, 2823. (12) Adams, T. M.; Dowling, M. F.; Abdel-Khalik, S. I.; Jeter, S. M. Applicability of traditional turbulent single-phase forced convection correlations to non-circular microchannel. Int. J. Heat Mass Transfer 1999, 42, 4411. (13) Adams, T. M.; Abdel-Khalik, S. I.; Jeter, S. M.; Qureshi, Z. H. An experimental investigation of single phase forced convection in microchannels. Int. J. Heat Mass Transfer 1998, 41, 851. (14) Xu, J.; Froment, G. Methane steam reforming, methanation and water-gas shift. Part IsIntrinsic kinetics. AIChE J. 1989, 35, 88.

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Received for review February 21, 2005 Revised manuscript received May 20, 2005 Accepted June 2, 2005 IE050215N