steam methane reforming reaction process intensification by using a reactor

Chemical Engineering Journal xxx (2012) xxx–xxx

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Chemical Engineering Journal

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Steam methane reforming reaction process intensification by using amillistructured reactor: Experimental setup and model validation for globalkinetic reaction rate estimation

M. Mbodji a,⇑, J.M. Commenge a, L. Falk a, D. Di Marco b, F. Rossignol b, L. Prost c, S. Valentin c,R. Joly c, P. Del-Gallo c

a Laboratoire Réactions et Génie des Procédés, CNRS-Université de Lorraine, ENSIC, 1 rue Grandville, BP 20451, 54001 NANCY Cedex, Franceb Laboratoire des Science des Procédés Céramiques et Traitements de Surface, 12 rue Atlantis, 87068 Limoges Cedex, Francec Air Liquide, Centre de Recherche Claude & Delorme, 1 chemin de la porte des Loges, BP 126, 78354 Jouy-En-Josas Cedex, France

h i g h l i g h t s

" Millistructured reactor is suitable for kinetic study of fast reactions." SMR process can be intensified with respect to energy efficiency and process size." SMR kinetics depending on catalyst microstructure is developed and validated." Highly-active Rh catalyst is suitable for industrial SMR process intensification." Hydraulic diameter of 400 lm is needed to suppress transport phenomena limitations.

a r t i c l e i n f o

Article history:Available online xxxx

Keywords:Microstructured reactorMethane reformingSyngasHydrogenProcess intensificationMicroreactor modelingKinetic data acquisition

1385-8947/$ - see front matter � 2012 Elsevier B.V. Ahttp://dx.doi.org/10.1016/j.cej.2012.07.117

⇑ Corresponding author.E-mail address: [email protected]

Please cite this article in press as: M. Mbodji etimental setup and model validation for global k

a b s t r a c t

In the frame of steam methane reforming process intensification, a highly active and stable catalyst basedon rhodium with catalyst formulation and structure adapted to millistructured reactors has been formu-lated. This catalyst has been tested in industrial conditions (800, 850 or 900 �C and 20 bars) on a singlechannel which is representative of one channel of a more complex millistructured SMR reactor. Then, adetailed mathematical model for acquisition of the global reaction kinetics with this new catalyst hasbeen developed and validated from experimental catalytic tests. The developed kinetics is dependentof the catalyst microstructure. This study presents the set-up, the model, the experimental catalytic runsand the global kinetics estimation protocol. It demonstrates, on one hand, that millistructured reactor issuitable for kinetic data acquisition and, on the other hand, the possibility of SMR process intensification,for improved energy efficiency and process size reduction.

� 2012 Elsevier B.V. All rights reserved.

1. Introduction

Steam methane reforming (SMR) of natural gas is the maincommercial process for synthesis gas production (H2, CO). In thisprocess, methane reacts with steam to produce a mixture of hydro-gen, carbon dioxide and carbon monoxide. This reaction is highlyendothermic and is performed in the presence of a catalyst suchas nickel or rhodium at high temperature (800–1000 �C), high pres-sure (20–40 bars) and steam-to-carbon ratio varying between 1.8and 4. In the classical process, a set of tubes filled with catalystis operated inside a furnace equipped with burners. These burnersprovide the heat needed for the reaction. The exit temperature of

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y.fr (M. Mbodji).

al., Steam methane reforminginetic reaction rate estimation

the process gas ranges from 700 to 950 �C. These conditions arelimited by the tube metallurgy. The reactor tube has a length of10–12 m and an internal diameter in the order of 10 cm. This pro-cess is well known and controlled. However, the overall efficiencyof the process is decreased by heat losses. The intensification of theSMR process by using microstructured reactors should enable onthe one hand to resolve this heat losses problem and on the otherhand to reduce substantially the size of process units, their ener-getic consumption and their environmental impact [1,2]. The highsurface-to-volume ratio of microstructured reactors provides ahighly efficient heat transfer and reduces the potential for temper-ature gradients in catalyst layers deposited on microchannel wallswhen performing highly endothermic reactions. Compared toconventional fixed-bed catalytic reactors, microstructured reactorsadvantages are considerable particularly in terms of yield,

reaction process intensification by using a millistructured reactor: Exper-, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117

http://dx.doi.org/10.1016/j.cej.2012.07.117

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Notations

Am active surface of active metal per mass of rhodium(m2

sma/grhodium)CT,g total concentration in the gas phase (mol/m3)Dh hydraulic diameter of the reactor (m)F0 total molar flow rate of reactants (mol/s)FCH4,0 molar methane flow rate at the reactor inlet (mol/s)Finert molar flow rate of inert species (mol/s)hloc heat-transfer coefficient between the gas and the walls

(W/m2 K)kd,j mass-transfer coefficient between the gas and the cata-

lytic wall (m/s)P total pressure (Pa)Pcg contact perimeter between the catalytic bar and the gas

(m)Pmg contact perimeter between the thermocouple and the

gas (m)Ps external reactor perimeter (m)Psc contact perimeter between the catalytic bar and the

reactor (m)Psg contact perimeter between the gas and the reactor walls

(m)R universal gas constant (J/mol K)r1 rate of the SMR reaction (mol/m2

SampleSurface s)

r2 rate of the WGS reaction (mol/m2SampleSurface s)

Rsc thermal resistance between the catalytic bar and thereactor (m2 K/W)

Rsm thermal resistance between the thermocouple and thereactor (m2 K/W)

Sco CO selectivity (–)Tg gas temperature (K)Ts reactor skin temperature (K)Tc catalyst temperature (K)uc gas velocity (m/s)XCH4 methane conversion (–)yg,j gas phase molar fraction of species j (–)yc,j molar fraction of species j in the catalyst (–)z axial position along the channel (m)

Greek notationsa ratio between the height and the width of the reactor (–

)l dynamic gas viscosity (Pa s)ti,j stoichiometric coefficient of species j in reaction i (–)k thermal conductivity of the gas (W/m/K)DrH850�C heat of reaction at 850 �C (J/mol)

Dimensionless numbersGzth thermal Graetz number (–)Gzm material Graetz number (–)Nu Nusselt number (–)Pr Prandtl number (–)Re Reynolds number (–)Sc Schmidt number (–)

2 M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx

selectivity to the desired product and safety. However, this changein production technology must be coupled to the development ofhighly active and stable new catalysts in order to ensure the sameconversion rate at lower residence times and catalyst formulationsadapted to microreactors. Microreactors are also characterized bythe use of very small reactants and catalyst quantities (usually inthe range 0.01–1 g); therefore, they appear to be a very good toolfor the acquisition of kinetic data and also for the determinationof catalyst behavior and activity [3,4].

1.1. Review of SMR reactors

Reactor miniaturization is known to improve heat and masstransfer, however, this strategy is not always sufficient for processintensification. Catalyst intensification is also needed to avoid hotspots [5]. Nickel is the most common industrial catalyst used forSMR owing to its robustness, its catalytic activity and its relativetolerance to poisons, such as sulfur, chloride, and heavy metals.Noble metals such as ruthenium, rhodium, and palladium are alsosuitable for SMR. Stefanidis and Vlachos [6] studied the intensifica-tion of steam reforming of natural gas and tested whether steamreforming on nickel is feasible by intensifying the process via min-iaturization. They found that the steam reforming reaction timescales for rhodium and nickel depend more on the reaction tem-perature than mixture composition. Over the temperature range1000–1500 K, the steam reforming on rhodium is faster than onnickel by a factor of 3 to 20. Below this range, steam reformingon rhodium is one order of magnitude faster than on nickel.

Zeppieri et al. [7] investigate the kinetics of methane steamreforming reaction over a rhodium–perovskite catalyst of formulaBaRhxZr(1�x)O3 at atmospheric pressure and in the temperaturerange 723–1023 K. Their results show that SMR reaction rate isfirst order with regard to methane and 0th order with regard tosteam. Methane conversion is proportional to the partial pressureof methane and the contact time. Results from Iglesia et al. [8] bear

Please cite this article in press as: M. Mbodji et al., Steam methane reformingimental setup and model validation for global kinetic reaction rate estimation

out these affirmations. Comparing the performances of a rhodium–perovskite catalyst to a commercial nickel-based catalyst, Zeppieriet al. [7] confirm that the rhodium–perovskite catalyst is the mostactive: high methane conversion close to the theoretical thermody-namic value is experimentally obtained with a low quantity of cat-alyst. Furthermore, carbon deposition is lower than on acommercial nickel-based catalyst.

Leventa et al. [9] studied SMR in a microreactor filled with anindustrial catalyst containing 15% nickel. Their experiments wereperformed in the temperature range 600–840 �C. The pressurerange was 2.5–9 bars, with hydrogen-to-methane ratio of 0.5–2and steam-to-methane ratio 2–3. They found that the increase ofH2-to-CH4 ratio in the feed enhances the catalyst activity. However,an increased steam-to-methane ratio in the feed moves the reform-ing reaction in the opposite direction. Steam acts as an inhibitor onthe catalyst activity and the reaction rate. They also observed thatthe smaller diameter of the microreactor enabled decreasing thecatalyst quantity and acquisition of reaction kinetics at hightemperatures up to 840 �C without reaching equilibrium.

Microreactors are increasingly used as tools for catalytic activitymeasurement. Peela et al. [10] studied steam reforming of ethanolover 2%Rh/20%CeO2/Al2O3 catalyst in a microchannel reactor. Theycompared microchannel reactor performance with that of apacked-bed reactor using 2%Rh/20%CeO2/Al2O3 catalyst at identicaloperating conditions and found the same activity but the selectiv-ity to desired product was higher in the microchannel reactor. TheH2 yield obtained in the microchannel reactor was 65 L/g/h as com-pared to 60 L/g/h in the packed-bed reactor. The high selectivity ofH2 is attributed to improved heat management in the microchan-nel, resulting in a more uniform temperature throughout the cata-lyst. The radial temperature gradient in the packed-bed reactor andin microchannel reactor by using 2D models for each type ofreactor was also investigated. The maximum temperature differ-ence in the packed bed reactor was about 15 K whereas that inthe microchannel reactor was only 0.3 K. These results show that



600 650 700 750 800 850

-100

-80

-60

-40

-20

0

20

40

Temperature [°C]

Rea

ctio

n G

ibbs

Fre

e E

nerg

y [k

J/m

ol]

SMRRMMethane crackingWGSBoudouardCO reduction

Fig. 1. Gibbs free energies of reactions (DrG) as a function of temperature.

M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx 3

microchannel reactors significantly reduce the temperature gradi-ent over the reactor cross-section due to their high heat-transfercoefficient.

Wang et al. [11] assessed methane steam reforming over Rh/MgO–Al2O3 catalysts in microchannel chemical reactors. Experi-mental results show that rhodium catalyst supported on MgO–Al2O3 is highly active and stable over a wide range of steam-to-car-bon ratios and resistant to coke formation. Methane steam reform-ing reaction rate on this catalyst in microchannel reactor wascompared to that of a conventional micro-tubular reactor. Resultsconfirm the performance enhancement in microchannel reactors.All of these studies show that currently methane steam reformingintensification is feasible. Indeed, reactor and catalyst intensifica-tion are increasingly controlled.

The present study is focused on syngas production by steammethane reforming in a millistructured reactor. Catalysts basedon Rh/Al2O3 enabling to reach high conversion at low residencetimes have been developed and tested. Experiments are conductedat 800 �C, 850 �C or 900 �C, 20 bars and a steam-to-methane ratioof 3. The main goal of this work is to determine SMR and WGSkinetics reactions rates from experimental catalytic tests. Theexperimental results coupled with a mathematical plug-flow reac-tor model taking into account heat and mass transfer between thereactant gas and the catalyst enables identification of the kineticparameters (activation energies and pre-exponential rate con-stants) of SMR reaction by minimizing the sum of squared differ-ence between measured methane conversion, outlet gastemperature and calculated values given by the reactor model.

1.2. Description of the reacting system

The following chemical reactions should be expected when per-forming steam methane reforming:

� the endothermic steam reforming reaction (SMR)

CH4 þH2O() 3H2 þ CO ð1Þ

� the reverse methanation reaction (RM)

CH4 þ 2H2O() 4H2 þ CO2 ð2Þ

� and the exothermic water gas shift (WGS)

COþH2O() H2 þ CO2 ð3Þ

The main drawback of SMR is the risk of carbon formation. Caremust be taken to avoid carbon formation due to:� methane cracking

CH4 () Cþ 2H2 ð4Þ

� the Boudouard reaction

2CO() Cþ CO2 ð5Þ

� and CO reduction

COþH2 () CþH2O ð6Þ

Table 1Reaction heats of steam reforming and carbon formation reactions.

Reaction Name DrH850�C (kJ/mol)

1 Steam reforming reaction 2262 Reverse methanation 1933 Water gas shift �334 Methane cracking 905 Boudouard �1696 CO reduction �135


In Table 1 are reported the heats of these six reaction. As steamreforming is the major reaction, the global system can be consid-ered as endothermic.

1.3. Thermodynamic analysis

The Gibbs free energies (DG) of the most significant reactionsoccurring during steam methane reforming are given as a functionof temperature in Fig. 1.

According to these thermodynamic data, in our operating con-ditions (above 800 �C and 20 bars), SMR, RM and methane crackingare the most favorable reactions. Carbon formation, harmful to theoperation of production units, can be limited by using an excess ofoxidizing agent as H2O. Carbon formation from methane crackingis catalyzed by chromium and iron and can be considered as aselectivity problem. It is usually resolved by using a catalyst anda reactor material on which carbon formation is unlikely.

1.4. Water gas shift reaction

The reactor model developed in this work takes into account theSMR and WGS reactions, which are the most-commonly consid-ered reactions when modeling steam methane reforming process.Thermodynamic analysis presented above shows that the watergas shift reaction is negligible in the operating conditions. Further-more, all experiments are carried out above 800 �C, and Gibbs freeenergy of the WGS reaction is positive for temperatures greaterthan 800 �C. The CO2 quantity recorded during catalytic tests wasnot significant, therefore only SMR reaction kinetic rate is studiedis this work.

Under these conditions, the estimation of the kinetics reactionrate then consists in finding the pre-exponential rate constantand the activation energy for the SMR reaction. The following sec-tions present the experimental set-up and the model developed fordata treatment.

2. Material and methods

2.1. Experimental test rig

The experimental set-up, on which catalytic tests have beenperformed, is shown in Fig. 2. In order to ensure a good mixtureof the reagents, a gas mixer and pre-heater is set before the reactor



Fig. 2. Picture of the experimental setup exhibiting the reactor in the open furnace.


entrance. The reactor consists of a rectangular channel (withdimensions of 5.5 mm width, 2.6 mm height and 47.75 cm length)within which a small bar of alumina (with dimensions of 5 mmwidth, 1.6 mm height and 20 cm length) coated with a rhodium-based catalyst is introduced. To provide the required heat to theendothermic reaction, the reactor is electrically heated. Two ovensare used to ensure the controlled heating of the system. CC2 de-notes the oven around the catalytic reactor itself, that containsthe catalytic coated sample and CC3 denotes the oven around thenon-catalytic part of the reactor. This oven CC3 is used to preventtemperature gradient in CC2 and provide controlled temperatureconditions to the reaction.

At the reactor outlet, the set-up is equipped with a condenserand a weighing system for measurement of the mass of condensedwater. In-line infrared analyzer is also used to analyze gasescomposition.

For gas temperature measurements, two thermocouples are setat the inlet and outlet of the reactor. Four thermocouples are alsoset on the reactor outer skin to measure the temperature profilealong the reactor. A mobile thermocouple is installed on the topwall of the inner channel in order to monitor the gas temperaturealong the reactor.

Fig. 3 illustrates a longitudinal view (Fig. 3a) and a cross-sectionalview (Fig. 3b) of the reactor with all geometric perimeters of interestthat will be considered in the model for heat and mass transfer.

For the reactor heat-transfer characterization, three heat-trans-fer parameters need to be determined:

� hloc: mean heat-transfer coefficient between the flowing gas andthe internal walls.� Rsm: contact thermal resistance between the thermocouple and

the reactor.

Fig. 3a. Longitudinal v


� Rsc: contact thermal resistance between the catalytic bar andthe reactor.

Experimental temperature measurements have been performedfor determining these three heat-transfer parameters, and will bediscussed further.

In this study, the part of conversion coming from the reactormetal alloy has been evaluated experimentally before and afterthe catalytic tests. Results showed that the non-catalytic reactoractivity has drastically evolved during the catalytic tests. Thus, afull reactor model considering the non-catalytic reactor activityand the catalyst activity has been developed. This is performedby coupling two reactor models in series. In the CC2 part of thereactor, the model considers the two active areas: the reactor wallsand the catalyst, whereas in the CC3 part, only the non-catalyticreactor activity of the walls is considered.

In order to estimate the fraction of conversion due to the reactorwalls, reactor activity is quantified by fitting experimental empty-reactor conversion after catalytic tests. To facilitate the readabilityof the present paper, the full reactor model is not presented in de-tails. Further, by using the full reactor model, it will be demon-strated that in the presence of catalyst sample, the non-catalyticreactor activity is negligible. The reactor model presented in thispaper concerns the CC2 part without considering the non-catalyticreactor activity.

2.1.1. Experimental catalytic tests conditionsThe following experimental conditions have been used to per-

form the catalytic tests. The reactor is fed with methane and steamwith a steam-to-carbon ratio of 3 at 800, 850 or 900 �C. The totalgas flow rate ranges from 0.0017 to 0.0079 mol/s in order to oper-ate with residence times between 40 and 200 ms in the CC2 part

iew of the reactor.



Mobile thermocouple: Tm

Reactor wall: Ts

Gas: Tg

Psm

Pmg

Psg

Ps

Mobile thermocouple: Tm

Catalyst sample: Tc

Reactor wall: Ts

Gas: Tg

Psm

Pmg

Psg

Pcg

Psc

Ps

Fig. 3b. Cross sectional view of the reactor.

Fig. 4. Catalyst holder.

Table 2Characteristics of tested catalysts.

Characteristics of tested catalysts Sample 1 Sample 2

Wash-coat thickness (lm) 2 10Am (m2

sma/grhodium) 80 235BET (m2/gcatalyst) 10 10Mass of catalyst (Mc) (mg) 4 22.4PR (rhodium quantity) (%) 20 1Dispersion (%) 18 53Rhodium particles size (nm) 6 2


where the catalyst sample is located. The residence time is com-puted at the reactor inlet as the ratio between the reactor volumeand the volume flow rate at the inlet gas temperature and pressure.The pressure for all experiments is set at 20 bars.

At the reactor exit, the gas is quickly cooled and passed througha gas–liquid separator where the unreacted water is separated andweighed by means of a weighing machine. The dry gas compositionis then determined by on-line infrared analyzer.

Carbon and hydrogen balances are carried out to check massbalances and to detect potential coke formation. For a given resi-dence time and temperature level (800, 850 or 900 �C), measure-ments have been performed during 72 h under reactionconditions (75% H2O, 25% CH4, and 20 bars). No catalyst deactiva-tion has been observed. Variable XCH4 represents the methane con-version and SCO the CO selectivity determined from the change ingas composition.

XCH4 ¼FCH4;0 � FCH4

FCH4;0

ð7Þ

SCO ¼FCO

FCO þ FCO2

ð8Þ

FCH4,0, denotes the methane flow rate at the reactor inlet. FCH4, FCO

and FCO2 respectively denote the methane, carbon monoxide andcarbon dioxide flow rates along the reactor.

2.1.2. Synthesis of catalystsCatalysts are made of rhodium metallic active nanoparticles

dispersed onto a commercial magnesium aluminate powder. First,the powder is treated by attrition, then it is impregnated with anexcess of aqueous rhodium nitrates solution. The mass of rhodiumnitrates is calculated to 20 wt.% rhodium in the final product forthe first sample and 1 wt.% for the second sample. The impregna-tion is conducted under heating at 150 �C and steering until wateris completely evaporated. Residues obtained are finally calcined inair to form the catalyst phase. For the experimental study, catalystsare deposited as layers with a thickness less than 12 lm on alu-mina substrates by dip coating.

2.1.3. CharacterizationThe morphology and the thickness of catalysts layers have been

evaluated using a Zeiss Ultra-55 scanning electron microscope be-fore and after ageing in a steam methane reforming atmosphere at850 �C. Samples have been observed at three different locations ofthe substrate: bottom, middle and head.

Temperature-programmed reduction and chemisorption mea-surements have been carried out on a Micromeritics AutoChem II2920 and an Asap 2020 on the catalyst powder before dip-coatingto control the catalyst activity.

2.1.4. Characteristics of catalyst samplesAn example of the tested catalyst holders is presented in Fig. 4.

The characteristics of the catalyst samples are summarized inTable 2.


2.2. Reactor model for kinetics study

In this section, the one-dimensional plug-flow reactor modeldeveloped for the reactor simulation is presented. As indicatedabove, the non-catalytic reactor activity is not considered in themodel version described in this paper. This model takes into ac-count the SMR and WGS reactions and will be used for kineticparameters identification from experimental tests. The assump-tions detailed below have been considered.

� Plug flow of the reactant gas.� The behavior of gases is modeled by the ideal gas law.� There is no reaction in homogeneous phase.� Reactions occur on the surface of the wash-coat deposited on

the sample holder.� There is no limitation by internal transfer in the wash-coat.� There is no heat transfer by radiation.

To describe the concentration and temperature profiles alongthe reactor, a one-dimensional plug-flow model including heatand mass transfer between the reactant gas and the catalyst hasbeen developed. The catalyst is supposed to be uniformly coatedon the catalytic bar. The heat is also provided uniformly throughall the walls of the reactor. The way this heat is transferred tothe reactants is modeled by a usual convection transfer law. Thespecific heat flux is one of the model parameters. All the heatand mass balances described below are written under steady-stateconditions. Mass-transfer coefficients are introduced in the modelto account for the species transport limitation between the bulkgas mixture and the catalytic active surface. As demonstrated byMladenov et al. [12], the introduction of mass-transfer coefficientsin the plug-flow reactor model improves its accuracy. However,they have to be used with caution since they are often based onempirical correlations. In this study, the external mass-transfer



Table 3Kinetic models of steam reforming of hydrocarbons.

Reference Form of the kinetics law

Bodrov [14] Langmuir–HinshelwoodKohmenko et al. [15] Temkin IdentityRostrup-Nielsen [16] Two-step kinetics, power lawTottrup [17] Pellet kinetics, power lawXu and Froment [18] Langmuir–HinshelwoodAparaicio [19] Microkinetic analysis


coefficient was evaluated by numerical simulation using FLUENT�,and will be presented further.

The concentration evolution of the considered species j (j = CH4,H2O, CO, H2, CO2) in the gas phase results from the gas convectionand the reaction at the catalytic wall. The net flux of component jfrom the bulk fluid to the wall is composed of a classical convecto-diffusive term and a net flux due to the reaction stoichiometry.Combining the mass balance for each species to the overall massbalance enables to describe the evolution of the gas phase compo-sition as:

dðyg;jÞdz

¼ kd;jPscP=ðRTgÞ

F0 þ Finert þ 2X1FCH4 ;0

� �ðyc;j � yg;jÞ ð9Þ

where yg,j denotes the gas phase molar fraction of species j, yc,j themolar fraction of species j in the catalyst, z the axial position alongthe channel, kd,j the mass-transfer coefficient between the gas andthe catalytic wall, P the total pressure, Tg the gas temperature, Finert

the molar flow rate of inert species, F0 the total molar flow rate ofreactants, and Psc the contact perimeter between the catalyst sur-face and the gas.

The mass balance in the catalyst layer is written as the equalitybetween the molar flux transferred from the gas and the flux con-sumed by chemical reactions:

kd;jCT;gðyg;j � yc;jÞ þ m1j � 2yg;j

� �r1 þ m2jr2 ¼ 0 ð10Þ

where r1 and r2 are the reaction rates of SMR and WGS 1 and 2,respectively. CT,g is the total concentration in the gas phase.

A heat balance on the gas phase enables to describe the evolu-tion of the gas temperature along the reactor with the followingrelation:

ucXcCT;gCpgdTg

dzþU ¼ 0 ð11Þ

where U denotes the heat transferred by convection between thegas and the catalyst sample, the gas and the mobile thermocoupleand the gas and the reactor walls.

U ¼ hloc½PcgðTg � TcÞ þ PsgðTg � TsÞ þ PmgðTg � TmÞ�

where hloc denotes the mean local heat-transfer coefficient by con-vection between the gas and the reactor internal elements. Pcg, Psg

and Pmg respectively correspond to the contact perimeters betweenthe gas and the catalyst, the gas and the reactor walls, and the gasand the mobile thermocouple.

Within the catalyst layer, the enthalpy balance is written byequalizing the heat flux provided by the furnace, the flux ex-changed with the gas phase, the source term related to the reform-ing reaction of methane and the Water Gas Shift reaction:Psc

RscðTc � TsÞ þ hlocPcgðTc � TgÞ þ r1DrH1Psc þ r2DrH2Psc ¼ 0 ð12Þ

where DrH1 and DrH2 denote the heat of SMR and WGS reactions,respectively.

To describe the pressure drop under laminar flow conditions,Shah and London [13] correlation is used:

dPdz¼ f Re

2luc

D2h

ð13Þ

f Re ¼ 24ð1� 1;3553aþ 1;9467a2 � 1;7012a3 þ 0;9564a4

� 0;2537a5Þwhere a denotes ratio between the height and the width of thereactor.

A heat balance on the mobile thermocouple and on the reactorenables to describe their temperature profiles as:

Psm

RsmðTs � TmÞ � hlocPmgðTm � TgÞ ¼ 0 ð14Þ


uPS ¼ hlocPsgðTs � TgÞ þPsm

RsmðTs � TmÞ þ

Psc

RscðTs � TcÞ

� �ð15Þ

where Ps denotes the external reactor perimeter and u the specificheat flux received by the reactor. This flux is one of the modelparameters.

2.3. Heat losses on the experimental device and boundary conditions ofthe reactor models

For a good estimation of kinetic parameters, it is very importantto know accurately, for each experiment, the heat received by thecatalytic surface. Experimental tests have been conducted in orderto estimate the heat losses in the experimental device. Resultsshow that heat losses are very large depending on the reactantgas residence time: heat losses range from 80% to 93% of the totalexperimental heat flux furnished by the electrical heat. Despitethese large values, it is still possible to determine the experimentalheat flux consumed by the endothermic SMR reaction for each test,by performing a heat balance based on the inlet and the outlet gastemperature and methane conversion.

In the model presented here, the specific heat flux u is consid-ered as the thermal boundary conditions. It is also possible to setthe experimental reactor temperature as a boundary condition.This can be done by replacing the Eq. (15) by the experimentalreactor temperature.

In the CC2 reactor part, several thermocouples provide theexperimental reactor temperature profile. For the CC3 part wherethere is no reactor thermocouple and catalyst sample, the reactortemperature is assumed to be equal to the temperature measuredby the mobile thermocouple. The kinetic parameters estimated foreach boundary condition will be compared.

2.4. Reaction rates

The kinetics of steam methane reforming reaction has beenstudied extensively by several groups. There is a general agree-ment on the first order kinetics with respect to methane, but theactivation energies vary between 20–160 kJ/mol. These differencesmight be explained by experimental inaccuracies due to transportrestrictions in the sense of diffusion and heat restrictions. Whilethe exact mechanism of the steam methane reforming reaction isstill under debate today, the most important steps are: (1) decom-position of methane on a metal surface to hydrocarbon fragmentsand carbon atoms, (2) dissociative adsorption of water to H and OHspecies (3) OH or O species combine with C to form CO. Several ki-netic studies on SMR reaction can be found in the literature. Table 3presents some of them [14–19].

Recently, Wei and Iglesia [20] proposed a simple equation forkinetics of steam methane reforming. They found that the activityat 600–700 �C only depends on the partial pressure of methane,implying that the rate determining step is the initial activation ofa C–H bond in methane.

The work presented in this paper is part of a preliminary designapproach of a millistructured reactor heat-exchanger for theproduction of syngas. To reach that goal, the reaction rate of the



0 0.02 0.04 0.06 0.08 0.1 0.12 0.14 0.16 0.18 0.2820

830

840

850

860

Reactor length [m]

Tem

pera

ture

[°C

]

CC2 part

Reactor Mobile thermocouple

0.2 0.25 0.3 0.35 0.4 0.45650

700

750

800

850

Reactor length [m]

Tem

pera

ture

[°C

]

CC3 part

Mobile thermocouple

Fig. 5. Experimental temperature profiles along the reactor with active catalystsample in the CC2 part.


main reaction (steam methane reforming: SMR) has to be mea-sured precisely in the same operating conditions as the future reac-tor. Therefore, it must be emphasized that a global kinetics modelis more appropriate than the microkinetics of the SMR reactionwith detailed reaction mechanism and determination of the limit-ing step. Such a lumped kinetics reaction rate will enable to prop-erly design a milli-structured exchanger reactor for syngasproduction at the industrial scale.

In order to express this overall reaction rate, the same formal-ism as used by Tonkovich et al. [21] to describe SMR reaction overa rhodium on Mg-spinel catalyst, is adapted in the model by addinga constant depending on the catalyst microstructure. Withoutgoing into details, the SMR and WGS reactions rates can be writtenas:

r1 ¼ Kpre exp 1 exp � Ea1

RTc

� �yc;CH4

yc;H2O � P2 yc;COy3c;H2

Keq1

!Kl

r2 ¼ Kpre exp 2 exp � Ea2

RTc

� �yc;COyc;H2O �

yc;CO2yc;H2

Keq2

� �Kl

where the reaction rates r1 and r2 are expressed in [mol/m2

SampleSurface/s].Kl is a constant depending on the catalyst microstructure

[m2ActiveMetal=m2

surface of holder] and might be expressed by first approx-imation as:

Kl ¼AmMcPR

PscL

where Am denotes the active surface of active metal per unit of massof rhodium (m2

sma/grhodium), Mc the mass of catalyst, PR the rhodiumquantity in the catalyst, (Psc � L) the surface of the holder on whichthe catalyst sample is coated.

Ea1 and Ea2 (J/mol) denote the activation energy of the SMR andWGS reactions, respectively.

Kpreexp1 and Kpreexp2 [mol/m2ActiveMetal/s] denote the pre-exponen-

tial rate constants of SMR and WGS reactions, respectively.

Keq1¼101;3252

�exp �26;830Tc

þ30:114� �

Equilibrium constant of SMR ½Pa2�

Keq2¼exp4400

Tc�4:036

� �Equilibrium constant of WGS reaction ½—�

As said previously, the WGS reaction can be neglected under theoperating conditions of this study. Estimation of the kinetics reac-tion rate then consists in finding the pre-exponential rate constantand the activation energy of the SMR reaction.

2.5. Reactor model resolution

The reactor model is a set of differential and algebraic equa-tions. The solver function ode15s available on MATLAB� is usedto solve this system. This solver uses the Gear method which isadapted to the resolution of stiff systems. After integration of theset of equations, mass and overall enthalpy balances are computedand satisfied with less than 0.001% and 1%, respectively.

From a purely numerical point of view, the reactor model de-scribed above is complete and ready to be computed. However,to make it more reliable and to correctly reproduce the experimen-tal results, it is useful to provide a good description of the heat andmass-transfer phenomena that occur between the reactant gas andthe catalyst. In the following section, heat and mass transfer in thereactor are investigated.


2.6. Heat transfer in the reactor

Microchannels are known as apparatuses that provide highheat-transfer coefficients, which makes them particularly suitablefor kinetics studies. Their remarkable heat-transfer capacity isdue to the fact that the heat-transfer coefficient is proportionalto the reciprocal of the channel hydraulic diameter.

An example of experimental temperature profiles along thereactor during catalytic tests in the presence of active catalyst sam-ple is shown in Fig. 5. In the CC2 part containing the catalyst sam-ple, the reactor temperature decreases along the reactor length as aresult of the heat consumption by the endothermic steam reform-ing reaction. From the end of the CC2 part to the middle of CC3, thereactor temperature is constant. This can be explained by the factthat heat consumption by SMR reaction is not significant becausemost of the methane quantity has already been consumed by reac-tion in the CC2 part. Reactor temperature decreases at the end ofthe CC3 part due to heat losses by natural convection betweenthe experimental setup and the surrounding air.

The reaction kinetics can be determined from the comparisonbetween the experimental tests and the calculated results com-puted from the heat and mass balances presented above. The res-olution of the system requires the knowledge of specificparameters as the following three heat-transfer parameters de-fined above: hloc, Rsm and Rsc.

These parameters have been determined from experimentaltemperatures measurements without chemical reaction.

For these tests, the reactor is fed with nitrogen. A heat flux is setat the external reactor walls. Once thermal steady state is reached,the experimental temperatures of the gas, the mobile thermocou-ple and the reactor are recorded. Heat balances on the gas, on themobile thermocouple and on the catalytic bar than enable to relatetheir temperatures as follows:

Heat balance on the gas:

dTg

dz¼ 1

FinertCT;g

PmgPmg

PsmRsm þ 1

hloc

þ hlocPsg þPcg

Pcg

PscRsc þ 1

hloc

24

35ðTs � TgÞ ð16Þ

Temperature of the mobile thermocouple

Tm ¼Ts þ

hlocPmgRsm

PsmTg

1þ hlocPmgRsm

Psm

ð17Þ



Table 4Experimental heat-transfer parameters.

Heat-transfer parameter hloc (W/m2 K) Rsm (m2 K/W) Rsc (m2 K/W)

Average value 442 0.0038 0.013Standard deviation 65 0.0005 0.016


Temperature on the catalytic surface

Tc ¼Ts þ

hlocPcgRsc

PscTg

1þ hlocPcgRsc

Psc

ð18Þ

It is important to point out here the fact that when both thermalresistances Rsm and Rsc are set equal to 0, the temperature of the cat-alytic bar and that of the mobile thermocouple are equal to thereactor temperature i.e. Tm = Tc = Ts. The three heat-transfer param-eters have been identified by minimizing the sum of squared differ-ences between measured and modeled temperatures. Fig. 6 showsan example of temperature profiles along the reactor after heat-transfer parameters identification. A mean difference of 2 �C be-tween modeled and experimental temperature has been obtained.

For all tests, the average values of the mean heat-transfer coef-ficient between the flowing gas and the reactor internal elements(catalytic bar, mobile thermocouple and reactor walls), the thermalresistance between the thermocouple and the reactor, and thethermal resistance between the catalytic bar and the reactor andtheir standard deviation are presented in Table 4. The relativelylarge error on the resistance Rsc indicates that this parameter isnot properly estimated, since the standard deviation on Rsc is largerthan its average value. This result is not surprising since there is nodirect temperature measurement on the catalytic bar. Convectiongas heat-transfer coefficient in microreactors depends on thehydraulic diameter and on the gas composition and usually rangesfrom 400 to 2000 W/m2. Kays and Crawford [22] proposed the fol-lowing correlation to estimate Nusselt number for fully-developedlaminar flow in rectangular ducts with constant heat fluxcondition:

Nu ¼ 8:235ð1� 1:883aþ 3:767a2 � 5:814a3 þ 5:361a4 � 2a5Þ

where a is ratio between the height and the width of the reactor,and Nu the Nusselt number defined as:

Nu ¼ hlocDh

k

For a mean temperature of 400 �C, hloc calculated from Kays andCrawford [22] correlation is 235 W/m2 K. The experimental resultsgive an average value of 442 W/m2 K with a standard deviation of65 W/m2 K. This difference can be explained by the fact that hloc

from the experimental tests is an average value from several exper-iments and by the assumption that considers that hloc is constant

0 0.02 0.04 0.06 0.08 0.1 0.12 0.14 0.16 0.18 0.2350

360

370

380

390

400

410

420

430

440

Tem

pera

ture

[°

C]

Reactor length [m]

Gas ModelOutlet gas Experimental

Mobile thermocouple Model

Mobile thermocouple Experimental

Reactor ExperimentalCatalytic bar Model

Fig. 6. Temperature profiles in the reactor without reaction for calibration of heat-transfer.


along the reactor. It is well known that hloc is constant only forfully-developed flows.

Thermal resistances, required for the reactor model accuracy,are also estimated by experimental temperature measurements.However, in order to have a good description of the heat transferin the reactor model, simulations with the commercial CFD pack-age FLUENT� have been performed without chemical reaction bysetting a constant wall-temperature boundary condition and feed-ing a N2 flow at inlet temperature. For each simulation, the localNusselt number is computed with the following relation:

Nu ¼ QDh

ðTw � TgÞk

where Q denotes the heat flux at the wall, Dh the hydraulic diame-ter, Tw and Tg the wall temperature and the mass-averaged gastemperature, respectively. Several simulations have been per-formed by varying the gas inlet velocity. They enabled descriptionof the Nusselt variation as a function of the thermal Graetz numberalong the reactor with this following relation:

Nu ¼ 4:58 expð0:003GzthÞ with Gzth ¼RePrDh

z

The variation of the Nusselt number as a function of the thermalGraetz number along the reactor is shown in Fig. 7. When the flowis fully developed, the Nusselt number tends towards a constant va-lue of 4.58. This result is in agreement with those obtained by Kaysand London [23]. The slight difference can be explained by reactorcross section that is slightly different from a perfect rectangle dueto the presence of the mobile thermocouple (see Fig. 3b).

2.7. Mass transfer in the reactor

Mass-transfer coefficients must also be preliminary determinedto study the kinetics of the reaction. It is difficult to perform reli-able mass-transfer coefficient measurement in microdevices andmass transfer has been evaluated here by CFD simulations using

0 0.05 0.1 0.15 0.2 0.25

4

6

8

10

12

14

1 / Graetz Number [-]

Nus

selt

Num

ber

[-]

CFD results with Fluent Nu = 4.58exp(0.003Gzth)

Fig. 7. Nusselt variation as a function of the thermal Graetz number along thereactor.



Inlet Oulet

Uniform wall temperature

Adiabatic wall temperatureTemperature profile

Fig. 8. Boundary conditions for mass-transfer study.


FLUENT�. To perform these simulations, the heat and mass transferare assumed analogous. Thus, numerical heat-transfer simulationswith appropriate boundary conditions have been performed in or-der to find a correlation which describes the analogous mass trans-fer. For simulation of the catalytic wall, a uniform wall temperatureboundary condition is used. As there is no mass transfer on theother walls, adiabatic boundary conditions are used (see Fig. 8).Male et al. [24] investigated mass transfer in a microreactor byusing a similar method.

Several simulations have been performed by varying the N2

velocity inlet. Sherwood number is computed by using the follow-ing relation:

Sh � Nu ¼ QDh

ðTc � TgÞk

Sherwood number at the reactor entrance varies strongly as a func-tion of the gas inlet velocity. However, for all simulations, the Sher-wood number tends towards an asymptotic value of 3.99 which isin very good agreement with literature. Indeed, Kays and London[23] reported an average Nusselt number of 3.9 in the case of a rect-angular channel having the same aspect ratio a (heigh/width) andboundary conditions. To consider the entrance effects on themass-transfer, the material Graetz number is introduced:

Gzm ¼ReScDh

z

Fig. 9 depicts the Sherwood number variation as a function of thematerial Graetz number.

When the material Graetz number is less than 10, the Sherwoodnumber is constant and tends towards its limiting value 3.99. ForGraetz numbers above 10, the entrance effects cannot be ne-glected. From these simulations results, the following correlation

0 0.05 0.1 0.15 0.2 0.25 0.3 0.35 0.4 0.45 0.5

4

6

8

10

12

14

1 / Graetz material Number [-]

She

rwoo

d N

umbe

r [-

]

CFD results with FluentSh = 3.97*exp(0.0023*Gzm)

Fig. 9. Sherwood number variation as a function of the material Graetz number.


has been established between Sherwood number and the materialGraetz number.

Sh ¼ 3:97 expð0:0023GzmÞ

This correlation is used in the reactor model to represent the exter-nal mass transfer between the gas and the catalytic surface.

3. Kinetic parameters identification

For each catalytic test, the gas phase molar fraction of H2, CO,CO2 and CH4 and the outlet gas temperature are measured and re-corded. The kinetic parameters are estimated by minimizing thesum of the squared difference F between measured methane con-version, outlet gas temperature and calculated values given by thereactor model. The following function F is minimized:

F ¼X Xmodel

CH4� Xexperiment

CH4

XexperimentCH4

!2

þTmodel

g � Texperimentg

Texperimentg

!20@

1A

Kinetic parameters determination then consists in solving a nonlin-ear optimization problem without constraints. The function Fmin-search available in MATLAB� optimization toolbox, based on theSIMPLEX method, is used to find kinetic parameters (activationenergies and pre-exponential rate constants). This parametric opti-mization is performed simultaneously on several experiments con-ducted at different residence times and temperature levels.

4. Impact of the non-catalytic reactor activity on the overallmethane conversion

In order to properly determine the reaction kinetics, we mustensure that the activity of the metallic walls of the reactor, esti-mated by the methane conversion, is negligible compared to theactivity of the catalyst holder. Therefore, experimental tests withan inert holder have been carried out before and after the catalytictests. Fig. 10 depicts the non-catalytic reactor conversion in pres-ence of an inert catalyst sample before and after the tests. It isimportant to precise here, that residence time is computed alongthe reactor by considering the CC2 and CC3 part.

One can note that the non-catalytic reactor activity has drasti-cally evolved during the catalytic tests. After several tests, the reac-tor intrinsic activity increases and is not negligible compared tosome experiments with active catalyst sample.

100 150 200 250 300 350 400 450 500 550 600 6500

10

20

30

40

50

60

70

80

90

Residence time [ms]

Non

-cat

alyt

ic r

eact

or a

ctiv

ity a

t 850

°C

Methane conversion before catalytic tests Methane conversion after catalytic tests

Fig. 10. Non-catalytic reactor conversion without catalyst before and after thecatalytic tests.



0.3 0.4 0.5 0.6 0.70.25

0.3

0.35

0.4

0.45

0.5

0.55

0.6

0.65

0.7

0.75

Model [-]

Exp

erim

ent [

-]

Methane conversion + or - 2 %

840 845 850840

841

842

843

844

845

846

847

848

849

850

Model [-]

Exp

erim

ent [

-]

Outlet gas temperature + or - 5°C

Fig. 11. Comparison between model and experiment results for the non-catalytic reactor activity. conversion (left) and outlet gas temperature (right).


As described above, to evaluate the influence of this non-catalytic activity, the full reactor model taking into account thenon-catalytic reactor activity has been used. The full reactor modelconsists in coupling two reactors in series. In the CC2 part, thereactor model considers the two active areas: the reactor wallsand the catalyst. In the CC3 part, only the non-catalytic reactoractivity is considered.

50 100 15020

30

40

50

60

70

80

90

Residence time [ms]

CH

4 Con

vers

ion

[%]

800°C

sample 1 sample 2 thermodynamic equilibrium

5020

30

40

50

60

70

80

90

Residenc

85

Fig. 12. Methane conversion as a function of


The activity of the non-catalytic walls of the reactor is quanti-fied by fitting experimental reactor activity after the catalytic tests.Comparison between model and experimental results for the non-catalytic reactor activity is shown in Fig. 11. The calculated valuesof the methane conversion and the gas temperature are in verygood agreement with the experimental ones.

100 150e time [ms]

0°C

50 100 15020

30

40

50

60

70

80

90

Residence time [ms]

900°C

the residence time and the temperature.



Table 5Kinetic parameters for SMR reaction by considering the non-catalytic reactor activity and comparison between model and experimental results.

Residence time (ms) Methane conversion (%) Outlet gas temperature (�C) Part of methane conversion due to the catalyst sample (%)

Model Experiment Model Experiment

40 Sample 2 (800 �C) 33 35 816 793 9060 40 41 807 795 94

100 51 51 806 799 97147 60 60 815 801 99

40 Sample 1 (800 �C) 30 30 776 777 9560 37 35 768 780 97

100 45 42 764 787 98147 50 50 785 791 99

40 Sample 1 (850 �C) 43 43 845 849 8660 51 50 840 852 91

100 65 64 852 856 96147 70 69 852 858 98

0 0.1 0.2 0.3 0.4 0.50

0.05

0.1

Reactor length [m]

Rea

ctio

n ra

tes

[mol

/m2 /s

]

Inert catalyst sample

CatalystReactor CC2 partReactor CC3 part

0 0.1 0.2 0.3 0.4 0.5

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

Reactor length [m]

Rea

ctio

n ra

tes

[mol

/m2 /s

]

Active catalyst sample

CatalystReactor CC2 partReactor CC3 part

Fig. 13. Comparison between SMR kinetic reaction rate on the catalyst sample and on the non-catalytic walls reactor (temperature conditions, illustrated in Fig. 5).


5. Results

5.1. Experimental results of the catalytic tests with catalyst sample

Here are presented the first tested samples. These samples en-able to validate the determination of the kinetic parameters. Boththese samples can be distinguished by their wash-coat thickness,rhodium quantity, rhodium particle size and dispersion. They havebeen tested at 800, 850 and 900 �C and for residence times be-tween 40 and 150 ms.

Fig. 12 shows methane conversion as a function of the residencetime and temperature for samples 1 and 2. Methane conversion in-creases with increasing residence time and/or temperature. How-ever, sample 2 is more active than sample 1 despite the fact thatsample 2 has less rhodium quantity. Indeed methane conversionwith sample 2 is greater than that obtained by sample 1. The goodperformances of sample 2 could be explained by the good disper-sion of the rhodium particles. Furthermore, the small rhodium par-ticle size provides to the catalyst a high specific surface.


5.2. Kinetics parameters identification taking into account the non-catalytic reactor activity

In order to quantify the kinetic reaction rate of the SMR reactionon the catalyst, the full reactor model considering the non-catalyticactivity of the reactor walls and the catalyst activity of the holderhas been used. The kinetics of the SMR reaction specific to thereactor walls was already evaluated in the presence of an inertsample.

Then, by using the full reactor model with several active areas,the kinetic parameters (Kpreexp1 and Ea1) of the catalytic SMR reac-tion can be identified. Table 5 shows the estimated kinetic param-eters, the part of conversion due to the catalyst sample and acomparison between experiment and model results in terms ofmethane conversion and outlet gas temperature. These resultsare obtained by setting a constant heat flux at the reactor walls.One can note a good agreement between model and experimentalresults despite the experimental measurements uncertainties.

The kinetics constants of the methane steam reforming are:



Table 6Comparison of identified kinetic parameters with literature results.

Kinetic parameters Kpreexp1 (mol/m2sma/s) Ea1 (J/mol)

This workBy considering the non-catalytic

reactor activity9.47 � 107 166,310

Without considering the non-catalyticreactor activity

1.68 � 108 165,740

Tonkovich et al. [21]Kinetic parameters Kpreexp1 (mol/m3

catalyst/s) Ea1 (J/mol)

1.275 � 108 169,500


� Kpreexp1 = 9.47 � 107 mol/m2sma/s.

� Ea1 = 166,310 J/mol.

Methane conversion due to the catalyst sample increases withthe residence times and the catalyst activity. Otherwise, these re-sults show that when an active catalyst sample is placed insidethe reactor, all methane conversion is due to the catalyst samplewhich confirms that the non-catalytic reactor activity is then neg-ligible. This can be explained by the fact that the catalyst sample islocated at the reactor entrance and its activation energy is almosttwice less than the non-catalytic reactor activation energy.

Experimental heat received by the reactor during the catalytictests was difficult to estimate due to the large heat losses. As pre-sented in the model equations, it is also possible to consider theexperimental heat provided to catalytic holder for the endothermicreaction from the thermal balance based on methane conversion,inlet and outlet gas temperature. The reactor model has been im-proved in order to avoid these uncertainties by setting an experi-mental reactor temperature as the thermal boundary condition.The kinetics values obtained with the new boundary conditionsare similar to those obtained with the first boundary condition.

SMR kinetic reaction rate on the catalyst and on the non-cata-lytic reactor walls is shown in Fig. 13. When an inert catalyst sam-ple is located in the CC2 part of the reactor, the SMR reaction rateon the catalyst is equal to 0 and all the methane conversion is dueto the non-catalytic reactor activity. By contrast, when an activecatalyst sample is used, methane conversion due to the non-cata-lytic reactor activity can be considered as negligible. Dependingon residence time and temperature level, methane conversiondue to the catalyst activity ranges from 86 to 99% of the overallmethane conversion.

Kinetic parameters were also estimated by considering that allmethane conversion is due to the catalyst i.e. by using the reactormodel presented previously. As can be seen in Fig. 14, experimen-tal results and model-predicted values are in perfect agreement.

The kinetics constants of the methane steam reforming are:

20 30 40 50 60 7020

25

30

35

40

45

50

55

60

65

70

75

Model

Exp

erim

ent

Methane conversion [%] + or - + 2 %

Fig. 14. Comparison between model and experiment by considering that all methane consample 2. Conversion (left) and outlet gas temperature (right).


� Kpreexp1 = 1.68 � 108 mol/m2sma/s.

� Ea1 = 165,740 J/mol.

6. Comparison with literature results

Tonkovich et al. [21] conducted steam methane reforming reac-tion by using a rhodium on Mg-spinel catalyst, and estimated theSMR kinetic reaction rate by fitting kinetic data. Their kineticparameters and those obtained in this work are summarized in Ta-ble 6. The pre-exponential constants are not directly comparable,due to the difference in the kinetic formulation. However, the acti-vation energies are in the same order of magnitude.

7. Discussion

The detailed mathematical model for acquisition of kinetic datadeveloped in this work enabled to find CH4 reforming kinetic reac-tion rate. A very good agreement between model and experimentalresults has been obtained. However, we noted some difficulties toestimate kinetic parameters from experimental tests conducted at900 �C on sample 1, 850 and 900 �C on sample 2. This is explainedby heat and external mass transfer limitations which appear when

760 780 800 820 840 860

760

780

800

820

840

860

Model

Exp

erim

ent

Outlet gas temperature [°C] + or - 10°C

version is due to the catalyst for tests at 800 and 850 �C on sample 1 and 800 �C on



0 0.05 0.1 0.15 0.210-1

100

101

102

Reactor length [m]

Cha

ract

eris

tic ti

me

[ms]

Hydraulic diameter 1 mm

ReactionExternal mass transfer

0 0.05 0.1 0.15 0.210-1

100

101

102

Reactor length [m] C

hara

cter

istic

tim

e [m

s]

Hydraulic diameter 0.4 mm

ReactionExternal mass transfer

Fig. 15. Characteristic times analysis.


performing SMR reaction on a highly active catalyst at high tem-perature in a microchannel reactor having a large hydraulic diam-eter (>1 mm). Further tests are required with reactor hydraulicdiameter below 400 lm for reduction of the heat and mass-trans-fer limitations.

Characteristic times of SMR reaction and external mass transferhave been investigated and are shown in Fig. 15. A steam-to-car-bon ratio of 3 has been used. The reactant gas temperature rangesfrom 650 �C to 900 �C. Characteristic times of reaction and externalmass transfer decrease along the reactor due to the increasing tem-perature. For a reactor with a hydraulic diameter of 1 mm, thereaction and external mass-transfer characteristic times are inthe same order of magnitude for temperatures near 780 �C. Fortemperatures greater than this value, heat and/or external mass-transfer limitations appear and become more and more pro-nounced when increasing temperature. A similar result was foundby Arzamendi et al. [25] who investigated steam methane reform-ing intensification by using a squared monolith channel and a nick-el-based catalyst. By varying channel sides between 0.35–2.8 mm,they showed that 0.7 mm is a sufficiently low dimension for SMRprocess intensification. In the case of a rhodium-based catalystwhich is more active than the nickel based catalyst, results showedthat, to eliminate heat or mass-transfer limitations and for processintensification, it is needed to use a hydraulic diameter below0.4 mm. The final module design still has to be chosen after eco-nomic assessment and by considering additional technical aspectsrelated to the mechanical resistance of the device or the manufac-turing of the microstructured system or the possibilities for cata-lyst coating inside the reactor.

The full reactor model enabled to estimate kinetic reaction rateof SMR from the catalytic tests in spite of the reactor activity andcomplex heat management inside the reactor. Currently, the sametests are conducted on a reactor coated with alumina in order tosuppress reactor activity. The fact that the part of methane conver-sion coming from the non-catalytic reactor is negligible in thepresence of active catalyst sample is studied experimentally andwill be the subject of another publication.


8. Conclusions

Steam methane reforming process intensification by using amillistructured reactor and a rhodium-based catalyst has beeninvestigated in this work. A detailed mathematical model for ki-netic reaction rate measurement from experimental catalytic testshas been developed in order to obtain the kinetics of the reactionswhich depends on the catalyst microstructure. A one-dimensionalheterogeneous plug-flow reactor taking into account heat andmass transfer between the flowing gas and the catalytic surfaceof the wash-coat has been chosen for the reactor model. In orderto increase the accuracy of the model, instead of using one of theavailable correlations, heat transfer has been characterized bymeasuring experimental reactor temperatures profiles. Numericalsimulations of heat and mass transfer with FLUENT� have beenperformed in order to find a correlation which describes preciselytransfer coefficients between the bulk of the flow and the surfacein the reactor model.

Two catalyst samples with different wash-coat thicknesses, rho-dium quantity, rhodium particle size and dispersion have beentested at 800, 850 and 900 �C and for residence times between40 and 150 ms. Catalytic tests performed on these samples showedthe importance of the catalyst characteristics on the performance.The catalytic performance is different as a function of the catalystdispersion in the wash-coat. Some of these tests also fulfill the con-ditions of kinetic parameters identification and enable to validatethe mathematical model for kinetic reaction rate estimation. Theidentified rhodium activation energies (166,310/165,740 J/mol)by considering or not the non-catalytic reactor activity are in goodaccordance with the literature value (169,500 J/mol).

To sum it up, this study demonstrates on one hand that rhodiumcatalyst is highly active, suitable and adapted to millistructuredreactor, and on the other hand that, for SMR process intensification,it is needed to reduce the reactor hydraulic diameter below 400 lmfor heat and mass-transfer limitations elimination. Experimentalresults showed that the single channel reactor is a very good toolfor the determination of catalyst behavior and activity, which is




representative of a more complex millistructured SMR reactor. Thevalidated reactor model is also an efficient tool to design and tostudy, thanks to kinetic parameters for the SMR reaction, theperformance of such a millistructured reactor/heat-exchanger.

Acknowledgements

The authors gratefully acknowledge the French Ministry ofEconomy, Finance and Industry and Air Liquide for funding thisstudy.

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steam methane reforming reaction process intensification by using a reactor

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