Ind. Eng. Chem. Res. 2005, 44, 759-768
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Dynamics of a Methanol Reformer for Automotive Applications A. Varesano, I. Guaglio, G. Saracco,* and P. L. Maffettone Dipartimento di Scienza dei Materiali ed Ingegneria Chimica, Politecnico di Torino, Corso Duca degli Abruzzi 24, I-10129 Torino, Italy
A model for a methanol reformer for automotive applications is presented. The steam reforming reactor is coupled with a burner that supplies heat. In the burner, methanol combustion takes place both homogeneously (in the gas phase) and heterogeneously (within a platinum catalyst layer deposited on the heat-exchange surface). The steam reformer is modeled as a pseudohomogeneous reactor. The reforming kinetics is validated by comparing theoretical results with experimental data found in the literature. The pressure inside both reactors is assumed to be constant, and the gas is in plug flow. The model consists of 19 partial differential equations and is solved numerically. An optimal reactor design is proposed to achieve maximum reactor compactness and methanol conversion both in the reformer and in the burner, as well as a temperature in the reformer always lower than 573K to avoid catalyst sintering. Furthermore, the reactor dynamics is analyzed for the sake of defining operability maps in terms of gas velocities of the feed streams, which are expected to support the choice of an optimal reactor control strategy to avoid methanol leakage and thermal “runaways”. Introduction A fuel cell (FC) is an electrochemical device that converts chemical energy into electrical energy. These systems are very efficient and environmentally clean; consequently, scientific and technological interest in them is rapidly growing. The implementation of these devices in automobiles is obviously attracting great interest as fuel cells fulfill restrictive low-emission regulations (e.g., refs 1-3). Automotive applications require fuel cells with high specific power, high efficiency, low costs, long durability, and compact dimensions. Proton-exchange-membrane fuel cells4 are typically implemented in cars because they are small and light, operate at relatively low temperature, and are operationally flexible. The required power supply can be obtained by stacking several elements (e.g., ref 5). A fuel cell is usually fed with hydrogen that can be either stocked or produced on board. The car manufacturers developing FC vehicles are considering a variety of fuel possibilities; hydrogen, in fact, can be produced by reforming various liquid fuels such as methanol, gasoline, or diesel oil. The on-board reformation of gasoline is inviting in that it would not require new distribution infrastructures. On the other hand, some manufacturers are pointing to methanol as a fuel because it is easier to stock and handle than hydrogen and is also more efficiently reformed than gasoline.6 Figure 1 shows how fuel-cell-powered vehicles based on either pure hydrogen or on-board-reformed methanol can outperform commonly adopted internal combustion engines in terms of tank-to-wheel efficiency. This holds especially in the medium-to-low power range were lightduty (passenger cars, vans) and heavy-duty (buses, trucks) vehicles are operated on average (see Figure 1). In the present paper, the reactor setup needed to reform methanol in automobile applications is studied. * To whom correspondence should be addressed. E-mail:
[email protected]. Tel.: +39-011-5644654. Fax: +39-0115644699.
Figure 1. Tank-to-wheel efficiencies of fuel cell based powertrains (hydrogen, methanol reformer) vs internal combustion engines (diesel, gasoline). Source: Ricardo Consulting Engineers Ltd., Shoreham-by-Sea, United Kingdom.
Catalytic steam reforming is a widespread application, and the interested reader can refer to Røstrup-Nielsen7 for a summary. Methanol reforming is an endothermic reaction that runs sufficiently fast at 250-300 °C on copper-based catalysts (e.g., ref 8). Usually, the reactors operate under pressure (3-15 bar). The required heat can be supplied either by using an external burner or by associating an exothermic oxidation reaction within the reformer (autothermal process). Although a kinetic scheme for the autothermal case is still lacking, the nonautothermal reactions have been extensively studied in the literature.9-14 Thus, we concentrate on the description of a nonautothermal methanol reformer. The typical configuration of a hydrogen production process based on a nonautothermic methanol steam reforming is sketched in Figure 2. A methanol-water stream is evaporated and then fed to the reformer. A burner coupled with the reformer supplies the required heat. The output stream is rich in CO and has to be purified because CO poisons the FC catalyst. Usually, this purification step is performed with a water-gas shift (WGS) reaction and/or the preferential CO oxida-
10.1021/ie0496378 CCC: $30.25 © 2005 American Chemical Society Published on Web 01/21/2005
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Figure 2. Schematic of a methanol steam reformer plant for automotive applications.
tion (PROX), and the purified stream is finally fed to the FC. The exhausted hydrogen exiting the cell is usually recycled to the burner. In the present paper, a burner/reformer reactor is modeled and numerically simulated. The aim of this work is to develop a dynamic model to characterize the transient response of these devices under changing loads. Steam reforming reactors for automotive applications are, in fact, typically operated under rapidly changing operating conditions,5,8,15-20 and thus, a dynamic description of the process could be of value in both design and control. The reactor design is performed here by maximizing hydrogen productivity under different operating conditions preventing catalyst from thermal deactivation. Indeed, economic considerations impose a high methanolto-hydrogen conversion, but CuO/ZnO/Al2O3 catalysts must be operated below 573 K (e.g., ref 15) to prevent deactivation. Thus, a safely designed operation must fulfill this constraint. As will be shown, the requirement of high productivity combined with controlled reformer temperature impose a well-defined ratio between the burner feed and the reformer feed, which could, in principle, be varied to ensure the instantaneous FC hydrogen demand. In the following sections, the model is first introduced, then validated, and finally used to define design and control strategies for the reactor setup. Model Description The model is constructed assuming a plate/plate geometry.21 The reformer is treated as a fixed-bed pseudohomogeneous reactor, whereas the burner is a catalytic combustor consisting of a catalytic porous layer adhering on the plate. The model has been constructed to qualitatively describe the reactor behavior, and thus several simplifications have been considered to avoid cumbersome analysis. Figure 3 shows a schematic of the elemental cell of the burner/reformer reactor. The actual device consists of a parallel set of equal elemental cells. The geometry is plate/plate, and a metallic wall separates the burner from the reformer. A similar configuration was first proposed by Jenkins.23 The reformer is a fixed bed and is treated as consisting of a single phase (pseudohomogeneous approach, e.g., refs 24, 25). This assumption is valid under the hypothesis that mass- and heat-transfer resistances within the catalyst and at the gas-solid interface are negligible. Internal mass- and heat-transfer resistances can be
Figure 3. Elemental cell of the burner/reformer setup. The coordinate system is reported in the figure. The length scales are distorted for the sake of clarity. The calculations were carried out by assuming the following dimensions: thickness of the metallic wall ) 10-3 m, thickness of the burner catalyst layer ) 5 × 10-3 m, length ) 0.5 m, width ) 0.2 m.
neglected if the catalytically active components are mostly deposited on the outer surface of the catalyst pellets, which is generally the case for reforming catalysts. The use of highly active catalysts increases the importance of the transport resistances between the solid and the fluid and, hence, might invalidate the applicability of the pseudohomogeneous models. Disregarding intraparticle profiles might lead to some underestimation of the maximum temperature in the reactor. A criterion for determining the onset of interphase heat-transfer limitation was derived by Mears26 for the Arrhenius type of reaction rate dependency on the temperature and under the assumptions of negligible direct thermal conduction between spherical particles and negligible interphase mass-transfer resistance. The criterion states that the actual reaction rate deviates less than 5% from the reaction rate calculated assuming identical solid-phase and bulk-fluid conditions if the following inequality is satisfied
|
|
RTR ∆HRrRrp < 0.15 U3T ER
where ∆HR, rR, and ER are the overall reaction enthalpy, rate, and activation energy, respectively, in the steam reforming section (to be evaluated on the grounds of the overall set of kinetics expressions and enthalpies listed in the next section for the various reactions of interest). Thorough calculations showed that, in the present investigation, the Mears criterion is satisfied at any operating conditions, provided the catalyst pellet diameter is lower than 2 mm, which is quite reasonable owing to the small thickness of the reformer (30 mm for the optimum design as later demonstrated). Conversely, in the burner, the catalyst is deposited as a thin washcoat layer (10-50 µm thick) on the wall. The flow is laminar with Reynolds numbers in the range of 10-1000. At the relatively high feed temperature adopted in the present investigation, the reactor is in the mass-transfer-controlled regime, and the ignition point is close to the inlet. The burner is therefore modeled as a heterogeneous reactor consisting of two phases: a solid phase at the wall where catalyst is present and a gas phase in the bulk. Gases are treated as ideal and flow co-currently in the burner and in the reformer. Plug flow is assumed everywhere. The system is considered as one-dimensional: only the longitudinal coordinate plays a role. Both the burner and the reformer are operated under constant, and in general different, pressures. Specific heats have been assumed
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to be temperature-independent. Longitudinal diffusion in the gas phases has been neglected with respect to convective flow. Longitudinal thermal diffusion takes place in the solid catalyst in the burner, as well as in the metallic plate separating the reformer from the burner. Mass- and heat-transfer coefficients are taken as constants, independent of axial position or temperature. Considering the small dimensions involved, the asymptotic values of the transport coefficients are appropriate,27 and thus, for the characteristic Peclet number, it turns out that a reasonable estimate for the masstransfer coefficient at the burner side is 0.00264 m/s, whereas for heat transfer, the value 45 W/(m2 K) was adopted. Conversely, at the reformer side, a wall-to-bed laminar heat-exchange coefficient was assumed equal to 470 W/(m2 K), in line with previous investigations by Trimm and co-workers.8 Finally, the heat-exchange coefficient between the burner catalyst layer and the metal plate was fixed at 10000 W/(m2 K). On the basis of these assumptions, the temperature difference between the reformer catalyst and the metal plate was always lower than 60 °C, which should affect the predictions based on the plug-flow model assumption to a minor though nonnegligible extent. Reaction Kinetics: The Reformer. There are conflicting views in the literature concerning the reaction mechanism for methanol steam reforming.9 The steam reforming reactions on CuO/ZnO/Al2O3 catalyst have been modeled by referring to three overall reactions k1
CH3OH + H2O {\ } CO2 + 3H2 k -1
k2
} CO + 2H2 CH3OH {\ k -2
k3
CO + H2O {\ } CO2 + H2 k-3
(1)
(2)
r2 ) k2
[
CR,2,0 -(E1/RTR) e CR,1,0 Fcat(1 - R) PR D1 Patm
k1 )
k2 )
A2e-(E2/RTR) F (1 - R) PR D2 cat Patm
( )
(4)
Jiang et al.11,12 suggested a kinetic scheme that entails negligible CO production. Their mechanism is suitable for low steam/methanol ratios that are not of interest here. Recently, Peppley et al.13,14 improved their model by accounting for all of the stages (eqs 1-3). That reaction mechanism is capable of describing the kinetics in a wide pressure range (1-35 bar). For comparison, spot calculations were performed with this scheme, but the predictions compare with the available experimental data less satisfactorily at a significantly higher computational cost. Hence, this kinetics was not considered further. Reaction Kinetics: The Burner. The burner is a catalytic (platinum) combustor. Only CH3OH and air are fed to the burner. We assume that the combustion follows the kinetics proposed by Ito et al.28 Indeed, to our knowledge, this is the only kinetic scheme on methanol catalytic combustion available in the literature. Two different reaction mechanisms are considered: one deals with the homogeneous combustion in the gas phase, and the other deals with the heterogeneous combustion in the catalytic wall. The overall reaction and reaction rates are k4,s,k4,g 3 CH3OH + O2 98 CO2 + 2H2O 2
r4,s ) k4,sCB,s,1 r4,g ) k4,gCB,g,1
(3)
Equation 1 is the steam reforming reaction, eq 2 is the methanol decomposition reaction, and eq 3 is the WGS reaction. Several schemes are reported in the literature dealing with methanol reforming on CuO/ZnO/Al2O3. Santacesaria and Carra`10 modeled the kinetics with a phenomenological Langmuir-Hinshelwood approach by assuming that only stages 2 and 3 are relevant; more precisely, the methanol decomposition was considered as irreversible and the WGS at equilibrium. Amphlett and co-workers studied the reaction mechanism at atmospheric pressure according to the approach of Santacesaria and Carra`. They also proposed a scheme for relatively high pressures15 that assumes a two-stage mechanism consisting of the steam reforming reaction (eq 1) and the methanol decomposition reaction (eq 2). In the present work, this mechanism is implemented because, as shown in a subsequent section, it proved capable of describing available experimental data5,17 measured under the operating conditions of interest here. The proposed kinetic expressions are
r1 ) -k1CR,1
( )] ( )
A1 + B1 ln
(5)
where k4,s refers to the heterogeneous reaction and k4,g to the homogeneous one. The two constants are given by
(
k4,s ) exp
(
k4,s ) exp
-5.9 × 104 + 22.45 RTB,s
)
-6.4 × 103 + 5.8 RTB,s
k4,g ) exp
(
)
for TB,s < 380 K
for 380 < TB,s < 830 K
)
-6.1 × 103 + 3.2 RTB,g
(6)
where TB,s and TB,g are the temperatures of the solid phase and of the gas, respectively. Balances. The following chemical compounds are present in the reformer: CH3OH, H2O, H2, CO, and CO2. The burner contains CH3OH, O2, N2, CO2, and H2O. In the model equations, the reactants are indicated by an index, as detailed in Table 1. The model equations are reported in Table 2. Mass balances are written for each chemical component in the three units: the reformer, the solid phase in the burner,
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Table 1. Reactants and Stoichiometric Coefficients reaction reactant
index (j)
1
2
4
CH3OH H2O CO2 CO H2 O2 N2
1 2 3 4 5 6 7
-1 -1 +1 0 +3 0 -
-1 0 0 +1 +2 0 -
-1 +2 +1 0 0 -1.5 -
and the gas phase in the burner. Heat balances are written for the same units and for the separating wall. Because gases are treated as ideal, the local volumetric flow rate depends on both the local temperature and the local gas-phase average molar mass. As already mentioned, the pressures in both reactors are assumed to be constant (negligible pressure drops). The average molar mass has been treated as a pseudo-steady-state quantity, and thus, it appears as a coefficient of the concentration time derivative but its value is continuously updated. In the heat equations, the gas specific heat does not depend on temperature and composition, whereas in the reformer heat balance, only the catalyst heat capacity, which is assumed to be constant, is considered. Moreover, an overall constant heat-exchange coefficient quantifies the heat lost to the surrounding environment. These assumptions, made for the sake of simplicity, should bring about errors that are significantly lower than the uncertainties in the kinetics expressions, as later discussed. The model consists of 19 partial differential equations (PDEs). The following 20 boundary conditions are considered
CR,j(0,t) ) CR,j,in(t)
for j ) 1-5
TR(0,t) ) TR,in(t) CB,g,j(0,t) ) CB,g,j,in(t)
for j ) 1-3, 6, 7
TB,g(0,t) ) TB,g,in(t)
| | | |
∂TB,s(x,t) ∂x
x)0
∂TB,s(x,t) ∂x
x)L
∂TW(x,t) ∂x
x)0
∂TW(x,t) ∂x
x)L
(15) (16) (17) (18)
)0
(19)
)0
(20)
Figure 4. Comparison of experimental data of Du¨sterwald et al.17 and model predictions in isothermal conditions. Upper graph: methanol conversion vs theoretical hydrogen productivity. Lower graph: CO selectivity vs theoretical hydrogen productivity. The lines are the theoretical predictions. The symbols are the experimental data. Solid lines and b, 533 K; dashed lines and O, 553 K; dot-dashed lines and 1, 573 K.
The set of 19 PDEs was integrated with two different approaches: a simple upwinded semiimplicit finitedifference algorithm and the method of lines based on Matlab ODE solvers. The results of the two codes perfectly overlap. The results presented below were obtained with a spatial discretization of 50 points as this value ensures numerical convergence and satisfactory accuracy. Reaction Mechanism Validation
)0
(21)
)0
(22)
No adjustable parameter is present in the model. Parameter values are reported in the Nomenclature section with reference to literature sources. It should be mentioned that the specific heat of the burner catalyst is assumed to be equal to that of Al2O3 as the presence of platinum in its structure is negligible. Finally, the specific heat of the reformer catalyst has also been assumed to be that of Al2O3 even though the quantity of other components is not negligible in this case. Reaction enthalpies are considered to be temperature-independent.
The assumed kinetics of methanol reforming reactions were qualitatively validated with reference to the experimental data proposed by Du¨sterwald et al.17 and Peters et al.,5 who examined the performance of different catalysts for fuel cell reformers. Their experiments were carried out at constant temperature under both transient and steady-state conditions. Conversion and CO selectivity were measured at steady state under different operating conditions. The experiments of Du¨sterwald et al.17 and Peters et al.5 were performed such that the heat needed for the endothermic reaction was provided by condensing steam. The experiments were devised so as to keep the reformer temperature at a constant value. The conversion and CO selectivity data were plotted versus theoretical specific (per unit mass of catalyst) hydrogen productivity. This quantity depends on the feed conditions. In such a case, the model reduces to eq 10 only as no energy balance is needed and the burner is absent. As a consequence, the param-
Ind. Eng. Chem. Res., Vol. 44, No. 4, 2005 763 Table 2. Balance Equations burner gas
burner solid reformer
Mass Balances ∂CB,g,j ∂(TB,gCB,g,j ) ) -ψB - τB(CB,g,j - CB,s,j) + G4,jr4,g ∂t ∂x
j ) 1-3, 6
(7)
∂CB,g,j ∂(TB,gCB,g,7) ) -ψB ∂t ∂x ∂CB,s,j ) ξh(CB,g,j - CB,s,j) + G4,jr4,g ∂t
j)7
(8)
j ) 1-3 6, 7
(9)
∂CR,j ∂(TRCR,j) G1,jr1 + G2,jr2 ) -ψR + ∂t ∂x R Heat Balances
j ) 1-5
(10)
burner gas ∂TB,g ∂TB,g ) -TB,gψB + ∂t ∂x burner solid ∂TB,s ) ∂t wall
Kcat,B
∂x2
∂2 T W
(11)
2fU2 U1 + Cgas + (T - TB,s) + r4,s∆H4 h sB(1 - B) W Ccat,B
(12)
(
+ ξh(TB,g - TB,s)
)
πe U1 + Cgas - (TB,g - Te) + r4,g∆H4 h B cjP,B,0FjB
)
(13)
f [U (T - TW) + U3(TR - TW)] - πe(TW - Te) ∂TW sW 2 B,s ∂x ) ∂t CW PR ∂TR 2fU3 (TR - TW) - πe(TR - Te) + r1∆H1 + r2∆H2 -ψR cjP,R,0R ∂TR R ∂x sR ) ∂t Ccat,R(1 - R) KW
reformer
∂2TB,s
(
-τB(TB,g - TB,s)
2
+
eters to be fixed are those appearing in eq 10: the reformer geometry and the operating conditions (vR,0, T, P). These data were determined on the basis of the experimental operating conditions reported in those papers. Model predictions and experimental data by Du¨sterwald et al.17 are shown in Figure 4. Figure 4 shows the methanol conversion and the CO selectivity versus the theoretical hydrogen production at different reformer temperatures. The agreement is generally good, and the same behavior was found with data taken from Peters et al.5 At low temperature, the conversion predictions behave better than selectivity ones, whereas the model predicts the CO selectivity very well at the highest temperature. Some spot calculations were performed with the kinetic model by Santacesaria and Carra` presented in ref 10, and the results were generally worse. As a result of the above considerations, one must conclude that the complex nature of the catalyst and the reaction process is the main reason for the remarkable differences of the various validated kinetic models published in the literature for a formally equivalent catalyst. The selected kinetics scheme by Amphlett et al.15 obviously cannot describe perfectly any set of experimental data published by others; however, it was found to be the one with the widest applicability among those considered. The reader should be aware that the conclusions drawn in the present work are strongly dependent on the validity of the model, which should carefully be checked on a case-by-case basis for the specific CuO/ZnO/Al2O3 catalysts of interest. Burner/Reformer Configuration In view of the qualitative and, to some extent, quantitative agreement with experimental data reported in the previous section, the reforming kinetics was implemented to describe typical nonisothermal operating conditions of the burner/reformer reactor.
(14)
Table 3. Feed Conditionsa
a
reactant
reformer
burner gas
CCH3OH CH2O CCO2 CCO CH2 CO2 CN2
0.0325 0.0487 0 0 0 -
0.0054 0 0 0 0.0102 0.0385
T ) 450 K for both the reformer and burner gas.
From now on, all parameter values are those reported in the Nomenclature section, and the feed conditions are those reported in Table 3. These feed conditions were derived from the specific investigations by Trimm and co-workers8 on the role of the feed composition in the specific productivity of hydrogen for the sake of maximizing the latter. A 1.5 H2O/CH3OH ratio was also selected as the best reformer feed composition by other researchers.5,20 Furthermore, feed temperatures lower than 450 K will possibly entail too low reforming kinetics,8 especially in the long term, when some catalyst deactivation is to be expected. The reformer/burner system is geometrically designed to optimize temperatures and conversions. In this regard, several configurations were analyzed in order to determine the optimal solution. Two geometrical parameters were investigated: the reformer and the burner height. The other dimensions were considered as fixed (reactor width, 0.2 m; reactor length, 0.5 m). For each height arrangement, the reformer feed velocity and the burner feed velocity were varied. Table 4 reports significant data for a selection of burner/reformer geometries; lines in italics refer to conditions with reformer temperatures exceeding the sintering limit (573 K). The acceptable methanol conversion is chosen to give a residual presence of methanol in the output stream of less than 6000 ppm as tolerated by the FC. The desired hydrogen productivity is set to 7.9 Nm3/h; if we suppose that two identical burner/reformer units
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Table 4. Performance Indicators of Reformer/Burner Configurationa SB (m)
VR (m/s)
VB (m/s)
0.002 0.003 0.004 0.005 0.0035 0.0045 0.005 0.003 0.0035 0.0045 0.005 0.0025 0.0033 0.004
0.5 0.5 0.5 0.5 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0 1.0
0.5 0.5 0.5 0.5 0.7 0.7 0.7 1.0 1.0 1.0 1.0 1.5 1.5 1.5
0.01 0.011 0.012 0.009 0.01 0.011 0.008 0.009 0.01 0.01167 0.013 0.0133 0.0076 0.008 0.0085 0.009 0.00533 0.006 0.0067 0.0068
0.3 0.3 0.3 0.3 0.3 0.3 0.3 0.3 0.3 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.4
0.25 0.25 0.25 0.3 0.3 0.3 0.35 0.35 0.35 0.3 0.3 0.3 0.5 0.5 0.5 0.5 0.75 0.75 0.75 0.75
0.012 0.014 0.016 0.0085 0.01 0.011 0.012 0.01 0.012 0.014 0.02 0.022 0.012 0.014 0.016 0.0085 0.01 0.011 0.012 0.0066 0.0072
0.25 0.25 0.25 0.25 0.25 0.25 0.25 0.25 0.25 0.25 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.4 0.5 0.5
0.25 0.25 0.25 0.4 0.4 0.4 0.4 0.3 0.3 0.3 0.25 0.25 0.4 0.4 0.4 0.6 0.6 0.6 0.6 0.3 0.3
0.01 0.02 0.022 0.024 0.026 0.028 0.01 0.015 0.02 0.015 0.02 0.025 0.020 0.024 0.014 0.016 0.018 0.020 0.022 0.024
0.1 0.1 0.15 0.15 0.15 0.15 0.2 0.2 0.2 0.2 0.2 0.2 0.25 0.25 0.25 0.25 0.25 0.3 0.3 0.3
0.3 0.3 0.15 0.15 0.15 0.15 0.4 0.4 0.4 0.3 0.3 0.3 0.25 0.25 0.4 0.4 0.4 0.3 0.3 0.3
reference methanol conversion (%) SR ) 0.01 m 39 62 78 89 56 68 73 60 68 80 84 58 76 84 SR ) 0.02 m 87 90 96 89 93 96 89 94 98 89 93 94 89 91 94 96 88 93 96 97 SR ) 0.03 m 86 93 97 89 97 99 100 85 92 97 88 94 85 91 96 85 91 95 98 82 93 SR ) 0.05 m 91 94 91 94 97 99 84 99 100 90 98 100 87 93 90 95 98 86 90 93
Tmax (K)
Yh (Nm3/h)
538 554 567 577 560 569 574 558 564 574 579 552 567 577
2.14 3.42 4.29 4.9 6.22 7.56 8.11 4.63 7.5 8.81 9.31 6.36 8.41 9.34
566 571 580 567 572 578 565 572 579 572 578 579 566 569 572 576 559 566 573 576
5.63 5.87 6.21 5.78 6.03 6.22 5.83 6.1 6.31 7.83 8.18 8.27 7.83 8.01 8.27 8.45 7.74 8.18 8.45 8.54
557 564 572 555 564 571 618 554 564 573 567 571 559 566 573 555 563 568 582 564 574
7.13 7.64 8.02 7.38 7.99 8.16 8.25 7.07 7.68 8.11 11.72 12.41 11.27 12.06 12.67 11.26 12.09 12.52 12.88 12.56 13.29
538 585 552 556 561 566 541 567 684 550 563 713 551 558 550 556 565 552 555 559
5.4 5.3 7.55 7.8 8.02 8.18 9.2 10.6 10.8 9.8 10.5 10.8 11.98 12.79 12.39 13.04 13.52 14.28 14.86 15.35
a Conditions entailing maximum reforming temperature limit of 573 K are in italics.
Figure 5. Steady-state temperature profiles for the optimal reformer/burner setup.
are stacked, the total productivity will be about 15.8 Nm3/h. This value corresponds to an electric power (of the cell) of about 60 kW. Configurations with sR ) 0.01 m give an acceptable hydrogen productivity (∼7.9 Nm3/ h) with a rather low methanol conversion. Configurations with sR ) 0.05 m provide an exceedingly high hydrogen yield when operated with typical feed speeds (0.2-0.5 m/s), whereas an optimal productivity is obtained when the reformer speed is lowered to 0.15 m/s. This configuration, however, is the least compact among those considered. The “optimal” geometry is considered the one that guarantees the highest methanol conversion with the lowest maximum temperature in the reformer and the most compact geometry (sR ) 0.03 m, sB ) 0.011 m). All subsequent calculations were carried out with this geometry. Figure 5 shows the steady-state temperature profiles in the system. The temperature of the gas phase in the burner quickly rises to its maximum value, which is attained in 0.08 m, and then slowly decreases in the rest of the reactor, reaching 556 K at the burner exit. The solid phases in the burner and the wall are always almost at thermal equilibrium. The reformer temperature is always below 573 K, which is the maximum temperature allowed to avoid catalyst thermal sintering.16 The reformer temperature is approximately constant in a large zone (from 0.1 to 0.3 m) because of the decrease of both the heat flow from the burner and the heat subtraction for the endothermic reforming reaction. The former is due, in fact, to the lowering of the burner temperature caused by methanol consumption; the latter is related to the progressive conversion of the methanol. From a practical point of view, it is worth noting that the exhaust gases exit at a temperature of about 537 K, a reasonable temperature if successive treatments are considered (e.g., WGS). A final comment is in order. The pseudohomogeneous assumption, of course, implies that the reformer temperature does not depend on any transverse coordinate. Now, the metallic wall temperature is higher than the highest allowed temperature in the reformer catalyst, and as a consequence, the catalyst in contact with the wall is obviously thermally degraded. However, in view of the qualitative approach of this work, it is assumed that the amount of catalyst suffering damage is negligible when the pseudohomogeneous temperature is below 573 K. Figure 6 shows the steady-state concentration profiles within the reformer at the same conditions as in Figure 5. Methanol rapidly decreases along the reactor: the
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Figure 6. Steady-state cencentration profiles in the reformer for the optimal reformer/burner configuration.
conversion is already close to completion in the first part of the reformer (0.3 m). In the last 0.2 m of the reformer, the methanol conversion passes from 95% to 98.5%. In this way, the discharged stream has a methanol concentration low enough to be fed to the FC, e6000 ppm.17 The FC catalyst, in fact, tolerates very low concentrations of CO (
