Heat Supply and Hydrogen Yield in an Ethanol Microreformer

ARTICLE pubs.acs.org/IECR

Heat Supply and Hydrogen Yield in an Ethanol Microreformer Andres M. Anzola,*,† Yanina M. Bruschi,† Eduardo Lopez,†,‡ Noemí S. Schbib,† Marisa N. Pedernera,† and Daniel O. Borio† † ‡

PLAPIQUI (UNS-CONICET), Camino La Carrindanga, km. 7, 8000 Bahía Blanca, Argentina Institut de Tecniques Energetiques, Universitat Politecnica de Catalunya, Avenida Diagonal 647, Ed. ETSEIB, 08028 Barcelona, Spain ABSTRACT: A simulation study of a microreactor for steam reforming of ethanol (ESR) over a washcoated Pd catalyst is reported in the present manuscript. Both co- and countercurrent configurations are considered for the streams of reactants and flue gas through microchannels of a square section. Two contiguous channels, assumed as representative of the whole reactor behavior, are simulated using a 1D pseudohomogeneous model. The microreactor performance is analyzed for different heat-supply strategies and channel sizes. The results demonstrate that the ESR in the microreactor is strongly controlled by the heat supply. The hydrogen yield has a marked dependence on the total amount of heat transferred from the flue gas to the reformate and the axial distribution of the heat flow. The cocurrent-flow configuration proved to yield the higher performance. While smaller channels work better for cocurrent operation, an optimum compromise value is found for the channel size in countercurrent flow.

1. INTRODUCTION The production of bioethanol by fermentation of sugar cane, corn, or other agricultural waste is an attractive alternative for the production of hydrogen or synthesis gas because of its high energy density. Bioethanol is easier to store, handle, and transport in a safe way than other fuels because of its lower toxicity and volatility. Some attention has been paid lately to processes for the production of hydrogen or synthesis gas from bioethanol by means of catalytic ethanol steam reforming (ESR).1-3 Several mechanisms and kinetics are proposed for ESR for different catalysts. The most common catalysts for steam-reforming processes are Ni-based. Several kinetic models have been proposed for ESR on a Ni catalyst.4-6 The Ni catalysts were also evaluated for ESR by means of both experimental and simulation studies.7-10 Simson et al.11 reported a kinetic and experimental study for ESR using a Rh/Pt washcoated monolith catalyst. Other studies have been performed for ESR on Pt/Ni,12 Co/Al2O3,13 CeZr-Co oxide catalysts14 and Ru catalysts.15 Lopez et al.16 carried out the kinetic modeling of ESR on a washcoated Pd-based catalyst. A set of three catalytic reactions were considered: ethanol reforming, methane reforming, and water gas shift. ESR on a Pd-based catalyst was also studied by Goula et al.17 and Scott et al.18 Several types of reactors can be used for hydrogen production, depending on the process scale, the process constraints, and the availability of raw materials. The steam reforming of natural gas or heavier hydrocarbons is industrially carried out in tubular reformers. Other designs, such as autothermal reformers, can be used in other large-scale applications such as methanol production or Fischer-Tropsch synthesis.19 Multichannel catalytic reactors have been proposed as an alternative to performing highly endothermic processes.20 In particular, microreactors provide a high surface-to-volume ratio and high heat- and mass-transfer coefficients, which make them appropriate for processes requiring high heat fluxes per unit area. r 2010 American Chemical Society

In the last years, microreactors have received considerable attention in the literature.21,22 Because of the small geometry and the high heat conductivity of the materials, laboratory-scale microreactors can be electrically heated and isothermal conditions can usually be assumed.23 However, under higher production-scale conditions, the heat supply should be provided through other sources, e.g., catalytic combustion in contiguous microchannels20,24 or convective heating using flue gas coming from an external combustion chamber.25 This work uses this last option by considering a combustion chamber upstream of the reactor. The generated heat in the chamber can not only be used to provide the reaction heat but also for earlier steps such as the evaporation and superheating of the feed stream up to the reaction temperature.9,26 The process can be improved by using ethanol as a fuel and avoiding the use of alternative fuels.9 The nonisothermal behavior of the ESR microreactor is simulated, and different flow configurations between the process gas and heating medium are analyzed.

2. MATHEMATICAL MODEL Figure 1 shows schemes of the microreactor under study.27 As shown, adjacent foils with square microchannels are used to circulate the reactive mixture (C2H5OH þ H2O) and a flue gas stream coming from an external combustion chamber, which supplies the heat required for the endothermic reforming reactions. Two flow configurations are analyzed: co- and countercurrent schemes. A Pd-based catalyst is assumed to be coated on Special Issue: IMCCRE 2010 Received: March 26, 2010 Accepted: July 15, 2010 Revised: July 14, 2010 Published: August 16, 2010 2698

dx.doi.org/10.1021/ie100739m | Ind. Eng. Chem. Res. 2011, 50, 2698–2705

Industrial & Engineering Chemistry Research

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Table 1. Scheme of Reactions (1)-(3) C2 H5 OHþH2 OfCH4 þCO2 þ2H2

ΔH R;2 ¼ 205:8 kJ=mol

ð2Þ

ΔH R;3 ¼ -41:17 kJ = mol

ð3Þ

CH4 þH2 O T COþ3H2 COþH2 OTCO2 þH2

ΔH R;1 ¼ 8:73 kJ=mol ð1Þ

the metallic microchannels. The system of reactions is presented in Table 1. Reaction (1) is assumed to be irreversible, whereas reactions (2) and (3) are equilibrium-limited.16 A 1D pseudohomogeneous model is selected to simulate steady-state operation of the microreactor, subject to the following assumptions: • Isobaric conditions: laminar flow through short channels without pellets ensures a low pressure drop. • Heat losses from the microreactor to the environment are neglected (the reactor is assumed to be properly isolated). • Axial dispersion phenomena are neglected, considering the high linear velocities selected. • The use of small channels supports the assumption of flat mass and temperature radial profiles. Temperature and composition gradients in the cross section are not considered. • Uniform flow distribution in all of the microchannels (a proper flow distributor is assumed22). • Two contiguous channels (reactants and flue gas) are modeled as representative of the whole microreactor. From these assumptions, the governing equations for the reaction and flue gas streams are given below:

Figure 1. Schemes of the ESR microreactor: (a) cocurrent; (b) countercurrent.

for i = 1-3 (reactions of Table 1) and j = Et, H2O, CH4, CO2, CO, H2. Flue Gas Side Energy Balance AT, H ½UaH ðTH -TÞ dTH ¼ ð(Þ N PH dz FH, j CPH, j

Reaction Side Mass Balances

i¼1

dFEt ¼ AT ð-r1 Þ dz

ð4Þ

dFH2 O ¼ AT ð-r1 -r2 -r3 Þ dz

ð5Þ

The signs - and þ in eq 11 correspond to cocurrent and countercurrent arrangements, respectively. Boundary Conditions

dFCO ¼ AT ðr2 -r3 Þ dz

ð6Þ

dFCO2 ¼ AT ðr1 þr3 Þ dz

ð7Þ

dFH2 ¼ AT ð2r1 þ3r2 þr3 Þ dz

ð8Þ

dFCH4 ¼ AT ðr1 -r2 Þ dz

ð9Þ

8 > < Fj ¼ Fj, in z ¼ 0 T ¼ Tin > : TH ¼ TH, in Countercurrent:

( z ¼0

Fj ¼ Fj, in T ¼ Tin

z ¼ L fTH ¼ TH, in 3 P AT ½ ðri Þð-ΔHR , i ÞþUaðTH -TÞ i¼1

N P j¼1

Fj Cp, j

ð11Þ

Cocurrent:

Energy Balance dT ¼ dz

for j ¼ H2 O, CO2 , O2 , N2

ð12Þ

ð13Þ

ð14Þ

for j ¼ Et, H2 O, CH4 , CO2 , CO, H2 .

ð10Þ

The symbols Fj and FH,j represent the molar flows of species j per channel, flowing on the reactants and flue gas sides, respectively. The intrinsic power law kinetic expressions (r1, r2, and r3) reported by Lopez et al.16 have been adopted in the simulations.

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Table 2. Geometrical Parameters channel length, L23 channel width = height, b23

0.08 m 200 μm

fin width, e23

100 μm

channels per foil

200

number of channels (reactant side), NC height of the stack

31 600 9.48 cm

thickness of the washcoated catalyst, wc23

1 μm

coated surface of the stack

1.51 m2

Table 3. Operating Conditions feed pressure, P0 total molar feed flow rate, Fin

0.1 MPa 10.24 Nm3/h

steam-to-carbon molar ratio, S/C total molar flue gas flow rate, FH,in

3 14.51 Nm3/h

Figure 2. Outlet ethanol conversion and reactant temperature as affected by the heat duty. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C; b = 200 μm. Cocurrent scheme.

flue gas molar fractions, % CO2

10.43

H2O

15.78

O2

3.05

N2

70.74

inlet temperature on the reaction side, Tin

600 °C

inlet temperature on the flue gas side, TH,in

600-1000 °C

The kinetic parameters were obtained for a steam-to-carbon ratio in the range of 2-4 and temperatures lower than 800 °C. The thermodynamic properties of the components are extracted from the literature.28,29 The global heat-transfer coefficient is calculated as follows: 1 1 e -1 -1 þ U ¼ þ ð15Þ h hH k where e/k f 0 because of the small thickness (100 μm) and high thermal conductivity of the metallic wall. The heat-transfer convective coefficients of each side are obtained from the Nusselt expression applicable for square-channel-structured reactors proposed by Cybulski et al.:30,31 hdh b 0:45 ¼ 2:978 1 þ 0:095Re Pr Nu ¼ ð16Þ L λ The reference size of the microchannels (b = 200 μm) and the washcoat thickness (wc = 1 μm) are based on those proposed by G€orke et al.23 Similar washcoat thicknesses have been reported elsewhere.21,25,32 For the base case, a total number of 31 600 channels are selected for both the reactants (158 foils of 200 channels) and the flue gas in order to obtain a hydrogen production typical for a 10 kWth fuel cell (FH2 = 3.36 Nm3/h). The geometric parameters and operating conditions used in the simulations are given in Tables 2 and 3, respectively. The feed corresponds to one typical of a ethanol steam reformer.2,6,23 Equations (4)-(11) were integrated by means of a Gear algorithm. The cocurrent scheme constitutes an initial value problem, while the countercurrent arrangement defines a boundary-value problem that was solved using the shooting method.

3. RESULTS AND DISCUSSION To analyze the influence of the heat supply on the reactor performance, feed conditions of the process stream are kept

constant. The total flow of ethanol fed to the microreactor (FEt,in = 1.46 Nm3/h = 65.33 mol/h) is selected so that the desired hydrogen production (FH2 = 3.36 Nm3/h) can be reached for a hydrogen yield ηH2∼ 2.30. Regarding the heating medium, both the total molar flow rate (FH,in) and the inlet temperature (TH,in) are varied. First, the behavior of the microreactor operating under a cocurrent-flow configuration between reactants and flue gas streams is studied. 3.1. Cocurrent-Flow Configuration. 3.1.1. Heat Supply and Hydrogen Production. Figure 2 shows the outlet ethanol conversion (XEt) and the outlet reactants temperature as a function of the total heat transferred from the flue gas stream to the process gas; i.e., Q = FHCp,H(TH,in - TH,out). In order to vary the total heat duty, two operating modes on the flue gas side are considered. Thus, curve A is obtained for a constant flue gas flow rate, FH = 14.52 Nm3/h, by changing the inlet temperature (TH,in) from 600 to 1000 °C. The results of curve B, conversely, correspond to a constant inlet temperature of the heating medium (TH,in = 800 °C) and a flow rate range of 0.8 < FH < 55.64 Nm3/h. Point 1 represents an operating condition for which a typical production for a 10 kWth fuel cell is attained (i.e., FH2 = 3.36 Nm3/h). Figure 3 shows the hydrogen yield (ηH2) and the CH4 yield (ηCH4) for the same operating conditions as those in Figure 2. As in other steam-reforming processes, the reactant conversion and hydrogen yield increase strongly with the heat supply. As shown in Figure 2, for Q > 1.4 kW, the conversion values are slightly higher for curve A, which reaches almost complete conversion for Q ∼ 2 kW. Contrarily, the outlet temperatures corresponding to curve A are lower than those of curve B, for Q > 1.4 kW. The behavior of the methane yield is not monotonic with respect to the heat supply. Both curves of ηCH4 (A and B) increase up to approximately the same point because of the predominance of reaction (1) over reaction (2). For Q ∼ 1.2 kW, the methane yield is a maximum. Beyond this condition, however, curve A drops more sharply than curve B because the equilibrium of reaction (2) is favored. Consequently, for Q > 1.4 kW, the high temperatures of the flue gas stream (curve A) lead to an improvement in the hydrogen yield for two main reasons: higher ethanol conversions and, in particular, lower CH4 contents at the reactor outlet. Because curves A and B shown in Figures 2 and 3 are not coincident, it is clear that the performance of the cocurrent microreactor is defined not only by the amount of transferred 2700



Figure 3. Hydrogen and methane yields as affected by the heat duty. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C; b = 200 μm. Cocurrent scheme.

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Figure 5. Ethanol conversion vs microchannel width (b) for different flue gas inlet temperatures. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C. Cocurrent scheme.

Figure 4. Ua calculated and the total heat duty vs microchannel width (b) for different flue gas inlet temperatures Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C. Cocurrent scheme.

Table 4. Constant Parameter Values for Reactants and Flue Gas Sides AT NC ¼ 12:45 cm2

total cross-sectional free area

L ¼ 8 cm u ¼ 740:73 cm=s uH ¼ 1289:85 cm=s RV ¼ 0:015 cmcat 3 =cmreactor 3

heat but also by the shape of the heat flux profile along the reactor length. For other operating conditions (e.g., lower TH,in values than 800 °C), the differences between curves A and B are even more accentuated (results not shown). 3.1.2. Design Considerations. One distinctive aspect of the microreactors that directly concerns the heat transfer is the size of the microchannels. In this section, an analysis of the impact of the microchannel size on the heat transfer for the cocurrent configuration is presented. As a common comparison basis, both the total cross-sectional area and the residence time of the process and flue gas streams are kept constant. A constant ratio between the catalyst volume and the reactor volume is specified. The adopted parameters are summarized in Table 4.

Figure 6. Hydrogen yield vs microchannel width (b) for different flue gas inlet temperatures. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C. Cocurrent scheme.

Simulations are carried out for different channel widths (b). For each value of b, the number of channels (NC) and the washcoat thickness (wc) are varied to satisfy the specified parameters given in Table 4. Figure 4 shows the evolution of the product Ua (kW/m3 3 K), i.e., the heat-transfer parameter given in the right-hand side of eqs 10 and 11, as a function of the channel width (b). The parameter Ua decreases sharply with b for two reasons: (i) the specific heat-transfer area depends inversely on the channel width (i.e., a = 2/b); (ii) for the studied operating conditions, the Nusselt number given by eq 16 shows a weak dependence on the Reynolds number; i.e., the heat-transfer coefficients (h) decrease with the hydraulic diameter (dh), for both reaction and flue gas sides (dh = b). Figure 4 also includes the resulting heat duties for three different flue gas inlet temperatures (right ordinate axis). As higher channel widths are employed, the cocurrent microreactor supplies smaller amounts of heat to the reactant stream. Consequently, the ethanol conversion (Figure 5) and hydrogen yield (Figure 6) drop as the channel width (b) is increased. As expected, the conversion and hydrogen yield increase for higher 2701



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Figure 9. Ethanol conversion vs microchannel width (b) for different flue gas inlet temperatures. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C. Countercurrent scheme.

Figure 7. Axial profiles of ethanol conversion, temperatures, and local heat fluxes for two microchannel width sizes, points 1 (b = 200 μm) and 2 (b = 2000 μm) of Figure 5. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C; TH,in = 800 °C. Cocurrent scheme.

Figure 8. Calculated Ua and total heat duty vs microchannel width (b) for different flue gas inlet temperatures. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C. Countercurrent scheme.

TH,in values. It is important to note from Figure 6 that the cocurrent microreactor leads to ηH2 values above 3, provided that small enough channels and flue gas temperatures higher than 900 °C are employed (note that the theoretical stoichiometric yield for ESR is ηH2 = 6 if ethanol, CH4, and CO are not present at the reactor outlet). To complete the analysis, two different designs of Figure 5 are analyzed (points 1 and 2). Figure 7 shows the axial profiles of the local heat flux q (kW/m), temperatures of the process and flue gas (T and TH), and ethanol conversion (XEt) for points 1 (b = 200 μm) and 2 (b = 2000 μm). Both designs present decreasing heat flux profiles along the reactor length (Figure 7a,b). However, the microreactor with the smaller width presents much

higher local heat fluxes near the reactor entrance (z < 0.01 m) because of the higher values of Ua (see Figure 4). These high local heat supplies cause the process stream to reach high temperatures near the entrance (Figure 7c). Downstream of this short initial zone, the temperature of both streams decreases and, as a consequence of the high heat-transfer rates, the temperature profiles approach each other. The high temperatures in the first reactor section lead to high reaction rates and a sharp increase of the ethanol conversion (Figure 7e). The axial profiles corresponding to point 2 (b = 2000 μm) are not so steep. As seen in Figure 7b, the heat flux distribution for point 2 is much more uniform than that of point 1. This can be attributed to the small value of Ua, despite the high differences between the temperatures of both streams. As a result, the profile of the process gas temperature is more flat along the microreactor (Figure 7d) and the ethanol conversion presents a more uniform growth (Figure 7f). It is important to note that a sharply decreasing heat flux profile (Figure 7a), with a very high heat supply at the entrance of the microreactor, is convenient to maximize the hydrogen yield. 3.2. Countercurrent-Flow Configuration. A similar analysis is carried out for the countercurrent-flow scheme. As in section 3.1, the geometrical parameters b, NC, and wc are varied to satisfy the relationships given in Table 4. Figures 8-10 are analogous to Figures 4-6 of the cocurrent scheme. The results, however, do not present the same trends. Figure 8 shows the influence of the channel width on the parameter Ua. The behavior of the heat-transfer parameter is similar to that described in Figure 4; i.e., a sharp decline in the values of Ua occurs as b increases. The heat duty, conversely, does not show a monotonous decrease with b, as in the cocurrent case. For each inlet temperature of the flue gas stream, the heat duty is maximized at an intermediate value of the channel width (b ∼ 700 μm). For the same operating conditions, the outlet conversion of ethanol presents a trend similar to that observed for the heat supply (Q). As seen in Figure 9, the conversion is maximum at intermediate values of the channel width. This optimum value of b shifts to the left as TH,in increases. Smaller and bigger channels than the optimum one seem to be inconvenient for the ESR. Besides, the ethanol conversions are lower than those corresponding to the cocurrent configuration (Figure 5). 2702


Industrial & Engineering Chemistry Research The hydrogen yield (Figure 10) shows a behavior similar to that of ethanol conversion. By a comparison of Figures 6 and 10, it is important to note that for equivalent operating conditions the hydrogen yields for countercurrent flow are always smaller than those obtained with a cocurrent flow. As shown in Figure 10, ηH2 never reaches the value of 3 under countercurrent flow, even though the reactor is being operated at optimal conditions. A deeper analysis of the results of Figures 8-10 can be carried out by inspection of the axial profiles of the three selected designs of Figure 9 (points 1-3). These points correspond to different channel widths of the microreformer (200, 1000, and 2000 μm) and TH,in = 800 °C. Figure 11 shows the axial profiles of ethanol conversion, temperatures of both streams (process and flue gas), and local heat fluxes. As can be seen, the heat flux profiles (Figure 11a-c) present an opposite trend with respect to the cocurrent scheme; i.e., increasing heat fluxes are observed for the three channel widths. As in the cocurrent case (Figure 7a,b), the heat flux distribution is steeper as smaller channels are employed. Regarding the

Figure 10. Hydrogen yield vs microchannel width (b) for different flue gas inlet temperatures. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C. Countercurrent scheme.

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temperature curves (Figure 11d-f), variation of the flue gas temperature along the axial coordinate is more pronounced as the channel width decreases. The inlet temperature TH,in is kept constant at 800 °C (z = 0.08 m). For the highest channel (b = 2000 μm), the outlet flue gas temperature is around 671 °C. At point 2 (b = 1000 μm), this outlet temperature decreases up to 600 °C (close to the feed temperature Tin) because of the higher heat-transfer area per unit volume. Finally, if the smallest channel width is selected (b = 200 μm), a temperature crossing occurs approximately in the middle of the reactor (Figure 11d). In the first half of the reactor length, the flue gas stream is acting as a coolant (because TH < T for z < 0.04 m, the local heat fluxes in Figure 11a are negative). Under these conditions, the microreactor shows two zones of high temperature (close to the inlet and outlet) and a cold zone in the middle. Consequently, the ethanol conversion profile (Figure 11g) is flat in the central zone, with two increasing sections close to z = 0 and L. Nevertheless, the outlet conversion is poor (see point 1 in Figure 9). Point 3 (Figure 11i), conversely, shows a uniform increase of conversion along the reactor; there is no temperature crossing, but the parameter Ua is too low and the heat duty is poor. Point 2 (Figure 11h) represents the intermediate case, for which the conversion is a maximum.

4. CONCLUSIONS The influence of different design and operating variables on the performance of a microreactor for ESR has been analyzed. The results demonstrate that the ethanol conversion, methane slip, and hydrogen yield are mainly controlled by the heat supply. However, the hydrogen production rate is determined univocally not only by the total amount of heat transferred to the reactants but also by the axial evolution of the local heat-transfer fluxes. Similarly to other reforming processes, e.g., natural gas steam reforming, the optimal heat flux for ESR is decreasing. To approximate this optimal heat supply policy, a cocurrent-flow configuration using a flue gas stream as the heating medium is a good choice.

Figure 11. Axial profiles of ethanol conversion, temperatures, and local heat fluxes for three microchannel width sizes, points 1 (b = 200 μm), 2 (b = 1000 μm), and 3 (b = 2000 μm) of Figure 9. Fin = 10.24 Nm3/h; S/C = 3; Tin = 600 °C; TH,in = 800 °C. Countercurrent scheme. 2703


Industrial & Engineering Chemistry Research Under cocurrent-flow operation, the microreactor shows a performance substantially improved with respect to the countercurrent scheme, particularly when small channel widths are selected. The axial distribution of the heat supply under countercurrent-flow operation is opposite to the optimal one; therefore, lower conversions and hydrogen productions are obtained for equivalent conditions. Nonmonotonous trends with respect to the specific heat-transfer area are obtained, which suggests that there are optimum values for the heat-transfer parameter Ua, which depends on the selected operating conditions.

’ AUTHOR INFORMATION Corresponding Author

*E-mail: [email protected].

’ ACKNOWLEDGMENT Support of this work through Universidad Nacional del Sur (UNS) and Consejo Nacional de Investigaciones Científicas y Tecnologicas (CONICET) is gratefully acknowledged. ’ NOMENCLATURE a = heat-transfer area per unit volume = 2bL/b2L, m2/m3 AT = cross-sectional area of channels b2, m2 b = width of the square channel, m Cp,j = specific heat of component j, J/mol 3 K dh = hydraulic diameter = 2b2/(b þ b), m e = fin width, μm Fj = molar flow of component j (reaction side), Nm3/h FH,j = molar flow of component j (flue gas side), Nm3/h h = heat-transfer coefficient, W/m2 3 K ΔHR,I = heat of reaction of reaction i, J/mol k = wall thermal conductivity, W/m 3 K L = channel length, m NC = number of microchannels N = number of components Nu = Nusselt number = hdh/λ, dimensionless P = total pressure (both sides), MPa Pr = Prandlt number = Cpμ/λ, dimenssionless Q = total heat flux, kW q = local heat flux, kW/m ri = reaction rate of reaction i, i = 1-3, mol/m3 3 s Re = Reynolds number = Fudh/μ, dimensionless RV = catalyst and reactor volume ratio, m3cat/m3reactor S/C = steam-to-carbon molar ratio, mol of H2O/mol of C T = temperature (reaction side), °C TH = temperature (flue gas side), °C u = average velocity, m/s U = overall heat-transfer coefficient, W/m2 3 K wc = thickness of the washcoated catalyst, μm XEt = ethanol conversion, dimensionless z = axial coordinate, m Greek letters

G = density, kg/m3 μ = viscosity, Pa 3 s ηi = yield of component i, dimensionless λ = thermal conductivity, W/m 3 K Subscripts

cat = catalyst Et = ethanol

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j = component j i = reaction i H = heating gas (flue gas) in = inlet out = outlet 0 = at the axial coordinate z = 0 L = at the axial coordinate z = L C2H5OH = ethanol CH4 = methane CO = carbon monoxide CO2 = carbon dioxide H2 = hydrogen H2O = water O2 = oxygen N2 = nitrogen

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