Aqueous phase reforming in a microchannel reactor: The effect of mass transfer on hydrogen selectivity

Article abstract of DOI:10.1039/c3cy00577a

Aqueous phase reforming of sorbitol was carried out in a 1.7 m long, 320 μm ID microchannel reactor with a 5 μm Pt-based washcoated catalyst layer, combined with nitrogen stripping. The performance of this microchannel reactor is correlated to the mass transfer properties, reaction kinetics, hydrogen selectivity and product distribution. Mass transfer does not affect the rate of sorbitol consumption, which is limited by the kinetics of the reforming reaction. Mass transfer significantly affects the hydrogen selectivity and the product distribution. The rapid consumption of hydrogen in side reactions at the catalyst surface is prevented by a fast mass transfer of hydrogen from the catalyst site to the gas phase in the microchannel reactor. This results in a decrease of the concentration of hydrogen at the catalyst surface, which was found to enhance the desired reforming reaction rate at the expense of the undesired hydrogen consuming reactions. Compared to a fixed bed reactor, the selectivity to hydrogen in the microchannel reactor was increased by a factor of 2. The yield of side products (mainly C3 and heavier hydrodeoxygenated species) was suppressed while the yield of hydrogen was increased from 1.4 to 4 moles per mole of sorbitol fed.

Full text of DOI:10.1039/c3cy00577a

Catalysis
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PAPER
Aqueous phase reforming in a microchannel reactor:
the effect of mass transfer on hydrogen selectivity
Cite this: Catal. Sci. Technol., 2013,
3, 2834
Maria Fernanda Neira D’Angelo, Vitaly Ordomsky, John van der Schaaf,
Jaap C. Schouten and T. Alexander Nijhuis*
Aqueous phase reforming of sorbitol was carried out in a 1.7 m long, 320 mm ID microchannel reactor
with a 5 mm Pt-based washcoated catalyst layer, combined with nitrogen stripping. The performance of
this microchannel reactor is correlated to the mass transfer properties, reaction kinetics, hydrogen
selectivity and product distribution. Mass transfer does not affect the rate of sorbitol consumption,
which is limited by the kinetics of the reforming reaction. Mass transfer significantly affects the
hydrogen selectivity and the product distribution. The rapid consumption of hydrogen in side reactions
at the catalyst surface is prevented by a fast mass transfer of hydrogen from the catalyst site to the gas
phase in the microchannel reactor. This results in a decrease of the concentration of hydrogen at the
catalyst surface, which was found to enhance the desired reforming reaction rate at the expense of the
undesired hydrogen consuming reactions. Compared to a fixed bed reactor, the selectivity to hydrogen
in the microchannel reactor was increased by a factor of 2. The yield of side products (mainly C3 and
heavier hydrodeoxygenated species) was suppressed while the yield of hydrogen was increased from
1.4 to 4 moles per mole of sorbitol fed.
Received 29th May 2013,
Accepted 8th August 2013
DOI: 10.1039/c3cy00577a
www.rsc.org/catalysis
chemistry and the extent of the undesired hydrogen-consuming
reactions become more significant as the size of the substrate
1 Introduction
increases.2 For practical implementation, however, the conversion
of larger carbohydrates is desired to minimize pre-treatment costs.
Glucose is an attractive reactant for APR,13–16 but its conversion to
hydrogen is additionally limited by homogeneous decomposition.
Alternatively, glucose can be partially hydrogenated in a fast and
nearly 100% selective step to sorbitol.17 This sugar–alcohol is not
sensitive to homogeneous decomposition and can be consequently
reformed to hydrogen with greater selectivities than glucose.18 In
this case, the selectivity is only restricted by undesired hydrogen-
consuming reactions on the catalyst surface.2
Various examples in the literature highlight the relevance of
the interface mass transfer rate, and thus the reactor design, on
the product selectivity.19,20 One example is a microreactor in
which gas and liquid are forced to flow through channels with
diameters of 10À4–10À3 m. At these scales, the hydrodynamics
differ from those in conventional systems. Of possible flow
regimes, the Taylor flow regime is one of the most stable.21–23
In the Taylor flow regime, gas bubbles and liquid slugs flow
through the channel as a plug flow in an alternating sequence.
The gas bubbles are longer than the channel diameter and a
thin liquid film separates the gas bubbles from the channel
walls, where the catalyst is deposited. This system benefits from
high diffusion rates, short diffusion paths and large interfacial
surface areas, resulting in excellent mass and heat transfer
The decentralized conversion of biomass to hydrogen is an
attractive alternative to feed highly efficient polymer electrolyte
membrane (PEM) fuel cells for power generation.1 Hydrogen
can be obtained from an aqueous solution of a (bio)carbohy-
drate over a Pt-based catalyst under mild conditions (ca. 200 1C,
30 bar) via aqueous phase reforming (APR):2–4
CnH2mOn + nH2O " nCO2 + (m + n)H2
(1)
This route5–7 gives the possibility to upgrade liquid waste
streams from bio-based industries into a nearly CO-free hydrogen
gas mixture with minimized energy input.8,9 The production of
hydrogen via APR is challenged by undesired side reactions like
the acid catalyzed dehydration of the substrate, followed by
hydrogenation over the metal surface. Rearrangement reactions
may also occur given the high functionality of the intermediate
products and the combined catalytic effect of the active metal, the
support and the liquid medium. These reactions represent an
overall consumption of hydrogen and result in the formation of a
significant amount of liquid phase water-soluble oxygenated
products, and eventually alkanes.10–12 The complexity of this
Eindhoven University of Technology, Department of Chemical Engineering and
Chemistry, PO Box 513, 5600 MB, Eindhoven, The Netherlands.
E-mail: T.A.Nijhuis@tue.nl; Tel: +31-40-2473671
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rates.24,25 Provided there is a good affinity between the liquid fashion as the washcoated layer. The resulting powder was
and the washcoated walls of the channel, this system offers an pressed and sieved to obtain particle size between 422 and
ideal wettability and a continuous renewal of reactants on the 600 mm.
catalyst surface.26 Despite these advantages, the application of
microreactor systems for heterogeneously catalyzed reactions is
still under development since some issues like catalyst loading,
catalyst stability and scaling up need to be solved to make these
processes industrially attractive.
2.2 Experimental setup
The experimental setup in Fig. 1 was used to conduct the
conversion of sorbitol to hydrogen by APR. The washcoated
microchannel is located in an oven with circulating hot air that
provides constant heating to the reactor to maintain isothermal
operation at 220 1C. The adiabatic temperature decrease
at 100% conversion is less than 8 K (calculated as DT = 6 Â
105 J molÀ1 sorbitol  10À3 mol sorbitol/mol water/76 J molÀ1
water K), which can be easily overcome due to the excellent heat
transfer characteristics of the microchannel reactor. Prior to
operation, the washcoated microchannel reactor was pressur-
ized, pre-heated and pre-wetted. A syringe pump (Teledyne
ISCO 500D) was used to feed the liquid solution of 1 wt%
sorbitol (Sigma) in water. During the reaction, the microchan-
nel reactor was operated under the Taylor flow regime by
co-feeding the liquid reactant and nitrogen. Both phases were
pre-mixed before the reactor inlet in a Y-mixer (VICI Valco). The
reaction products were quenched downstream and collected in
a pressurized stainless steel vessel where the gas phase was
separated from the liquid phase. The pressure of the system
(i.e. 35 bar) was controlled by a back pressure regulator con-
nected to the gas line that exited the steel vessel. To ensure
proper control of the pressure, an argon flow was fed to this
vessel, also serving as an external standard to quantify the gas
composition.
We have recently reported that APR of ethylene glycol can be
successfully conducted in a catalytically washcoated micro-
channel reactor that operates under the Taylor flow regime
with continuous phase-transfer hydrogen separation.27 The
focus of that study, as initial proof of concept, was the prepara-
tion of the catalyst as well as the evaluation of its activity and
stability during APR. In that study, a visible increase of the
hydrogen selectivity compared to a conventional fixed bed
reactor was detected. This was attributed to the high mass
transfer rates and to the hydrodynamic characteristics of the
microchannel reactor. The use of a microchannel reactor is
expected to be even more beneficial during APR of larger
carbohydrates like sorbitol with stronger proclivity for undesired
hydrogen-consuming reactions than ethylene glycol.
In this study, we investigate the benefits of using a micro-
channel reactor for the production of hydrogen via APR of
sorbitol.
C6H14O6 + 6H2O " 6CO2 + 13H2
(2)
The implications of the mass transfer effects of the microchannel
reactor for the hydrogen selectivity and the reaction kinetics are
investigated. As a model system, APR of sorbitol is conducted in a
1.7 m long, 320 mm ID microchannel reactor with a 5 mm Pt-based
washcoated catalyst layer as well as in a conventional fixed bed
reactor. The performance of the reactors is evaluated in terms
of sorbitol conversion, hydrogen selectivity and the observed
product distribution, which is analyzed according to the reaction
pathways earlier suggested in the literature.10,28 In this way, we
are able to correlate mass transfer characteristics of the micro-
channel reactor with the hydrogen selectivity and the reaction
kinetics.
The identification and quantification of the liquid phase
reaction products were done by a combination of HPLC (column:
Metacarb 67H, detector: RI), GC (column: Cp-sil5, detector: FID) and
GC-MS (column: DB-200, detector: MS) analysis. The composition of
the gas phase was monitored on-line with GC (columns: Poraplot/
Molsieve 5A, detectors: TCD/FID).
The composition of gas and liquid was used to determine
the conversion of sorbitol and the selectivity towards the main
products when varying the liquid flow rate (0.75–9 ml hÀ1) and
the gas to liquid ratio at the microchannel inlet (RGL
=
0–2 mgas3/mliquid3). These results were compared to those
obtained in analogous experiments conducted in a stainless
steel tubular reactor (10 cm long, 2.5 mm ID) with a porous bed
2 Experimental
2.1 Catalyst preparation
A 5 mm boehmite layer was deposited on the walls of a 1 m long,
320 mm ID fused silica microchannel (Varian) by following a
static washcoating methodology, fully described in a previous
work.27 Pt was loaded on the support by impregnation using a
solution of Pt(C5H7O2)2 in Cl2CH2 with the desired concen-
tration to obtain 0.7 mg Pt per 1.7 m or 2 wt% Pt/AlO(OH).
Activation of this layer was done by calcination at 350 1C for 6 h
and reduction at 220 1C with hydrogen for 2 h prior to the
reaction.
The catalytic activity of the washcoated microchannel
was compared to that of a reference bulk catalyst in a fixed
bed reactor. For the preparation of this catalyst, the AlO(OH)
precursor was calcined, impregnated and activated in the same
Fig. 1 Experimental setup for APR of sorbitol.
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of 110 mg of bulk catalyst (2 wt% Pt/AlO(OH)). The porosity of
the bed was ca. 0.3. This reactor was exchanged by the wash-
coated microchannel in the same experimental setup depicted
in Fig. 1. In these experiments, the fixed bed reactor was placed
vertically and operated downflow. After pre-wetting of the
catalyst with the liquid reactant, a nitrogen flow was fed co-
currently in the same fashion as previously done with the
washcoated microchannel reactor. The liquid and gas flow
rates were adjusted according to the amount of Pt in the reactor
to maintain a comparable WHSV and equal RGL as those in the
microchannel. If not indicated otherwise, the gas stream fed is
pure nitrogen. An extended series of experiments were conducted
in the fixed bed reactor to investigate the effect of the hydrogen
partial pressure on the reaction rate and the hydrogen selectivity.
In these experiments, various mixtures of nitrogen and hydrogen
Fig. 2 Sorbitol conversion versus residence time for the fixed bed reactor (open
circles), with 2 mg Pt loading; and the microchannel reactor (full circles), with
0.7 mg Pt loading. Reaction conditions: 220 1C, 35 bar, 1 wt% of sorbitol in water,
were pre-mixed with the liquid stream with a constant gas to
R
GL = 2mgas3/mliquid3 at the reactor inlet, liquid flows were adjusted in each case
3
liquid ratio of RGL = 2mgas3/mliquid
.
to achieve analogous WHSV.
2.3 Conversion, selectivity and yield
The following expressions were used to calculate the conversion
Table 1 Kinetics and mass transfer coefficients
of sorbitol (XSB), the selectivity to hydrogen on the basis of the
Steps
Parameters Fixed bed reactor Microchannel
sorbitol converted (SH ), the selectivity to hydrogen on the basis
2
of the gas phase products (SH ,gas), and the yield of carbon Mass transfer of sorbitol, FMT
1
2
10À1
10À3
100
sÀ1
—
a
External L–S
Intra-particle
(kLSbaLS SB
)
products (YC ):
fR
10À5
i
XSB = (FSB,in À FSB,out)/FSB,in  100%
SH = FH /(13  FSB,in  XSB/100)  100%
Surface reactions, Fk and Fk
H
cR
Reforming
kR
10À8
10À3
10À8
10À3
sÀ1
sÀ1
d
Hydrogenation kH
2
2
Mass transfer of hydrogen, FMT
2
100
10À2
10À1
10À2
—
b
SH ,gas = FH /(13/6 Â FC,gas) Â 100%
Intra-particle
External S–L
External S–G
fH
(kSLaSL H
(kSGaSG H
2
2
a
2 e
)
)
10À2
—
sÀ1
sÀ1
2
YC = FC /(FC,prods) Â 100%
i
i
a
Liquid–solid mass transfer coefficient estimated from correlations of
where FSB,in and FSB,out are the molar flows of sorbitol in the
liquid phase at the inlet and outlet respectively, FH is the molar
Wakao et al.31 for the fixed bed reactor and Kreutzer et al.24 for the
b
microchannel. Thiele modulus calculated assuming the first order
2
reaction as f = L(k/De)0.5, where L = particle diameter/6 for the bulk
catalyst, L = coating thickness for the washcoated microchannel, and De
flow of hydrogen produced, FC,gas is the molar flow of carbon
atoms in the gas phase products (i.e. alkanes and CO2), FC,i is
the molar flow of carbon atoms in the component i and FC,prods
is the total molar flow of carbon atoms in the products. Notice
that two definitions for the selectivity to hydrogen are given (i.e.
SH and SH ,gas). The first definition is preferred in this work
c
is the effective diffusivity coefficient. Reaction rate constant estimated
d
from experimental data. Reaction rate constant approximated using
the kinetic model of Brahme and Doralswamy.32,33 e Solid–gas mass
transfer coefficient estimated from correlations of Kreutzer et al.24 for
the microchannel.
2
2
since we observed a significant amount of products in the
liquid phase at all space velocities studied. The second defini-
tion is widely used in the literature10,29 and has been calculated
in this study for the sake of comparison.
to liquid ratio of RGL = 2mgas3/mliquid3. The differences in terms of
hydrodynamics of the two reactor types seem not to influence the
conversion of sorbitol. Intra-particle mass transfer does not play a
role in the overall rate of sorbitol consumption and liquid–solid
external mass transfer is orders of magnitude faster than the
reforming reaction rate (Table 1). Therefore, the bulk catalyst used
in the fixed bed reactor and the catalyst washcoat deposited on the
microchannel walls are comparable in terms of activity. The char-
acterization of these catalysts by TEM analysis confirms the same
result: both presented a similar Pt particle size (i.e. 1.5 nm), which is
generally a good indicator of a comparable catalytic activity.
3 Results and discussion
3.1 Conversion of sorbitol
Fig. 2 shows the conversion of sorbitol in the washcoated
microchannel and the fixed bed reactor as a function of the
residence time (calculated on the basis of the total Pt loading).
Both reactor types present a similar rate of consumption of
sorbitol. The two reactors were operated in the same fashion.
They were equally pre-heated, pre-wetted and they were main-
3.2 Selectivity to hydrogen
tained under analogous reactor conditions. Notice that in both The selectivity towards hydrogen as a function of conversion
cases, nitrogen was pre-mixed with the liquid reactant in a gas and the reactor type is shown in Fig. 3. In the case of the fixed
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Fig. 3 Selectivity to hydrogen versus sorbitol conversion for the fixed bed
reactor (open circles), with 2 mg Pt loading; and the microchannel reactor (full
circles), with 0.7 mg Pt loading. Reaction conditions: 220 1C, 35 bar, 1 wt% of
3
sorbitol in water, RGL = 2mgas3/mliquid at the reactor inlet, liquid flows were
adjusted in each case to achieve analogous WHSV.
Fig. 4 Schematic representation of the mass transfer and reaction steps
towards the formation of hydrogen from sorbitol. FMT is the mass transfer rate
1
of sorbitol from the liquid to the catalyst site, Fk is the reforming reaction rate, Fk
R
H
bed reactor the selectivity is relatively low (ca. 20%) in the range of
is the hydrogenation reaction rate, and FMT2 is the mass transfer rate of hydrogen
sorbitol conversion studied. Notice that these selectivity values from the catalyst site to the gas phase. (CHO)I, (CHO)II and (CHO)III are inter-
mediate polyhydroxy species.
were calculated on the basis of the total sorbitol converted (SH ).
2
The hydrogen selectivity based on the carbon in the gas phase
(SH ,gas) ranges from 63 to 78% in the case of the fixed bed reactor,
2
very close to the values found in the literature.29 In this study, we hydrogen produced (FMT + Fk ), or in other words, by the ratio
2
H
have considered that it is more appropriate to use the first FMT /(FMT + Fk ). Therefore, the increase in hydrogen selectivity
2
2
H
definition of hydrogen selectivity since a significant amount of in the case of the microchannel can be explained on the basis of
the side products of this reaction are in the liquid phase within the excellent mass transfer compared to the fixed bed reactor.
the range of space velocities studied. These products are mostly The large interfacial mass transfer rate in the microchannel
formed by dehydration and hydrogenation of sorbitol and the reactor is the result of an increased mass transfer coefficient
intermediate compounds, thus representing consumption of both and a large interfacial surface area.
sorbitol and hydrogen (i.e. increase of conversion and decrease of
In principle, the increase in hydrogen selectivity in the
hydrogen selectivity). As the sorbitol conversion increases, the microchannel reactor may be ascribed to additional factors derived
hydrogen consuming reactions contribute to a larger extent. from the distinct hydrodynamics of this reactor compared to the
Therefore, there is a decreasing trend in the hydrogen selectivity conventional fixed bed reactor. For example, fixed bed reactors
as the conversion increases. For ca. 20% o XSB 4 70% the typically present problems of catalyst wettability,30 probably not
selectivity to hydrogen is nearly constant. This tendency can be present in the microchannel reactor.26 However, an insufficient
explained by the formation of hydrogen from intermediates wettability of the catalyst is not expected in any of the two reactors.
(i.e. glycerol), which confirms that the generation of hydrogen Given the hydrophilic nature of the catalyst support and the liquid
proceeds through several consecutive steps.
medium, it is expected that the catalyst pores are rapidly filled with
Fig. 3 shows a clear superiority of the microchannel over the liquid by capillary forces during the pre-wetting procedure.
fixed bed reactor in terms of hydrogen selectivity. The selectivity is Besides, any difference in catalyst wettability would have been
nearly double in the whole conversion range. In the microchannel, reflected in the trend of sorbitol conversion of the two reactor types
it is possible to reach a larger sorbitol conversion with a greater (Fig. 2). Therefore, the enhanced mass transfer of hydrogen will be
hydrogen selectivity. As a result, the maximum production of the most important difference between the two reactors and we
hydrogen per sorbitol feed in the microchannel significantly will discuss the results accordingly. As not only the mass transfer
exceeds that of the fixed bed reactor (i.e. 4.0 vs. 1.4 moles of of hydrogen is enhanced but also that of the liquid components,
hydrogen per mole of sorbitol feed, respectively).
the rapid surface renewal of the liquid film in the washcoated
As schematically depicted in Fig. 4, once hydrogen is formed microchannel reactor in the Taylor flow regime will also influence
over the catalyst surface due to the reforming reaction, two the hydrogen selectivity. The concentration of water-soluble
competitive follow-up routes are possible: (1) hydrogen con- oxygenates on the catalyst surface determines the rate of
sumption in side reactions; and (2) the removal of hydrogen hydrogen-consuming side reactions, previously referred as Fk
from the catalyst surface via mass transfer to the gas phase,
where it cannot further react. The selectivity to hydrogen is
.
H
3.3 Mass transfer limitations
given by the relative amount of hydrogen that is transferred to Table 1 shows the reaction rate constants and the mass transfer
the gas phase (FMT ) with respect to the total amount of coefficients that quantify the speed of each step during the
2
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transformation of sorbitol to hydrogen (Fig. 4). These are the
mass transfer of sorbitol from the liquid to the catalyst active
site (FMT ), the surface reforming reaction that produces hydrogen
1
(Fk ), the reactions that consume hydrogen (Fk ), and the mass
R
H
transfer of hydrogen from the catalyst active site to the gas phase
(FMT ). While the slow kinetics of the reforming reaction are the
2
limiting factor for the sorbitol conversion rate (FMT c Fk ), the
1
R
high rate of hydrogen consumption is highly competitive with the
rate of mass transfer of hydrogen from the catalyst active site to
the gas phase (FMT E Fk ). Therefore, the excellent mass transfer
Fig. 5 Formation of the Taylor flow regime in a glass micro-mixer prior to the
washcoated microchannel reactor.
2
H
properties of the microchannel do not have a significant effect on
the conversion rate (Fig. 2) but strongly influence the selectivity of
the reaction (Fig. 3). By improving the mass transfer rate, the
hydrogen-consuming reactions are depressed and the selectivity
towards hydrogen is increased.
hydrogen partial pressure has an inhibiting effect on the
kinetics of the reforming reaction. The negative reaction order
of hydrogen concentration on the reforming kinetics, which
can be attributed to a decrease in the availability of the catalyst
active sites, might also explain the saturating tendency
observed in Fig. 2 as the conversion increases. As the hydrogen
partial pressure increases with increasing residence time, the
reforming rate decreases.
The higher selectivity to hydrogen observed in the micro-
channel reactor compared to the fixed bed reactor resulted in
higher concentrations of hydrogen in the gas phase (Fig. 7). A
stronger inhibiting effect of the hydrogen concentration on the
reaction kinetics would be expected in the case of the micro-
channel reactor. Nevertheless, no significant difference was
observed in the sorbitol conversion achieved by the two reactors
(Fig. 2). This suggests that the higher mass transfer coefficient
in the microchannel reactor leads to a reduction of the hydro-
gen concentration over the catalyst surface, thus minimizing
the inhibiting effect of hydrogen on the reforming rate.
According to Fig. 6, the hydrogen partial pressure also
presents a negative effect on the selectivity towards hydrogen.
According to the values in Table 1, the mass transfer rate of
hydrogen is improved in the microchannel reactor compared to
the fixed bed reactor due to an increase of both intra-particle
mass transfer rate (diffusion of hydrogen through the catalyst
particle) and external mass transfer rate (from the catalyst
surface to the gas phase). The intra-particle mass transfer
limitations of hydrogen were quantified in this study by estimating
the Thiele modulus (fH), although this term is typically used when
the diffusion of the reactant and the chemical reaction occur as
two steps in series. In this case, the intra-particle diffusion and the
hydrogenation reaction are two parallel steps, but still the Thiele
modulus can be used as an indicator of the ratio between the rate
of hydrogenation and the rate of hydrogen diffusion through the
pores. A value of fH 4 0.3 indicates that the intra-particle mass
transfer of hydrogen is slower than the hydrogenation reaction,30
thus predicting a more severe effect of hydrogenation (i.e. lower
hydrogen selectivity and higher yield of hydrogenated products).
The intra-particle hydrogen diffusion limitation is prevented by the
thin catalytic layer in the microchannel reactor, whereas the larger
catalyst particle size in the fixed bed reactor results in a Thiele
modulus fH c 0.3. Smaller particle sizes in the fixed bed reactor
would increase the internal diffusion rate but would represent an
increase in pressure drop, which might become significant in
larger units where longer tubes and larger liquid flow rates are
used. The external mass transfer rate of hydrogen from solid to gas
is also intensified in the microchannel reactor under the Taylor
flow regime. This is due to the large and uniformly distributed
contact area between the catalytic layer on the microchannel walls
and the gas bubbles (Fig. 5), and due to the large mass transfer
coefficient of hydrogen from the catalyst surface to the gas bubbles
through the thin liquid film formed between these two phases.21,24
The high external mass transfer rate of hydrogen has however a
milder effect on the hydrogen selectivity, since it is higher than the
hydrogenation rate for both reactors.
The negative values of SH observed at hydrogen pressures
2
above 4 bar are explained by a net mass transfer of hydrogen
from the gas phase to the catalyst surface where it is consumed
in side reactions. The total amount of hydrogen consumed is
greater than the amount of hydrogen produced.
3.4 Effect of hydrogen partial pressure
The effect of the hydrogen partial pressure on the reforming
rate and on the selectivity of APR was examined by co-feeding
mixtures of hydrogen and nitrogen at the inlet of the fixed bed
reactor. The mass transfer coefficients were constant and only
the hydrogen partial pressure was varied. Fig. 6 shows that the
Fig. 6 Effect of hydrogen partial pressure on the sorbitol conversion (open
circles) and on the selectivity to hydrogen (full circles) in the fixed bed reactor.
Reaction conditions: 220 1C, 35 bar, 1 wt% of sorbitol in water, liquid flow rate of
3
7 ml hÀ1, RGL = 2mgas3/mliquid at the reactor inlet, with increasing hydrogen
concentration in the inlet gas feed from 0 to 100%.
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microchannel is not successful. The gas produced is highly
concentrated in hydrogen, thus limiting the possibility of
hydrogen removal from the catalyst surface via mass transfer
to the gas phase. In this case, the performance of the micro-
channel reactor is analogous to that of the fixed bed reactor.
The effect of adding nitrogen as a stripping agent is beneficial
for both reactors. The conversion is increased due to a kinetic
effect of reducing hydrogen partial pressure. The selectivity also
rises due to an increase of the driving force for the mass
transfer. The increase in selectivity is larger in the case of the
microchannel reactor due to an increase in the mass transfer
coefficient.
3.6 Reaction mechanism
Fig. 7 Hydrogen partial pressure at the outlet of the reactor versus residence
time for the fixed bed reactor (open circles), with 2 mg Pt loading; and the
microchannel reactor (full circles), with 0.7 mg Pt loading. Reaction conditions:
So far we have simplified the surface chemistry to two types of
reactions: one that produces hydrogen (i.e. the reforming
reactions) and another that consumes hydrogen. Earlier inves-
tigations on APR10,28 reveal a more complicated chemistry
comprising many chemical reactions and intermediates (Fig. 9).
The general belief is that hydrogen is produced by initial dehydro-
genation of the substrate, followed by decarboxylation or
decarbonylation and water gas shift on the metal sites (i.e.
the reforming pathway).11 This pathway leads to the production
of hydrogen, CO2 and a series of smaller carbohydrates that are
sensitive to further reforming reactions. In a hydrogen rich
medium, direct hydrogenolysis of the internal C–C bonds of
sorbitol may also occur34 on the metal sites. Due to the large
yields of glycerol, we believe that direct hydrogenolysis of the
intermediate C–C bond is one of the preferential routes. In
addition to C–C cleavage reactions, sorbitol and the smaller
carbohydrates undergo dehydration of the C–O bonds, cata-
lyzed by the acid support (i.e. AlO(OH)) or the liquid medium,
followed by hydrogenation of the dehydrated intermediates on
the metal sites.11,35 Successive dehydration/hydrogenation of
the intermediates leads to a wide variety of water soluble
hydrodeoxygenated species bearing from 1 to 6 carbon atoms
with a reduced OH/C ratio. The bi-functionality of the catalyst
also favors rearrangement reactions of the substrate and the
intermediate products towards the production of organic acids,
aldehydes, ketones and ring compounds like furan deriva-
tives.29 Further dehydration and hydrogenation of these species
results in the formation of C1 to C6 alkanes.
3
220 1C, 35 bar, 1 wt% of sorbitol in water, RGL = 2mgas3/mliquid at the reactor
inlet (no hydrogen added), liquid flows were adjusted in each case to achieve
analogous WHSV.
3.5 Effect of nitrogen addition
According to our analysis, the selectivity towards hydrogen
depends on the mass transfer rate of hydrogen from the catalyst
surface to the gas phase. This was enhanced by a combination
of two factors: increase of the mass transfer coefficient (micro-
channel operating under the Taylor flow regime vs. fixed bed
reactor); and increase of the driving force for the mass transfer
by a decrease of the hydrogen concentration in the gas phase
(addition of nitrogen). The combined effect of these factors is
observed in Fig. 8, where the conversion and the selectivity to
hydrogen have been compared for two different gas to liquid
ratios at the inlet. If only liquid is fed to the microchannel
reactor and the conversion to gaseous products is limited,
the development of the desired Taylor flow regime in the
3.7 Product distribution
Fig. 10 shows the distribution of carbon within the gas and the
liquid phase products at 25% of sorbitol conversion for the two
reactor types. In both cases, the largest fraction in the gas phase
is CO2, which is the product of the reforming pathway. No CO
was detected in any of the experiments conducted. This may be
due to the low reaction temperature and the high concentration
of water that shift the water-gas-shift reaction towards the
formation of CO2. It may also indicate that CO2 is mainly
produced by decarboxylation, rather than decarbonylation
and water-gas-shift. According to the results of the present
study, it is not possible to elucidate the preferential route for
CO2 production. Further research should be conducted to
Fig. 8 Effect of the gas to liquid ratio at the reactor inlet (RGL) on the sorbitol
conversion (black area) and on the selectivity to hydrogen (gray area) for the
fixed bed reactor and the microchannel reactor. Reaction conditions: 220 1C,
35 bar, 1 wt% of sorbitol in water, liquid flow rate of 15 ml hÀ1 (fixed bed reactor)
and 6 ml hÀ1 (microchannel reactor), no hydrogen added.
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Fig. 9 Reaction pathway for the aqueous phase reforming of sorbitol.
microchannel reactor than in the fixed bed reactor, suggesting
that the reforming pathway is facilitated in the former com-
pared to the latter. This difference is attributed to the increased
mass transfer rate of the hydrogen from the catalyst active site
to the inert gas phase in the microchannel reactor, which aids
the rapid turnover of the catalyst sites and therefore favors the
rate of C–C cleavage reactions. On the other hand, the yield of
heavier liquid products, in particular C3 species, is remarkably
higher in the case of the fixed bed reactor. A breakdown of the
yield of the main components of the C3 fraction is given in
Fig. 11. Formation of glycerol, the major side product in this
fraction, is clearly more pronounced in the fixed bed reactor.
This can be explained by a more rapid hydrogenolysis of the
sorbitol due to a higher concentration of hydrogen at the
surface of the bulk catalyst in the fixed bed reactor.
Fig. 10 Yield of carbon containing products in the fixed bed reactor (black area)
and the microchannel reactor (gray area) at 25% conversion. Gas-C1: CH4; gas-
C2: C2H6; gas-C3: C3H8; liquid-C1: methanol; liquid-C2: ethanol, ethyleneglycol,
Further conversion of glycerol can explain the production of
the remaining components of the C3 fraction. Earlier studies
glycolaldehyde and acetic acid; liquid-C3: see Fig. 11; liquid-C4: 1,2-butanediol; on glycerol conversion propose the formation of glyceraldehyde
2,3-butanediol, butyric acids, 3-hydroxydihydro-2-furanone and butyrolactone;
liquid-C5: 1,2-pentanediol, pentanoic acid, 5-hydroxymethyldihydrofuran-2-one,
2-methyltetrahydrofuran, xylofuranose; and liquid-C6: 2-hexanol, isosorbide and
sorbitan. Unbalance below 5%.
unravel this question, which is outside the scope of this paper.
Minor amounts of C1–C3 alkanes were found, while only traces
of C4–C6 alkanes were detected (i.e. yields lower than 0.05%).
In the liquid phase, the C3 fraction is noticeably larger than the
rest, followed by C6, C5 and C4 species. In these fractions, a
significant amount of hydrodeoxygenated species (e.g. 1,2-pro-
panediol, butanediols, furan derivatives, isosorbide) are found.
These results suggest a significant rate of hydrogenolysis of the
intermediate C–C bond of sorbitol towards C3 products, as well
as a significant rate of the dehydration–hydrogenation reac-
tions due to a highly acidic (catalyst support/liquid phase) and
hydrogen rich environment. The conversion of these species to
alkanes is expected to occur at larger residence times.29
Fig. 11 Breakdown of yield of carbon containing products in the C3 fraction in
Fig. 10 also highlights a different chemistry in each reactor
type. The yield of CO2 is distinctively higher in the case of the
the fixed bed reactor (black area) and the microchannel reactor (gray area) at
25% conversion.
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Catalysis Science & Technology
via dehydrogenation of glycerol on the active metal and the 1,2-propanediol via hydrogenolysis of C–C and C–O bonds. There
production of hydroxyacetone via dehydration on the acid is a consistency between the products that are enhanced when
sites.36 The relative amounts of glyceraldehyde and hydroxy- increasing the hydrogen partial pressure and those that were
acetone are nearly the same in both reactor types, suggesting an favored in the fixed bed reactor compared to the microchannel
analogous balance between dehydrogenation and dehydration reactor. This suggests that the main difference between these
reactions (i.e. comparable balance metal/acid sites) in both reactors is the concentration of hydrogen at the catalyst surface,
cases. This is in line with the initial statement of the two and therefore its availability to participate in hydrogenolysis
catalytic systems being comparable in terms of their reactivity. reactions. This is in line with the previous argument of the
These two intermediates may undergo a sequence of hydro- microchannel reactor offering a more rapid removal of hydrogen
genation reactions towards 1,2-propanediol. Compared to the from the catalyst surface and thus favoring the hydrogen producing
microchannel reactor, in the fixed bed reactor the yield of reactions while preventing hydrogen consuming reactions.
1,2-propanediol relative to the total yield of C3 products is
3.8 Practical considerations
significantly higher, whereas that of the intermediates (i.e.
glyceraldehyde and hydroxyacetone) is lower. This indicates Whereas the technology of fixed bed reactors is well estab-
that the selectivity for hydrogenation reactions is remarkably lished, for a practical implementation of the microreactor
higher in the fixed bed reactor. An important amount of concept, several considerations should be taken into account.
propionic acid is also detected in both reactors. This suggests Since the catalyst is immobilized on the microchannel walls,
a significant contribution of the rearrangement reactions to the the recovery and regeneration of the catalyst is difficult. There-
overall reaction network. This acid may be produced by the fore, the catalyst stability is an important requirement. In this
reaction between glyceraldehyde and water, with a subsequent study, a washcoated microchannel was tested for 3 weeks of
sequence of dehydration and hydrogenation reactions. The time on stream. This test included numerous runs under
yield of this product with respect to the total yield of C3 various reaction conditions, as well as several intermediate
products is slightly larger in the case of the fixed bed reactor.
checks under the same reaction conditions to identify any loss
We have previously presented that the addition of hydrogen of activity. During this time, the catalytic activity remained
at the inlet of the reactor represented a decrease in the reaction constant. Further discussion on the stability of this catalytic
rate and the hydrogen selectivity. This was attributed to the system under APR conditions is given in our previous work.27
preferential hydrogenation of the substrate and its intermediates The use of nitrogen as a stripping agent introduces some
rather than the desired reforming pathway with increasing concerns regarding the production costs, which can be miti-
hydrogen partial pressures. The yield of the most significant gated by recirculation of this gas after the hydrogen content is
carbon containing products obtained at different hydrogen consumed, for example, in a PEM fuel cell. Combustion of the
partial pressures is given in Fig. 12. A higher hydrogen concen- alkane fraction of the gaseous products can be conducted on
tration is detrimental for the conversion to CO2. Instead, the the outer side of the microchannels to supply the heat of the
addition of hydrogen enhances the conversion to glycerol and reaction. Due to the large surface to volume ratio, microchan-
nel reactors are also very efficient heat exchangers.37 In a final
hydrogen generation unit, the desired production scale can be
achieved by parallelization of identical microchannels in a
stack configuration with intermediate heating. Nevertheless,
numbering up of multiphase microchannel reactors is a field
still under development. Promising results have been recently
published by our group with gas–liquid applications that can
meet a small/medium production scale.38 Scaling up of cataly-
tic washcoated microchannel reactors for this application is a
work in progress.
4 Conclusions
The Pt-based washcoated microchannel reactor for aqueous
phase reforming of sorbitol, combined with nitrogen stripping,
has excellent mass transfer from the catalyst active site to the
inert gas phase, which aided the rapid removal of hydrogen
from the catalyst surface and limited its availability to partici-
pate in undesired hydrogenation reactions. Compared to a
Fig. 12 Effect of hydrogen partial pressure on the yield of CO2, glycerol and
fixed bed reactor, the selectivity to hydrogen obtained in the
microchannel reactor was increased by a factor of 2 regardless
1,2-propanediol in the fixed bed reactor. Reaction conditions: 220 1C, 35 bar,
3
1 wt% of sorbitol in water, liquid flow rate of 7 ml hÀ1, RGL = 2mgas3/mliquid
of the sorbitol conversion. The yield of hydrogen was increased
from 1.4 to 4 moles per mole of sorbitol fed. The mass transfer
at the inlet, with increasing hydrogen concentration in the inlet gas feed from
0 (light gray area) to 100% (black area).
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11 L. Vilcocq, A. Cabiac, C. Especel, S. Lacombe and D. Duprez,
Catal. Today, 2012, 189, 117–122.
12 D. J. M. de Vlieger, B. L. Mojet, L. Lefferts and K. Seshan,
J. Catal., 2012, 292, 239–245.
13 G. Wen, Y. Xu, Z. Xu and Z. Tian, Catal. Commun., 2010, 11,
522–526.
14 A. Yousuf, Waste Manage., 2012, 32, 2061–2067.
15 S. Murat Sen, C. A. Henao, D. J. Braden, J. A. Dumesic and
C. T. Maravelias, Chem. Eng. Sci., 2012, 67, 57–67.
16 S. G. Wettstein, D. M. Alonso, E. I. Gu¨rbu¨z and
J. a. Dumesic, Curr. Opin. Chem. Eng., 2012, 1, 218–224.
17 J. W. Han and H. Lee, Catal. Commun., 2012, 19, 115–118.
18 R. R. Davda and J. a. Dumesic, Chem. Commun., 2004, 36–37.
19 J. H. Yang, H.-J. Kim, D. H. Chun, H.-T. Lee, J.-C. Hong, H. Jung
and J.-I. Yang, Fuel Process. Technol., 2010, 91, 285–289.
20 A. N. Tsoligkas, M. Simmons, J. Wood and C. Frost, Catal.
Today, 2007, 128, 36–46.
21 M. T. Kreutzer, F. Kapteijn, J. A. Moulijn and J. J. Heiszwolf,
Chem. Eng. Sci., 2005, 60, 5895–5916.
22 P. Angeli and A. Gavriilidis, Proc. Inst. Mech. Eng., Part C,
2008, 222, 737–751.
23 M. J. F. Warnier, M. H. J. M. Croon, E. V. Rebrov and
J. C. Schouten, Microfluid. Nanofluid., 2009, 8, 33–45.
24 M. T. Kreutzer, P. Du, J. J. Heiszwolf, F. Kapteijn and
J. A. Moulijn, Chem. Eng. Sci., 2001, 56, 6015–6023.
25 E. V. Rebrov, A. Berenguer-Murcia, A. E. H. Wheatley, B. F. G.
Johnson and J. C. Schouten, Chim. Oggi, 2009, 27, 4–7.
26 N. Shao, W. Salman, A. Gavriilidis and P. Angeli, Int. J. Heat
Fluid Flow, 2008, 29, 1603–1611.
27 M. F. N. D’Angelo, V. Ordomsky, V. Paunovic, J. van der
Schaaf, J. C. Schouten and T. A. Nijhuis, ChemSusChem,
2013, DOI: 10.1002/cssc.201200974.
28 T. Jiang, T. Wang, L. Ma, Y. Li, Q. Zhang and X. Zhang, Appl.
Energy, 2012, 90, 51–57.
properties of the microchannel reactor did not affect the rate of
sorbitol consumption, which is limited by the kinetics of the
reforming reaction. Hydrogen inhibits the reforming reaction
and favors side – hydrogen consuming – reactions. A decrease
of the hydrogen partial pressure is beneficial in both the
microchannel reactor and the fixed bed reactor. These benefits
are more significant for the microchannel reactor, since it
combines the effect of a reduced hydrogen partial pressure in
the gas phase with a high mass transfer coefficient, resulting in
a greater reduction of the hydrogen concentration at the
catalyst. This has significant implications for the product
distribution. The analysis of the products reveals an important
contribution of the hydrogenolysis of the C–C and C–O bonds
of sorbitol and its intermediates as an important source of
hydrogen selectivity loss. The most significant products of the
hydrogen consuming reactions are water-soluble C3 species
like glycerol and 1,2-propanediol, followed by heavier hydro-
deoxygenated compounds in the C6, C5 and C4 fractions. The
formation of these species rises when the hydrogen partial
pressure increases. Since the hydrogen concentration at the
catalyst surface is lower in the microchannel reactor, the yield
of the hydrogenated products is severely depressed while the
products of the reforming route (i.e. CO2) are remarkably
increased.
Acknowledgements
This research has been performed within the framework of the
CatchBio program. The authors gratefully acknowledge the
support of the Smart Mix Program of the Netherlands Ministry
of Economic Affairs and the Netherlands Ministry of Education,
Culture and Science. Yagmur Karakus is acknowledged for her
contribution to this work.
29 A. V. Kirilin, A. V. Tokarev, L. M. Kustov, T. Salmi,
J.-P. Mikkola and D. Y. Murzin, Appl. Catal., A, 2012,
435–436, 172–180.
30 V. V. Renade, R. V. Chaudgari and P. R. Gunjal, Trickle Bed
Reactors, Elsevier, 2011, p. 285.
31 N. Wakao, S. Kaguei and T. Funazkri, Chem. Eng. Sci., 1979,
34, 325–336.
32 E. Crezee, Appl. Catal., A, 2003, 251, 1–17.
References
1 Y. Wang, K. S. Chen, J. Mishler, S. C. Cho and X. C. Adroher,
Appl. Energy, 2011, 88, 981–1007.
2 R. D. Cortright, R. R. Davda and J. A. Dumesic, Nature, 2002,
418, 964–967.
3 J. W. Shabaker, R. R. Davda, G. W. Huber, R. D. Cortright
and J. A. Dumesic, J. Catal., 2003, 215, 344–352.
4 G. W. Huber and J. A. Dumesic, Catal. Today, 2006, 111, 33 P. H. Brahme and K. Doralswamy, Ind. Eng. Chem. Process
119–132. Des. Dev., 1976, 15, 130.
5 N. C. Juben, W. H. George and A. D. James, Angew. Chem., 34 M. Banu, P. Venuvanalingam, R. Shanmugam, B. Viswanathan
Int. Ed., 2007, 46, 7164–7183. and S. Sivasanker, Top. Catal., 2012, 55, 897–907.
6 D. O. Ozgu¨r and B. Z. Uysal, Biomass Bioenergy, 2011, 35, 35 R. Weingarten, G. A. Tompsett, W. C. Conner Jr and
822–826. G. W. Huber, J. Catal., 2011, 279, 174–182.
7 S. Kandoi, J. Greeley, D. Simonetti, J. Shabaker, J. A. Dumesic and 36 D. L. King, L. Zhang, G. Xia, A. M. Karim, D. J. Heldebrant,
¨
M. Mavrikakis, J. Phys. Chem. C, 2011, 115, 961–971.
8 R. R. Davda, J. W. Shabaker, G. W. Huber, R. D. Cortright
and J. A. Dumesic, Appl. Catal., B, 2005, 56, 171–186.
9 G. W. Huber, J. W. Shabaker, S. T. Evans and J. A. Dumesic,
Appl. Catal., B, 2006, 62, 226–235.
X. Wang, T. Peterson and Y. Wang, Appl. Catal., B, 2010, 99,
206–213.
37 E. V. Rebrov, J. C. Schouten and M. H. J. M. de Croon, Chem.
Eng. Sci., 2011, 66, 1374–1393.
38 M. Al-Rawashdeh, J. Zalucky, C. Mu¨ller, T. A. Nijhuis,
V. Hessel and J. C. Schouten, Ind. Eng. Chem. Res., 2013,
DOI: 10.1021/ie4009277.
10 G. W. Huber, R. D. Cortright and J. a. Dumesic, Angew.
Chem., Int. Ed., 2004, 43, 1549–1551.
c
2842 Catal. Sci. Technol., 2013, 3, 2834--2842
This journal is The Royal Society of Chemistry 2013