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chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339
Contents lists available at ScienceDirect
Chemical Engineering Research and Design
j o u r n a l h o m e p a g e : w w w . e l s e v i e r . c o m / l o c a t e / c h e r d
Thermally integrated bio-syngas-production for
biorefineries
T. Bludowsky , D.W. Agar
Technical University Dortmund, Department of Biochemical and Chemical Engineering, Laboratory of Technical Chemistry B,
Emil-Figge-Str. 66, 44227 Dortmund, Germany
a b s t r a c t
Maximising thethermal efficiencyof a biorefinery is a major challengein theproduction of economicallycompetitive
biomass-based chemicals. This paper compares different bio-syngas to methanol routes with respect to their energy
demands and proposes a novel heat integration concept.
Previous studies on biorefineries have tended to focus on the chemical transformations involved. The defunction-
alisation of biofeedstocks required to eliminate the excess oxygen they contain is also a very energy-consuming
process, which exerts a considerable influence on the overall yields which can be achieved. In this article we deal
with the less appreciated issue of thermal integration by analysing two principle routes for the generation of bio-
synthesis gas for methanol synthesis: a conventional high temperature biomass gasification process and a scheme
based on aqueous-phase reforming (APR). Low temperature gasification processes below 250 C, such as APR, permit
oneto use the heat liberated in the methanol synthesis for the endothermic synthesis gas production step. The com-
position of the resultant synthesis gas must be modified slightly to meet the demands of the methanol synthesis
reaction using a low temperature retro-water-gas shift reaction in a special adsorptive reactor. The results indicate
that the low temperature arrangement has an edge in terms of the net energy consumption for a given methanol
production and the analysis reveals topics of interest for future research in this area.
2009 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.
Keywords: Biorefinery; Thermal integration; Synthesis gas; Aqueous-phase reforming; Methanol synthesis
1. Introduction
Thedepletion of fossil fuel reserves makes it necessary to find
alternatives for both energy generation and the production
of organic chemicals. The use of biomass has been proposed
as a promising, environmentally compatible and sustainablesolution to both problems. Whilst there are a variety of other
options available for energy generation, e.g. atomic, solar and
wind power, biomass alone can serve as a carbon source for
the production of chemicals once fossil fuels are exhausted.
It is thus necessary to develop new industrial manufacturing
processes for chemicals based on renewable resources.
The amounts of biomass required provide a further argu-
ment for itsuse in chemical production rather than forenergy
generation. Fig. 1 illustrates that the overwhelming propor-
tion of fossil fuels is used for the generation of electrical and
thermal energy or in the transport sector, whilst a relatively
Corresponding author. Tel.: +49 (0) 231 755 5332; fax: +49 (0) 231 755 2698.E-mail address: thomas.bludowsky@bci.tu-dortmund.de(T. Bludowsky).Received13 October 2008; Receivedin revisedform 24 February 2009; Accepted 5 March 2009
small fraction is employed for chemical production. Since the
amount of agricultural or other suitable land available for
biomass harvesting is limited and must also serve the needs
of food production, it can be easily calculated that it is unfea-
sible to meet even the present demands for energy production
using biomass-based fuels, even using second generation lig-nocellulosic biofuels. On the other hand the use of renewable
biomass resources for the chemical industry is a much more
viableproposition, offering little competition to other land use
needs. Of the 11,050 m2 of usable land available pro capita
in 2050, woefully inadequate for energy requirements, it has
been estimated that around 400 m2 would suffice to cover the
needs of chemical production (Pfennig, 2007).
The National Renewable Energy Laboratory (NREL, 2008)
has designated the biorefinery concept, analogous to todays
petrochemical refineries, as the most promising approach for
the production of biomass-based chemicals. Biorefineries use
0263-8762/$ see front matter 2009 The Institution of Chemical Engineers. Published by Elsevier B.V. All rights reserved.doi:10.1016/j.cherd.2009.03.012
http://www.sciencedirect.com/science/journal/02638762mailto:thomas.bludowsky@bci.tu-dortmund.dehttp://dx.doi.org/10.1016/j.cherd.2009.03.012http://dx.doi.org/10.1016/j.cherd.2009.03.012mailto:thomas.bludowsky@bci.tu-dortmund.dehttp://www.sciencedirect.com/science/journal/026387628/8/2019 Bludowsky e Agar (2009)
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chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339 1329
Fig. 1 Consumption of fossil fuels in 2006 (IEA, 2008).
physical, chemical, biological and thermal treatments and
processes to fractionate and transform the biomass into plat-
form chemicals that can be used for biobased commodity
chemical production (Kamm et al., 2006). Due to thehigh func-
tional natureof the moleculespresent in the biomass, allthese
processes have the elimination of unwanted functional groups
as a common objective.
There are two major strategies for defunctionalising
biomolecules. The first is to breakdown the biomass par-
tially into its underlying building blocks, a concept which
has the advantage of exploiting the natural synthesis poten-tial available. Werty and Petersen (2004) have identified 12
attractive building blocks for the further synthesis of high-
value biobased chemicals. Even so, the resulting molecules
still often have to be defunctionalised further prior to their
use for commodity chemical production and their integration
into established chemical manufacturing network structures
is a daunting challenge.
Thesecondstrategy is to transformthe biomass completely
into the C1 feedstock synthesis gas. On one hand this con-
cept has the disadvantage that a lot of energy is needed to
rupture carbon chains, which must then be reconstructed
with considerable effort in the subsequent synthesis and pro-
cessing steps. On the other hand this technique representsa relatively straightforward front-end substitution yielding a
feedstock that can be used for conventional thermochemi-
cal syntheses in existing plants to produce a wide variety of
commodity chemicals. Moreover synthesis gas or syngas gen-
eration is perhaps more appropriatefor decentralised biomass
processing and can be easily transported over pipelines to
larger plants for further conversion. As a consequence of
these advantages many biorefinery concepts incorporate syn-
gas generation, either as a major pathway in its own right or
as an essential component of defunctionalisation strategies
(Kamm and Kamm, 2007).
Depending on the raw material, pyrolysis, gasification or
steam reforming are commonly used to manufacture bio-syngas. The problem with such endothermic processes is the
high temperature level required for conversion using con-
ventional technologies. A low temperature process would
be more suitable for thermal integration with downstream
exothermic processingsteps,such as methanol synthesis (MS)
or FischerTropsch-synthesis (FTS). Todays petrochemical
plants face a similar predicament: the initial functionalisation
of alkanes to carbon monoxide or alkenes in steam reform-
ers or steam crackers necessitate temperatures of around
9001000 C, which can only be provided through the combus-
tion of by-products or some of the fossil fuel feedstock itself.
This represents a considerable sacrifice in terms of yields and
efficiencies and requires costly heat-integrated reactor sys-
tems.
Fig. 2 depicts the temperature levels required for vari-
ous processes generating synthesis gas from biomass. It can
be recognised, that the operating temperatures for pyroly-
Fig. 2 Temperature levels of various syngas generation
processes.
sis, gasification and steam reforming lie well above those of
the typical chemical processes for converting synthesis gas to
chemical. A direct thermal coupling between the exothermicMS or FTS and the endothermic synthesis gas generation is
thus not possible. The heat of reaction needed for pyrolysis
or gasification therefore has to be supplied externally by com-
bustion as in the conventional petrochemical plants described
above. Ideally, providing the heat required for synthesis gas
generation directly from the heat of reaction liberated by the
MS or FTS could permit drastic improvements in the overall
efficiency of biomass utilisation. From Fig. 2 it can be seen that
the only gasification process enabling an expedient thermal
integration of this kind is that of aqueous-phase reforming
(APR).
Cortright et al. (2002) have demonstrated, that carbo-
hydrate monomers can be converted under relatively mildconditions (225 C, 30bar) in the presence of a catalysts into
carbon dioxide and hydrogen by means of APR. Supported
Pt/Al2O3 or Raney NiSn, which suppress the chemically unde-
sirable methanation reaction, can be used as catalysts (Davda
et al., 2005; Shabaker et al., 2004). Since the reaction takes
place in the liquid phase, the energy-intensive evaporation of
any water present can be dispensed with in APR, making it
especially suitable for the gasification of moist biomass. A fur-
ther advantage is that the gas is generated at higher pressure,
thus saving the compression energy that would otherwise be
need for most downstream processing. The major disadvan-
tage of APR is that it is only able to convert the carbohydrate
monomers derived from cellulose and hemicellulose and, incontrast to high temperature gasification, it cannot process
the substantial lignin fraction of the biomass. A pre-treatment,
hydrolysis and fractionation of the biomass would thus be
necessary before APR could be applied (Huber and Dumesic,
2006).
By way of an example, the APR of glucose, the constituent
monomer of cellulose, can be considered as a benchmark:
C6H12O6 + 6H2O 6CO2 + 12H2 H0R = +627kJ/mol (1)
Methanation occurs as an unwanted consecutive side-
reaction:
CO2 + 4H2 CH4 + 2H2O H0R = 165kJ/mol (2)
In order to be able to use the synthesis gas produced in
the methanol or FischerTropsch syntheses, it is necessary to
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1330 chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339
raise the concentration of carbon monoxide in the gas. This
can be done using the retro-water-gas shift reaction (RWGS),
which takes place on a CuO/ZnO/Al2O3 catalyst.
CO2 +H2 CO+H2O H0R = +41.2 kJ/mol (3)
The maximal conversion in the RWGS is dictated by thechemical equilibrium of reaction (3). Due to the endother-
mic nature of the reaction, conversion of 50% or more can
only be attained at temperatures in excess of 800 C in con-
ventional reactors. Normal operation of an RWGS reactor is
thus unsuitable for an energy-efficient, thermally integrated
carbon monoxide production, since the high reaction tem-
perature would render thermal coupling with the MS or FTS
impossible,thus negating the advantageof the aqueous-phase
reforming in this respect.
In order to permit a reaction temperature for the RWGS
which is as low as possible whilst still permitting a rea-
sonable conversion, Carvill proposed a Sorption-Enhanced
Reaction Process (Carvill et al., 1996; Nararaj et al., 1996),an in situ adsorption of the water vapor formed takes place
simultaneously with the catalytic reaction. The removal of
the water vapor from the reaction mixture shifts the position
of the chemical equilibrium toward the increased production
of carbon monoxide. In this manner, one can achieve almost
complete carbon dioxide conversion at temperatures as low
as 250 C, if desired.
The adsorptive reactor for RWGS comprises a mixed
fixed-bed of catalyst and adsorbent pellets. A standard low
temperature water-gas shift catalyst of CuO/ZnO/Al2O3 can be
used for this purpose, for example, as it provides an accept-
able reaction rate for the reverse reaction at the temperature
envisaged (Amadeo and Laborde, 1995). The adsorbent mustpossess a high adsorption capacity at the reaction tempera-
ture of around 250 C, take up water vapor selectively and be
stable under cyclic operation under the prevailing conditions.
These specifications are best met by zeolites, the well-defined
porenetworks of which, for example in theNaX und3A forms,
provide selective water adsorption characteristics (Richrath,
2007). Apart from this, the non-linear adsorption isotherms of
the pertinent zeolites exhibit high adsorption capacities even
at low water vapor partial pressures.
In the concept presented by Carvill et al. (1996) there are
four further phases in addition to the reaction/adsorptionperiod, for the regeneration of the adsorbent (Fig. 3). The first
stage in desorption is the depressurisation of the reactor (Step
2) and subsequent stripping of the fixed-bed with an inert gas
(Step 3). The reactor is then flushed with theproduct gas (Step
4) and the pressure raised back up to the operating level for
the reaction phase (Step 5).
The main shortcomingof thisprocess is naturallythe cyclic
operation, which means that at least two reactors must be
operated in parallel to ensure continuous operation of the
plant.
The modified synthesis gas obtained from RWGS can
then be fed to the methanol and FischerTropsch synthe-
ses. Methanol is one of the most important basic chemicalsmanufactured globally in terms of sheer volume (41106 t/a
in 2007) (PCI - Ockerbloom & Co, 2008). The first industrial
methanol synthesis was implemented by BASF in 1923 on a
Zn/Cr2O3-Katalysator at temperaturesbetween 300and 450C
and pressures of 250350 bar. Today the process is mostly car-
ried out on a Cu/ZnO-based catalyst at lower pressures (50100
bar) and temperatures (230300 C) (Fiedler et al., 2000). The
primary reactions occurring in the methanol synthesis reac-
tor are the parallel hydrogenation of carbon monoxide (4) and
carbon dioxide (5) to methanol and the reverse water-gas shift
reaction (3).
CO+ 2H2 CH3OH H0R = 90kJ/mol (4)
CO2 + 3H2 CH3OH+H2O H0R = 49kJ/mol (5)
Fig. 3 Cycle phases of sorption enhanced RWGS adapted from Carvill et al. (1996).
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chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339 1331
Fig. 4 Two-platform concept of a conventional LCF biorefinery, adapted from Kamm and Kamm (2007).
The most widespread production processes for methanolare the quasi-isothermal Lurgi process and the adiabatic ICI
process (Cheng, 1994). For the thermal coupling sought here,
only the quasi-isothermal multitubular reactor need be con-
sidered further, since the staged adiabatic reactor does not
permit such an effective removal of the heat of reaction. For
optimalconversionthe synthesisgas ratio(H2CO2)/(CO+CO2)
should be slightly above 2 in the methanol synthesis feed
(Cheng, 1994).
Methanol is itself a feedstock for a whole series of further
chemicals. The largest amounts are used in the manufac-
ture of formaldehyde (35%), MTBE (9%) and acetic acid (9%)
(Spath and Dayton, 2003). By means of the methanol to
olefin (MTO) process developed by Mobil (Stcker, 1999) or themethanol to propylene (MTP) process recently presented by
Lurgi (2008) methanol can also serve as a basic chemical for
polymer production. Hydrocarbon fuels may also be gener-
ated from methanol with thehelp of themethanol to gasoline
(MTG) technology (Tabak and Yurchak, 1990). For this rea-
son methanol represents a critical link in synthesis gas based
product network structures.
The direct synthesis forfuels from synthesis gascan be car-
ried out directly using the FischerTropsch process, which was
discovered in the first half of the twentieth century and devel-
oped for large-scale production during the Second World War.
Iron, cobalt or ruthenium can be used as catalysts (Huber et
Fig. 5 Process concept for a heat-integrated biorefinery.
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1332 chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339
al., 2006) and FTS can be practised in one of two basic variants:
the high temperature FTS (330350 C) producing mostly short
chain hydrocarbons (gasoline) and light olefins in a fluidised-
bed reactor usually based on the design of SASOL (Kaneko
et al., 2001) or the low temperature (220250 C) slurry bub-
ble column reactor technology (Gttel et al., 2007), in which
waxes and long chain hydrocarbons predominate. The FTS (6)
is strongly exothermic and thus eminently suitable as a heatsource for the APR process.
CO+ 2H2 (1/n)CnH2n +H2O H0R < 0 (6)
A drawback of FTSis the inherent chain growth mechanism
according to AndersonSchulzFlory-probability distribution
(ASF) (Gttel et al., 2007), so that complex product mixtures
are always obtained, which must then be fractionated prior
to subsequent processing for chemical production. Attempts
to overcome the intrinsic ASF-distribution have so far proved
fruitless (Huber et al., 2006).
To help visualise the concept being proposed for thermalintegration the dual platform concept (sugar and syngas) for
a lignocellulosic feedstock biorefinery) (LCF) as described by
Kamm and Kamm (2007) will be considered (Fig. 4). In this
concept some of the biomass is broken down into the basic
component fractions of cellulose, hemicellulose and lignin,
from which further platform chemicals are then recovered
as described earlier. The remaining biomass feedstock is con-
vertedto synthesis gasby gasification processes, which is then
either used in chemical production or for the associated gen-
eration of energy to cover the biorefinery requirements. In this
process a portion of the biomass, usually the chemically less
accessible lignin, must be burnt to heat the gasification reac-
tor, which is operated at around 800
C. The heat of reactionfrom the MS or the FTS cannot be used for the purpose, since
the temperature level is too low. Of course it can nevertheless
be used elsewhere within the biorefinery, e.g. for thermohy-
drolysis.
The modification of the standard LCF biorefinery concept
proposed here is illustrated in Fig. 5. The biomass is first
fractionated into lignin, cellulose and hemicellulose. The last
two components are then broken down into their monomeric
carbohydrate constituents using biological and/or chemical
processes. The sugars thus produced can either serve as sub-
strates for further chemical or biological transformations to
chemical intermediates or, alternatively, be converted into
synthesis gas with the help of the APR process. The synthesisgas composition required is adjusted by means of the adsorp-
tive RWGS reactor. The methanol formed in the following
synthesis step can either be used directly or further converted
to whateverproducts aredesired in the manner described ear-
lier. None of the processes involved are actually new, but their
combination in this way permits one to utilise the heat of
reaction liberated in the methanol reactor at c. 250 C for the
gasification in the APR reactor at 225 C. The lignin arising in
this process can either be directly converted to chemicals or
used to meet other energy demands within the biorefinery.
2. Balance equations
In this section the process concept for synthesis gas produc-
tion from biomass proposed in the previous section will be
compared with a conventional gasification scheme. For this
purpose a flowsheet for the APR, RWGS and MS arrangement
will be developed as will a corresponding system comprising
a conventional high temperature gasification with a down-
stream methanol synthesis.
All the mass and energy balances were implemented in
ASPEN PLUSTM software. Glucose was used as a model com-
ponent for biomass, since it is the underlying building block
of cellulose and its properties are well known. The ther-
modynamic data required were calculated with the help ofthe Predictive SoaveRedlichKwong model (Holderbaum and
Gmehling, 1991). Both processes involving synthesis gas gen-
eration and methanol synthesis are designed to produce one
tonne of methanol per hour as a benchmark. All thermal and
electrical inputs and outputs given thus refer to this pro-
duction rate. To simplify matters, the pressure losses in the
individual unit operations have been neglected.
2.1. Gasification processes
For the high temperature gasification of biomass to gener-
ate synthesis gas, only processes yielding a very low level of
inert gases need be considered (Bridgwater, 1984). Air-blown
gasification can thus be rejected because of the high nitro-
gen concentrations in the resulting synthesis gases. Thusonly
either indirectly heated or oxygen-blown gasifications remain
as options. For the purposes of the balance equations an indi-
rect gasification process was selected, since it provided the
most appropriate comparison with the APR. In the indirect
gasification, the synthesis gas generation and the combustion
take place in spatially segregated reactors, so that the synthe-
sis gas is uncontaminated by the flue gas from the air-blown
combustion and thus contains no nitrogen.
In order to attain a high level of hydrogen in the product
gas, a steam gasification was chosen. Theoperatingconditions
were selected to reflect the measurements of Herguido et al.
(1992) in which the gasification product gas was reported to
exhibit a H2/CO ratioof 2. This ratiocan be manipulated easily
by varying the amount of steam fed, thus rendering further
adjustment through the water-gas shift reaction superfluous.
In accordance with the data provided by Herguido et al.
(1992) the gasification was considered to be carried out at
800 C and atmospheric pressure. The gasification (7), water-
gas shift (8) and the methanation (9) reactions were taken into
account.
C6H12O6 6CO+ 6H2 H0R = +610kJ/mol (7)
CO+H2 CO2 +H2O H0R = 41.2 kJ/mol (8)
CO+ 2H2 CH4 +H2O H0R = 206kJ/mol (9)
The product gas composition given by Herguido et al.(1992)
corresponds to typical results for the steam gasification of
biomass and is given in Table 1. To fulfil the mass balance
requirements with this composition the following overallreac-
Table 1 Product composition for steam-gasification(Herguido et al., 1992).
Component Mol%
H2 51
CO 24
CO2 19
CH4 6
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chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339 1333
Fig. 6 Gasification process with methanol synthesis.
tion equation was derived (10).
C6H12O6 + 1.7H2O 6.32H2 + 2.92CO+ 2.39CO2 + 0.69CH4
H0R = +446kJ/mol (10)
For purposes of comparison with this gasification process,
the product gas composition from an operational technical-
scale process employing a fast internal circulating bed (FICFB)
biomass gasification in a power station at Gssing, Austria
(Hofbauer et al., 2002; Wiese, 2008) with a water-gas shift con-
version was also used for one set of balance calculations.
Since only minimal deviations were observed from the gasifi-
cation process selected, the results for the steam gasificationdescribed by Herguido et al. (1992) can probably be considered
representative for atmospheric pressure steam gasification
processes.
It was further assumed that the feedstock is entirely con-
verted to synthesis gas in the gasification. The formation of
coke, tar and oil was neglected. This assumption is justified,
since the gasification process under consideration is oper-
ated with external heating and the amount of biomass to
be burnt to meet the heat demand is calculated retrospec-
tively. In the actual indirectly heated gasification, the coke
and tars formed would be burnt, thus reducing the amount
of biomass needed for this purpose. As an indication of the
expected product distribution for atmospheric gasification at800 C the results of Bridgewater (1994), showing a ratio of
Gas:Chars:Tars = 0.85:0.1:0.05, can be used.
The overall energy balance of the gasification process con-
siders the feedstock heating and evaporation, the gasification
process itself, the cooling of the gaseous products, the CO 2removal process, the compression of the synthesis gas up
to the methanol synthesis pressure, the methanol synthesis
and the cooling of the final product stream. For reasons of
brevity, further details of the process and the heat exchange
networks will notbe presentedin detail. Theprocess flowsheet
is depicted in Fig. 6.
The carbon dioxide separation is not explicitly modelled
in the energy balance. It is assumed that the energy require-
ments for this step are similar in both processes and thus
cancel one another out. In practice one can use a simple non-
regenerative water scrubber for removing carbon dioxide from
pressurised gas streams, as in the processingof methanefrom
biogas plants (Kapdi et al., 2005). In the literature, the energydemand for CO2 removal using mixtures of secondary and ter-
tiary amines in a two stage absorber is estimated to be around
30 MJ/kmol CO2.
The compression of the synthesis gas up to the methanol
synthesis pressure of 50 bar was assumed to be carried out
in two stages with intermediate cooling. The heat dissipation
into the cooling system was assumed to be at a low temper-
ature level and thus non-recoverable. The energy balance for
the compression assumed an isentropic process with an esti-
mated efficiency of 72%.
The methanol synthesis was carried out at 250 C and
50 bar. An overall conversion of 100% was assumed, which in
practice in almost achieved with the help of a recycle stream(Fiedler et al., 2000). By-product formation and the resulting
purification of the raw methanol that would be necessarywere
not considered. For balancing purposes only the formation of
methanol via CO (4) was considered.
3. APR process
The flowsheet proposed for the APR process is based on the
values provided in the literature by Cortright et al. (2002).
The reaction takes place at 225 C and 50 bar. The pressure
selected was somewhat higher than that given in the litera-
ture cited to save on subsequent compression of the gas for
the methanol synthesis. The level of conversion of glucose to
gaseous products was set at 50%. Other liquid phase reactions
were neglected, since it was assumed that the soluble organic
compounds thus formed would be converted to synthesis gas
during subsequent passes of the recycle stream through the
reactor, thus fulfilling the overall reaction equation. The aque-
ous reforming (1) and methanation (2) reactions were taken
into consideration. The recycle flow couldbe ascertainedusing
thespecified inlet concentration of theAPR reactor of 1% m/m
Table 2 Product composition after aqueous-phasereforming.
Component Mol%
H2 57
CO2 36
CH4 7
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1334 chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339
Fig. 7 APR + RWGS + MeOH-synthesis process.
glucose. The assumed composition of the synthesis gas gener-
ated isgivenin Table2, which is based on themeasurements of
Cortright et al.(2002). In orderto obtain results consistent with
the reaction system given above, the values were modifiedslightly.
One thus obtains the following overall reaction for the APR
(11):
C6H12O6 + 4.08H2O 8.16H2 + 5.04CO2 + 0.96CH4
H0R = +385kJ/mol (11)
As a simplification, the adsorptive RWGS reactor was mod-
elled as a normal stoichiometric reactor with subsequent
water vapor removal. The reverse water-gas shift reaction (3)
achieves a carbon dioxide conversion of nearly 100% during
the reaction/adsorption phase (Carvill et al., 1996). For thesimulation, continuous operation of the RWGS was assumed,
which in practice would be realised by tandem operation of at
least two reactors in parallel. The heat of adsorption has been
neglected in the calculations, as it is anticipated that it would
be consumed regeneratively during the subsequent adsorbent
regeneration through decompression. The energy required for
compression during the regenerative cycle of the RWGS reac-
tor was not taken into account, since it is almost negligible by
virtue of the low volumetric flows involved.
The assumption of ignoring the adsorption/desorption
enthalpies in the retro-water-gas shift (RWGS) reactor can
be justified briefly as follows: in comparison to the adsorp-
tion enthalpy released (HR Ads =75kJ/mol, Carvill et al.,1996) the endothermic reaction enthalpy of RWGS is lower
(HR RWGS = +41 kJ/mol), and thus the overall adsorptive reac-
tion process is only mildly exothermic. For this modest
amount of heat, it is anticipated that the fixed-bed can serve
as a regenerative heat store over the reactive-desorptive cycle
(Nieken and Watzenberger, 1999). In the desorptive phase, the
regeneratively stored heat is then consumed bythe desorption
process. The temperature excursions associated with such
internal regenerativeheat recovery can be minimised by incor-
porating additional thermal ballast in the fixed-bed or byusing
phase change materials to enhance isothermal heat storage
capacities. Nevertheless, theheat forthe RWGSreactionstill of
course has to be provided regardless of this regenerative heat
exchange process. Assuming only minor temperaturechanges
during cyclic operation, the driving force for desorption is the
depressurisation (e.g. from 50bar to atmosphere pressure). In
our model we have assumed two RWGS reactors operated in
parallel, which require the same period for adsorption and
desorption. These two reactors are treated as a single unit
operation and therefore in our model only the overall external
heat input needed to cover the reaction enthalpy is supplied.For the methanol synthesis and all other unit operations
thesame conditionsand assumptions wereemployed as in the
high temperature gasification process described previously.
The flowsheet is depicted in Fig. 7.
4. Results
The solution of the mass balance equations for the processes
presented in the previous sections yielded the results given in
Table 3. It can be seen that for the APR + RWGS-process more
glucose must be provided for the same amount of methanol.
This is a consequence of thelarger amount of methaneformed
as a by-product in the APR reaction. The amount of carbon
dioxide purged from both processes is similar, but neverthe-
less slightly higher in the APR process because the H2/CO2ratio in the synthesis gas formed is somewhat less than twice
the H2/CO ratio in the high temperature gasification process.
Furthermore the much greater water feed to the APR process
is apparent, that largely reappears as condensate product.
It would be desirable to cut the recycle flow rate in the
APR reactor by operating at higher glucose concentrations, to
improve process economics by reducing equipment dimen-
sions. However, Davda and Dumesic (2004) and Davda et al.
(2005) report that the hydrogen selectivity diminishes dra-
matically and hydrocarbon levels rise when 10% m/m glucose
feed concentration is used. They attribute this phenomenon
Table 3 Mass balances.
Gasification process APR + RWGS-process
Stream Inlet [kg/h] Stream Inlet [kg/h]
Water 328 Water 4504
Glucose 1925 Glucose 2068
Gasification process APR+ RWGS-process
Stream Outlet [kg/h] Stream Outlet [kg/h]
CO2 1124 Condensed water 3662MeOH 1000 CO2 1171
Off-gas 129 RWGS-water 562
MeOH 1000
Off-gas 177
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Table 4 Heat exchange and energy demands.
Gasification process APR + RWGS-process
Unit Cold streams [kW] Unit Cold streams [kW]
HeatEx3 1783 HeatEx1 1444
Gasification 1385 APR 2857
HeatEx2 279
RWGS reactor 337
Gasification process APR + RWGS-process
Unit Hot streams [kW] Unit Hot streams [kW]
HeatEx4. 994 Condenser (1) 3064
MeOH-reactor 884 MeOH-reactor 886
Product cooler (3) 438 Product cooler (2) 437
Gasification process APR + RWGS-process
Unit Electrical power [kW] Unit Electrical power [kW]
Compressor 1 263 Pump 27
Compressor 2 465
to undesirable liquid phase reactions and propose the hydro-
genation of glucose to sorbitol in an additional upstream
reactor to counteract this difficulty, since sorbitol exhibits
higher hydrogen selectivities at higher inlet concentrations in
the APR.
When one considers the heat and energy consumptions of
the two processes given in Table 4 one is struck by the higher
heat energy requirement of the APR reactor compared to the
high temperature gasification, although the reaction enthalpy
for the APR is somewhat lower. This apparent contradiction
is due to the evaporation of water within the reactor. The
pressure of the APR process (50 bar) lies well above the boil-ing pressure at the prevailing temperature of 225 C (25.5bar),
but the fraction of water vapor in the gas following phase
separation in nevertheless about 56%, due to the high vapor
pressure of water at this temperature. Overall the APR reactor
consumes 1268 kW for the endothermic reaction and 1589kW
for water evaporation. It can be seen that, without this latent
heat effect, the methanol reactor could supply roughly 70%
of the heat needed for the APR reaction. However, it must be
remembered that heat input is also required for the RWGS
reactor, albeit much less than for the APR reactor.
A pinch analysis was carried out for both processes, using
the ASPEN PinchTM software. Calculation of the composite
curves was done by the method presented from Lindhoff
(Linnhoff and Hindmarsh, 1983), assuming a minimal tem-
perature difference of 10 C between hot and cold streams in
heat exchangers. The most important thermal properties andstreams for both processes can be found in Table 5.
In the composite curves for the gasification shown in Fig. 8,
it is apparent that theheat of reaction must be supplied exter-
nally. For the reaction temperature level of 800C, combustion
Table 5 Thermal properties.
APR + RWGS-process
Stream m [kg/h] Vapor fraction Cp [kJ/kg K] H [MJ/kg] T [K]
Inlet water 4504 0 4.26 15.88 25 C
Inlet glucose 2068 0 3.06 8.20 25 C
APR reactor inlet 287110 0 4.84 14.91 225
CAPR reactor outlet 287110 0.023 4.79 14.87 225 C
RWGS reactor inlet 1740 1 2.75 6.97 250 C
RWGS reactor outlet gas 1177 1 2.82 3.01 250 C
RWGS reactor outlet H2O 562 0 5.25 14.89 250 C
MeOH-reactor outlet 1177 1 2.51 5.72 250 C
Off-gas 177 1 2.61 4.70 25 C
MeOH 1000 0 2.78 7.48 25 C
Gasification process
Stream m [kg/h] Vapor fraction Cp [kJ/kg K] H [MJ/kg] T [K]
Inlet water 328 0 4.26 15.88 25 C
Inlet glucose 1925 0 3.06 8.20 25 C
Gasification inlet 2253 1 2.39 6.47 800 C
Gasification outlet 2253 1 2.25 4.65 800 C
MeOH-reactor inlet 1129 1 2.92 2.90 250 C
MeOH-reactor outlet 1129 1 2.56 5.72 250 C
Off-gas 119 1 3.45 4.31 25 C
MeOH 1000 0 2.788 7.48 25 C
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Fig. 8 Composite curvesgasification process.
is the only suitable process which can cover this requirement.
The waste heat from the methanol synthesis is only partially
utilised in the process considered. In practice the relatively
high temperature level of 250 C would certainly make it anattractive proposition for alternative thermal applications in
the biorefinery, for example biomass solubilisation. From a
general pointof view onlyrelatively littlepoorly utilisable ther-
mal energy arises in this process. The pinch temperature for
this process has a value of 245 C.
For the composites curves of the APR process (Fig. 9)
it should be noted, that the entire heat production in the
methanol synthesis can be consumed in the APR reaction.
However, this amount of heat is entirely inadequate to meet
all the thermal needs of the APR. The evaporation of water in
the APR reactor leads on one hand to a substantial increase in
the external heat requirement and on the other necessitates
an almost equally large increase in the cooling capacity, sinceonly a small fraction of the heat of condensation can be effec-
tively utilised within the process. Since the internal process
heat source has been exhausted, the RWGS reactor must be
externally heated, since no other heat source is available for
this purpose. The pinch temperature in this case is 220 C.
A comparison of the two processes reveals that under the
prevailing conditions the APR process has a higher exter-
nal heat demand than the gasification process, despite the
thermal coupling with the methanol synthesis. The intended
effect of saving thermal inputs cannot be achieved, since too
much water is being evaporated in the APR reactor. How-
ever, examining the electrical power requirements leads one
to revise this evaluation somewhat. The compression of the
Fig. 9 Composite curvesAPR+ RWGS-process at 50 bar.
Fig. 10 Composite curvesAPR at 100 bar + RWGS-process.
synthesis gas from the high temperature gasification process
requires a much greater power input than feeding the glu-
cose solution into the pressurised APR system. Assuming a
power station efficiency of 40% for electrical power genera-tion means that the electrical energy must be weighted by a
factorof 2.5higher thanthe thermal energy. Determining over-
all energy requirement in this manner (i.e. thermal+ electrical
power inputs) manifests a clear advantage for the APR process
over the gasification (2442kW vs. 3790 kW), even allowing for
the increased cooling capacity for the former.
In order to suppress the evaporation of water in the APR
one could employ higher operating pressures in the APR reac-
tor. Increasing the liquid feed pumping power to realise this
would probably present little difficulty. However, Davda et al.
(2005) report an increase in hydrogenation activity at higher
pressures, with the result that selectivities tend to shift from
hydrogen to alkanes. Ignoring this chemical effect and assum-ing the same gas composition as before leads to the composite
curves for operation at 100bar shown in Fig. 10. It is clear that
in this case only a relatively small part of the heat require-
ment of the APR reactor arises due to evaporative water losses
with the product gases, and, as a consequence, about 50%
of APR heat demand can now be covered by the heat from
the methanol synthesis. This cuts the external heat input
necessary by roughly 40% (816 kW) in comparison to the gasi-
fication, whilst the pump power required increases from 27
to 41 kW. This measure is still energetically attractive when
the electrical power is weighted 2.5 times more than thermal
power, reflecting the efficiency of electrical power generation.
In practice, the improvement in the thermal balance at higherpressure must be set off against the increased alkane concen-
trations in the mass balance. Unfortunately we are unaware
of any literature providing reliable data on gas composition
for APR at 100 bar. An alternative solution might be to try and
recover the heat of condensation from the water vapor lost
with the gasby some form of vapor recompression. This would
most likely entail considerable equipment costs and detract
from the simple elegance of the APR process.
In the mass balancefor the high temperature gasification it
was assumed thatthe process wasexternally heated. In actual
operation, the heat supply to the gasification is obtained from
combustion of part of the biomass feedstock or other fuel. For
thecase in hand 354 kg/h of dry glucose must beburnt to cover
the heat demand of the gasification (neglecting heat transfer
losses). The true amount required will be considerably higher,
since the flue gases containing a large portion of the heat of
combustion will leave the combustion reactor and only some
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of the heat of combustion will thus be imparted to the heat
transport medium, such as sand. The residual heat in the flue
gases can, of course, be used elsewhere in the biorefinery, but
is lost as far as the gasification is concerned.
For a rough estimate of the variable operating costs (energy
and raw materials) for the APR process at 50bar and gasifi-
cation with methanol synthesis under the conditions given,
reference prices for the energy and material inputs must bedefined. The price for glucose from lignocellulose was set at
10 Eurocent/kg (Huber et al., 2006). Electric energy was priced
at 10 Eurocent/kWh and thermal energy at 4 Eurocent/kWh.
The costs for cooling were neglected, because it was assumed
that most waste heat could be used elsewhere within the
biorefinery (for example, in the hydrolysis of cellulose and
hemicellulose). The cost of the water feed was taken as 1 D/t.
For the gasification it was assumed that the enthalpy for the
reaction would be supplied by the combustion of 354kg/h
glucose. In the APR process, the condensate was recycled
as water feed. With theses assumptions, the APR process
appears to have a slight economic edge over the gasification
option with costs of (305 D/(htMeOH)) against (324 D/(htMeOH)).As expected, the largest contribution to the simplified oper-
ating costs comes from the biomass feedstock in both cases.
The second most important contribution differs for the two
processes, however: heating costs for the APR process and
compression outlays for the gasification process. It is there-
fore not the thermal integration that leads to the advantage
of the proposed APR scheme, but rather achieving the system
pressure by pumping liquids in place of more energetically
extravagant gas compression. The benefit of thermal integra-
tion only becomes apparent at higher operating pressure, at
which water evaporation is suppressed. In the case studied
here the operating costs would drop to 258 D/(htMeOH) for APR
operation at 100bar.
5. Discussion
The process proposed here for the generation of bio-syngas
envisages the production of synthesis gas in a biorefinery at
much lower temperature levels than conventional high tem-
perature gasification processes and would permit thermal
integration with exothermic downstream processes, such as
the methanol synthesis. Up to 70% of the heat of reaction
needed for the APR could be met with heat released from the
methanol synthesis. Nevertheless, it should be clearly recog-
nised, that the heat balance for this process is very adversely
affected by the high energy demand imposed by water evap-oration in the APR reactor, a point which seems to have been
overlooked in the past. Failure to resolve this issue means
that APR offers little advantage thermally over a conventional
gasification. Only when water evaporation is suppressed at
higher operating pressure do the advantages of thermal cou-
pling between the methanol synthesis and APR emerge.
When one considers the entire energy demand, the APR
process has the advantage that it is able to furnish the oper-
ating pressure for the methanol synthesis more favourably
through pumpingliquids, whichrequiresless electrical energy
than compressing gases. This leads to a slight edge over the
gasification with respect to theoperating costs forthe test case
considered here. In order to assess to process economics more
objectively one needs to consider the investment costs for
the two processes as well. On one hand the APR reactors will
be much larger than those for gasification, due to the inher-
ently slower reaction kinetics, and the need for at least two
adsorptive RWGS reactors will increase the investment costs
still further. On the other hand the immense costs needed
for the large compressors in the gasification can be saved. As
there is little, if any, industrial experience with APR or adsorp-
tive RWGS reactors and since the data available, on long-term
catalyst performance with closed recycle loops for example,
is incomplete it is difficult to provide a reliable assessment
of this technology. In terms of the technical risk, the wellknown and industrially applied gasification processes have a
clear lead when it comes to questions of scale-up and similar
issues.
One possibility of overcoming the disadvantage of high
compression costs in the gasification process would be to
operate the gasification under pressure in the range of the
pressure required for the downstream synthesis. For the indi-
rectly heated steam gasification with circulating fluidised bed
andan additionalcombustion chamber (e.g.the FICFB process)
the technical challenges for an operation at higher pressure
have not yet been surmounted and an implementation in
the foreseeable future is unlikely (Ciferno and Marano, 2002).
An alternative would be an autothermal oxygen-blown gasi-fication under pressure, as is being presently studied in the
CHRISGAS project (Albertazzi et al., 2005). The introduction of
the biomass feedstock into the pressurised reactor represents
a difficult, but by no means impossible hurdle (Freihling et al.,
2007). Even if the technical problems of an oxygen-blown pres-
surised gasification can be overcome, it would still necessary
to compare the energy requirements for generating pres-
surised oxygen, by the liquefaction of air for example, with
that of the compression of the synthesis gas from an atmo-
spheric gasification. Bisio et al. (2002) give the energy demand
for cryogenic oxygen recovery (99% O2) as 1100kWh/tO2 . If
the energy demand of our gasification example were to be
met by internal combustion with oxygen, the resultant energyrequirement for oxygen production would be 415 kW h/tMeOH)
as opposed to a compression energy of 728 kW h/tMeOH for
atmospheric operation, i.e. a saving of 43%.
The carbon balance of both processes studied is poor. For
the APR process only 45% of the carbon in the biomass feed-
stock is to be found in the methanol product and for the
gasification only 41%. 39% of the carbon fed to the APR pro-
cess is lost in the form of carbon dioxide, and this rises to 49%
for the high temperature gasification. The balance of the car-
bon is present in the form of the by-product methane. The
carbon loss is mainly a consequence of the fact that one
must addwaterin the reforming step to attain the syngas ratio
of C:H:O = 1:4:1 needed for methanol. The oxygen thus intro-duced must be eliminated as CO2. The carbon balance could
be improved dramatically if hydrogen could be introduced
directly to adjust the synthesis gas composition. Carbon diox-
ide could then be completely converted to synthesis gas using
RWGS and the costs for its separation could be avoided. Eco-
logically, of course, thiswould only make sense if thehydrogen
was derived from non-fossil, renewable energy sources, such
as solar or wind energy, which seems a distant prospect at
present. Should such a hydrogen source be available, it would
especially favour the oxygen-blown gasification, since the
oxygen from the electrolytic hydrogen generation could be
fed directly to the gasification, elegantly circumventing the
need forair separation processes (Dietenberger and Anderson,
2007). The advantage of the APRprocess in thisinstance would
be that the RWGS reactor is already present and just needs to
be enlarged to accommodate a larger throughput, whereas for
the gasification this step must be grafted on. Furthermore, the
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1338 chemical engineering research and design 8 7 ( 2 0 0 9 ) 13281339
introduction of hydrogen in this manner would lead to the
liberation of more heat in the methanol reactor, which could
then make a larger contribution to heat integration.
When one considers the feasible feedstocks for the APR
and gasification processes, it is obvious that APR is more
suitable for biomass with a higher moisture, since in the gasi-
fication this would entail a significant additional loading of
the combustion process through the vaporisation enthalpy.For example, Schuster et al. (2001) have calculated that in
the FICFB process the entire product gas must be burnt
to cover the heat demands of gasification when the mois-
ture content of the biomass reaches 66%. The APR process,
on the other hand, can tolerate biomass of any moisture
content.
A drawback of the process proposed with respect to the
gasification is that the lignin present in the biomass must be
separated beforehand, because the conditions in the APR are
insufficient to achieve its conversion. Since the options for
chemical utilisation of any lignin recovered are severely lim-
ited at present (Kamm et al., 2006), the largest portion can
only be utilised thermally for energy generation. Whilst theheat of combustion of the lignin separated in a biorefinery
must be used directly at high temperatures for the gasification
process, the APR option enables one to exploit these high tem-
peraturesfor generating electricity and then usingthe residual
low temperature waste heat for heating the APR and RWGS
reactors. By this improved combined heat and power arrange-
ment, the overall energy efficiency of the lignin combustion
could probably be enhanced.
In contrast to gasification, APR is to some extent in
direct competition to biotechnological fermentation pro-
cesses, which can also convert carbohydrate monomers to
simplermolecules. Thelow energy demandand thedirect pro-
duction of a chemical (e.g. ethanol or methane) represent themajor advantages of the biological processes. Fermentation
processes often suffer from the extreme substrate selectivity
of the microorganisms and low space-time yields, however.
Chemical processes are able to treat a broader spectrum of
feedstocks and do it more rapidly. For example the bakers
yeast used for alcoholic fermentation can convert the glucose
derive from cellulose, but not the pentoses from hemicellulose
hydrolysis (Westermann et al., 2007). There are developments
underway to improvematters, butyields andreactionratesare
still very low ( Jeffries, 2006). Superior space-time yields rep-
resent a clear advantage of the APR process over competing
biological process that is unlikely to be endangered. Biologi-
cal systems may nevertheless indicate the asymptotic carbonefficiencies which may be achieved, for example around 67%
for the fermentative production of ethanol or methane, since
evolutionhas compelled the microorganisms involvedto max-
imise the energy yield available from such reactions.
The APR process presented can of course be used for
thermal integration in conjunction with the FischerTropsch
process. This would be particularly interesting for the pro-
duction of gasolines for the transport sector, where the broad
product spectrum is less of an issue. It has even proved pos-
sible to combine the APR and FTS steps in a single reactor
providingexcellentthermal integration (Simonettiet al., 2007).
However, Spath and Dayton (2003) and Hamelinck et al. (2004)
report that the price for biomass-based FTS-diesel is 4050%
higher than that for methanol or hydrogen from renewable
feedstocks. In view of this, it wouldappear to make more sense
to use such fuels in cars driven by fuel cells, even when the
changes necessary to the present distribution infrastructure
and the power trains of existing automobiles are taken into
consideration.
To summarise, the proposed APR process with RWGS and
methanol synthesis can make a contribution to the improved
energy efficiency of a biorefinery, by enabling synthesis gas
production at low temperatures, which open up new possibil-
ities for thermal integration. A prerequisite for the large-scale
implementation of such technology is further research toclarify the considerable risks still posed by the long-term
operation of the APR and the cyclic operation of a RWGS
reactor. The investigation of APR processes under higher pres-
sure would seem to offer a particularly promising avenue for
research.
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