Bludowsky e Agar (2009)

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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: [email protected](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:[email protected]://dx.doi.org/10.1016/j.cherd.2009.03.012http://dx.doi.org/10.1016/j.cherd.2009.03.012mailto:[email protected]://www.sciencedirect.com/science/journal/02638762


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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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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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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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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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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.


Unit Cold streams [kW] Unit Cold streams [kW]

HeatEx3 1783 HeatEx1 1444

Gasification 1385 APR 2857

HeatEx2 279

RWGS reactor 337


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


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


11/12


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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