Chemical Engineering Journal xxx (2012) xxx–xxx
Contents lists available at SciVerse ScienceDirect
Chemical Engineering Journal
journal homepage: www.elsevier.com/locate/cej
Steam methane reforming reaction process intensification by using a
millistructured reactor: Experimental setup and model validation for global
kinetic 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, France
Laboratoire des Science des Procédés Céramiques et Traitements de Surface, 12 rue Atlantis, 87068 Limoges Cedex, France
c
Air Liquide, Centre de Recherche Claude & Delorme, 1 chemin de la porte des Loges, BP 126, 78354 Jouy-En-Josas Cedex, France
b
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 reactor
Methane reforming
Syngas
Hydrogen
Process intensification
Microreactor modeling
Kinetic data acquisition
a b s t r a c t
In the frame of steam methane reforming process intensification, a highly active and stable catalyst based
on rhodium with catalyst formulation and structure adapted to millistructured reactors has been formulated. This catalyst has been tested in industrial conditions (800, 850 or 900 °C and 20 bars) on a single
channel which is representative of one channel of a more complex millistructured SMR reactor. Then, a
detailed mathematical model for acquisition of the global reaction kinetics with this new catalyst has
been developed and validated from experimental catalytic tests. The developed kinetics is dependent
of the catalyst microstructure. This study presents the set-up, the model, the experimental catalytic runs
and the global kinetics estimation protocol. It demonstrates, on one hand, that millistructured reactor is
suitable for kinetic data acquisition and, on the other hand, the possibility of SMR process intensification,
for improved energy efficiency and process size reduction.
Ó 2012 Elsevier B.V. All rights reserved.
1. Introduction
Steam methane reforming (SMR) of natural gas is the main
commercial process for synthesis gas production (H2, CO). In this
process, methane reacts with steam to produce a mixture of hydrogen, carbon dioxide and carbon monoxide. This reaction is highly
endothermic and is performed in the presence of a catalyst such
as nickel or rhodium at high temperature (800–1000 °C), high pressure (20–40 bars) and steam-to-carbon ratio varying between 1.8
and 4. In the classical process, a set of tubes filled with catalyst
is operated inside a furnace equipped with burners. These burners
provide the heat needed for the reaction. The exit temperature of
⇑ Corresponding author.
E-mail address: mamadou.mbodji@ensic.inpl-nancy.fr (M. Mbodji).
the process gas ranges from 700 to 950 °C. These conditions are
limited by the tube metallurgy. The reactor tube has a length of
10–12 m and an internal diameter in the order of 10 cm. This process is well known and controlled. However, the overall efficiency
of the process is decreased by heat losses. The intensification of the
SMR process by using microstructured reactors should enable on
the one hand to resolve this heat losses problem and on the other
hand to reduce substantially the size of process units, their energetic consumption and their environmental impact [1,2]. The high
surface-to-volume ratio of microstructured reactors provides a
highly efficient heat transfer and reduces the potential for temperature gradients in catalyst layers deposited on microchannel walls
when performing highly endothermic reactions. Compared to
conventional fixed-bed catalytic reactors, microstructured reactors
advantages are considerable particularly in terms of yield,
1385-8947/$ - see front matter Ó 2012 Elsevier B.V. All rights reserved.
http://dx.doi.org/10.1016/j.cej.2012.07.117
Please cite this article in press as: M. Mbodji et al., Steam methane reforming reaction process intensification by using a millistructured reactor: Experimental setup and model validation for global kinetic reaction rate estimation, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Notations
Am
R
r1
active surface of active metal per mass of rhodium
(m2sma /grhodium)
total concentration in the gas phase (mol/m3)
hydraulic diameter of the reactor (m)
total molar flow rate of reactants (mol/s)
molar methane flow rate at the reactor inlet (mol/s)
molar flow rate of inert species (mol/s)
heat-transfer coefficient between the gas and the walls
(W/m2 K)
mass-transfer coefficient between the gas and the catalytic wall (m/s)
total pressure (Pa)
contact perimeter between the catalytic bar and the gas
(m)
contact perimeter between the thermocouple and the
gas (m)
external reactor perimeter (m)
contact perimeter between the catalytic bar and the
reactor (m)
contact perimeter between the gas and the reactor walls
(m)
universal gas constant (J/mol K)
rate of the SMR reaction (mol/m2SampleSurface s)
r2
rate of the WGS reaction (mol/m2SampleSurface s)
Rsc
thermal resistance between the catalytic bar and the
reactor (m2 K/W)
CT,g
Dh
F0
FCH4,0
Finert
hloc
kd,j
P
Pcg
Pmg
Ps
Psc
Psg
selectivity to the desired product and safety. However, this change
in production technology must be coupled to the development of
highly active and stable new catalysts in order to ensure the same
conversion rate at lower residence times and catalyst formulations
adapted to microreactors. Microreactors are also characterized by
the use of very small reactants and catalyst quantities (usually in
the range 0.01–1 g); therefore, they appear to be a very good tool
for the acquisition of kinetic data and also for the determination
of catalyst behavior and activity [3,4].
1.1. Review of SMR reactors
Reactor miniaturization is known to improve heat and mass
transfer, however, this strategy is not always sufficient for process
intensification. Catalyst intensification is also needed to avoid hot
spots [5]. Nickel is the most common industrial catalyst used for
SMR owing to its robustness, its catalytic activity and its relative
tolerance to poisons, such as sulfur, chloride, and heavy metals.
Noble metals such as ruthenium, rhodium, and palladium are also
suitable for SMR. Stefanidis and Vlachos [6] studied the intensification of steam reforming of natural gas and tested whether steam
reforming on nickel is feasible by intensifying the process via miniaturization. They found that the steam reforming reaction time
scales for rhodium and nickel depend more on the reaction temperature than mixture composition. Over the temperature range
1000–1500 K, the steam reforming on rhodium is faster than on
nickel by a factor of 3 to 20. Below this range, steam reforming
on rhodium is one order of magnitude faster than on nickel.
Zeppieri et al. [7] investigate the kinetics of methane steam
reforming reaction over a rhodium–perovskite catalyst of formula
BaRhxZr(1x)O3 at atmospheric pressure and in the temperature
range 723–1023 K. Their results show that SMR reaction rate is
first order with regard to methane and 0th order with regard to
steam. Methane conversion is proportional to the partial pressure
of methane and the contact time. Results from Iglesia et al. [8] bear
Rsm
Sco
Tg
Ts
Tc
uc
XCH4
yg,j
yc,j
z
thermal resistance between the thermocouple and the
reactor (m2 K/W)
CO selectivity (–)
gas temperature (K)
reactor skin temperature (K)
catalyst temperature (K)
gas velocity (m/s)
methane conversion (–)
gas phase molar fraction of species j (–)
molar fraction of species j in the catalyst (–)
axial position along the channel (m)
Greek notations
a
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 numbers
Gzth
thermal Graetz number (–)
Gzm
material Graetz number (–)
Nu
Nusselt number (–)
Pr
Prandtl number (–)
Re
Reynolds number (–)
Sc
Schmidt number (–)
out these affirmations. Comparing the performances of a rhodium–
perovskite catalyst to a commercial nickel-based catalyst, Zeppieri
et al. [7] confirm that the rhodium–perovskite catalyst is the most
active: high methane conversion close to the theoretical thermodynamic value is experimentally obtained with a low quantity of catalyst. Furthermore, carbon deposition is lower than on a
commercial nickel-based catalyst.
Leventa et al. [9] studied SMR in a microreactor filled with an
industrial catalyst containing 15% nickel. Their experiments were
performed in the temperature range 600–840 °C. The pressure
range was 2.5–9 bars, with hydrogen-to-methane ratio of 0.5–2
and steam-to-methane ratio 2–3. They found that the increase of
H2-to-CH4 ratio in the feed enhances the catalyst activity. However,
an increased steam-to-methane ratio in the feed moves the reforming reaction in the opposite direction. Steam acts as an inhibitor on
the catalyst activity and the reaction rate. They also observed that
the smaller diameter of the microreactor enabled decreasing the
catalyst quantity and acquisition of reaction kinetics at high
temperatures up to 840 °C without reaching equilibrium.
Microreactors are increasingly used as tools for catalytic activity
measurement. Peela et al. [10] studied steam reforming of ethanol
over 2%Rh/20%CeO2/Al2O3 catalyst in a microchannel reactor. They
compared microchannel reactor performance with that of a
packed-bed reactor using 2%Rh/20%CeO2/Al2O3 catalyst at identical
operating conditions and found the same activity but the selectivity to desired product was higher in the microchannel reactor. The
H2 yield obtained in the microchannel reactor was 65 L/g/h as compared to 60 L/g/h in the packed-bed reactor. The high selectivity of
H2 is attributed to improved heat management in the microchannel, resulting in a more uniform temperature throughout the catalyst. The radial temperature gradient in the packed-bed reactor and
in microchannel reactor by using 2D models for each type of
reactor was also investigated. The maximum temperature difference in the packed bed reactor was about 15 K whereas that in
the microchannel reactor was only 0.3 K. These results show that
Please cite this article in press as: M. Mbodji et al., Steam methane reforming reaction process intensification by using a millistructured reactor: Experimental setup and model validation for global kinetic reaction rate estimation, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117
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1.2. Description of the reacting system
The following chemical reactions should be expected when performing steam methane reforming:
the endothermic steam reforming reaction (SMR)
CH4 þ H2 O () 3H2 þ CO
ð1Þ
the reverse methanation reaction (RM)
CH4 þ 2H2 O () 4H2 þ CO2
40
20
Reaction Gibbs Free Energy [kJ/mol]
microchannel reactors significantly reduce the temperature gradient over the reactor cross-section due to their high heat-transfer
coefficient.
Wang et al. [11] assessed methane steam reforming over Rh/
MgO–Al2O3 catalysts in microchannel chemical reactors. Experimental results show that rhodium catalyst supported on MgO–
Al2O3 is highly active and stable over a wide range of steam-to-carbon ratios and resistant to coke formation. Methane steam reforming reaction rate on this catalyst in microchannel reactor was
compared to that of a conventional micro-tubular reactor. Results
confirm the performance enhancement in microchannel reactors.
All of these studies show that currently methane steam reforming
intensification is feasible. Indeed, reactor and catalyst intensification are increasingly controlled.
The present study is focused on syngas production by steam
methane reforming in a millistructured reactor. Catalysts based
on Rh/Al2O3 enabling to reach high conversion at low residence
times have been developed and tested. Experiments are conducted
at 800 °C, 850 °C or 900 °C, 20 bars and a steam-to-methane ratio
of 3. The main goal of this work is to determine SMR and WGS
kinetics reactions rates from experimental catalytic tests. The
experimental results coupled with a mathematical plug-flow reactor model taking into account heat and mass transfer between the
reactant gas and the catalyst enables identification of the kinetic
parameters (activation energies and pre-exponential rate constants) of SMR reaction by minimizing the sum of squared difference between measured methane conversion, outlet gas
temperature and calculated values given by the reactor model.
0
-20
-40
-60
-80
-100
600
SMR
RM
Methane cracking
WGS
Boudouard
CO reduction
650
700
750
800
850
Temperature [°C]
Fig. 1. Gibbs free energies of reactions (DrG) as a function of temperature.
In Table 1 are reported the heats of these six reaction. As steam
reforming is the major reaction, the global system can be considered as endothermic.
1.3. Thermodynamic analysis
The Gibbs free energies (DG) of the most significant reactions
occurring during steam methane reforming are given as a function
of temperature in Fig. 1.
According to these thermodynamic data, in our operating conditions (above 800 °C and 20 bars), SMR, RM and methane cracking
are the most favorable reactions. Carbon formation, harmful to the
operation of production units, can be limited by using an excess of
oxidizing agent as H2O. Carbon formation from methane cracking
is catalyzed by chromium and iron and can be considered as a
selectivity problem. It is usually resolved by using a catalyst and
a reactor material on which carbon formation is unlikely.
ð2Þ
1.4. Water gas shift reaction
and the exothermic water gas shift (WGS)
CO þ H2 O () H2 þ CO2
ð3Þ
The main drawback of SMR is the risk of carbon formation. Care
must be taken to avoid carbon formation due to:
methane cracking
CH4 () C þ 2H2
ð4Þ
the Boudouard reaction
2CO () C þ CO2
ð5Þ
and CO reduction
CO þ H2 () C þ H2 O
ð6Þ
Table 1
Reaction heats of steam reforming and carbon formation reactions.
Reaction
Name
DrH850°C (kJ/mol)
1
2
3
4
5
6
Steam reforming reaction
Reverse methanation
Water gas shift
Methane cracking
Boudouard
CO reduction
226
193
33
90
169
135
The reactor model developed in this work takes into account the
SMR and WGS reactions, which are the most-commonly considered reactions when modeling steam methane reforming process.
Thermodynamic analysis presented above shows that the water
gas shift reaction is negligible in the operating conditions. Furthermore, all experiments are carried out above 800 °C, and Gibbs free
energy of the WGS reaction is positive for temperatures greater
than 800 °C. The CO2 quantity recorded during catalytic tests was
not significant, therefore only SMR reaction kinetic rate is studied
is this work.
Under these conditions, the estimation of the kinetics reaction
rate then consists in finding the pre-exponential rate constant
and the activation energy for the SMR reaction. The following sections present the experimental set-up and the model developed for
data treatment.
2. Material and methods
2.1. Experimental test rig
The experimental set-up, on which catalytic tests have been
performed, is shown in Fig. 2. In order to ensure a good mixture
of the reagents, a gas mixer and pre-heater is set before the reactor
Please cite this article in press as: M. Mbodji et al., Steam methane reforming reaction process intensification by using a millistructured reactor: Experimental setup and model validation for global kinetic reaction rate estimation, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Fig. 2. Picture of the experimental setup exhibiting the reactor in the open furnace.
entrance. The reactor consists of a rectangular channel (with
dimensions 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 mm
width, 1.6 mm height and 20 cm length) coated with a rhodiumbased catalyst is introduced. To provide the required heat to the
endothermic reaction, the reactor is electrically heated. Two ovens
are used to ensure the controlled heating of the system. CC2 denotes the oven around the catalytic reactor itself, that contains
the catalytic coated sample and CC3 denotes the oven around the
non-catalytic part of the reactor. This oven CC3 is used to prevent
temperature gradient in CC2 and provide controlled temperature
conditions to the reaction.
At the reactor outlet, the set-up is equipped with a condenser
and a weighing system for measurement of the mass of condensed
water. In-line infrared analyzer is also used to analyze gases
composition.
For gas temperature measurements, two thermocouples are set
at the inlet and outlet of the reactor. Four thermocouples are also
set on the reactor outer skin to measure the temperature profile
along the reactor. A mobile thermocouple is installed on the top
wall of the inner channel in order to monitor the gas temperature
along the reactor.
Fig. 3 illustrates a longitudinal view (Fig. 3a) and a cross-sectional
view (Fig. 3b) of the reactor with all geometric perimeters of interest
that will be considered in the model for heat and mass transfer.
For the reactor heat-transfer characterization, three heat-transfer parameters need to be determined:
hloc: mean heat-transfer coefficient between the flowing gas and
the internal walls.
Rsm: contact thermal resistance between the thermocouple and
the reactor.
Rsc: contact thermal resistance between the catalytic bar and
the reactor.
Experimental temperature measurements have been performed
for determining these three heat-transfer parameters, and will be
discussed further.
In this study, the part of conversion coming from the reactor
metal alloy has been evaluated experimentally before and after
the catalytic tests. Results showed that the non-catalytic reactor
activity has drastically evolved during the catalytic tests. Thus, a
full reactor model considering the non-catalytic reactor activity
and the catalyst activity has been developed. This is performed
by coupling two reactor models in series. In the CC2 part of the
reactor, the model considers the two active areas: the reactor walls
and the catalyst, whereas in the CC3 part, only the non-catalytic
reactor activity of the walls is considered.
In order to estimate the fraction of conversion due to the reactor
walls, reactor activity is quantified by fitting experimental emptyreactor conversion after catalytic tests. To facilitate the readability
of the present paper, the full reactor model is not presented in details. Further, by using the full reactor model, it will be demonstrated that in the presence of catalyst sample, the non-catalytic
reactor activity is negligible. The reactor model presented in this
paper concerns the CC2 part without considering the non-catalytic
reactor activity.
2.1.1. Experimental catalytic tests conditions
The following experimental conditions have been used to perform the catalytic tests. The reactor is fed with methane and steam
with a steam-to-carbon ratio of 3 at 800, 850 or 900 °C. The total
gas flow rate ranges from 0.0017 to 0.0079 mol/s in order to operate with residence times between 40 and 200 ms in the CC2 part
Fig. 3a. Longitudinal view of the reactor.
Please cite this article in press as: M. Mbodji et al., Steam methane reforming reaction process intensification by using a millistructured reactor: Experimental setup and model validation for global kinetic reaction rate estimation, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Psm
Mobile thermocouple: Tm
Gas: Tg
Pmg
Psg
Catalyst sample: Tc
Pcg
Reactor wall: Ts
Psc
Ps
Psm
Mobile thermocouple: Tm
Pmg
Gas: Tg
Psg
Reactor wall: Ts
Ps
Fig. 3b. Cross sectional view of the reactor.
where the catalyst sample is located. The residence time is computed at the reactor inlet as the ratio between the reactor volume
and 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 through
a gas–liquid separator where the unreacted water is separated and
weighed by means of a weighing machine. The dry gas composition
is then determined by on-line infrared analyzer.
Carbon and hydrogen balances are carried out to check mass
balances and to detect potential coke formation. For a given residence time and temperature level (800, 850 or 900 °C), measurements have been performed during 72 h under reaction
conditions (75% H2O, 25% CH4, and 20 bars). No catalyst deactivation has been observed. Variable XCH4 represents the methane conversion and SCO the CO selectivity determined from the change in
gas composition.
X CH4 ¼
SCO
F CH4;0 F CH4
F CH4;0
F CO
¼
F CO þ F CO2
ð7Þ
ð8Þ
FCH4,0, denotes the methane flow rate at the reactor inlet. FCH4, FCO
and FCO2 respectively denote the methane, carbon monoxide and
carbon dioxide flow rates along the reactor.
2.1.2. Synthesis of catalysts
Catalysts 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 an
excess of aqueous rhodium nitrates solution. The mass of rhodium
nitrates is calculated to 20 wt.% rhodium in the final product for
the first sample and 1 wt.% for the second sample. The impregnation is conducted under heating at 150 °C and steering until water
is completely evaporated. Residues obtained are finally calcined in
air to form the catalyst phase. For the experimental study, catalysts
are deposited as layers with a thickness less than 12 lm on alumina substrates by dip coating.
2.1.3. Characterization
The morphology and the thickness of catalysts layers have been
evaluated using a Zeiss Ultra-55 scanning electron microscope before and after ageing in a steam methane reforming atmosphere at
850 °C. Samples have been observed at three different locations of
the substrate: bottom, middle and head.
Temperature-programmed reduction and chemisorption measurements have been carried out on a Micromeritics AutoChem II
2920 and an Asap 2020 on the catalyst powder before dip-coating
to control the catalyst activity.
2.1.4. Characteristics of catalyst samples
An example of the tested catalyst holders is presented in Fig. 4.
The characteristics of the catalyst samples are summarized in
Table 2.
Fig. 4. Catalyst holder.
2.2. Reactor model for kinetics study
In this section, the one-dimensional plug-flow reactor model
developed for the reactor simulation is presented. As indicated
above, the non-catalytic reactor activity is not considered in the
model version described in this paper. This model takes into account the SMR and WGS reactions and will be used for kinetic
parameters identification from experimental tests. The assumptions 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 along
the reactor, a one-dimensional plug-flow model including heat
and mass transfer between the reactant gas and the catalyst has
been developed. The catalyst is supposed to be uniformly coated
on the catalytic bar. The heat is also provided uniformly through
all the walls of the reactor. The way this heat is transferred to
the reactants is modeled by a usual convection transfer law. The
specific heat flux is one of the model parameters. All the heat
and mass balances described below are written under steady-state
conditions. Mass-transfer coefficients are introduced in the model
to account for the species transport limitation between the bulk
gas mixture and the catalytic active surface. As demonstrated by
Mladenov et al. [12], the introduction of mass-transfer coefficients
in the plug-flow reactor model improves its accuracy. However,
they have to be used with caution since they are often based on
empirical correlations. In this study, the external mass-transfer
Table 2
Characteristics of tested catalysts.
Characteristics of tested catalysts
Sample 1
Sample 2
Wash-coat thickness (lm)
Am (m2sma /grhodium)
BET (m2/gcatalyst)
Mass of catalyst (Mc) (mg)
PR (rhodium quantity) (%)
Dispersion (%)
Rhodium particles size (nm)
2
80
10
235
10
4
20
18
6
10
22.4
1
53
2
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
coefficient was evaluated by numerical simulation using FLUENTÒ,
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 convection
and the reaction at the catalytic wall. The net flux of component j
from the bulk fluid to the wall is composed of a classical convectodiffusive term and a net flux due to the reaction stoichiometry.
Combining the mass balance for each species to the overall mass
balance enables to describe the evolution of the gas phase composition as:
dðyg;j Þ
P=ðRT g Þ
ðy yg;j Þ
¼ kd;j Psc
F 0 þ F inert þ 2X 1 F CH4 ;0 c;j
dz
ð9Þ
where yg,j denotes the gas phase molar fraction of species j, yc,j the
molar fraction of species j in the catalyst, z the axial position along
the channel, kd,j the mass-transfer coefficient between the gas and
the catalytic wall, P the total pressure, Tg the gas temperature, Finert
the molar flow rate of inert species, F0 the total molar flow rate of
reactants, and Psc the contact perimeter between the catalyst surface and the gas.
The mass balance in the catalyst layer is written as the equality
between the molar flux transferred from the gas and the flux consumed by chemical reactions:
kd;j C T;g ðyg;j yc;j Þ þ
m1j 2yg;j r1 þ m2j r2 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 evolution of the gas temperature along the reactor with the following
relation:
uc Xc C T;g C pg
dT g
þU¼0
dz
ð11Þ
where U denotes the heat transferred by convection between the
gas and the catalyst sample, the gas and the mobile thermocouple
and the gas and the reactor walls.
U ¼ hloc ½Pcg ðT g T c Þ þ Psg ðT g T s Þ þ Pmg ðT g T m Þ
where hloc denotes the mean local heat-transfer coefficient by convection between the gas and the reactor internal elements. Pcg, Psg
and Pmg respectively correspond to the contact perimeters between
the gas and the catalyst, the gas and the reactor walls, and the gas
and the mobile thermocouple.
Within the catalyst layer, the enthalpy balance is written by
equalizing the heat flux provided by the furnace, the flux exchanged with the gas phase, the source term related to the reforming reaction of methane and the Water Gas Shift reaction:
Psc
ðT c T s Þ þ hloc Pcg ðT c T g Þ þ r 1 DrH1 Psc þ r2 DrH2 Psc ¼ 0
Rsc
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:
dP
2luc
¼ f Re 2
dz
Dh
ð13Þ
f Re ¼ 24ð1 1; 3553a þ 1; 9467a2 1; 7012a3 þ 0; 9564a4
Psm
Psc
ðT s T m Þ þ
ðT s T c Þ
Rsm
Rsc
ð15Þ
where Ps denotes the external reactor perimeter and u the specific
heat flux received by the reactor. This flux is one of the model
parameters.
2.3. Heat losses on the experimental device and boundary conditions of
the reactor models
For a good estimation of kinetic parameters, it is very important
to know accurately, for each experiment, the heat received by the
catalytic surface. Experimental tests have been conducted in order
to estimate the heat losses in the experimental device. Results
show that heat losses are very large depending on the reactant
gas residence time: heat losses range from 80% to 93% of the total
experimental heat flux furnished by the electrical heat. Despite
these large values, it is still possible to determine the experimental
heat flux consumed by the endothermic SMR reaction for each test,
by performing a heat balance based on the inlet and the outlet gas
temperature and methane conversion.
In the model presented here, the specific heat flux u is considered as the thermal boundary conditions. It is also possible to set
the experimental reactor temperature as a boundary condition.
This can be done by replacing the Eq. (15) by the experimental
reactor temperature.
In the CC2 reactor part, several thermocouples provide the
experimental reactor temperature profile. For the CC3 part where
there is no reactor thermocouple and catalyst sample, the reactor
temperature is assumed to be equal to the temperature measured
by the mobile thermocouple. The kinetic parameters estimated for
each boundary condition will be compared.
2.4. Reaction rates
The kinetics of steam methane reforming reaction has been
studied extensively by several groups. There is a general agreement on the first order kinetics with respect to methane, but the
activation energies vary between 20–160 kJ/mol. These differences
might be explained by experimental inaccuracies due to transport
restrictions in the sense of diffusion and heat restrictions. While
the exact mechanism of the steam methane reforming reaction is
still under debate today, the most important steps are: (1) decomposition of methane on a metal surface to hydrocarbon fragments
and carbon atoms, (2) dissociative adsorption of water to H and OH
species (3) OH or O species combine with C to form CO. Several kinetic studies on SMR reaction can be found in the literature. Table 3
presents some of them [14–19].
Recently, Wei and Iglesia [20] proposed a simple equation for
kinetics of steam methane reforming. They found that the activity
at 600–700 °C only depends on the partial pressure of methane,
implying that the rate determining step is the initial activation of
a C–H bond in methane.
The work presented in this paper is part of a preliminary design
approach of a millistructured reactor heat-exchanger for the
production of syngas. To reach that goal, the reaction rate of the
Table 3
Kinetic models of steam reforming of hydrocarbons.
0; 2537a5 Þ
where a denotes ratio between the height and the width of the
reactor.
A heat balance on the mobile thermocouple and on the reactor
enables to describe their temperature profiles as:
Psm
ðT s T m Þ hloc Pmg ðT m T g Þ ¼ 0
Rsm
uPS ¼ hloc Psg ðT s T g Þ þ
ð14Þ
Reference
Form of the kinetics law
Bodrov [14]
Kohmenko et al. [15]
Rostrup-Nielsen [16]
Tottrup [17]
Xu and Froment [18]
Aparaicio [19]
Langmuir–Hinshelwood
Temkin Identity
Two-step kinetics, power law
Pellet kinetics, power law
Langmuir–Hinshelwood
Microkinetic analysis
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
main reaction (steam methane reforming: SMR) has to be measured precisely in the same operating conditions as the future reactor. Therefore, it must be emphasized that a global kinetics model
is more appropriate than the microkinetics of the SMR reaction
with detailed reaction mechanism and determination of the limiting step. Such a lumped kinetics reaction rate will enable to properly design a milli-structured exchanger reactor for syngas
production at the industrial scale.
In order to express this overall reaction rate, the same formalism as used by Tonkovich et al. [21] to describe SMR reaction over
a rhodium on Mg-spinel catalyst, is adapted in the model by adding
a constant depending on the catalyst microstructure. Without
going into details, the SMR and WGS reactions rates can be written
as:
!
y y3
Ea1
2 c;CO c;H2
yc;CH4 yc;H2 O P
r 1 ¼ K pre exp 1 exp Kl
RT c
K eq1
yc;CO2 yc;H2
Ea2
yc;CO yc;H2 O Kl
r 2 ¼ K pre exp 2 exp RT c
K eq2
where the reaction rates r1 and r2 are expressed in [mol/
m2SampleSurface /s].
Kl is a constant depending on the catalyst microstructure
[m2ActiveMetal =m2surface of holder ] and might be expressed by first approximation as:
Kl ¼
Am M c PR
Psc L
where Am denotes the active surface of active metal per unit of mass
of rhodium (m2sma /grhodium), Mc the mass of catalyst, PR the rhodium
quantity in the catalyst, (Psc L) the surface of the holder on which
the catalyst sample is coated.
Ea1 and Ea2 (J/mol) denote the activation energy of the SMR and
WGS reactions, respectively.
Kpreexp1 and Kpreexp2 [mol/m2ActiveMetal /s] denote the pre-exponential rate constants of SMR and WGS reactions, respectively.
K eq1 ¼ 101;3252
26;830
exp þ 30:114 Equilibrium constant of SMR ½Pa2 Tc
2.6. Heat transfer in the reactor
Microchannels are known as apparatuses that provide high
heat-transfer coefficients, which makes them particularly suitable
for kinetics studies. Their remarkable heat-transfer capacity is
due to the fact that the heat-transfer coefficient is proportional
to the reciprocal of the channel hydraulic diameter.
An example of experimental temperature profiles along the
reactor during catalytic tests in the presence of active catalyst sample is shown in Fig. 5. In the CC2 part containing the catalyst sample, the reactor temperature decreases along the reactor length as a
result of the heat consumption by the endothermic steam reforming reaction. From the end of the CC2 part to the middle of CC3, the
reactor temperature is constant. This can be explained by the fact
that heat consumption by SMR reaction is not significant because
most of the methane quantity has already been consumed by reaction in the CC2 part. Reactor temperature decreases at the end of
the CC3 part due to heat losses by natural convection between
the experimental setup and the surrounding air.
The reaction kinetics can be determined from the comparison
between the experimental tests and the calculated results computed from the heat and mass balances presented above. The resolution of the system requires the knowledge of specific
parameters as the following three heat-transfer parameters defined above: hloc, Rsm and Rsc.
These parameters have been determined from experimental
temperatures measurements without chemical reaction.
For these tests, the reactor is fed with nitrogen. A heat flux is set
at the external reactor walls. Once thermal steady state is reached,
the experimental temperatures of the gas, the mobile thermocouple and the reactor are recorded. Heat balances on the gas, on the
mobile thermocouple and on the catalytic bar than enable to relate
their temperatures as follows:
Heat balance on the gas:
2
dT g
1
P mg
4
¼
F inert C T;g Pmg Rsm þ
dz
Psm
The reactor model is a set of differential and algebraic equations. The solver function ode15s available on MATLABÒ is used
to solve this system. This solver uses the Gear method which is
adapted to the resolution of stiff systems. After integration of the
set of equations, mass and overall enthalpy balances are computed
and satisfied with less than 0.001% and 1%, respectively.
From a purely numerical point of view, the reactor model described above is complete and ready to be computed. However,
to make it more reliable and to correctly reproduce the experimental results, it is useful to provide a good description of the heat and
mass-transfer phenomena that occur between the reactant gas and
the catalyst. In the following section, heat and mass transfer in the
reactor are investigated.
P sc
Rsc þ h1
5ðT s T g Þ
ð16Þ
loc
hloc Pmg Rsm
Tg
Psm
h P R
1 þ loc mg sm
Psm
Ts þ
Tm ¼
ð17Þ
Temperature [°C]
860
CC2 part
850
840
Reactor
Mobile thermocouple
830
820
0
0.02
0.04
0.06
0.08
0.1
0.12
0.14
0.16
0.18
0.2
Reactor length [m]
Temperature [°C]
2.5. Reactor model resolution
þ hloc P sg þ Pcg
Temperature of the mobile thermocouple
4400
K eq2 ¼ exp
4:036 Equilibrium constant of WGS reaction ½—
Tc
As said previously, the WGS reaction can be neglected under the
operating conditions of this study. Estimation of the kinetics reaction rate then consists in finding the pre-exponential rate constant
and the activation energy of the SMR reaction.
1
hloc
3
P cg
850
800
CC3 part
750
700
Mobile thermocouple
650
0.2
0.25
0.3
0.35
0.4
0.45
Reactor length [m]
Fig. 5. Experimental temperature profiles along the reactor with active catalyst
sample in the CC2 part.
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Temperature on the catalytic surface
Table 4
Experimental heat-transfer parameters.
hloc Pcg Rsc
Ts þ
Tg
Psc
Tc ¼
hloc Pcg Rsc
1þ
Psc
ð18Þ
It is important to point out here the fact that when both thermal
resistances Rsm and Rsc are set equal to 0, the temperature of the catalytic bar and that of the mobile thermocouple are equal to the
reactor temperature i.e. Tm = Tc = Ts. The three heat-transfer parameters have been identified by minimizing the sum of squared differences between measured and modeled temperatures. Fig. 6 shows
an example of temperature profiles along the reactor after heattransfer parameters identification. A mean difference of 2 °C between modeled and experimental temperature has been obtained.
For all tests, the average values of the mean heat-transfer coefficient between the flowing gas and the reactor internal elements
(catalytic bar, mobile thermocouple and reactor walls), the thermal
resistance between the thermocouple and the reactor, and the
thermal resistance between the catalytic bar and the reactor and
their standard deviation are presented in Table 4. The relatively
large error on the resistance Rsc indicates that this parameter is
not properly estimated, since the standard deviation on Rsc is larger
than its average value. This result is not surprising since there is no
direct temperature measurement on the catalytic bar. Convection
gas heat-transfer coefficient in microreactors depends on the
hydraulic diameter and on the gas composition and usually ranges
from 400 to 2000 W/m2. Kays and Crawford [22] proposed the following correlation to estimate Nusselt number for fully-developed
laminar flow in rectangular ducts with constant heat flux
condition:
2
3
4
5
Nu ¼ 8:235ð1 1:883a þ 3:767a 5:814a þ 5:361a 2a Þ
where a is ratio between the height and the width of the reactor,
and Nu the Nusselt number defined as:
Nu ¼
hloc Dh
k
For a mean temperature of 400 °C, hloc calculated from Kays and
Crawford [22] correlation is 235 W/m2 K. The experimental results
give an average value of 442 W/m2 K with a standard deviation of
65 W/m2 K. This difference can be explained by the fact that hloc
from the experimental tests is an average value from several experiments and by the assumption that considers that hloc is constant
Heat-transfer parameter
hloc (W/m2 K)
Rsm (m2 K/W)
Rsc (m2 K/W)
Average value
Standard deviation
442
65
0.0038
0.0005
0.013
0.016
along the reactor. It is well known that hloc is constant only for
fully-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 transfer
in the reactor model, simulations with the commercial CFD package FLUENTÒ have been performed without chemical reaction by
setting a constant wall-temperature boundary condition and feeding a N2 flow at inlet temperature. For each simulation, the local
Nusselt number is computed with the following relation:
Nu ¼
QDh
ðT w T g Þk
where Q denotes the heat flux at the wall, Dh the hydraulic diameter, Tw and Tg the wall temperature and the mass-averaged gas
temperature, respectively. Several simulations have been performed by varying the gas inlet velocity. They enabled description
of the Nusselt variation as a function of the thermal Graetz number
along the reactor with this following relation:
Nu ¼ 4:58 expð0:003Gzth Þ with Gzth ¼
RePr Dh
z
The variation of the Nusselt number as a function of the thermal
Graetz number along the reactor is shown in Fig. 7. When the flow
is fully developed, the Nusselt number tends towards a constant value of 4.58. This result is in agreement with those obtained by Kays
and London [23]. The slight difference can be explained by reactor
cross section that is slightly different from a perfect rectangle due
to the presence of the mobile thermocouple (see Fig. 3b).
2.7. Mass transfer in the reactor
Mass-transfer coefficients must also be preliminary determined
to study the kinetics of the reaction. It is difficult to perform reliable mass-transfer coefficient measurement in microdevices and
mass transfer has been evaluated here by CFD simulations using
440
Gas Model
Outlet gas Experimental
Mobile thermocouple Model
430
12
Mobile thermocouple Experimental
410
Reactor Experimental
Nusselt Number [-]
Temperature [°C]
420
Catalytic bar Model
400
390
380
10
8
6
370
360
350
CFD results with Fluent
Nu = 4.58exp(0.003Gz th)
14
4
0
0.02
0.04
0.06
0.08
0.1
0.12
0.14
0.16
0.18
0.2
Reactor length [m]
Fig. 6. Temperature profiles in the reactor without reaction for calibration of heattransfer.
0
0.05
0.1
0.15
0.2
0.25
1 / Graetz Number [-]
Fig. 7. Nusselt variation as a function of the thermal Graetz number along the
reactor.
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Temperature profile
has been established between Sherwood number and the material
Graetz number.
Adiabatic wall temperature
Oulet
Inlet
Sh ¼ 3:97 expð0:0023Gzm Þ
This correlation is used in the reactor model to represent the external mass transfer between the gas and the catalytic surface.
Uniform wall temperature
3. Kinetic parameters identification
Fig. 8. Boundary conditions for mass-transfer study.
FLUENTÒ. To perform these simulations, the heat and mass transfer
are assumed analogous. Thus, numerical heat-transfer simulations
with appropriate boundary conditions have been performed in order to find a correlation which describes the analogous mass transfer. For simulation of the catalytic wall, a uniform wall temperature
boundary condition is used. As there is no mass transfer on the
other walls, adiabatic boundary conditions are used (see Fig. 8).
Male et al. [24] investigated mass transfer in a microreactor by
using a similar method.
Several simulations have been performed by varying the N2
velocity inlet. Sherwood number is computed by using the following relation:
Sh Nu ¼
QDh
ðT c T g Þk
Sherwood number at the reactor entrance varies strongly as a function of the gas inlet velocity. However, for all simulations, the Sherwood number tends towards an asymptotic value of 3.99 which is
in very good agreement with literature. Indeed, Kays and London
[23] reported an average Nusselt number of 3.9 in the case of a rectangular channel having the same aspect ratio a (heigh/width) and
boundary conditions. To consider the entrance effects on the
mass-transfer, the material Graetz number is introduced:
Gzm ¼
ReScDh
z
Fig. 9 depicts the Sherwood number variation as a function of the
material Graetz number.
When the material Graetz number is less than 10, the Sherwood
number is constant and tends towards its limiting value 3.99. For
Graetz numbers above 10, the entrance effects cannot be neglected. From these simulations results, the following correlation
CFD results with Fluent
Sh = 3.97*exp(0.0023*Gzm)
X
0
@
X model
X experiment
CH4
CH4
X experiment
CH4
10
8
6
!2
þ
T model
T experiment
g
g
T experiment
g
!2 1
A
Kinetic parameters determination then consists in solving a nonlinear optimization problem without constraints. The function Fminsearch available in MATLABÒ optimization toolbox, based on the
SIMPLEX method, is used to find kinetic parameters (activation
energies and pre-exponential rate constants). This parametric optimization is performed simultaneously on several experiments conducted at different residence times and temperature levels.
4. Impact of the non-catalytic reactor activity on the overall
methane conversion
In order to properly determine the reaction kinetics, we must
ensure that the activity of the metallic walls of the reactor, estimated by the methane conversion, is negligible compared to the
activity of the catalyst holder. Therefore, experimental tests with
an inert holder have been carried out before and after the catalytic
tests. Fig. 10 depicts the non-catalytic reactor conversion in presence of an inert catalyst sample before and after the tests. It is
important to precise here, that residence time is computed along
the reactor by considering the CC2 and CC3 part.
One can note that the non-catalytic reactor activity has drastically evolved during the catalytic tests. After several tests, the reactor intrinsic activity increases and is not negligible compared to
some experiments with active catalyst sample.
90
Methane conversion before catalytic tests
Methane conversion after catalytic tests
80
12
Sherwood Number [-]
F¼
Non-catalytic reactor activity at 850°C
14
For each catalytic test, the gas phase molar fraction of H2, CO,
CO2 and CH4 and the outlet gas temperature are measured and recorded. The kinetic parameters are estimated by minimizing the
sum of the squared difference F between measured methane conversion, outlet gas temperature and calculated values given by the
reactor model. The following function F is minimized:
70
60
50
40
30
20
10
4
0
100
0
0.05
0.1
0.15
0.2
0.25
0.3
0.35
0.4
0.45
0.5
150
200
250
300
350
400
450
500
550
600
650
Residence time [ms]
1 / Graetz material Number [-]
Fig. 9. Sherwood number variation as a function of the material Graetz number.
Fig. 10. Non-catalytic reactor conversion without catalyst before and after the
catalytic tests.
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
0.75
850
Methane conversion
+ or - 2 %
849
0.65
848
0.6
847
Experiment [-]
Experiment [-]
0.7
0.55
0.5
0.45
846
845
844
0.4
843
0.35
842
0.3
841
0.25
0.3
0.4
0.5
0.6
0.7
Outlet gas temperature
+ or - 5°C
840
840
845
Model [-]
850
Model [-]
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 noncatalytic activity, the full reactor model taking into account the
non-catalytic reactor activity has been used. The full reactor model
consists in coupling two reactors in series. In the CC2 part, the
reactor model considers the two active areas: the reactor walls
and the catalyst. In the CC3 part, only the non-catalytic reactor
activity is considered.
850°C
800°C
900°C
sample 1
90
sample 2
thermodynamic equilibrium
90
CH4 Conversion [%]
The activity of the non-catalytic walls of the reactor is quantified by fitting experimental reactor activity after the catalytic tests.
Comparison between model and experimental results for the noncatalytic reactor activity is shown in Fig. 11. The calculated values
of the methane conversion and the gas temperature are in very
good agreement with the experimental ones.
90
80
80
80
70
70
70
60
60
60
50
50
50
40
40
40
30
30
30
20
50
100
150
Residence time [ms]
20
50
100
150
Residence time [ms]
20
50
100
150
Residence time [ms]
Fig. 12. Methane conversion as a function of the residence time and the temperature.
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Table 5
Kinetic 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
60
100
147
Sample 2 (800 °C)
33
40
51
60
35
41
51
60
816
807
806
815
793
795
799
801
90
94
97
99
40
60
100
147
Sample 1 (800 °C)
30
37
45
50
30
35
42
50
776
768
764
785
777
780
787
791
95
97
98
99
40
60
100
147
Sample 1 (850 °C)
43
51
65
70
43
50
64
69
845
840
852
852
849
852
856
858
86
91
96
98
0.8
Catalyst
Reactor CC2 part
Reactor CC3 part
Catalyst
Reactor CC2 part
Reactor CC3 part
0.7
Reaction rates [mol/m2/s]
Reaction rates [mol/m2/s]
0.6
0.1
Inert catalyst sample
0.05
0.5
Active catalyst sample
0.4
0.3
0.2
0.1
0
0
0
0.1
0.2
0.3
0.4
0.5
Reactor length [m]
0
0.1
0.2
0.3
0.4
0.5
Reactor length [m]
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
5.2. Kinetics parameters identification taking into account the noncatalytic reactor activity
5.1. Experimental results of the catalytic tests with catalyst sample
Here are presented the first tested samples. These samples enable to validate the determination of the kinetic parameters. Both
these samples can be distinguished by their wash-coat thickness,
rhodium quantity, rhodium particle size and dispersion. They have
been tested at 800, 850 and 900 °C and for residence times between 40 and 150 ms.
Fig. 12 shows methane conversion as a function of the residence
time and temperature for samples 1 and 2. Methane conversion increases with increasing residence time and/or temperature. However, sample 2 is more active than sample 1 despite the fact that
sample 2 has less rhodium quantity. Indeed methane conversion
with sample 2 is greater than that obtained by sample 1. The good
performances of sample 2 could be explained by the good dispersion of the rhodium particles. Furthermore, the small rhodium particle size provides to the catalyst a high specific surface.
In order to quantify the kinetic reaction rate of the SMR reaction
on the catalyst, the full reactor model considering the non-catalytic
activity of the reactor walls and the catalyst activity of the holder
has been used. The kinetics of the SMR reaction specific to the
reactor walls was already evaluated in the presence of an inert
sample.
Then, by using the full reactor model with several active areas,
the kinetic parameters (Kpreexp1 and Ea1) of the catalytic SMR reaction can be identified. Table 5 shows the estimated kinetic parameters, the part of conversion due to the catalyst sample and a
comparison between experiment and model results in terms of
methane conversion and outlet gas temperature. These results
are obtained by setting a constant heat flux at the reactor walls.
One can note a good agreement between model and experimental
results despite the experimental measurements uncertainties.
The kinetics constants of the methane steam reforming are:
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
Kpreexp1 = 9.47 107 mol/m2sma /s.
Ea1 = 166,310 J/mol.
Table 6
Comparison of identified kinetic parameters with literature results.
Kinetic parameters
Methane conversion due to the catalyst sample increases with
the residence times and the catalyst activity. Otherwise, these results show that when an active catalyst sample is placed inside
the reactor, all methane conversion is due to the catalyst sample
which confirms that the non-catalytic reactor activity is then negligible. This can be explained by the fact that the catalyst sample is
located at the reactor entrance and its activation energy is almost
twice less than the non-catalytic reactor activation energy.
Experimental heat received by the reactor during the catalytic
tests was difficult to estimate due to the large heat losses. As presented in the model equations, it is also possible to consider the
experimental heat provided to catalytic holder for the endothermic
reaction from the thermal balance based on methane conversion,
inlet and outlet gas temperature. The reactor model has been improved in order to avoid these uncertainties by setting an experimental reactor temperature as the thermal boundary condition.
The kinetics values obtained with the new boundary conditions
are similar to those obtained with the first boundary condition.
SMR kinetic reaction rate on the catalyst and on the non-catalytic reactor walls is shown in Fig. 13. When an inert catalyst sample is located in the CC2 part of the reactor, the SMR reaction rate
on the catalyst is equal to 0 and all the methane conversion is due
to the non-catalytic reactor activity. By contrast, when an active
catalyst sample is used, methane conversion due to the non-catalytic reactor activity can be considered as negligible. Depending
on residence time and temperature level, methane conversion
due to the catalyst activity ranges from 86 to 99% of the overall
methane conversion.
Kinetic parameters were also estimated by considering that all
methane conversion is due to the catalyst i.e. by using the reactor
model presented previously. As can be seen in Fig. 14, experimental results and model-predicted values are in perfect agreement.
The kinetics constants of the methane steam reforming are:
This work
By considering the non-catalytic
reactor activity
Without considering the non-catalytic
reactor activity
Tonkovich et al. [21]
Kinetic parameters
Kpreexp1 (mol/m2sma /s)
Ea1 (J/mol)
9.47 107
166,310
1.68 108
165,740
Kpreexp1 (mol/m3catalyst /s)
Ea1 (J/mol)
1.275 108
169,500
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 reaction by using a rhodium on Mg-spinel catalyst, and estimated the
SMR kinetic reaction rate by fitting kinetic data. Their kinetic
parameters and those obtained in this work are summarized in Table 6. The pre-exponential constants are not directly comparable,
due to the difference in the kinetic formulation. However, the activation energies are in the same order of magnitude.
7. Discussion
The detailed mathematical model for acquisition of kinetic data
developed in this work enabled to find CH4 reforming kinetic reaction rate. A very good agreement between model and experimental
results has been obtained. However, we noted some difficulties to
estimate kinetic parameters from experimental tests conducted at
900 °C on sample 1, 850 and 900 °C on sample 2. This is explained
by heat and external mass transfer limitations which appear when
75
860
70
65
840
60
Experiment
Experiment
55
50
45
820
800
40
35
780
30
25
20
20
Methane conversion [%]
+ or - + 2 %
30
40
50
Model
60
70
760
Outlet gas temperature [°C]
+ or - 10°C
760
780
800
820
840
860
Model
Fig. 14. Comparison between model and experiment by considering that all methane conversion is due to the catalyst for tests at 800 and 850 °C on sample 1 and 800 °C on
sample 2. Conversion (left) and outlet gas temperature (right).
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M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
2
2
10
10
Reaction
External mass transfer
Characteristic time [ms]
Characteristic time [ms]
Reaction
External mass transfer
1
10
0
10
1
10
0
10
Hydraulic diameter 1 mm
Hydraulic diameter 0.4 mm
-1
10
-1
0
0.05
0.1
0.15
0.2
Reactor length [m]
10
0
0.05
0.1
0.15
0.2
Reactor length [m]
Fig. 15. Characteristic times analysis.
performing SMR reaction on a highly active catalyst at high temperature in a microchannel reactor having a large hydraulic diameter (>1 mm). Further tests are required with reactor hydraulic
diameter below 400 lm for reduction of the heat and mass-transfer limitations.
Characteristic times of SMR reaction and external mass transfer
have been investigated and are shown in Fig. 15. A steam-to-carbon ratio of 3 has been used. The reactant gas temperature ranges
from 650 °C to 900 °C. Characteristic times of reaction and external
mass transfer decrease along the reactor due to the increasing temperature. For a reactor with a hydraulic diameter of 1 mm, the
reaction and external mass-transfer characteristic times are in
the same order of magnitude for temperatures near 780 °C. For
temperatures greater than this value, heat and/or external masstransfer limitations appear and become more and more pronounced when increasing temperature. A similar result was found
by Arzamendi et al. [25] who investigated steam methane reforming intensification by using a squared monolith channel and a nickel-based catalyst. By varying channel sides between 0.35–2.8 mm,
they showed that 0.7 mm is a sufficiently low dimension for SMR
process intensification. In the case of a rhodium-based catalyst
which is more active than the nickel based catalyst, results showed
that, to eliminate heat or mass-transfer limitations and for process
intensification, it is needed to use a hydraulic diameter below
0.4 mm. The final module design still has to be chosen after economic assessment and by considering additional technical aspects
related to the mechanical resistance of the device or the manufacturing of the microstructured system or the possibilities for catalyst coating inside the reactor.
The full reactor model enabled to estimate kinetic reaction rate
of SMR from the catalytic tests in spite of the reactor activity and
complex heat management inside the reactor. Currently, the same
tests are conducted on a reactor coated with alumina in order to
suppress reactor activity. The fact that the part of methane conversion coming from the non-catalytic reactor is negligible in the
presence of active catalyst sample is studied experimentally and
will be the subject of another publication.
8. Conclusions
Steam methane reforming process intensification by using a
millistructured reactor and a rhodium-based catalyst has been
investigated in this work. A detailed mathematical model for kinetic reaction rate measurement from experimental catalytic tests
has been developed in order to obtain the kinetics of the reactions
which depends on the catalyst microstructure. A one-dimensional
heterogeneous plug-flow reactor taking into account heat and
mass transfer between the flowing gas and the catalytic surface
of the wash-coat has been chosen for the reactor model. In order
to increase the accuracy of the model, instead of using one of the
available correlations, heat transfer has been characterized by
measuring experimental reactor temperatures profiles. Numerical
simulations of heat and mass transfer with FLUENTÒ have been
performed in order to find a correlation which describes precisely
transfer coefficients between the bulk of the flow and the surface
in the reactor model.
Two catalyst samples with different wash-coat thicknesses, rhodium quantity, rhodium particle size and dispersion have been
tested at 800, 850 and 900 °C and for residence times between
40 and 150 ms. Catalytic tests performed on these samples showed
the importance of the catalyst characteristics on the performance.
The catalytic performance is different as a function of the catalyst
dispersion in the wash-coat. Some of these tests also fulfill the conditions of kinetic parameters identification and enable to validate
the mathematical model for kinetic reaction rate estimation. The
identified rhodium activation energies (166,310/165,740 J/mol)
by considering or not the non-catalytic reactor activity are in good
accordance with the literature value (169,500 J/mol).
To sum it up, this study demonstrates on one hand that rhodium
catalyst is highly active, suitable and adapted to millistructured
reactor, and on the other hand that, for SMR process intensification,
it is needed to reduce the reactor hydraulic diameter below 400 lm
for heat and mass-transfer limitations elimination. Experimental
results showed that the single channel reactor is a very good tool
for the determination of catalyst behavior and activity, which is
Please cite this article in press as: M. Mbodji et al., Steam methane reforming reaction process intensification by using a millistructured reactor: Experimental setup and model validation for global kinetic reaction rate estimation, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117
14
M. Mbodji et al. / Chemical Engineering Journal xxx (2012) xxx–xxx
representative of a more complex millistructured SMR reactor. The
validated reactor model is also an efficient tool to design and to
study, thanks to kinetic parameters for the SMR reaction, the
performance of such a millistructured reactor/heat-exchanger.
Acknowledgements
The authors gratefully acknowledge the French Ministry of
Economy, Finance and Industry and Air Liquide for funding this
study.
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Please cite this article in press as: M. Mbodji et al., Steam methane reforming reaction process intensification by using a millistructured reactor: Experimental setup and model validation for global kinetic reaction rate estimation, Chem. Eng. J. (2012), http://dx.doi.org/10.1016/j.cej.2012.07.117