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Chemical Engineering Journal 166 (2011) 1116–1125

Contents lists available at ScienceDirect

Chemical Engineering Journal

journa l homepage: www.e lsev ier .com/ locate /ce j

microreactor modeling, analysis and optimization for methane autothermaleforming in fuel cell applications

.H. Akbari ∗, A.H. Sharafian Ardakani, M. Andisheh Tadbirenter for Fuel Cell Research, School of Mechanical Engineering, Shiraz University, Shiraz, 71348-51154, Iran

r t i c l e i n f o

rticle history:eceived 16 September 2010eceived in revised form3 December 2010ccepted 13 December 2010

eywords:icroreactor

a b s t r a c t

Hydrogen production through autothermal reforming (ATR) of hydrocarbons, such as methane, is anattractive option for mobile applications of hydrogen fuel cells. In the present study, a numerical inves-tigation of catalytic autothermal reforming of methane in a surface microreactor is presented. Themicroreactor is intended for use in a micro-fuel processor device for a low-power PEM fuel cell in amobile application. A three-dimensional ATR reactor model is developed to simulate the flow and sur-face reactions in a microchannel of square cross-section. A four-reaction mechanism is implemented tosimulate the surface reactions on a Ni/Al2O3 catalyst, and a global mechanism is used to model gas-phase

utothermal reformingethane

urface reactionFD modelinguel cells

methane oxidation. The governing equations in the model include conservations of mass, momentum,energy and chemical species. The simulation results reveal the dependency of hydrogen yield on spacevelocity (SV), air/fuel molar ratio (A/F), water/fuel molar ratio (W/F), and the feed gas temperature. Theoptimum conditions for the highest hydrogen yield were obtained in our simulations as: space velocityof 50,000 h−1, A/F of 1.0, W/F of 3.0 and feed gas temperature of 600 ◦C. The carbon monoxide productionat these conditions is small enough to make the generated synthesis gas suitable for fuel cell applicationsafter further treatment downstream of the ATR microreactor.

. Introduction

Considering the rapid depletion of non-renewable naturalesources and their environmental impact, renewable energyesources are sought as their substitutes particularly for fossil fuels.ydrogen is considered as a primary energy carrier in an economyased on such renewable energy resources. This has considerablyaised the interest in the development of low-cost, efficient, andompact hydrogen production technologies [1–3].

All fuel cell types provide electrical power by consuming pureydrogen or a hydrogen-rich synthesis gas generated from other

uels. There is a potential advantage in the fact that hydrogen ishe preferred fuel for all fuel cell types: hydrogen can be generatednd used to substitute hydrocarbons to fuel different applicationsncluding fuel cells and internal combustion engines. On the otherand, there are drawbacks in using hydrogen as the primary energyarrier, including the high cost required for hydrogen storage and

elivery at the present time. In addition, hydrogen storage can-ot be considered as a viable option for many mobile fuel cellpplications due to its very low energy density [4]. However, on-emand reforming of other fuels to hydrogen is a practical option to

∗ Corresponding author. Tel.: +98 917 308 8424; fax: +98 711 647 3511.E-mail address: h-akbari@shirazu.ac.ir (M.H. Akbari).

385-8947/$ – see front matter © 2010 Elsevier B.V. All rights reserved.oi:10.1016/j.cej.2010.12.044

© 2010 Elsevier B.V. All rights reserved.

tackle this problem at the present. Currently three major thermo-chemical reforming techniques are available to produce hydrogenfrom hydrocarbon fuels, i.e., steam reforming (SR), partial oxida-tion (POX), and autothermal reforming (ATR). SR has been usedextensively in different industries for hydrogen production, but thisprocess is strongly endothermic. Thus the overall configuration ofSR with heat exchangers makes the reforming system very bulkyand heavy, which is not suitable for a mobile fuel cell system.

Another technique that has been proposed to avoid heat transferproblems in SR is partial oxidation (POX) of the hydrocarbon fuel.This process is exothermic with easy start ups even in the absenceof a catalyst. However, high carbon monoxide concentrations areproduced in POX which must be removed from the synthesis gasotherwise it will poison the catalyst of certain types of fuel cells.

To avoid this problem, autothermal reforming (ATR) is sug-gested in which the thermal effects of POX and SR reactions arecombined by feeding the fuel, water, and air together into thereactor. The two processes (SR and POX) occur simultaneouslyin the presence of a catalyst in the reactor. The thermal energygenerated by POX is absorbed by SR, hence, the overall process

is thermally neutral. Moreover, the temperature of the process islowered, which is favorable to the water–gas shift reaction whichconsumes the generated carbon monoxide and produces morehydrogen. Advantages of an ATR reactor also include its small size,relative simplicity, easy operation, quick response, and inexpen-

M.H. Akbari et al. / Chemical Engineerin

Nomenclature

cp constant pressure specific heat (kJ/kg K)D binary diffusion coefficient (m2/s)E activation energy (kJ/kmol)�H reaction enthalpy (kJ/kmol)k thermal conductivity (W/m K) or kinetic rate con-

stantK adsorption constant in reactions R2–R4ko pre-exponent coefficient of kinetic rate constantKC adsorption constant in reaction R1Ke equilibrium constant in reactions R2–R4Ko pre-exponent coefficient of adsorption constant in

reactions R2-R4KC

o pre-exponent coefficient of adsorption constant inreaction R1

MW molar mass (kg/kmol)p partial pressure (bar)P pressure (Pa)r species reaction rate (kmol/kgcat h)R total reaction rate (kmol/kgcat h)Ru universal gas constant (8.315 kJ/kmol K)T temperature (K)u x-velocity component (m/s)�V velocity vector (m/s)v y-velocity component (m/s)w z-velocity component (m/s)x x-direction coordinate (m)y y-direction coordinate (m)Y species mass fractionz z-direction coordinate (m)

Greek symbols� reaction effectiveness factor� stoichiometry coefficient� kinematic viscosity (N s/m2)ω species surface reaction rate (kmol/m2 s)� density (kg/m3)�cat catalyst surface density (kg/m2)� species mole fraction

Sub- or superscriptscat catalysteff effectivein inleti ith species or reactionj jth species or reactionmix gas mixtureout outletvol volumetric reaction

sbt[

ctotas

given dimension) would not have any significant effect on the prod-ucts. Furthermore, the hydraulic diameter was selected based on

ive material requirements. For these reasons, the catalytic ATR haseen the focal point of much research in recent years as a feasibleechnique for hydrogen production in mobile fuel cell applications5–7].

Trimm and Lam [8] studied ATR of methane on Pt/Al2O3 fibreatalyst. The kinetics of the reactions were measured at tempera-ures around 800 K. Xu and Froment [9] presented intrinsic kinetics

f methane SR, methanation and the water–gas shift (WGS) reac-ion on Ni/Al2O3 catalyst for initial temperature range between 573nd 823 K. The kinetics they developed was implemented exten-ively in subsequent investigations.

g Journal 166 (2011) 1116–1125 1117

Ma et al. [10] presented a kinetic rate of total combustion ofmethane based on Pt/�-Al2O3 catalyst for initial temperature rangeof 663–723 K. de Groote and Froment [11] simulated the catalyticPOX of methane to synthesis gas on a Ni catalyst in an adiabaticreactor with various inlet air/fuel (A/F) and water/fuel (W/F) molarratios. They presented the kinetic rates for eight reactions alongwith a series of effectiveness factors associated with these reac-tions.

Dias and Assaf [12] examined ATR of methane over Ni/�-Al2O3catalyst on a thin layer packed bed. The effects of adding small quan-tities of noble metals to the base catalyst were also investigatedat the temperature range of 673–873 K. This study indicated thatthe effect of noble metals was significant and increased hydrogenproduction up to 25%.

Halabi et al. [13] presented a one-dimensional model for het-erogeneous autothermal reforming of methane over Ni/Al2O3 in afixed-bed reactor with feed gas temperature and pressure of 773 Kand 1.5 bar. The optimum conditions were located in a space veloc-ity (SV) range from 1050 to 14,000 h−1, steam/carbon molar ratioof 4.5–6.5 and oxygen/carbon molar ratio of 0.45–0.55 to obtainmaximum yield of hydrogen.

Most of the previous investigations were carried out for packed-bed reactors (with volumetric reactions). This type of reactor causeshigh pressure losses which prohibit their application for mobiledevices. In the present study, a microchannel reactor is consid-ered for catalytic ATR of methane in the presence of a Ni-basedcatalyst on its internal surfaces. A three-dimensional, steady-state, adiabatic model is developed. The reactor performance isexamined with respect to major parameters including feed gastemperature, air/methane and water/methane molar ratios andspace velocity. The microchannel reactor is of a square cross sec-tion with 340-�m sides and 8.5 mm length. These dimensionswere selected based on practical constraints and requirements.The synthesis gas generated by an assembly of such ATR microre-actors will be fed to water–gas shift and preferential oxidation(PROX) reactors downstream. Ultimately, the generated hydrogen-rich stream would be fed to a PEM fuel cell for a portable electronicdevice.

Our simulation results indicate that one such microchannel ATRreactor will produce enough synthesis gas to generate 0.01 W ofelectrical power (in a PEM fuel cell) at optimum operating condi-tions of the reactor. Therefore, to power a 100 W portable device,for instance, a total number of 10,000 of these microreactors willbe required to be assembled in parallel (an array of 100 × 100microchannels). Such assembly will take up an overall space ofabout 35 mm × 35 mm × 10 mm as part of a complete micro-fuelprocessor device. One advantage of a microchannel surface reactorof this type is its negligible pressure loss, in addition to its smallsize and weight [14].

2. Numerical model

2.1. Model geometry

The simulated microreactor consists of a channel with squarecross section of 340-�m sides and 8.5 mm length. As mentionedpreviously, these dimensions were selected based on practical con-straints and requirements. More specifically, our simulation resultsindicate that any further increase in the reactor length (for the

our practical constraints in the fabrication and coating of suchmicrochannels. The schematic geometry and dimensions of thechannel are presented in Fig. 1.

1118 M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125

2

pp7a

tIhts(m

2

oabC

E

∇

s

V

m

∇

ω

f

Y

oc

(

Table 1List of autothermal reactions.

Description Reaction �H298 K (kJ/mol)

R1: Total combustion CH4 + 2O2 → CO2 + 2H2O −802R2: Steam reforming CH4 + H2O ↔ CO + 3H2 206R3: Water gas shift CO + H2O ↔ CO2 + H2 −41R4: Steam reforming CH4 + 2H2O ↔ CO2 + 4H2 165R5: Partial oxidation CH4 + 0.5O2 → CO + 2H2 −36

Fig. 1. Schematics and dimensions of the microchannel reactor.

.2. Model assumptions

The flow regime is assumed to be steady, laminar and incom-ressible. The gas mixture is treated as an ideal gas with variableroperties based on the local composition. The mixture consists ofspecies: CH4, O2, CO, CO2, H2, H2O and N2. Furthermore, Dufour

nd Soret effects are neglected in the species mass diffusion.A three-dimensional model has been constructed and used in

his analysis because of the nature of the phenomena involved here.t must be noted that the channel is of a square cross section andence no axis of symmetry exists1. Moreover, the bulk flow is inhe axial direction, while mass transfer occurs mainly by diffu-ion in the transverse directions towards/from the reactor wallswhere reactions occur). Therefore, a one- or two-dimensional

odel would not be adequate to model this problem accurately.

.3. Governing equations

The governing equations include the ideal gas mixture equationf state, as well as the conservations of mass, momentum, energynd chemical species. Since the microchannel geometry resem-les a rectangular cube, the governing equations are stated in theartesian coordinate system.

The continuity equation for a 3D steady-state flow is given byq. (1).

· (� �V) = 0 (1)

The Navier–Stokes equation, given below, represents the con-ervation of momentum for the given conditions.

� · ∇(� �V) = −∇p + ∇ · (�∇ �V) (2)

The mass balance for the ith species is given by Eq. (3), with theolar production rate given by Eq. (4).

· (� �VYi) = ∇ · (Deff,i �∇Yi) + ωiMWi (3)

i = �cat

∑4

j=1(�ij Rj �j) (4)

Eqs. (3) and (4) are used for 6 species, and the nitrogen massraction is calculated from the following equation.

N2 = 1 −∑

i,i /= N2

Yi (5)

Since the chemical reactions occur only on the internal wallsf the microchannel, the source term in Eq. (3) is zero for internalomputational cells.

1 A square channel has two planes of symmetry, but has no axis of symmetryunlike a circular channel with an axis of symmetry).

R6: Partial combustion CH4 + O2 → CO2 + 2H2 −71R7: Dry reforming CH4 + CO2 → 2CO + 2H2 247R8: Boudouard reaction 2CO ↔ C + CO2 −172R9: Decomposition CH4 ↔ C + 2H2 75

The conservation of energy is given by Eq. (6) below.

∇ · (�cp �VT) = ∇ · (k∇T) + �cat

∑4

i=1(�Hi Ri �i) (6)

For the same reason stated above, the source term in this equa-tion is also non-zero only on the internal walls of the microchannel.

Fluid density is computed from the ideal gas equation fora multi-component mixture. However, since the temperatureand pressure have negligible variations along the simulated ATRmicroreactor, the temperature and pressure are taken at their inletvalues in the following equation.

� = P MWmix

RuT(7)

The mixture molar mass in the above equations is computedfrom the following.

MWmix =∑

i�i MW = 1∑

iYi/MWi(8)

The binary mass diffusivity is calculated by the relationshipsgiven by Reid et al. [15]. The effective mass diffusion coefficient iscomputed from the following equation.

Deff,i = 1 − �i∑∀jj /= i�j/Dij

(9)

The constant pressure specific heat, dynamic viscosity and ther-mal conductivity of the gas mixture are computed by Eqs. (10)–(12).

cp,eff =∑

icp,iYi (10)

�eff =∑

i�iYi (11)

keff =∑

iki�i (12)

Methane conversion is given by the following equation.

CH4 conversion(%) =�in

CH4− �out

CH4

�inCH4

× 100 (13)

2.4. Chemical kinetics

Catalytic autothermal reforming of methane involves manyreactions with different rates, some of which are reported by Hal-abi et al. [13] and listed in Table 1. The complex nature of thedetailed mechanism and limitations imposed by numerical calcu-lations are good reasons to consider only reactions with significantrates. According to [4,13], four reactions are proven to have highrates in comparison with others. Therefore, only these four majorreactions are considered in the present model. These reactions are

listed as follows.

(1) The exothermic total combustion of methane,

R1 : CH4 + 2O2 → CO2 + 2H2O, H298 = −802 kJ/mol (14)

M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125 1119

Table 2Kinetic parameters.

Reaction koj (kmol/kg cat h) Ej (kJ/kmol)

17 −1.5

(

(

(

steTcL

Wkkr

R

R

R

R

Q

Hotac

Toftr

TE

Table 4Adsorption constants.

Species Eq. no. Koi (/bar) �Hi (kJ/kmol)

O2 18 5.08 × 104 (bar0.5) 66,200CH4 18 4.02 × 105 103,500H 22 6.12 × 10−9 −82,900

R1 5.852 × 10 (bar ) 204,000R2 4.225 × 1015 (bar0.5) 240,100R3 1.955 × 106 (bar−1.0) 67,130R4 1.020 × 1015 (bar0.5) 243,900

2) The endothermic steam reforming of methane to carbonmonoxide,

R2 : CH4 + H2O ↔ CO + 3H2, H298 = 206 kJ/mol (15)

3) The water–gas shift reaction,

R3 : CO + H2O ↔ CO2 + H2, H298 = −41 kJ/mol (16)

4) The steam reforming of methane directly to carbon dioxide.

R4 : CH4 + 2H2O ↔ CO2 + 4H2, H298 = 165 kJ/mol (17)

Xu and Froment [9] suggested a kinetic rate model for methaneteam reforming over Ni-based catalyst, which has been vastlyested under lab-scale conditions. Also, Trimm and Lam [8] and Mat al. [10] reported kinetic rate equations for methane combustion.he model proposed in [8] is for methane combustion over Pt-basedatalyst at a fixed temperature, while that in [10] is based on theangmuir–Hinshelwood model adjusted for Ni-based catalyst.

In the present study, the kinetic rate equations for SR andGS reactions are adopted from [9], and the methane combustion

inetic rate from [10]. Eqs. (18)–(22) below show the correspondinginetic rates for reactions given by Eqs. (14)–(17) in a temperatureange of 400–600 ◦C.

1 =k1pCH4 p1/2

O2

(1 + KCCH4

pCH4 + KCO2

p1/2O2

)2

(18)

2 = k2

p2.5H2

(pCH4 pH2O −

p3H2

pCO

Ke2

)× 1

Q 2r

(19)

3 = k3

pH2

(pCOpH2O − pH2 pCO2

Ke3

)× 1

Q 2r

(20)

4 = k4

p3.5H2

(pCH4 p2

H2O −p4

H2pCO2

Ke4

)× 1

Q 2r

(21)

r = 1 + KCOpCO + KH2 pH2 + KCH4 pCH4 + KH2OpH2O

pH2

(22)

ere kj = koj × e(−Ej)/RuT in Eqs. (18)–(21) is the kinetic rate constantf reaction j (j = reactions R1–R4), in which k1 is determined fromhe data of [10], while k2 to k4 are from [9]. These kinetic data (kojnd Ej) are given in Table 2. Kej in Eqs. (19)–(21) is the equilibriumonstant of reaction j (j = reactions R2–R4) which can be found in

able 3. KC = KC × e(−�HiC

)/RuT in Eq. (18) is the adsorption constant

i oi

f species i (i = CH4, O2) in the combustion reaction R1 and can beound in Table 4. Furthermore, Ki = Koi × e(−�Hi)/RuT in Eq. (22) ishe adsorption constant of species i (i = CO, H2, CH4, H2O) in theeforming reactions R2–R4 which can also be found in Table 4.

able 3quilibrium constants.

Reaction Equilibrium constant Kej

R2 5.75 × 1012 exp(−11476/T) (bar2)R3 1.26 × 10−2 exp(4639/T)R4 7.24 × 1010 exp(−21646/T) (bar2)

2

CH4 22 6.65 × 10−4 −38,280H2O 22 1.77 × 105 bar 88,680CO 22 8.23 × 10−5 −70,650

Once the reaction rates are calculated from Eqs. (18)–(22), it willbe possible to obtain the rate of formation or consumption of eachspecies by summing up the reaction rates of that species in all thereactions. de Groote and Froment [11] suggested corrected reac-tion rates to reduce the rates obtained by Eqs. (18)–(22) in order toaccount for the mass transport intraparticle transformations. Thesecorrected rates are estimated by multiplying an effectiveness fac-tor to each reaction rate. Considering the particle diameter in ourcase (assumed to be 20 �m), the effectiveness factors for reactionsR1–R4 are �1 = 0.64, �2 = 0.76, �3 = 0.99 and �4 = 0.70, respectively.Therefore, the final equations for the calculation of the reaction ratefor each species are summarized as follows:

rCH4 = −�1R1 − �2R2 − �4R4rO2 = −2�1R1rCO2 = �1R1 + �3R3 + �4R4rH2O = 2�1R1 − �2R2 − �3R3 − 2�4R4rCO = �2R2 − �3R3rH2 = 3�2R2 + �3R3 + 4�4R4

(23)

The above surface reactions occur on the microreactor walls,while gas-phase methane oxidation may take place within the reac-tor volume. Hence, a global mechanism for the volumetric methaneoxidation in the gas phase is also considered in this study. One of themost frequently used global mechanisms for this reaction in engi-neering applications is given by Westbrook and Dryer [16]. Thissingle-step mechanism is given in Eq. (24), and is implemented inthe present work.

R1,vol = A exp(−Ea

RuT

) p−0.3CH4

p1.3O2

RuT(24)

The values of parameters A and Ea/Ru in the above equation are1.3 × 1011 s−1 and 24,358 K, respectively.

2.5. Boundary conditions

Velocity, pressure, temperature and the composition of the inletgas mixture are specified at the microchannel inlet. The channelwalls are assumed adiabatic, but are permeable to species masstransfer due to surface reactions. The adiabatic assumption of thewalls is justified considering the fact that in an actual assemblyof such microreactors, one single channel is surrounded by similarchannels that are operating at the same conditions. Therefore, noheat will be transferred from one channel to the adjacent one.

Considering the Knudsen number of the flow, the flow regime iscontinuum and hence the no-slip condition is used on the microre-actor walls. The outflow mixture is discharged to atmosphericpressure, and the gradients of all other dependent variables areset to zero at the microreactor outlet.

2.6. Solution method

In order to solve the non-linear, coupled governing equations,a computational fluid dynamics (CFD) code was developed basedon the finite volume method using the SIMPLE algorithm for thecoupling of pressure and velocity domains [17]. To linearize the

1120 M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125

Table 5Operating conditions for the baseline case.

Parameter Value

Space velocity (h−1) 20,000Molar A/F ratio 3.5

cf[pe

sattnoest

3

3

svaipwadraarddoorsceortCft

3

Tpuzo

Molar W/F ratio 1.0Inlet temperature (K) 873Inlet pressure (Pa) 101,325

onvective terms the power-law scheme is used, and a central dif-erence scheme is implemented to linearize the diffusive terms18]. The algebraic system of equations is solved using the Jacobianoint-to-point iteration method, except for the pressure correctionquation which is solved by the Gauss–Seidel iteration method [19].

To save computational time, the governing equations are firstolved without considering reactions and property changes. Afternumber of iterations, this solution is used as an initial guess for

he reactive flow calculations. Moreover, thermo-physical proper-ies of the gas mixture are updated every 50 iterations to reduceumerical fluctuations. In order to stabilize numerical calculationsf the reactive flow field, under-relaxation factors are used for allquations which are gradually increased up to 0.95. This promotestability of the numerical calculations, while secures accuracy ofhe numerical results.

. Results and discussion

.1. Validation

The present model has been developed for methane ATR in aurface microreactor (as opposed to a packed-bed reactor witholumetric reactions); as a result, there has been difficulty incquiring compatible data, either experimental or numerical, forts direct validation. Only a qualitative validation was thereforeossible in this case. In order to qualitatively validate the presentork, the experimental results reported by Dias and Assaf [12]

re used here. They performed experiments for methane ATR atifferent temperatures on a Ni/Al2O3 catalyst in a packed-bedeactor with 9 mm diameter and 10 mm length. The methane,ir and water volume flow rates were 3.33 × 10−4, 8.9 × 10−4

nd 1.332 × 10−3 m3/s, respectively, corresponding to CH4:O2:H2Oatio of 1:0.5:4. Numerical simulations are performed for identicalimensions and conditions using the present CFD code, with theifference that in our simulations the catalytic reactions occur onlyn the reactor wall surfaces. The simulation results for the variationf H2 and CO versus inlet temperature are compared with the cor-esponding experimental data in Fig. 2. The general trends are veryimilar, but the simulated surface reactor shows lower conversionompared with the actual packed-bed reactor. This was obviouslyxpected, since the two reactors have the same dimensions andperating conditions but the packed-bed reactor possesses moreeaction sites. Nonetheless, it is observed in both sets of resultshat increasing temperature leads to an increase in both the H2 andO production. This qualitative validation of the simulation results

rom the present CFD code, gives us confidence in the accuracy ofhe developed model.

.2. Baseline simulation

The operating conditions for the baseline case are specified in

able 5. Shown in Fig. 3(a) are contours of the axial velocity com-onent in the mid x–z plane. Note the difference in the length scalesed for the axial direction (x) and the two lateral directions (y and) in this figure. The calculation results indicate that fully devel-ped flow is reached at a distance 10 times the hydraulic diameter

Fig. 2. Comparison of the present simulation results with experimental data: outletmole fractions of (a) H2 and (b) CO.

of the channel, and that the peak value of this velocity compo-nent increases to twice the value of the inlet velocity. Both of thesecharacteristics are compatible with theoretical and experimentalexpectations for the given flow conditions. The gage pressure con-tours along the microchannel are plotted in Fig. 3(b). The pressureloss along the microreactor is seen to be negligible compared withwhat would occur in a similar packed-bed reactor. This negligi-ble pressure loss (less than 3 Pa) also confirms the validity of theconstant-pressure assumption made for density calculations.

Species mole fractions and temperature variations (averagedat each cross section) along the microchannel are presented inFig. 4 for the baseline conditions. It is observed that most of themethane consumption occurs in the first quarter of the microchan-nel length which results from high kinetic rates of total combustionof methane (R1) and then steam reforming of methane to CO (R2).Oxygen is consumed very rapidly at the entrance of the microchan-nel; this is caused by two effects: low mole fraction of the inletoxygen and high kinetic rate of methane combustion (R1) whichis about two orders of magnitude larger than other kinetic ratesin the entrance region. CO2 mole fraction keeps rising along themicrochannel length, although it reaches 75% of its final value atthe first 5% of the channel length due to high kinetic rates of reac-tions R1 and R2 (Eqs. (14) and (15)). The subsequent slow increase

in CO2 mole fraction is because of the activation of WGS and SRreactions.

It is also observed in Fig. 4 that water mole fraction increasesat the entrance region of the microreactor resulting from methanecombustion, but it decreases consequently because water is con-

M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125 1121

Fm

sCa(aitatr

Fm

ig. 3. Contours of (a) the axial velocity component, and (b) gage pressure in theid x–z plane along the microchannel for the baseline conditions.

umed in SR and WGS reactions to produce hydrogen, CO and CO2.arbon monoxide mole fraction increases along the microchannellmost monotonically. This is because CO is produced in SR reactionR2) at a moderate rate, while it is consumed in WGS reaction (R3)t a lower rate. However, the SR kinetic rate is dominant in compar-

son with WGS reaction; thus CO mole fraction shows an increasingrend along the microreactor. The variation of H2 mole fraction islso shown in Fig. 4. Hydrogen production is the main objective ofhe ATR reactor which appears in reactions R2–R4 (SR and WGSeactions). The hydrogen mole fraction increases to 85% of its final

ig. 4. Variations of the species mole fractions and temperature along theicrochannel for the baseline conditions.

Fig. 5. Corrected rates of reactions R1–R4 versus microchannel relative length forthe baseline conditions.

value in the middle of the microchannel, while it keeps an increas-ing trend at a lower rate in the second half of the microchannellength.

Also shown in Fig. 4 is the temperature variation along themicrochannel. It is observed that the temperature varies very lit-tle (less than 2 ◦C) along the channel, which is consistent with thenature of the overall process being autothermal.

The corrected rates of reactions R1–R4 (R1�1–R4�4) along themicrochannel are plotted in Fig. 5 for the baseline case. It is observedthat reaction R1 is practically completed in the first 20% of the chan-nel length, while other reactions continue to occur along the entirereactor length. It is interesting to note that the kinetic rate of theWGS reaction (R3) is orders of magnitude less than that for theother reactions in the entrance region of the reactor. This figurealso shows that in most of the reactor length, the corrected kineticrates of the SR and WGS reactions are sorted in the following order:R2�2 > R3�3 > R4�4.

Shown in Fig. 6 is a comparison of the simulation results forthe species mole reactions along the microreactor with and with-out considering methane gas-phase oxidation reaction. The figureclearly indicates that the effect of the gas-phase reaction is quitenegligible in comparison with the surface reactions considered hereas the two sets of results are almost identical.

The simulation results for the baseline conditions show that theoutlet H2 mole fraction is 18.5% (dry basis), while the outflow alsocontains 6.5% CO. It is noted that the microreactor considered in thisinvestigation is of a surface reaction type; hence a lower hydrogen

Fig. 6. Effect of methane gas-phase oxidation reaction on the species mole fractionsalong the microchannel for the baseline conditions.

1122 M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125

Ft

pt

atftsiCrpe

3

gatsseof

3

fwTimrfib

acsrf

v

ig. 7. Effect of space velocity on outlet species mole fractions (other parameters atheir baseline values).

roduction is expected compared with a packed-bed reactor withhe same space velocity and operating conditions.

Equilibrium conditions for the baseline pressure, temperaturend inlet mole fractions were obtained using the CHEMKIN collec-ion. The equilibrium mole fractions (dry basis) at these conditionsor CH4, H2, CO and CO2 are 1.7%, 32.3%, 6.5% and 9.8%, respec-ively. A comparison of these mole fractions with the correspondingimulation results for the catalytic surface reactor is interesting; itndicates that lower amounts of hydrogen and similar amounts ofO are produced in the surface reactor compared with the equilib-ium conditions. This comparison has been made to verify that theredicted hydrogen fraction from the present simulation does notxceed that at equilibrium conditions.

.3. Parametric study

In this section, the influence of different parameters on hydro-en yield and carbon monoxide production of the microreactorre investigated. These parameters include space velocity, feed gasemperature, air/CH4-supplied molar ratio (A/F), and water/CH4-upplied molar ratio (W/F). The microreactor dimensions are theame as those in the baseline case. The inlet volume flow rate inach case is calculated from the specified space velocity (SV). More-ver, species mole fractions are calculated from A/F and W/F ratiosor each case.

.3.1. Effects of space velocityPresented in Fig. 7 are the variations of the outlet species mole

ractions for different space velocities among 1000–100,000 h−1

ith other parameters at their baseline values (given in Table 5).hese results indicate that the outlet mole fraction of methanencreases as SV is increased (its conversion is decreased). This is

ainly because of a reduction in the kinetic rate of the steameforming reactions R2 and R4 with increasing SV. The outlet moleraction of oxygen is always almost zero. This is because of its lownlet amounts and also the very high kinetic rate of methane com-ustion.

Fig. 7 also shows that the outlet CO2 mole fraction decreasess SV is increased. Although the production of carbon dioxide ishiefly due to total combustion of methane, reaction R4 has also

ome contribution in the CO2 production. As mentioned before, thiseaction is suppressed as SV in increased; hence CO2 outlet moleraction decreases with an increase in SV.

The outlet mole fraction of water vapor reaches a minimumalue at a space velocity of 10,000 h−1 as SV is changed in a

Fig. 8. Effect of inlet temperature on outlet dry mole fractions of (a) H2 and (b) CO(other parameters at their baseline values).

range among 1000–100,000 h−1. The outlet hydrogen mole fractionreaches its peak value of 19.2% (dry basis) at SV of 50,000 h−1, whilethe mole fractions of both CO and CO2 show decreasing trends withan increase in SV.

3.3.2. Effects of feed gas temperatureVariations of the outlet mole fractions of hydrogen and carbon

monoxide with the feed gas temperature are presented in Fig. 8,with other parameters at their baseline values (given in Table 5).These results show that the mole fractions of both hydrogen andcarbon monoxide increase as the inlet temperature is increased ina range of 400–600 ◦C. This is mainly due to an increase in the rateof SR reactions R2 and R4 with temperature, as can be inferred fromtheir activation energy values given in Table 2. Higher temperaturesalso hinder WGS reaction (being an exothermic reaction) which fur-ther increases CO content. In fact, the CO production increases morethan 4 times as the temperature is increased beyond 500 ◦C up to600 ◦C, while the hydrogen mole fraction increases almost linearlyin the temperature range of 400–600 ◦C. The hydrogen mole frac-tion is 18% at a temperature of 600 ◦C which is about 4.5 times morethan that at 400 ◦C.

Shown in Fig. 9 is the outlet temperature versus the inlet tem-perature with other parameters fixed at their baseline values. It isseen that the variation of the outlet temperature is linear, and thatits value is very close to the inlet temperature as expected in anATR reactor.

3.3.3. Effects of A/F and W/F molar ratiosPresented in Fig. 10(a) is the variation of methane outlet mole

fraction (dry basis) versus A/F ratio, with W/F ratio as a parameter(other parameters are at their baseline values, given in Table 5).Fig. 10(b) shows methane conversion corresponding to the above

M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125 1123

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ig. 9. Effect of inlet temperature on outlet temperature (other parameters at theiraseline values).

onditions. It is seen that by increasing air or water content ofhe feed gas mixture, methane outlet fraction will decrease, i.e.,ts conversion will increase.

Increasing A/F ratio provides more oxygen for the methaneombustion reaction (R1) to occur which also consumes more

ethane. Clearly, increasing the inlet water content promotes SR

eactions (R2 and R4), both of which also consume methane. It islso observed that at high A/F ratios, methane outlet fraction islmost independent of W/F ratio. This is because with enough oxy-

ig. 10. Effect of A/F ratio on (a) outlet methane mole fraction and (b) methaneonversion, for different W/F ratios (other parameters at their baseline values).

Fig. 11. Effect of A/F ratio on outlet mole fractions of (a) H2 and (b) CO, for differentW/F ratios (other parameters at their baseline values).

gen, most of the methane is consumed by the combustion reaction(R1) due to its kinetic rate being orders of magnitude higher thanthe other reactions. In this case, water content of the mixture willhave an insignificant effect on the overall reactor product, since SRand WGS reactions (R2–R4) are deactivated.

The outlet hydrogen and CO mole fractions (dry basis) versusA/F ratio are plotted in Fig. 11, with W/F ratio as a parameter. Itis observed in Fig. 11(a) that by decreasing A/F ratio or increasingW/F ratio the outlet hydrogen mole fraction is increased, which isvery desirable for fuel cell applications. In fact, the maximum outletH2 mole fraction is 38.5% (dry basis) which occurs at A/F ratio of1.0, W/F ratio of 3.0, and other parameters at the given baselineconditions (SV of 20,000 h−1 and feed gas temperature of 600 ◦C).

By decreasing the inlet air content the combustion reaction ofmethane (R1) is demoted and less methane is consumed by thisreaction. This will promote the SR reactions directly, and the WGSreaction indirectly, both of which will result in generation of morehydrogen. Furthermore, increasing the inlet water content will pro-mote the same reactions which will further increase hydrogenproduction.

Fig. 11(b) shows that increasing either A/F ratio or W/F ratio willdecrease the outlet CO content; this is also desirable in fuel cellapplications. For the given range of parameters simulated here, thelowest mole fraction of CO is 2.3% which is obtained at A/F ratio of5.0, W/F ratio of 3.0 and other parameters at their baseline values.

Increasing the inlet air content will consume more methane bythe combustion reaction, leaving less methane available to the SRreactions. This in turn will give less chance for CO generation. Fur-thermore, as the inlet water content is increased, both the SR and

1124 M.H. Akbari et al. / Chemical Engineering Journal 166 (2011) 1116–1125

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Table 6Optimum operating conditions and the corresponding outlet mole fractions.

Parameter Value

Space velocity (h−1) 50,000Molar A/F ratio 1.0Molar W/F ratio 3.0Inlet temperature (K) 873Inlet pressure (Pa) 101,325Outlet H2 mole fraction (%) 38.9

ig. 12. Effect of A/F ratio on outlet temperature for different W/F ratios (otherarameters at their baseline values).

GS reactions (R2–R4) are promoted. However, the WGS reactions more sensitive to this increase, which causes a reduction in theutlet CO content.

The outlet temperature variations versus A/F ratios for different/F ratios are presented in Fig. 12. It is observed that at higher

/F ratios the outlet temperature decreases. This decrease of theutlet temperature is because of the increased mass flow rate whichutbalances the increased heat release from the oxidation reaction.

.3.4. Optimum operating conditionsBased on the simulation results presented above, one can sug-

est the following regarding the optimum operating conditions forhe ATR reactor modeled here. It appears that an optimum valueor the space velocity exists for every given set of conditions. Forxample, for the baseline conditions this value is obtained around0,000 h−1 in our simulations. Also, since higher water contentsoth promote hydrogen production and reduce CO generation, a/F ratio of 3.0 is suggested.Selection of an optimum value for A/F ratio may seem cumber-

ome. On the one hand a higher value of A/F reduces CO content,nd on the other hand a lower value of A/F increases hydrogeneneration. However, considering the relative effects of A/F ration the generated hydrogen and CO as well as the relative impor-ance of higher hydrogen generation versus lower CO content, itppears that A/F ratio must be selected in favor of higher hydrogeneneration. Thus, an A/F ratio of 1.0 may be suggested for the mostesirable synthesis gas product. One should note that the outletole fraction of hydrogen is more sensitive to A/F ratio than that of

O. For example, it is observed that at W/F of 3.0, Fig. 11 shows out-et hydrogen mole fractions of 38.5% and 14% at A/F ratios of 1.0 and.0, respectively, while these values are 5.8% and 2.2% for CO. Theossibility of coke formation may be of concern at low A/F ratios.owever, the suggested temperature range is not high enough toromote the possibility of such reactions.

Fig. 8 shows that both the outlet hydrogen and CO fractions arencreased with the feed gas temperature. However, the increase inydrogen generation is more considerable than that in CO (the for-er is 1.2% and the latter 0.55%). Therefore an inlet temperature of

00 ◦C is suggested for the given conditions to improve the quality

f the generated synthesis gas.

For the given range of parameters, the optimum operating con-itions, as well as the corresponding outlet mole fractions areummarized in Table 6. It is observed that at these conditions theutlet H2 and CO mole fractions are 38.9% and 4.8% (dry basis),

Outlet CO mole fraction (%) 4.8Outlet CO2 mole fraction (%) 21.6Outlet CH4 mole fraction (%) 13.1

respectively. The generated synthesis gas will need further treateddownstream of the ATR microreactor in separate WGS and PROXreactors before being fed to a low-power PEM fuel cell for a mobileapplication.

4. Conclusions

In this study, catalytic autothermal reforming of methane issimulated in a rectangular surface microreactor using a three-dimensional CFD model. The main global surface reactions areincluded in the model that consist of methane total combustion,steam reforming of methane to CO and CO2, and the water–gas shiftreaction. In addition, the gas-phase reaction of methane oxidationis considered in the simulations. The governing equations includeconservation of mass, momentum, energy and chemical species, aswell as the ideal gas mixture equation of state. A CFD code based onthe finite volume method was developed to solve the non-linear,coupled governing equations.

The microchannel is of a square cross section with 340-�msides and 8.5 mm length. Simulation results are first presented fora baseline case. The influence of the main contributing factors onthe performance of the autothermal microreactor is then investi-gated. These parameters include space velocity, air/fuel molar ratio,water/fuel molar ratio, and the feed gas temperature. The followingconclusions are drawn from the present study:

• Temperature variation along the microreactor is less than 4 ◦C;consequently, the process is effectively isothermal.

• The volumetric gas-phase reaction of methane oxidation is neg-ligible in comparison with the surface reactions.

• Hydrogen production reaches its maximum value at a spacevelocity of 50,000 h−1 for the given baseline conditions.

• Methane conversion increases with increasing A/F ratio or W/Fratio; it reaches its peak value at A/F ratio of 5.0 almost indepen-dent of W/F ratio.

• Hydrogen production is promoted by increasing W/F ratio ordecreasing A/F ratio.

• Increasing either A/F ratio or W/F ratio hinders CO generation.• The optimum operating conditions for the given range of param-

eters are: SV of 50,000 h−1, A/F ratio of 1.0, W/F ratio of 3.0, andfeed gas temperature of 600 ◦C.

• At these conditions, the outlet hydrogen and CO mole fractionsare 38.9% and 4.8% (dry basis), respectively.

Such ATR surface microreactor will be an essential componentin the development of a complete micro-fuel processor for fuel cellapplications in portable electronic devices.

Acknowledgment

The second author would like to thank Prof. H. Emdad for provid-ing insightful assistance during the development of the CFD code.

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