Category Archives: Handbook of biofuels production

Design methods

The design methods for each type of gasifier would be different, as they differ in their operation. We know every reaction is governed by three main parameters: time, temperature, and turbulence. So before going into detail in designing we must be clear about which parameters among these three are most important for each type of the gasifier.

• Fixed bed gasifier: residence time

• Fluidized bed gasifier: turbulence (better mixing)

• Entrained bed gasifier: temperature.

Being that fixed bed gasifiers are governed by residence time does not mean that other factors do not play a role. In fixed bed gasifiers, it is very difficult to create turbulence and it is even more difficult to maintain a uniform temperature difference. Thus, by maintaining higher residence time in the reactor one can get higher conversion. Similarly, in the case of the fluidized bed gasifier, better mixing of the fuel with the bed material can result in higher conversion as the residence time is very small. In the case of the entrained bed gasifier, a very high temperature is maintained, as the residence time is low and mixing is poor. Thus, innovation in entrained bed gasifiers is always focused on how to create better mixing/turbulence or increase residence time in the reactor, while in fluidized bed gasifiers it is focused mainly on increasing the residence time and maintaining uniform temperature. However, in the case of fixed bed gasifiers, new ideas are being researched to develop uniform temperature distribution. In the design process, these key points should always be kept in mind.

There are different approaches to design but every design method consists of two major sections:

1 Mass balance: This section identifies the amount of mass that is going into and out of the system.

2 Energy balance: This section identifies the amount of energy that is going into and out of the system.

The idea is to first define the system and then identify the correct mass and energy balance for that system.

446 Handbook of biofuels production

Upgrading of biomass-to-liquids-Fischer-Tropsch products

Summarizing the above, there are at present two catalyst systems available for large — scale commercial plants — cobalt-based and iron-based — and two operating modes of the FT process — low and high temperature. The iron catalyst produces gaseous and gasoline range products when operated in the high-temperature range, usually in fluid catalyst bed reactors. In the low-temperature range, both iron and cobalt catalysts produce a large amount of high boiling, waxy products and straight-run diesel and naphtha. The wax is then upgraded to lower boiling range products and normally distilled to yield highly paraffinic, zero sulphur and zero aromatic middle distillate diesel fuels, with naphtha as a co-product. Typical carbon number distribution of HTFT and LTFT products is given in Table 19.3 (de Klerk, 2008).

Table 19.3 The carbon number distribution of high temperature Fischer-Tropsch (HTFT) and low temperature Fischer-Tropsch (LTFT) products, excluding C1-C2 hydrocarbons


HTFT (Synthol)

LTFT (Arge)

Carbon number distribution (mass %)

C3-C4, LPG



C5-C10, naphtha



C11-C22, distillate



C22 and heavier



Aqueous products



Compound classes


> 10%

Major product


Major product

> 10%



< 1%


5-15 %


S — and N-species




Major by-product

Major by-product

Source: de Klerk, 2008.

As the focus of the BTL process, so far, has been to maximize the production of premium BTL-FT fuels, in this section, we will focus on the technologies for upgrading the FT waxes originating from the LTFT process mode to FT diesel and gasoline by hydrocracking and catalytic cracking, respectively. The upgrading of the FT naphtha co-product to gasoline will also be discussed.

Synthetic natural gas from lignocellulosic biomass

Biomass can be converted into a gas very similar to natural gas. This gas is called SNG or Substitute Natural Gas (bioSNG). It can be used as natural gas in any of its applications such as the production of power, heat and syngas, for example chemicals. Furthermore, SNG from lignocellulosic biomass can also be used as advanced biofuel. The availability of an existing natural gas infrastructure in countries like The Netherlands and the variety of potential applications are attractive arguments for the production of BioSNG.

The production of SNG from biomass starts with thermal gasification at temperatures of at least 800°C, where the biomass is converted into a combustible gas. Subsequent gas cleaning and upgrading results in two separate products, methane (CH4) and carbon dioxide (CO2). The methane is upgraded to the specification of the natural gas grid. The pure CO2 by-product can be wasted, but can also be sequestered in, e. g., underground geological formations to turn the whole biomass value chain into a net CO2-extraction process beyond CO2- neutral. The general process from biomass to SNG is shown in Fig. 21.4.


21.4 General process of biomass to SNG.

At the moment, there is no commercial plant producing bioSNG. Developments, however, have been started at the Paul Scherrer Institute (PSI) in Switzerland and the Energy Research Centre of The Netherlands (ECN) in The Netherlands, both have recognised that the gasifier’s choice is crucial for the overall efficiency of the process (Meijden et al., 2009; Rauch, 2009). Both have chosen the concept of indirect gasification to obtain an essentially N2-free gas with relatively high methane content. Additionally, indirect gasifiers do not need pure oxygen and therefore can do without an expensive and energy-consuming air separation unit. There are also some important differences which are summarised in Table 21.1. Included in the table is the process applied by DakotaGas in the US to produce SNG from lignite, in operation since 1984 (Stern, 2006).

As with almost all bioenergy processes, costs are mainly determined by the biomass costs, in particular at larger scales (Zwart et al., 2009a, 2009b; Zwart et al., 2006b). Therefore, SNG production costs are calculated for biomass prices of 0 and 2 €/GJth (e. g. locally available biomass) as well as of 4 and 6 €/GJth (e. g. biomass delivered at the gate of larger power plants). In Fig. 21.5, also reference is made to typical (commodity) prices for natural gas (grid as well as compressed), biogas, and biodiesel (the reference for the current European transportation fuel market), as valid on the Dutch market in 2007. With the natural gas commodity price as reference, it can be calculated that at sufficiently large scale, the cost of avoided CO2-emission can be below 60 €/ton CO2, even at a reasonable biomass price of 4 €/GJ.

Table 21.1 Main characteristics of SNG production processes




Suitable for biomass?



No (unless mixed with coal)

Air separation







Atmospheric indirect gasifier ‘FICFB’ (Ichernig et al., 2008)

Atmospheric indirect gasifier ‘MILENA’ (Meijden et al., 2008)

Pressurised fixed bed updraft Lurgi

Main gas cleaning

RME tar scrubber (Zwart, 2009)

OLGA tar removal (Zwart et al., 2009)



Fluidised bed process

Multiple fixed bed process

Multiple fixed bed process


10-50 MW

100+ MW

~3 GW

Main products

bioSNG and heat

bioSNG and CO2

SNG, CO2, tars

Energy efficiency solid fuel to SNG




— SNG production costs (biomass 0€/GJ) Commodity price natural gas

— SNG production costs (biomass 2€/GJ) — Compressed natural gas

— SNG production costs (biomass 4€/GJ) — Biogas (with Dutch subsidies)

SNG production costs (biomass 6€/GJ) — Biodiesel


Scale of SNG plant (MWth)

21.5 SNG production costs for different scales and biomass costs.

Production of biofuels via catalytic cracking

J. A. MELERO, A. GARCIA and M. CLAVERO, Rey Juan Carlos University, Spain

Abstract: This chapter highlights the feasibility of fluid catalytic cracking (FCC) units for the production of biofuels from different biomass feedstocks. Special attention will be focused on catalytic cracking of triglycerides, which are probably the most suitable feedstocks for their processing in FCC units since they possess density, viscosity and hydrogen/carbon ratio quite similar to those found in vacuum or hydrotreated gasoil usually fed to this refinery conversion unit. Likewise, we will comment on the influence of physicochemical properties of the different biomass feedstocks on the overall refinery facilities upstream FCC unit.

Key words: fluid catalytic cracking, triglycerides, thermal cracking, bio-oils.

15.1 Introduction

One promising alternative for the production of biofuels is the processing of biomass (cellulosic biomass and triglyceride-based biomass) in conventional oil refineries (Huber and Corma, 2007; Lappas et al, 2009). This alternative involves the co-feeding of biomass-derived feedstocks with typical petroleum feedstocks in conventional refining units. This strategy has significant advantages as compared with conventional processes of biofuels production. Petroleum refineries are already built, and hence, the use of existing infrastructure for the production of biofuels would require little capital investment (Huber and Corma, 2007; Holmgren et al.,

2007) . Moreover, a wide range of biofuels might be obtained, not only in the range of gasoline and diesel but also in the range of kerosene or fuel oil. The European Commission has set a goal that by 2020, 10% of transportation fuels in the European Union (EU) will be from renewable sources. Co-feeding biomass-derived molecules into a petroleum refinery could rapidly decrease our dependence on petroleum feedstocks and allow reaching the target of a more sustainable transport.

Several options are available for converting biomass-derived feedstocks into biofuels in a petroleum refinery: (1) Thermal (visbreaking and cocker units) and catalytic [fluid catalytic cracking (FCC) unit] cracking, (2) hydrotreating and (3) hydrocracking. Hydrogen-based processes are typically more expensive than cracking because they require hydrogen, and this consumption is even higher when biomass feedstocks are processed. Likewise, there are other drawbacks that limit the co-processing of biomass in hydrogen-based units, such as the poisoning of catalysts by water coming from hydrodeoxygenation reactions and the low quality of the resulting hydrogenated product to be used as diesel (mainly bad cold properties). Both issues require additional conditioning steps, and hence modification of refinery unit. Cracking reactions in a petrol refinery can be carried out in presence of catalyst (FCC unit) and in its absence (thermal units). Thermal units are not considered of interest for the production of biofuels since the resulting organic liquid product (OLP) contains a high content of oxygenated compounds independently of biomass feedstocks, and this reduces its interest as fuel transport. In contrast, catalytic cracking is faster and more selective than thermal cracking that allows working under milder reaction conditions, and hence minimizing yield towards gases, coke and heavy fractions and maximizing the production of liquid fraction suitable for use as transport fuel. Moreover, the presence of the catalyst shows a great ability to remove the oxygen-containing compounds and convert them into CO, CO2, H2O and a mixture of free oxygen hydrocarbons, although the extent of the oxygen removal is strongly dependent on the features of the initial feedstocks, as will be discussed in this chapter. A simplified reaction pathway for cracking reaction is outlined in Eq. 15.1.

CxHyOz — a Cx-b-d-e Hy-2C Oz-2b-c-d + b CO2 + c H2O + d CO + e C [15.1]

FCC is the most widely used process for the conversion of crude oil into gasoline and other hydrocarbons because of its flexibility to changing the feedstocks and product demands. The FCC process consists of three main steps: reaction process, separation of the products and regeneration of the spent catalysts. In the first step, a hot particulate catalyst is contacted with hydrocarbon feedstocks in a riser reactor to crack it, thereby producing cracked products and spent coked catalyst. After the cracking reaction takes place, the catalyst is largely deactivated by coke. Thus, at the end of the riser reactor, the spent catalyst is separated from the hydrocarbon products, stripped and sent to a fluidized bed regenerator to burn the coke and reactivate the catalyst. The hot catalyst is then recycled to the riser reactor for additional cracking and products are separated in a distillation column. A variety of process configurations and catalysts have been developed for the FCC process. FCC catalysts usually contain mixtures of a Y zeolite within a silica-alumina matrix, a binder, clay and some additives. Using FCC units for biomass conversion does not require any modification in the catalyst or the process itself. Moreover, the co-processing of renewable feedstocks in the FCC unit might involve some other process benefits such as an increase in the coke production, which could help to maintain the thermal balance between the reactor and the regenerator in the FCC unit; higher olefin production in the gas fraction, which favours the application of these compounds to produce polymers, alkylates and tertiary ethers; an increase in the amount of gasoline and in its octane number due to enhancement of aromatization reactions and olefins production and a decrease in the heavy fractions with a low commercial value obtained usually in the FCC unit.

Renewable feedstocks suitable to be fed in FCC units include highly oxygenated biomass such as bio-oils, glycerol, lignin and sugars, as well as triglycerides with

low oxygen content. Figure 15.1 schematizes the different routes to produce biofuels by means of catalytic cracking. The main challenge of this catalytic process is the removal of oxygen from biomass and enriching the hydrogen content of the final hydrocarbon product in order to improve their fuel properties. Chen et al. (1986) have defined the effective hydrogen index, (H/Ceff), where H, C, O, N and S correspond to the moles of hydrogen, carbon, oxygen, nitrogen and sulphur, respectively, which are present in the feed (Eq. 15.2).

(H/C)eff = H — 2°C 3N ~ 2S [15.2]

As seen in Fig. 15.1, this index for highly oxygenated feedstocks is clearly lower than 1, which means that these feedstocks are mainly formed by hydrogen — deficient molecules. This index for a mixture of hydrocarbons ranges from 2 (liquid alkanes) to 1 (for benzene). In contrast, triglyceride-based biomass (non­edible vegetable oils and animal fats as well as waste cooking oil) shows hydrogen index of ca. 1.5, which is quite similar to that of a mixture of hydrocarbons. These different values induce distinct chemistry involved in cracking process which will result in different product distribution. Likewise, other physical properties such as viscosity can affect dramatically the catalytic performance in the FCC unit.


Nevertheless, for the co-processing of renewable materials in a refinery, it is also necessary to take into account other important issues upstream FCC unit. The stability of refining streams in the storage, pre-heating or separation devices of a refinery is well known, as well as the compatibility with the materials of the

different systems. However, this behaviour is still unknown for biomass feedstocks and their mixtures with petrol feedstocks. Stability problems during their storage might occur as a consequence of low thermal and oxidative stability of renewable raw materials as well as corrosion problems might arise from the presence of free fatty acids. Likewise, stability and corrosion of these mixtures under higher temperature, similar to that found in feed lines and heat exchangers prior to the FCC reactor system, must also be taken into consideration.

The direct hydrolysis of ethylene to ethanol

Ethylene (C2H4) reacts with water to form ethanol via a catalytic addition reaction. The yield of ethanol production is determined by the equilibrium of the reaction:

C2H4 (g) + H2O (g) — C2H5OH (g) AH = -43.4 kJ/mol

The catalyst used is phosphoric acid (silica gel based), which sets some demanding standards concerning corrosion of the equipment.

The equilibrium reaction is influenced by temperature, pressure and the ratio of water to ethylene. Normal process conditions are equimolar concentrations, 250-300°C and 5-8 MPa, resulting in a conversion degree of only 7-22%. A lower temperature favours the ethanol production, but also favours the side reaction to diethyl ether:

C2H5OH (g) + C2H4 (g) — C2H5OC2H5 (g)

Too high pressure is also not favourable because higher alcohols will be formed:

C2H4 (g) — C4H8 (g) + H2O (g) — C4HPH (g)

17.2.4 The indirect hydrolysis of ethylene

The indirect hydrolysis of ethylene takes place with the aid of sulphuric acid. Ethylene is dissolved in concentrated sulphuric acid. Addition of water leads to the production of ethanol and some diethyl ether as side product:

C2H4 + H2SO4 — C2H5OSO3H AH = -60 kJ/mol

C2H4 + C2H5OSO3H — c2h5oso2oc2h5

Hydrolysis after addition of water:

C2H5OSO3H + H2O — C2H5OH + H2SO4 C2H5OSO2OC2H5 + H2O — c2h5oh + c2h5oso3h

Подпись: C2H5OSO2OC2H5 + C2H5OH Подпись: C2H5OC2H5 + C2H5OSO3H

Side reaction with water to form diethyl ether:

Dependent on reaction conditions 5-10% diethyl ether is formed during reaction. The use of concentrated sulphuric acid sets high standards for the equipment due to the corrosive character.

Conclusion is that for the chemical routes the direct route is not economically attractive and that the indirect routes are the best up to now, although relatively cumbersome. This is the reason that the biological route of sugars to ethanol is economical, interesting and also remains to be a good alternative for ethanol production.

Newest developments focus on a combination of gasification and biological fermentation processes. After gasification, anaerobic bacteria such as Clostridium ljungdahlii are used to convert the CO, CO2 and H2 into ethanol. Higher rates are obtained because the process is limited by the transfer of gas into the liquid phase instead of the rate of substrate uptake by the bacteria.2

image210 image211 Подпись: AG° = -216 kJ/mole ethanol AH° = -331kJ/mole ethanol AG° = -97.1 kJ/mole ethanol AH° = -349 kJ/mole ethanol AG° = -135 kJ/mol acetic acid AG° = -54.8 kJ/mol ethanol

Subsequent conversion of the formed synthesis gas to ethanol brings the formation of ethanol from a diversity of biomass sources into reach. Two routes can be followed in this respect: catalytic conversion of the synthesis gas via the routes shown in Table 17.1 or biological route via direct fermentation. For smaller scale installations (<100 Kton) this last route seems to be interesting compared to the catalytic route. The catalytic route needs a catalyst which is always sensitive to deactivation via pollution in the feedstock. Also catalysts are relatively expensive. This leads to high investment cost for cleaning (on ppm level) of synthesis gas and thus economically attractive at large scale processes only. The direct biological route seems more promising for smaller scale systems because they can endure pollution of the synthesis gas and low cost fermenters can be used at ambient process conditions. Fermentation tests for ethanol production from synthesis gas have been done in various reactor types. Phillips and others (1994) used a stirred batch reactor. Klasson and others (1990) used several continuous reactors, namely a stirred-tank reactor, a packed bubble column and a trickle-bed reactor. The processes take place at 37°C, and the pH is controlled. A frequently used bacterium is Clostridium ljungdahlii. This bacterium produces acetic acid as a side-product.

In the ideal case (no side-products) this results in the following overall theoretical reaction when starting from pinewood:

CH134Oa66 + 0.17O2 + 0.17H2O ^ 0.28C2H5OH + 0.44CO2

AH° = -59 kJ/mole wood

Overall combustion energy efficiency in the ideal case is (LHVethanol/LHVwood) = 0.28×1233/410 = 84.2%. This means that this route has potential as a possible route for ethanol production from wood.

It has also been shown (Durre, 2007) that it is possible to produce butanol. Durre mentions that ‘butanol has advantages over ethanol, such as higher energy content, lower water absorption, better blending ability and use in conventional combustion engines without modification. Like ethanol, it can be produced

fermentatively or petrochemically. (……. ) The best-studied bacterium to perform

a butanol fermentation is Clostridium acetobutylicum. Its genome has been sequenced, and the regulation of solvent formation is under intensive investigation. This opens the possibility to engineer recombinant strains with superior biobutanol — producing ability’. It is also possible to produce butanol from grass and straw via an enzymatic way (www. biobutanol. nl as of January 2010). However, for this route the same disadvantages are valid as for the enzymatic route to ethanol: only the biodegradable fractions can be converted to alcohols.

The question is how far the ideal route can be approached and at what costs. For this reason the next section describes a conceptual design and simulation of a wood to ethanol plant via gasification and direct fermentation. The design is based on literature data and performed with the use of the software package Aspen Plus (Van Kasteren et al, 2005). In the following sections the design assumptions, the input composition, the reactor section and the purification section are described in detail. Other process components are discussed in the general process description.

Chemical thermodynamics

The thermodynamic calculations are done with a Gibbs free energy minimization model13 using the predictive Soave-Redlich-Kwong Equation of state to calculate the required fugacity coefficients.14 In the thermodynamic calculations the non — gasified part of the feedstock remains as solid carbon. Biomass is taken as C6H10O4 in these calculations. As mentioned before, a thermodynamic analysis gives good insight in the possible product yields, because most reforming catalysts are actually designed to obtain chemical equilibrium.

Figure 20.2 shows the carbon decomposition boundaries for several steam over carbon rations (S/C = 1, 2, 3) against the background of the phase diagram of water. Operating points located above the carbon boundary lines give thermodynamic coke, while points below do not. Obviously, the absence of thermodynamic coke does not give much information about kinetic coke. On the other hand, if thermodynamics predicts coke, there is bound to be coke in practice. In


20.2 Thermodynamic carbon deposition boundaries for S/C = 1, 2 and 3. Biomass = C6H10O4.

Figure 20.2, the operating regimes for steam reforming and reforming in hot compressed water are also depicted.

From thermodynamic point of view, steam reforming of biomass can be done without coke formation already for S/C = 1 at temperatures above ~700°C. Reforming below 700°C, thus including pre-reforming toward methane, is certainly free of thermodynamic coke for S/C > 2. Reforming in hot compressed water will produce thermodynamic coke for concentrated feedstock solutions of 50 wt% organics or more (S/C = 1, ~ 57 wt% organics). Above 450°C feeds of up to 40 wt% organics (S/C = 2, ~40 wt% organics) can be handled. For the whole hot compressed region it holds that feeds below 30 wt% (~ S/C = 3), organics do not produce thermodynamic coke. Dry reforming of biomass (reaction equation [20.2]) always produces thermodynamic coke. To avoid coke formation, dry reforming should be combined with steam reforming.

The carbon distribution of the product gas and the hydrogen yield are depicted in Figure 20.3 for relevant conditions for steam reforming and reforming in hot compressed water. The carbon distribution is given as fraction of the total carbon content of the gas and the hydrogen yield is given as fraction of the maximal amount of hydrogen that can be produced according to:

C6H10O4 + 8H2O ^ 6CO2 + 13H2 [20.8]

The data for reforming in hot compressed water are given for 250 bar, temperatures between 250°C and 700°C and 10 wt% and 20 wt% organics. More concentrated feeds turned out to be very susceptible to coking in practice; more diluted feeds suffer from a too low energetic efficiency. Steam reforming is evaluated between 500°C and 1000°C, 1 bar and 30 bar for S/C = 1 to 12.

For reforming in hot compressed water it can be seen that thermodynamics dictate a CH4/CO2-rich gas below 400°C while gas mixtures containing CH4, CO2, and H2 are obtained at higher temperatures. H2/CO2 gas can be only achieved thermodynamically at high temperature (>600°C) and for unrealistic low reactant concentrations (<2 wt%). There are some attempts reported6,15 to decrease catalytically the methane formation rate via C-O bond cleavage and hydrogenation by poisoning while maintaining the high rates of C-C bond cleavage and shift for hydrogen production. Gas produced by reforming in hot compressed water typically has a (very) low CO content because of the high water concentration in combination water-gas-shift activity.

At 30 bar, steam pre-reforming (~500°C) creates according to thermodynamics CH4 and CO2, while at 1 bar already quite some hydrogen is produced. Complete methane conversion is obtained at moderate S/C (2-3) for 1 bar at 700°C and for 30 bar 900°C is required. The H2/CO and CO/CO2 ratio can be easily manipulated with the steam over carbon ratio. For typical CH4 steam reforming conditions (S/C = 3, 30 bar) the gas yields are also presented in Figure 20.3. The differences between CH4 and biomass can be explained by the fact that biomass contains ‘internal’ water in its molecular structure: C6H10O4 = C6H2(H2O)4.

image142,image146 image143,image147

Temperature [°С] Steam over carbon [-]

20.3 Carbon distribution and hydrogen yield of the product gas for relevant steam reforming and reforming in hot compressed water conditions as predicted by thermodynamics. Biomass = C6H10O4. For 30 bar and S/C = 3 also the lines for methane steam reforming are given (dotted lines).

Enzymatically catalyzed process

The lipase-catalyzed methyl esterification of the fatty acids present in canola oil deodorizer distillates (CODD) was studied by Ramamurthi et al. (1991).

CODD was esterified to methyl esters, using immobilized lipase Randozyme SP-382 as catalyst. Conversion of the FFA up to 96% was achieved without the use of vacuum or a dehydrating agent.

It was found that three variables, namely moisture content of the enzyme, reaction time and the amount of molecular sieves, did not exhibit any profound effect on the conversion rate. On the contrary, the ratio of the reactants had a significant effect on the conversion equilibrium and showed a high interaction effect along with the temperature. High conversion (>90%) was obtained at combinations of both high temperature (70°C) and low ratio of reactants (1.2) and for combinations of low temperature (50°C) and high ratio of reactants (2.0). It was observed that higher concentrations of enzymes could compensate the negative effect of increased temperature. The conversion of acylglycerols was not investigated in this study, since the esterification was considered as a preliminary step for the recovery of tocopherols and sterols.

Facioli and Barrera-Arellano (2001) investigated the enzymatic esterification of the FFA from SODD with ethanol using immobilized fungal lipase (LipozymeIM) as biocatalyst. SODD contained 47% FFA, 26% neutral oil and 26% unsaponifiable matter. The best conversion was above 88% with no tocopherol losses.

The esterification of S ODD with butanol, using Mucor miehei lipase as biocatalyst and SC-CO2, has been described by Nagesha et al. (2004). The feedstock was preliminary filtered in order to remove sediments and sterols and enzymatic hydrolyzed to FFA using immobilized lipase (Candida rugosa) in SC-CO2 reactor unit. Hydrolyzed SODD containing <88% FFA was further enzymatically esterified with M. miehei in presence of butanol, with a maximum yield of 95% FABE. The content of acylglycerols was not affected by esterification. The high content of residual glycerides (3%) present in the final FABE impeded its direct use as biodiesel.

Wang et al. (2006) described a process for simultaneous conversion of FFA (28%) and acylglycerols (60%) from SODD to alkyl esters using a mixture of two enzymes (3% Lipozyme TL IM and 2% Novozym 435) in the presence of tert — butanol as co-solvent. It was found that the negative effects on the enzyme stability caused by the excessive methanol ratio and by-product glycerol could be completely eliminated by using tert-butanol. The lipase activity remained stable after 120 cycles. The maximum yield of FAME (84%) was achieved with an increase of tert-butanol content up to 80% (based on the oil weight). However, a further increase of the solvent resulted in a decrease of the methyl esterification (ME) yield which was explained by the dilution effect on reactants.

Fine-porous silica gel and molecular sieves (3А) were found to be effective to improve biodiesel yield by controlling the water concentration formed as a by-product during the reaction. A conversion yield of 97% could be achieved when the 3А molecular sieves quantity was 10 times the maximal water weight (calculated from FFA) and 93% with less than 10 times silica gel as adsorbent. However, more than 10 times silica gel led to a decrease in the ME yield, which was explained by the reduced availability of methanol for the methanolysis due to its absorbance by silica.

Du et al. (2007) investigated the enzymatic esterification of SODD containing 28% FFA, 60% TAG and 6% tocopherols. The reaction was a lipase-mediated methanolysis using Novozym 435 as catalyst, at 40°C in a solvent-free medium. The enzyme kept its activity after being reused for 10 cycles, each cycle being 24 h. The highest biodiesel yield of 95% was achieved by adding tenfold molecular sieves (3А). The investigation of the lipase to methanol tolerance revealed that the lipase could maintain its stability and activity in the presence of methanol at even a three molar concentration. This tolerance was attributed to the presence of other compounds than TAG, namely FFA, sterols and tocopherols. A linear relationship between the FFA content and the lipase tolerance to methanol was observed but the presence of sterols and tocopherols showed no effect.

Modelling and optimization of the process

Despite the multitude of studies on fermentative hydrogen production, the kinetic models which have already been developed or used to describe the process are limited. This is due to the fact that hydrogen production and metabolic products distribution is affected by many factors and up to now, the role of each is not well understood. So, there is a lack of models which incorporate important parameters such as pH, hydrogen partial pressure and regulation mechanisms like the ratio of NADH/NAD+, influencing hydrogen production and products’ stoichiometry.

The majority of the researchers have used simple models in order to describe their experimental data. For example many of them have used the modified Gompertz equation developed by Zwietering et al. (1990) to predict hydrogen evolution in batch tests, using different substrates and inocula (pure or mixed cultures) (Lay et al, 1999; Chen et al., 2002; Wu and Lin, 2004; Fang et al.,

2006) . However, this equation cannot be applied in continuous systems, and it cannot predict the concentrations of substrates utilized and those of metabolites produced along with hydrogen.

Recently, researchers have used more complicated models, such as modified versions of the IWA Anaerobic Digestion Model No.1 (ADM1) (Batstone et al., 2002). The latter is a widely applicable mathematical model, which was developed for describing the anaerobic digestion process. The application of ADM1 to non — methanogenic systems demands modifications, since the initial model structure uses constant-stoichiometry to describe product generation from carbohydrates fermentation as well as excludes lactate and ethanol — two important metabolic products — from its structure. Lin et al. (2007) used modified ADM1 to describe glucose metabolism and products distribution (butyrate, acetate and ethanol) by selected clostridium species in batch cultures. Rodriguez et al. (2006a; 2000b) proposed an initial model to mechanistically describe formation of products in anaerobic fermentations and the predictions of this model were integrated in ADM1 as a variable stoichiometry function. Penumathsa et al. (2008) modified ADM1 in order to apply it to continuous bio-hydrogen production systems using a variable stoichiometry approach derived from experimental information. The simulation results obtained, provided good predictions of the dynamics in a continuous bio-hydrogen production reactor fed with sucrose, over a wide range of influent substrate concentrations. However, the modified ADM1 cannot predict and simulate the distribution of products from glucose metabolism under different environmental conditions. So, the induction of a proper regulating mechanism to regulate the fractionation of monosaccharides depending on hydrogen partial pressure, temperature, pH, etc. should make the model more robust and reliable for describing continuous fermentative hydrogen production systems.

Advantages of gasification technology

Advantages of gasification over combustion are as follows.

1. Gasification is a thermo-chemical conversion process, where the feed is converted into more valuable, environmentally friendly gaseous products that can be used for chemical, fuel, and energy production. The gaseous product can be converted to hydrogen or liquid fuels by reforming or by Fischer — Tropsch synthesis, respectively. The objective of combustion, on the other hand, is to thermally destruct the feed material and produce heat.

2. There is higher potential for overall energy efficiencies and conversion of difficult-to-handle feed materials into a gaseous fuel that can be handled with greater ease in conventional equipment designed for natural gas. For example, a producer gas flame can easily be burned with low NOx emissions, a gas flame can be easily directed to a certain heating zone, and each burner can be controlled easily.

3. The volume of gas produced is much lower in gasification than in combustion, thus a relatively smaller unit is required for the gas cleaning process.

4. The solid by-products of gasification are char (low-temperature gasification) and slag (high-temperature gasification). Char is used for various applications in the form of activated carbon, while slag, considered as non-hazardous, can be used as admix for road construction material. The by-product from combustion is mainly bottom ash, consisting of the mineral matters and unreacted carbon. The bottom ash is found to have a leaching property, thus it is considered to be hazardous. So, the solid by-products from gasification are useful and also environmentally friendly, while combustion produces a hazardous by-product.

Production of biofuels via Fischer-Tropsch synthesis: biomass-to-liquids

A. LAPPAS and E. HERACLEOUS, CPERI — Chemical Process Engineering Research Institute, Greece

Abstract: The production of synthetic fuels from biomass via Fischer — Tropsch (FT), otherwise known as biomass-to-liquids (BTL) process, constitutes one of the most promising routes for tomorrow’s fuels. In this chapter, basic topics, as well as current advances in the production of FT biofuels, are discussed. Starting with a short discussion on biomass gasification and syngas conditioning, the main types of FT reactors and catalysts, along with the different technologies for upgrading FT liquids to premium fuels are thoroughly discussed. Closing, recent advances in the commercialization of the BTL process are presented, along with a discussion on the advantages and limitations of this process and its outlook in the future fuels market.

Key words: biomass-to-liquids (BTL) process, Fischer-Tropsch (FT) synthesis, biomass gasification, Fischer-Tropsch reactors, upgrading of BTL-FT products.

19.1 Introduction

Growing environmental and security of supply concerns are the main drivers that bring about changes to fuel products. European Union (EU) policies on local air quality, climate change and sustainability, applied via Fuel Directives or Emission Directives, have strongly influenced research efforts and advances in conventional fossil, synthetic and bio-origin fuels. These, in combination with the depletion of the crude oil reserves, have rendered the production of hydrocarbons via the Fischer-Tropsch (FT) synthesis as one of the most promising routes for tomorrow’s fuels. According to a recent study (Takeshita and Yamaji, 2008), ‘FT synfuels become a major alternative energy carrier and have a noticeable share in the global final energy mix regardless of CO2 policy.’

The production of fuels via FT involves the conversion of the feedstock to synthesis gas (carbon monoxide and hydrogen) and subsequent synthesis of hydrocarbons via the FT synthesis reaction:

CO + 2 H2 ^ ‘-CH2-’ + H2O [1.1]

where ‘-CH2-’ represents a product consisting mainly of paraffinic hydrocarbons of variable chain length.

Generally, the FT process is operated in the temperature range of 150-300°C to avoid high methane by-product formation. Increased pressure leads to higher conversion rates and also favours the formation of desired long-chain alkanes. Typical pressures are in the range of one to several tens of atmospheres. The FT hydrogenation reaction is catalyzed mainly by Fe and Co catalysts, while the size and distribution of the hydrocarbon products of the reaction is generally governed by the Anderson-Schulz-Flory (ASF) chain polymerization kinetics model (Bartholomew, 1990).

One of the most important advantages of FT is its versatility concerning both feedstock and products. The FT process can produce hydrocarbons of different lengths from syngas originating from any carbon-containing feedstock, such as coal, natural gas and biomass. Depending on the feedstock, the process is referred to as CTL (coal-to-liquids), GTL (gas-to-liquids) or BTL (biomass-to-liquids). Moreover, synthetic fuels have distinct environmental advantages over conventional crude-refined fuels since they are virtually free of sulphur, nitrogen and aromatics. At the same time, they are largely compatible with current vehicles and fully blendable with conventional fuels and can thus be handled by existing fuel infrastructure. However, both the high energy demands and the large capital cost of FT plants contribute to the high price of synthetic FT fuels, and as a consequence, the economic viability of the FT process largely depends on the price of crude oil.

The FT process is not a new concept. It was first developed in Germany in the 1930s, as Germany was very poor in oil resources and needed, during the Second World War, to develop an independent source of transportation fuels based on their abundant coal resources (Davis, 2002). The exploitation of the vast oil reserves of the Middle East after the Second World War made the FT process uneconomical and interest decreased, with the exception of South Africa. South Africa has vast coal deposits, and the high oil prices combined with the oil embargo during the 1970s led to the great development of the FT process from SASOL (South African Synthetic Oil Limited) (Overett et al, 2000). The technical advances in the FT process and the increasing crude oil prices in combination with the depletion of the crude reserves have led, in the last few decades, to a renowned worldwide interest in the FT process. The FT process has already been commercialized on a large scale. Sasol Synfuels currently operates two CTL plants, processing 45 million tonnes of coal per year and fulfilling about 28% of South Africa’s diesel and petrol needs (Dry, 2002). Since 1993, Shell in Malaysia (Bintulu) and PetroS A in South Africa (Mossel Bay) have been operating industrial FT synthesis facilities, which produce liquid fuels from synthesis gas that originally comes from natural gas (GTL). Shell is currently constructing a new GTL plant in Qatar, which will be the world’s largest plant converting natural gas into 140 000 barrels per day of clean-burning liquid transport fuel and other products (Shell, 2009). A similar plant is also being built by Sasol and Qatar Petroleum in Qatar in the Persian Gulf.

This renewed interest in the FT process during the 1980s and 1990s was initiated based on the depletion of crude reserves, the subsequent increase of the crude oil price and the worldwide existence of much larger reserves of natural gas and coal. Today, global warming and the universal efforts for CO2 emissions reduction rekindle the interest in FT technology, as high-quality clean biofuels, compatible with existing infrastructure and vehicle technology, can be produced via the FT process using a wide variety of biomass resources. Materials foreseen to be used in the BTL process include wood and forest residues, agricultural residues and by-products, bagasse, lignocellulosic feedstock from processing residues (paper slurry, black liquor, etc.) and energy crops, with wood being the most commonly considered biomass feed.

The use of renewable resources as feedstock, with all associated environmental advantages, undoubtedly gives synthetic fuels a new dynamic. The production of synthetic fuels from biomass comprises of the three basic steps of all FT processes: gasification of the feedstock (in this case biomass) for production of synthesis gas (CO and H2) and gas cleaning/conditioning, FT synthesis for middle distillates production and upgrading of the FT liquids to high-quality fuel products. However, the development of a commercial BTL process is currently hindered by the fact that, in contrast to GTL, for which industrial synthesis gas production processes have been well known and used for several decades, there is at present no industrial unit for biomass gasification in existence. Closest to commercialization is CHOREN, a German-based technology company that has operated a BTL demonstration plant since 2005 and is currently constructing the first commercial BTL plant, employing their patented biomass gasification process and the Shell SMDS (Shell Middle Distillate Synthesis) FT process.

Research is actively ongoing on all the three steps of the process in an effort to improve the overall efficiency, with special focus on the biomass gasification step and subsequent gas conditioning prior to the FT reactor in order to meet the strict FT gas purification requirements. Several different types of gasification technology [e. g. fixed bed, circulating fluidized bed (CFB), entrained flow gasifiers, etc.] and operation modes have been considered and assessed and will be discussed later in the chapter.

In the next paragraphs, an overview of the basic topics, including current up-to-date advances in the production of biofuels via FT synthesis, will be discussed. Starting with a short discussion on biomass gasification, including types of gasifiers and gas cleaning techniques, we will then thoroughly describe the main types of reactors and catalytic materials currently employed for FT, followed by a comprehensive discussion on the different processes and technologies for the upgrading of the FT liquids to premium fuel products. In Section 19.3, we will give a description of the final BTL fuel products and their properties. Closing, the most recent advances in the commercialization of the BTL process will be presented, along with a discussion on the advantages and limitations of this process and its outlook in the future fuels market.