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membrane reactor
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kne
J.
cien
cked bed membrane reactors, has been compared for ultra-pure
profiles have been compared. The extent of mass and heat transfer limitations in the
different reactors has been evaluated, and strategies to decrease (or avoid) these limita-
different process steps such as feed gas preheating and pre-
tion) unit is used to achieve the desired hydrogen purity. In
view of thermodynamic limitations and the high endo-
thermicity of steam reforming, heat transfer at high temper-
atures (850950 C) is required, where excess of steam is usedto avoid carbon deposition (typical feed H2O/CH4 molar ratios
process units can be strongly decreased and the total required
based membrane reactors for pure hydrogen production
have been proposed in literature for different reaction
systems such as methanol reforming [8,9], ethanol reforming
[10], and autothermal reforming [11]. In particular, Gallucci
and Basile [9] have demonstrated the feasibility of packed bed
* Corresponding author. Tel.: 31 53 489 2370; fax: 31 53 489 2882.
Avai lab le at www.sc iencedi rect .com
w.
i n t e r n a t i on a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 0E-mail address: [email protected] (F. Gallucci).treatment (for example hydrodesulphurisation), primary and
secondary reformers (often multi-tubular fixed-bed reactors)
and high and low temperature shift converters, CO2 removal
and methanation units. Often a PSA (Pressure Swing Adsorp-
reactor volume can be largely reduced, while higher methane
conversions and hydrogen yields beyond thermodynamic
equilibrium limitations can be achieved, at lower tempera-
tures and with higher overall energy efficiencies [47]. Pd-The recent advances in Polymer Electrolyte Membrane Fuel
Cells (PEMFC) for small or medium scale applications make
the production of ultra-pure hydrogen a challenging topic in
energy conversion. On an industrial scale, most of the
hydrogen is currently produced via steam reforming of
methane (SRM). The traditional SRM process consists of
complex heat integration and the associated uneconomical
downscaling make this route inefficient. A high degree of
process integration and process intensification can be
accomplished by integrating hydrogen perm-selective
membranes in the steam reformer [2,3]. Via the integration
of hydrogen perm-selective membranes, the number ofAccepted 11 February 2010
Available online 15 March 2010
Keywords:
Hydrogen production
Membrane reactors
Fluidized bed
Heat and mass transfer limitations
1. Introduction0360-3199/$ see front matter 2010 Profesdoi:10.1016/j.ijhydene.2010.02.050tions have been proposed for the fluidized bed membrane reactor concept. The results
show that the packed bed membrane reactor requires in some conditions double
membrane area with respect to the fluidized bed membrane reactor.
2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.
25) [1]. For the production of ultra-pure hydrogen for small
scale application, the large number of process units withReceived in revised form
8 February 2010hydrogen production via methane reforming. Using detailed theoretical models, the
required membrane area to reach a given conversion and the prevailing temperatureReceived 16 November 2009 fluidized bed and paArticle history: In this theoretical work the performance of different membrane reactor concepts, bothTheoretical comparison of pacmembrane reactors for metha
Fausto Gallucci*, Martin Van Sintannaland,
Fundamentals of Chemical Reaction Engineering Group, Faculty of S
Enschede, The Netherlands
a r t i c l e i n f o a b s t r a c t
journa l homepage : wwsor T. Nejat Veziroglu. Pued bed and fluidized bedreforming
A.M. Kuipers
ce and Technology, (IMPACT) University of Twente,
e lsev ie r . com/ loca te /heblished by Elsevier Ltd. All rights reserved.
self supportedmembrane reactors for different fuels. Simakov
and Sheintuch [11] have developed a small scale autothermal
membrane reformer by coupling an exothermic reaction
(carried out in a separate compartment of the reactor) with the
endothermic methane reforming with hydrogen recovery
through Pd-based membranes.
Both packed bed membrane reactors [12,13] and fluidized
bed membrane reactors [1416] have already been presented
in literature for the reforming ofmethane and advantages and
disadvantages of both concepts have already been discussed.
However a direct quantitative comparison of the two concepts
at the same conditions is lacking. In this paper a direct
comparison between the two concepts is performed for ultra-
pure hydrogen production via methane reforming using
detailed theoretical models. The extent of mass and heat
transfer limitations in the different reactors is evaluated, and
strategies to decrease (or avoid) these limitations are
proposed.
2. Reactor configurations
order to increase the hydrogen permeation through the
membranes. With the fluidized bed membrane reactor
a virtually isothermal condition can be achieved and bed-to-
membrane mass transfer limitations are largely avoided. On
the other hand, bubble-to-emulsion phase mass transfer
limitations and the extent of gas back-mixing could deterio-
rate its performance. In particular, the use of membranes
inside the reactor could decrease the extent of back-mixing
and can also help in decreasing the bubble diameter,
enhancing the bubble-to-emulsion phase mass transfer. With
the help of a two-phase phenomenological reactor model, the
effect of bubble-to-emulsion phase mass transfer limitations
on the temperature profiles and reactor performance. The
influence of the reactor and particle dimensions is investigated.
i n t e r n a t i o n a l j o u rn a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 0 71432.1. Fluidized bed membrane reactor concept
A schematic representation of the considered fluidized bed
membrane reactor configuration is shown in Fig. 1. The
methane steam reforming takes place in a fluidized bed
operated in the bubbling regime. Pure hydrogen is recovered
via dead-end Pd-based membranes inserted into the fluidized
bed. The hydrogen recovery through the membranes shifts
both the methane reforming and water gas shift reactions
towards the products resulting in higher conversion and
hydrogen yield compared with a conventional reformer.
A pressure difference between the reaction side (fluidized bed)
and the permeation side (membrane lumen) is applied inFig. 1 Scheme of the fluidized bed membrane reactor.3. Reactor models
In both reactor concepts the reactions considered are the
following:
CH4 H2O5CO 3H2 (1)
COH2O5CO2 H2 (2)Where the reaction rate expressions are taken from Numa-
guchi et al. [17]:
r1 k1pCH4pH2O p3H2pCO=Keq;1
p1:596H2O
(3)and gas back-mixing are quantified and a possible strategy to
decrease these limitations is proposed.
2.2. Packed bed membrane reactor concept
A typical tube-in-tube packed bed membrane reactor config-
uration is considered (see Fig. 2). The catalyst is packed in the
tube side of the membrane while pure hydrogen is recovered
in the shell side of the reactor. Also in this case, a pressure
difference between the reaction side and the permeation side
is applied.
The reactor has been studied with both a 1D model and
adetailed2Dmodel inorder to identify theextentofwall-to-bed
heat transfer limitations and the bed-to-membrane mass
transfer limitations (concentrationpolarization) and their effectFig. 2 Scheme of the packed bed membrane reactor.
the bubble phase, distributed according to the local bubble
fraction. The gas extracted from the emulsion phase is
Bubble phase component mass balances1
s sXnc
(8)
i n t e r n a t i on a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 07144subsequently instantaneously replenished via exchange
from the bubble phase (to maintain the emulsion phase at
minimum fluidization conditions) (following Deshmukh
et al. [19,20]).
A uniform temperature is assumed throughout an entiresection of the fluidized bed, assuming no heat losses to the
surroundings (adiabatic conditions) and no heat transfer
limitations between the bubble phase and the emulsion
phase [21,22].
The mass and heat balance equations are as follows:
Total mass balance
usb;n1ATrb;n1 usb;nATrb;n use;n1ATre;n1 use;nATre;n
Xnc n
f00membranei;mol Mw;iAmembraneeb;nr2 k2pCOpH2O pH2pCO2=Keq;2
pH2O
(4)
The contribution by homogeneous gas phase reactions can
safely be neglected for this reaction system.
Themembranes considered in this study are Pd-based thin
dense layer supported membranes prepared by electroless
plating by ECN (Energy research Center of the Netherlands).
The hydrogen permeation rate through the palladium
membranes follows the Richardson equation:
JH2 P0e$eEa=RT$hplumenH2
0:5pshellH2
0:5i(5)
where the values of the apparent activation energy Ea and pre-
exponential factor Pe0 (in agreement with [18]) are
12,540 Jmol1 and 2.21 1003mol s1m2 Pa0.5, respectively.
3.1. Fluidized bed membrane reactor model
A typical 1D two-phase model for a membrane assisted
fluidized bed reactor has been used for the simulation of the
fluidized bed membrane reactor based on the following main
assumptions:
Dead-end hydrogen perm-selective membranes are inte-grated in the reactor.
The reactor consists of two-phases, viz. the bubble phaseand the emulsion phase.
The gas flowing through the emulsion phase is consideredto be completely mixed in each section and at incipient
fluidization conditions.
The bubble phase gas is assumed to be in plug flow (i.e., largenumber of CSTRs), where the bubble size and the bubble rise
velocity change for each section.
The heterogeneous reactions (methane steam reformingand water gas shift reactions) take place only in the emul-
sion phase, assuming that the bubble phase is free of cata-
lyst particles.
Gas removed from the fluidized bed via membranes isassumed to be extracted from both the emulsion phase andi1
f00membranei;mol Mw;iAmembrane1 eb;no 0 61 Note that
Transfer term
Q use;n1ATre;n1 use;nATre;n Xnci1
f00membranei;mol Amembrane1 eb;n
Xnci1
Kbe;i;nVb;nrb;nwb;i;n we;i;n
9where
use;nAT ue;nAT1 eb;nusb;0AT utotATeb;0use;0AT utotAT1 eb;0
Energy balance (in case of energy supply inside the reactor)
Xnci1
HTfeedi
usb;n0ATrb;i;n0 use;n0ATre;i;n0
Xnci1
HTouti
usb;nNATrb;i;nN use;nNATre;i;nN
(Xnc
i1HTouti
f00membranei;mol Mw;iATeb;n
f00membranei;mol Mw;iAT1 eb;n)
E 0 10
where E depends on the kind of energy supply used (see e.g.
Gallucci et al. [15]). All the parameters used are described
elsewhere [15].
3.2. Packed bed membrane reactor 1D and 2D models
The axial temperature and concentration profiles in both
reaction and permeation compartments were modeled with
a 1D reactor model. The mass and energy conservation
equations read:
Mass conservation equations:
vrsgu
sg
vz
CMSD 2pris2cell pr20
J (11)ub;n1ATrb;n1 ub;nATrb;n i1
Kbe;i;nVb;nrb;n wb;i;n we;i;n
Xnci1
f00membranei;mol Mw;iAmembraneeb;n
we;i;nSFQ wb;i;nSFQ 0 7Emulsion phase component mass balances1
use;n1ATre;n1 use;nATre;n Xnci1
Kbe;i;nVb;nrb;nwb;i;n we;i;n
Xnci1
f}membranei;mol Mw;iAmembrane1 eb;n
0@Xnrxn
j1nj;irj
1AVe;nrp;n1 ee we;i;nSFQ wb;i;nSFQ 0SFx x if x > 00 if x 0 :
0 lz
i n t e r n a t i o n a l j o u rn a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 0 7145vrtgu
tg
vz
CMSD2riJ (12)
rsgusg
vwsj;gvz
vvz
rsgD
sz
vwsj;gvz
rsj;g (13)
rtgutg
vwtj;gvz
vvz
rtgD
tz
vwtj;gvz
!61 3tg
dtp
jtj
1wsj;g
CMsD
2riJ jtj rtj;s 0
(14)
Energy conservation equations:
3sgr
sgC
sp;g rsbulkCsp;g
vTsvt
rsgusgCsp;gvTs
vz vvz
lseff
vTs
vz
Xj
rsj;gHsj;g
2pr0s2cell pr20
astwTtw Ts (15)
3tgr
tgC
tp;g rtbulkCtp;g
vTtvt
rtgutgCtp;gvTt
vz vvz
lteff
vTt
vz
Xj
rtj;gHtj;g
2riattw
Ttw Tt (16)
rtwg Ctwp;g
vTtw
vt vvz
ltweff
vTtw
vz
2r0r20 r2i
astwTs Ttw
2rir20 r2i
attwTt Ttw (17)
The remaining constitutive correlations needed to close the
model are summarised elsewhere (see Smit et al. [23])
The 2D model consists of a pseudo-homogeneous, 2D
reactor model consisting of the total gas phase continuity and
NavierStokes equations augmented with gas phase compo-
nent mass balances and the overall energy balance (see e.g.
Tiemersma et al. [12]). The model is based on a standard
dispersion model which describes the gas phase mass and
energy transport as convective flowwith superimposed radial
and axial dispersion.
The following assumptions have been made in this model:
The gas bulk can be described as an ideal Newtonian fluid. The porosity profiles have been accounted in the model.
The model equations in 2D axisymmetrical cylindrical co-
ordinates are:
Continuity equation
v3rg
vt V$
3rgu
0 (18)
Total momentum balance equation
v
vt
3rgu
V$
3rguu
3Vp b3rgu V$
3sg 3rgg (19)
Friction coefficient
b 1501 32
33
mg
rgd2p 1:751 3
333jujdp
(20)
wherejuj u2r u2z
q(21)where source terms equals:
Sr;i 1 3rSXnrj1
rjDHj for j 1;2;.;nr (28)
4. Results and discussion
All the simulations for both the FBMR and PBMR have been
performed without sweep gas and considering vacuum in the
permeation side. A comparison between the two reactor
concepts is carried out based on the membrane area required
for a target conversion, because it is anticipated that
membrane costs is the most important parameter in deter-
mining the reactor investment costs. A first comparison has
been made between the fluidized bed membrane reactor
model and the 1D packed bed reactor model at ideal condi-
tions: isothermal conditions and absence of mass and heat
transfer limitations, i.e., the number of grid cells of the 1D
model is set equal to number of CSTRs in theMAFBmodel. The
results show that as expected in these conditions the two
reactors give identical performance in terms of membrane
area required for a given conversion. In this way it has been
verified that the two models are working properly. The
following simulations have been performed with a heat flux
through the reactor walls. The main difference between the
fluidized bed and the packed bed membrane reactors is
related to the heat management. In fact, for the fluidized bed
membrane reactor it is well known that a virtually isothermal
condition can be achieved while for the packed bed
membrane reactor a temperature drop in the first part of therg Mgp
RTgideal gas (22)
Newtonian fluid
sg lg 23mg
V$uI mg
hVu VuT
i(23)
Porosity profile (Hunt and Tien [24])
3r 30 1 30exp 6r0 r
dp
(24)
Component mass balance
v
vt
3rgwi
V$
3rguwi
V$
rgDi$Vwi
Sr;i
with Di Dr;i 00 Dz;i
(25)
where source terms equals:
Sr;i 1 3rSMiXnrj1
ni;jrj for i 1;2;.;nc (26)
Energy balance
3rgCp;g 1 3rSCp;S
vTvt
Cp;gV$3rguT
V$li$VT Sh with li
lr 0
(27)reactor is always observed irrespective of the profile of
temperature at the reactor wall.
z, m
0. .2 0. .6 0. .0 1. .4 1.0 0 4 0 8 1 2 1 6 1.8 2.0
Tem
peratu
re, K
860
880
900
920
940
960
980
dp = 0.0005 m
dp = 0.0015 m
dp = 0.0025 m
Fig. 3 Axial temperature profile in a packed bed
membrane reactor.
z, m
0. .2 0. .6 0. .0 1. .4 1.0 0 4 0 8 1 2 1 6 1.8 2.0
Tem
peratu
re, K
860
880
900
920
940
960
980
Wall temperatureReaction temperature
Pre-reforming zone
Fig. 5 Axial temperature profile in a packed bed
membrane reactor with pre-reforming zone.
i n t e r n a t i on a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 07146A typical result for the axial temperature profile in the
packed bed membrane reactor is reported in Fig. 3, indeed
showing a temperature drop of 80100 K in the first part of the
reactor which can give stability and sealing problems for the
membrane. In fact, the membrane material should withstand
a large axial temperature gradient, which might cause the
detachment of the Pd-based layer from the support with
consequent loss in perm-selectivity. Moreover, the first part of
the membrane is not effectively used since it is working at
a relatively low temperature which, following Richardsons
equation, results in a lower hydrogen permeation flux. The
decrease in temperature at the beginning of the reactor also
gives a decrease in the reaction rate. The result is an increase
of the membrane area required for a specified conversion
compared to the case of constant temperature. In particular,
the membrane area required increases by about 21% ifcompared with an isothermal operation (which is only
possible in a fluidized bed membrane reactor). The effect of
the particle size on the temperature profile is negligible as also
z, m
0. .2 0. .6 0. .0 1. .4 1.0 0 4 0 8 1 2 1 6 1.8 2.0
CH
4 co
nversio
n, -
0.0
0.2
0.4
0.6
0.8
1.0
dp = 0.0005 m
dp = 0.0015 m
dp = 0.0025 m
Fig. 4 Methane conversion for different particle size in
a packed bed membrane reactor.indicated in the same figure. By changing the particle size the
combination of the temperature change and the change in the
effectiveness factor is counterbalanced so that the final
conversion is practically the same for different particle
diameters (see Fig. 4).
A way to overcome the problem of the temperature drop is
the use of a pre-reforming zone (in our case about 2025 cm)
where no membrane is applied. In this case (pre-reforming
section 25 cm) the membrane is used at an almost constant
temperature (maximum temperature difference 28 K) so that
the stability problems are prevented and the membrane is
effectively used resulting in a lower membrane area needed
for a given conversion (i.e., slightly longer packed bed, but
smaller membrane area). The results are shown in Fig. 5. In
these conditions the increase of membrane area with respect
to an ideal fluidized bed membrane reactor is 13%.
Another difference that can occur between a packed bed
and a fluidized bed is mass transfer limitations between thebed and the membrane wall which are present in the packed
bed but not in the fluidized bed. To investigate the extent and
r/R, -
0. .2 0. .6 0.0 0 4 0 8 1.0
H2 w
eig
ht fra
ctio
n, -
0.01
0.02
0.03
0.04
0.05
0.06
z/L = 0.1
z/L = 0.5
z/L = 0.8
Fig. 6 Radial profile of the H2 weight fraction for the
isothermal operation mode.
Fig. 8 reports the relative H2 weight fraction (defined as the
actual H2 weight fraction divided by the H2 weight fraction at
the catalyst center) as a function of the radial position. The
results reported in the Fig. 8 suggest that at higher membrane
permeabilities mass transport limitations to the membrane
Fig. 9 Schematic representation of the membrane reactor
concept with bubble increasing in size.
r/R, -
0. .2 0. .6 0.0 0 4 0 8 1.0
H2 w
eig
ht fractio
n, -
0.025
0.030
0.035
0.040
0.045
0.050
0.055
0.060
Original H2permeation rate
2 * Original H2
permeation rate
5 * Original H2
permeation rate
z/L = 0.1
Fig. 7 Radial profile of the H2 weight fraction for the
isothermal operation mode at different hydrogen
i n t e r n a t i o n a l j o u rn a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 0 7147the influence of the concentration polarization a 2D model
was used.
First the radial H2 concentration profiles have been calcu-
lated at different axial positions at isothermal conditions. As
can be seen in Fig. 6, radial concentration profiles are present
but not really pronounced. It can be concluded that for the
present membranes and for small membrane diameters (1 cm
in the simulation shown in the figure), the bed-to-wall mass
transfer limitations have a negligible influence on the
membrane area.
In view of further developments and optimization of Pd-
based membranes, higher membrane fluxes will become
possible in the near future. Whether concentration polariza-
tion will occur with increased permeability was investigated
numerically. Simulation results where the membrane
permeability was increased with a factor of 2 and 5. As shown
permeabilities.in Figs. 7 and 8, an increase of hydrogen permeability causes
a significant increase of the concentration polarization, even
in membrane tubes with a diameter as small as 1 cm.
r/R, -
0. .2 0. .6 0.0 0 4 0 8 1.0
Re
la
tiv
e H
2 w
eig
ht fractio
n, -
0.0
0.2
0.4
0.6
0.8
1.0
Original H2
permeation rate
2 * Original H2
permeation rate
5 * Original H2
permeation rate
z/L = 0.1
Fig. 8 Relative H2 weight fraction for the isothermal
operation mode at different hydrogen permeabilities.wall will negatively affect the reactor performance resulting in
an increased H2 slip through the reactor exhaust. The 2D
model can be further applied to quantify the effects of mass
Factor multiplying the mass transfer coefficient, -
1e+0 1e+1 1e+2 1e+3 1e+4 1e+5
CH
4 co
nversio
n, -
0.945
0.950
0.955
0.960
0.965
0.970
0.975
0.980
Mass transfer limitations calculated
as Fluidized Bed without internals
No mass transfer limitations
Fig. 10 Effects of bubble-to-emulsion phase mass transfer
limitations on the conversion (FBMR).
itself increases by increasing the reactor length as schemati-
cally indicated in Fig. 9.
As a result of this bubble increase, themethane conversion
decreases as indicated in Fig. 10. The figure shows that the
methane conversion decreases by increasing the mass
transfer limitations. In case of mass transfer limitations
calculated as a fluidized bed reactor without internals (worst
case) the methane conversion decreases tremendously as
compared with the case without mass transfer limitations
(previously indicated as ideal condition for fluidized bed
Membrane area, m2
3 4 5 6 7 8
CH
4 co
nversio
n, -
0.950
0.955
0.960
0.965
0.970
0.975
0.980
Fig. 11 Membrane area needed for a given conversion in
Table 1 Comparison between staged fluidized bed and1D packed bed.
T, C P, bar Fluidized bed (5 stages) Packed bed (1D)
700 20 3.24 3.94
i n t e r n a t i on a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 07148transfer limitations at different membrane diameters, and
different operating conditions.
Concerning the fluidized bed membrane reactor, an
important transfer limitation affecting its performance is the
mass transfer limitation between the bubble phase and the
emulsion phase. In fact, the gas transported inside the bubbles
should be exchanged with the emulsion phase to react. A high
mass transfer limitation (low mass transfer coefficient)
between the bubble phase and the emulsion phase results in
a larger gas slip via the bubble phase and a lower conversion
degree. In our fluidized bed membrane reactor model the
bubble-to-emulsion phases mass transfer coefficient is
calculated with the equations derived for a fluidized bed
without internals. Although the internals (solid membranes)
should enhance the mass transfer characteristics of the bed,
at the moment, reliable equations for bubble-to-emulsion
phases mass transfer coefficient for fluidized bed with
inserts are not available.
For a fair comparison with the packed bed membrane
reactor, the fluidized bed has been simulated with the same
case of mass transfer limitations (FBMR).membrane area but also with the same bed length. As
a matter of fact, the bubble-to-emulsion phase mass transfer
limitation increases with increasing bubble diameter, which
Number of stages
2 4 6 8 10 12
CH
4 co
nversio
n, -
0.945
0.950
0.955
0.960
0.965
0.970
0.975
0.980
Fig. 12 Conversion reached for a given area in case of
mass transfer limitations for different stages (FBMR).reactor). The figure also shows that by improving the mass
transfer by a factor of 10 results in a conversion close to the
ideal case without mass transfer limitations.
In order to achieve the same conversion degree of a fluid-
ized bed membrane reactor without mass transfer limitations
the membrane area installed in the reactor needs to be
increased as indicated in the following Fig. 11.
The membrane area required in case of mass transfer
limitations increases 2.4 times with respect to the case
without limitations as reported in the figure.
However, it has to be pointed out that the use ofmembranes
inside the bed leads to a decrease of the bubble size (both
because of gas extraction through themembranes and because
of break-up of bubbles by the solid membrane tubes) and
a consequent decrease of the mass transfer limitations. Fig. 10
shows that a decrease of 10 times in the mass transfer limita-
tions is enough to reach the limit conversion required. Thus,
a more detailed experimental work on the determination of
the bubble-to-emulsion phase mass transfer coefficient in
a fluidized bed with internals should be carried out.
On the other hand, even considering the worst case
(bubble-to-emulsion phase mass transfer coefficient equal to
a fluidized bedwithout internals) themass transfer problem in
the fluidized bed can be easily circumvented. In fact, themass
transfer resistance is higher when the bubble diameter
H4 C
on
ve
rs
io
n, -
0.7
0.8
0.9
1.0
Staged Fluidized bed
2D isothermal modelz/L, -
0.05 0.10 0.15 0.20 0.25
C
0.5
0.6
Fig. 13 Comparison between a staged fluidized bed and
a packed bed with 2D model. 5 bar reaction pressure.
i n t e r n a t i o n a l j o u rn a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 0 7149becomes larger, and the bubble diameter increases with
increasing of the bed height, so that we can reduce the bubble
diameter by inserting stagers such as meshing wires at
different reactor heights (i.e., staging the fluidized bed reactor).
InFig. 12 the conversion inafluidizedbedwithmass transfer
limitations is shown for different numbers of stages. For these
simulations the axial position of the stagers has not been opti-
mized. This means that the distance between two stages is
constant for one simulation and it is given by dividing the total
lengthof thereactorby thenumberofstages. Ina futureworkan
optimization of the axial position for stagers will be carried out.
The area used in this simulation was kept the same as was
needed in the case of no mass transfer limitations. From the
figure it can be seen that the conversion required can be
achieved already with 34 stages. Thus, dividing the reactor in
different stages completely circumvents the problems ofmass
transfer limitation for the fluidized bed membrane reactor.
A direct comparison (in terms of membrane area required)
between theperformancesof a stagedfluidizedbedmembrane
reactor and a packed bed membrane reactor simulated in
1D (no effects of bed-to-wall mass transfer limitations) is
reported in Table 1 (for a 50 Nm3/h hydrogen production).
The packed bed membrane reactor needs around 22%
larger membrane area if compared with a staged fluidized bed
membrane reactor. Moreover, the membrane in the packed
bed is exposed to an axial temperature variation of about 100 K
in the first 25 cm of reactor length, with possible stability and
sealing.
The fluidized bed becomes even more advantageous
compared topackedbedmembranereactorwhen theeffects of
concentration polarization on the performances of packed bed
are considered. A comparison in terms ofmethane conversion
between the staged fluidized bed membrane reactor and the
packed bed membrane reactor simulated with the 2D
isothermal model (effect of bed-to-wall mass transfer) is
reported in Fig. 13. The effect of bed-to-wall mass transfer
limitations in the packed bed reactor results in a decrease of
methane conversion compared to the fluidized bedmembrane
reactor. An overview of themembrane reactor increase due to
the effects of bed-to-wall mass transfer limitations is reported
inTable 2. In theworst case (complete conversion requiredand
1 bar reaction pressure) the packed bed membrane reactor
Table 2 Increase of membrane area with respect to theFBMR due to bed-to-membrane mass transfer limitation.
Conversiondegree
Preaction 1bar
Preaction 5bar
Preaction 10bar
0.975 85.4% 64.5% 60.0%
1 98.7% 81.3% 79.5%requires almost double themembrane areawith respect to the
staged fluidized bed membrane reactor.
Finally, we can state that the more evident advantages of
a fluidized bed reactor with respect to a packed bed membrane
reactor are: constant temperature along the reactor and better
heat integration (see Gallucci et al. [14,15]), no mass transfer
limitationbetweenthefluidizedbedand themembranesurface.
Some disadvantages such as the erosion problems and
horizontal membrane sealing should be further studied
experimentally.The authors are grateful to the Dutch Ministry of Economic
affairs for financial support of this work in the EOS program
(Project EOSLT05010).
Table of symbols
AT Area of bed cross section, m2
Amembrane,n Membrane surface area per cell, n, m2
CSTR Continuously stirred tank reactor,
dp Particle diameter, m
Cp Heat capacity, J/(kg K)
D Dispersion coefficient, m2/s
Dg Gas diffusivity, m2/s
eb Bubble phase fraction,
ee Emulsion phase fraction,
Ea Activation energy for hydrogen permeation, J/mol
g Gravitational acceleration (9.81), m/s2Hj Enthalpy of specie j, J/mol
HTi;x Enthalpy of component i at temperature T at position
x, J/mol
J Permeation flux through membrane, mol/(m2 s)
jj Mass flux component j, mol/(m2 s)
ki Reaction rate constant for ith reaction
Kbe,i,n Bubble-to-emulsion phase mass transfer coefficient
for component I in cell, n, s1
Keq,i Equilibrium constant for jth reaction [depending on
the reaction]5. Conclusions
In this work, two different membrane reactor concepts for
the H2 production via methane steam reforming have been
compared via detailed models. It has been shown that both
concepts may suffer from mass transfer limitations. For the
fluidized bed membrane reactor the mass transfer limitations
occur between the bubble phase and the emulsion phase. The
effect of these mass transfer limitations on the membrane
area required is quite significant. However, these mass trans-
fer limitations can be easily circumvented by staging the
fluidized bed with consequent break-up of bubbles and
decrease of mass transfer limitations. For the packed bed
membrane reactor, the mass transfer limitations occur
between the catalytic bed and the membrane area (concen-
tration polarization). These mass transfer limitations cannot
be easily avoided (not even with membrane tube diameters as
small as 1 cm), and the packed bedmembrane reactor requires
in some cases double the membrane area with respect to the
staged fluidized bed operated at the same conditions. More-
over, a 2025%moremembrane area is required by the packed
bed (with respect to the fluidized bed) because of the temper-
ature profile prevailing in the packed bed.With the advance of
the development of H2 perm-selective membranes with ever-
increasing permeabilities, the advantages of fluidized bed
membrane reactor become more and more pronounced.
AcknowledgmentMw[i] Molar mass for component i, kg/mol
CMD Average molar mass, kg/mol
Greek2
i n t e r n a t i on a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 7 1 4 2 7 1 5 07150a heat transfer coefficient, J/(m K s)
b friction factor,
r Density, kg/m3
3 Porosity,
3e Emulsion phase porosity,
l Thermal conductivity, J/(m K s)
leff Effective thermal conductivity, J/(m K s)
mg Viscosity of gas, Pa s
sg Stress tensor, kg/(m s)
f00membranei;mol Molar flux component i through themembrane percell, mol/(m2 s)
Subscripts
0 Reactor inlet,
b Bubble phase,
e Emulsion phase,
g gas phase,
i Component i,
j Number of reaction,
n Number of CSTRs for emulsion or bubble phase,
r radial co-ordinate,
s solid phase,
z axial co-ordinate,
r e f e r e n c e s
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Theoretical comparison of packed bed and fluidized bed membrane reactors for methane reformingIntroductionReactor configurationsFluidized bed membrane reactor conceptPacked bed membrane reactor concept
Reactor modelsFluidized bed membrane reactor modelPacked bed membrane reactor 1D and 2D models
Results and discussionConclusionsAcknowledgmentTable of symbolsReferences