Puigjaner L. (ed.) Syngas from Waste - Emerging Technologies
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4.3 Influence of Operating Conditions on Reactor Performance
According to thermodynamics, increasing the ratio H
2
O/CO, the CO conversion
fraction increases (see Fig. 1). The model was used to explore the effect of
increasing the H
2
O/CO ratio on the reactor performance. Figure 4 shows the result.
The curve is not monotone, but a maximum in CO conversion is obtained at H
2
O/
CO = 2.8.
The model indicates that the kinetic slows down when the H
2
O/CO ratio
exceeds a maximum value. This is an effect of the expression rate of the WGS
reaction written according to the Langmuir–Hinshelwood mechanism. The
experimental verification of the existence of a maximum of CO conversion frac-
tion as a function of H
2
O/CO ratio is a validation of the mechanism.
The reactor performance can be improved by setting different conditions, like
higher pressure and higher temperature. Figures 5 and 6 show the effect on these
parameters of the reaction zone pressure and reactor external wall temperature,
respectively. These results indicate that a combination of high temperature and
high pressure improves the reactor performance.
The WGS reaction is exothermic, and the equilibrium is shifted to lower CO
conversion as temperature increases. But the hydrogen removal sustains the
0.80
0.82
0.84
0.86
0.88
0.90
0.92
0.94
0.96
0.98
1.00
1.0 1.5 2.0 2.5 3.0 3.5
CO conversion fraction
H
2
O/CO
Fig. 4 CO conversion
fraction as a function of H
2
O/
CO ratio
0.70
0.75
0.80
0.85
0.90
0.95
1.00
0.70
0.75
0.80
0.85
0.90
0.95
1.00
600 620 640 660 680 700
H
2
recovered fraction
CO conversion fraction
Reactor external wall tem
p
erature (K)
CO conversion fraction
H
2
recovered fraction
Fig. 5 CO conversion
fraction and H
2
recovered
fraction as a function of
temperature of the external
wall of the reactor
158 A. Di Donato
reaction overcoming the thermodynamic limit. Figure 5 shows the effect of tem-
perature on CO conversion fraction and H
2
recovered fraction. The presence of the
membrane allows greater hydrogen production in the reaction zone. But the effect
on the H
2
separation efficiency is negligible. A better exploitation of the higher CO
conversion is obtained increasing the pressure in the reaction zone that means to
increase the driving force for the hydrogen separation and consequent higher H
2
flow rate through the membrane. Figure 6 shows the effect of pressure on CO
conversion fraction and H
2
recovered fraction. This result means that increasing
reaction rate is a necessary but not sufficient condition. The process effectiveness
strongly depends on the operating pressure.
4.4 Influence of Catalytic Bed Characteristics on Reactor
Performance
The evaluation of the effect of parameters such as catalyst particle size, void
fraction, catalyst efficiency, membrane thickness on CO conversion fraction and
H
2
recovered fraction is necessary to define the requirements for catalyst and
membrane.
Figures 7 and 8 show the CO conversion fraction as a function of catalyst
efficiency and void fraction in the catalytic bed, respectively. To obtain a CO
conversion fraction higher than 0.8 the catalyst efficiency must be higher than 0.3
and the void fraction lower than 0.6. This requirements must be obtained designing
the catalytic bed, in terms of particle size and shapes.
4.5 Performance Evaluation of Different Reactor Configurations
A typical application of this type of model is the simulation of reactor configu-
ration. The next two examples concern the geometry of the system reactor plus
0.70
0.75
0.80
0.85
0.90
0.95
1.00
0.70
0.75
0.80
0.85
0.90
0.95
1.00
1000 1500 2000 2500 3000 3500
H
2
recovered fraction
CO conversion fraction
Reaction side
p
ressure (kPa)
CO conversion fraction
H
2
recovered fraction
Fig. 6 CO conversion
fraction and H
2
recovered
fraction as a function of
pressure in the reaction zone
H
2
Production and CO
2
Separation 159
membrane. In the adopted tube-in-tube configuration, the catalytic bed can be in
the external zone or in the internal zone, as schematically depicted in Fig. 9.
Both configurations are reported in the literature. Here the model has been used
to compare the performances of the two configurations in similar conditions.
Figure 10 shows the difference in terms of CO conversion fraction.
0.00
0.10
0.20
0.30
0.40
0.50
0.60
0.70
0.80
0.90
1.00
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Catalyst efficiency
CO conversion fraction
Fig. 7 CO conversion
fraction as a function of
catalyst efficiency
0.50
0.55
0.60
0.65
0.70
0.75
0.80
0.85
0.90
0.95
1.00
0.3 0.4 0.5 0.6 0.7 0.8
Void fraction
CO conversion fraction
Fig. 8 CO conversion
fraction as a function of void
fraction in the catalytic bed
catalysed WGS reaction
H
2
H
2
catalysed WGS reaction
Fig. 9 Different
configurations of the
WGSMR
160 A. Di Donato
The configuration with the catalyst in the internal tube is clearly more efficient.
This is a consequence of the lower energy dissipation when the reaction zone is the
internal tube, which causes a higher temperature in the reaction zone. As seen
above, both CO conversion factor and H
2
recovered factor increase as the tem-
perature increases. The corresponding H
2
flow rate increase is of the order of 10%.
The second example shows the use of the model to optimise the length of the
membrane. The high cost of the membrane requires an analysis of the most
suitable dimensions and placement, without reducing its performance. The ques-
tion is if the membrane must have the same length as the reactor or it can be
shorter. Figure 11 schematically shows the configuration. Simulations with the
WGSMR model using different membrane lengths (M) inside the reactor (length
L = 10 m) were performed. The CO conversion fractions for the four calculated
M/L ratios are plotted in Fig. 12. The red curve is the X
CO
as a function of the
reactor length when the membrane inside the reactor has the same length of
the reactor itself. The blue curves have been calculated for membranes of different
lengths inside the reactor. The black curve is the X
CO
of the WGS reactor without
the membrane.
A membrane with a length of 40% of the total length of the reactor gives almost
the same X
CO
of the membrane as long as the reactor. A longer membrane does not
produce a significant X
CO
improvement.
0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1.0
012345678910
Reactor len
g
th (m)
CO conversion fraction
internal catalyst
external catalyst
sweep
catalyst
sweep
catalyst
Fig. 10 CO conversion
fraction for the two WGSMR
configurations along the
reactor length
Fig. 11 Schematic
representation of the tube-
in-tube reactor with the
membrane shorter than the
reactor
H
2
Production and CO
2
Separation 161
In conclusion the model suggests to build the WGSMR using the H
2
selective
membrane for only the final 60% of the internal tube, with a corresponding
reduction of the membrane cost. The rest of the tube is an impermeable tube.
4.6 Definition of a Working Point
The reactor working point is a set of operating conditions. Some of these condi-
tions, like temperature and pressure, could be imposed by operational constraints
of catalyst and membrane, and could be different from the optimum values. Once
defined the applicable temperature and pressure, the model can calculate the
syngas flow rate to be fed into the reactor to obtain the target value of hydrogen
flow rate.
Figure 13 shows an example. CO conversion fraction, H
2
recovered fraction
and produced H
2
flow rate (in moles per second) are calculated as a function of the
gas inlet flow rate. The calculated curves allow the identification of a working
point to obtain the above defined target of 180 m
3
STP h
-1
H
2
production
(equivalent to 2.2 mol s
-1
). With the set of operating conditions used for
0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1.0
0246810
Reactor len
g
th (m)
CO conversion fraction
M/R=0.8
M/R=0.2
M/R=0.4
M/R=0.6
Fig. 12 CO conversion
fraction calculated along the
reactor length in the
following cases: reactor
without H
2
selective
membrane (black curve),
reactor with the H
2
selective
membrane as long as the
whole reactor (red curve),
reactor with H
2
selective
membrane shorter than the
whole reactor (blue curves)
0.0
0.5
1.0
1.5
2.0
2.5
3.0
0.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1.0
012345678
H
2
produced (mol/s)
CO conversion and H2 recovered fractions
Gas inlet flow rate (mol/s)
H
2
recovered fraction
CO conversion fraction
H
2
produced
Fig. 13 CO conversion
fraction, H
2
recovered
fraction and produced H
2
flow rate as a function of gas
inlet flow rate
162 A. Di Donato
the calculation, a syngas flow rate of 3.3 mol s
-1
must be fed to the reactor
(i.e., *270 m
3
STP h
-1
).
5 Feasibility and Foreseen Technological Solutions
for Catalyst, Membrane and Reactor Configuration
Catalysts for the WGS reaction have been studied for over a century. A large
variety of catalyst can be used, including reducible oxides, oxidable metals and
noble metals dispersed in oxides.
The most common catalyst at higher temperature, used since 1914, is a mixture
of iron oxide and chromium oxide (Fe
2
O
3
and Cr
2
O
3
). Metallic Ni dispersed in an
oxide substrate (e.g., CaO, Al
2
O
3
and Ce
2
O
3
) is also extensively studied.
To improve the performance, especially in terms of long term stability and
resistance against polluting species such as sulphur and nitrogen compounds, a
large number of catalysts are continuously studied, the most present in the liter-
ature being copper, zinc, nickel and noble metals (Au, Pt, Pd) deposited on oxides
supports. Experimental and modelling studies dedicated to WGSMR report the use
of both commercial and purposely developed catalysts. The main issues in the
current research are the experimental determination of kinetics laws and param-
eters and the improvement of the stability and life of catalysts [26].
A recent review on WGS catalyst can be found in Smith et al. [27]. A sys-
tematic analysis of the molecular mechanisms of the WGS reaction has been
proposed by Callaghan [28].
On the basis of the transport mechanism of chemical species inside the wall,
the membranes for H
2
separation can be basically divided in two classes: porous
and dense membranes.
In porous membranes molecular hydrogen diffuses through the pores from the
side at higher concentration to the side at lower concentration. Permeability and
selectivity depend on pore size and pore structure, according to complex mecha-
nisms. Roughly speaking, permeability increases and selectivity decreases as pore
size increases. Membrane materials and pore structure (which means preparation
techniques) must be selected taking into account the requirements in terms of
hydrogen purity and hydrogen flow rate. An acceptable compromise seems to be
microporous membranes with pore diameter of the order of one nanometer.
Materials for hydrogen selective membranes working at high temperatures
(500–900C) are carbon, and ceramic (silica, alumina, zirconia, titania and
zeolites).
Dense membranes are made of ceramics or metals. In both cases the separation
is based on a mechanism of bulk dissolution/diffusion. In ceramic membranes the
hydrogen is transported as an ion (proton/electron diffusion); in metallic mem-
branes the hydrogen is dissolved and transported in atomic form in solid solution.
Ceramic membranes are made of oxides of Ce, Sr, Ba or mix. Metallic membranes
are based on thin layer of alloys of Pd with Ag or Cu, deposited on porous ceramic
H
2
Production and CO
2
Separation 163
substrates. Dense membranes are characterised by a very high selectivity. Critical
issues are physical and chemical stability of the materials.
At present metallic membranes are the most developed. Pd alloy-based tubular
membranes of length of the order of one meter are reported [29]. Pd-based
membranes are commercialised for hydrogen purification. Ceramic dense and
microporous membranes exist only as laboratory prototypes.
Metallic membranes at high hydrogen selectivity are typically made by
depositing the selective materials as a thin layer on a substrate, in flat or tubular
configurations. Several flat parallel membranes can be assembled to maximise the
packing density (i.e., membrane surface per reactor volume). Barriers can be used
to guide the gas flux in order to eliminate stagnant zones and preferential paths.
Tubular membranes can be deposited in the internal or external side of a tubular
support. Several tubes can be assembled in a single separation unit.
A clear description of the principle and technological application of selective
membrane for H
2
and CO
2
separation can be found in Kluiter [30] and in Ocwig
and Menoff [31].
WGSMR have been realised with both flat and tubular H
2
selective membranes,
although the second option is more popular and is adopted in a larger number of
applications. A recent review on patents for hydrogen production using membrane
reactors has been made by Gallucci et al. [32].
In the simplest configuration the catalyst is placed in the volume where the
feeding gas travels inside the reactor. Other configurations, with the catalyst on
the membrane surface or inside the membrane itself, have been experimented.
These configurations increase the technological problems of manufacturing and
handling of the reactor (for example, replacement of the exhaust catalyst)
without clear benefits in terms of efficiency and stability. The separated hydro-
gen is recovered by suction. The H
2
extraction can be enhanced by using a gas
sweeping the permeation zone. Inert gas or steam have been used for this
purpose.
Good sealing of the membrane, mechanical stability at the temperature and
pressure in service must be guaranteed. Tubular membranes seem more able to
satisfy these requirements. The manufacturing technique and the requirement of
good sealing at high temperatures imposes a practical lower limit to the diameter
of the tubes, preventing the use of capillary systems. Typical tubular membranes
have diameters of the order of centimetres.
6 Conclusions
The technological evolution of the WGSMR for H
2
production and CO
2
separation
from syngas is mature enough to allow the realisation of a pilot plant. The
knowledge about both WGS catalysed reaction and membranes for selective
hydrogen separation can be applied in a unit at pilot scale to be integrated in a
gasification plant.
164 A. Di Donato
The performances of a WGSMR, in terms of CO conversion and H
2
production
is the global result of the mutual interaction of chemical kinetic mechanisms and
hydrogen transport through the membrane, strongly depending on pressure, tem-
perature and gas distribution inside the reactor. These conditions depend, in turn,
on geometry of the reactor, characteristics of the catalytic bed, characteristics of
the membrane and operating conditions. To analyse the effect of these variables on
the reactor performance, a model based on appropriate physical laws and reliable
data of catalytic bed efficiency and membrane permeability is necessary.
In this chapter a mathematical model of a WGSMR has been presented,
developed to assist the basic design of a pilot scale reactor. The results of the
calculations confirmed the feasibility of the pilot reactor, supplied the operating
conditions to obtain the target values of hydrogen flow rate, and gave important
insights into reactor configuration and catalytic bed requirement for the engi-
neering of the reactor. A pilot plant, integrated in an industrial gasification plant,
will allow to investigate the reliability and profitability of the technology.
At the end of the chapter it is worth noting that a reactor model is only as good
as the experimental data used. In the presented model, catalyst efficiency and
membrane permeability are measured data. Although the qualitative trends
resulting from the simulations remain valid, the quantitative evaluations strongly
depend on the reliability of these data.
Acknowledgments I am grateful to Michele De Santis, Enrico Malfa, Stefano Martelli, Patrizia
Miceli and Ali Smith for revisions and corrections. My special thanks to Paolo Granati for
discussions and suggestions.
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