In case we can use more than just the component that is suitable as biofuel, and not waste all other components, we reduce the amount of biomass. We also reduce the land surface we require to substitute fossil resources, the amount of water and the amount of minerals which would be required if all of the applications would be produced from individual biomass resources. By closing loops in short circles, as is the case for water and minerals that are separated from the biomass fractions in small processing units, we prevent the loss of these valuable fertilisers. By small scale pre-processing we reduce the need for transportation.

More crop residues that now are wasted can be valorised. Small scale processing means that less capital investment is required to start a business. More people can start up their business but certainly more work can be done by farmers themselves. People will be less dependent on large companies and can work for their own future, invest in equipment and education of their children.

Small scale operations will eventually become available for poor developing countries because the absolute amount of money that is required is modest even under African conditions (Goense, 2006).

Since more raw materials will be available to be processed, and more people can afford to build a factory, the total volume of fossil resources that can be substituted under economic conditions will grow rapidly.

As it has already been mentioned, reactor configuration is considered to be crucial for the overall performance of fermentative hydrogen production process. It is presumed that it influences the reactor’s microenvironment, microbial population, hydrodynamic behavior, substrate-consortia contact, etc. (Venkata, 2009). In general, reactors for fermentative hydrogen production can operate in either batch or continuous mode. Batch mode fermentative hydrogen production has been shown to be more suitable for research purposes (Chen et al., 2002; Lee et al., 2002), but any industrially feasible process would most likely have to be performed on a continuous or at least semi-continuous (fed or sequencing batch) basis.

CSTR, is the most commonly used continuous reactor system, offering simple construction, ease of operation and effective homogenous mixing as well as temperature and pH control. In a conventional CSTR, biomass is well suspended in the mixed liquor, which has the same composition as the effluent. However, in this type of reactor, biomass has also the same retention time (SRT) as the HRT, and thus, its concentration in the mixed liquor as well as the hydrogen production is limited, since high dilution rates might cause biomass washout. However, it was recently found that hydrogen-producing biomass in a CSTR could be self­granulated or flocculated under proper conditions (Fang et al., 2002b; Zhang et al, 2004). Another approach to increase the biomass concentration in a CSTR is to immobilize biomass in biofilms or artificial granules made of various support materials such as cuprammonium rayon (Kim, 2002), polyvinyl alcohol (Kim et al, 2003, 2005), polyacrylamide and anionic silica sol (Kim et al., 2003,

2005) .

Another category of continuous flow reactors are the systems characterized by physical retention of the microbial biomass, which offer several advantages compared to the conventional CSTR systems. In these systems, the SRT is independent of HRT due to physical retention of the microbial biomass inside the reactor, allowing high cell concentrations and thus high hydrogen volumetric production rates with relatively small reactor volumes. Physical retention of microbial biomass could be accomplished by several different means, including the use of naturally forming flocs or granules of self-immobilized microbes, microbial immobilization on inert materials, microbial-based biofilms or retentive membranes (Hallenbeck and Ghosh, 2009). However, a potential problem with these types of reactors is the loss of hydrogen through the formation of methane due to extended retention of the biomass inside the reactor, permitting the establishment of slow-growing methanogenic populations. Different types of reactor used for continuous hydrogen production, are presented in Table 13.6. Up to now, a comparative study of reactor performance in terms of hydrogen productivity could not be done, since the operational parameters along with reactor configuration, in all these studies, are different.

Type of reactor Microorganisms Feedstock H2 production rate Maximum H2 yield References Continuous stirred tank reactor (CSTR) Mixed culture Glucose 0.54 L/L/d 1.7 mol/mol glucose Lin and Chang, 1999 Upflow anaerobic sludge blanket reactor (UASB) Sludge from wastewater treatment plant Sucrose 6.67 L/L/d 1.5 mol/mol sucrose Chang and Lin, 2004 Packed bed reactor (PBR) Anaerobic sludge Sucrose 5.35 L/L/d 0.7 mol/mol sucrose Li ef a/., 2006 Anaerobic sequencing batch reactor (ASBR) Sludge from wastewater treatment plant Glucose 5.52 L/L/d 73.8 mL/gCOD Cheong ef a/., 2007 Fixed bed bioreactor with activated carbon (FBBAC) Sludge from wastewater treatment plan Sucrose 31.68 L/L/d Chang ef a/., 2002 Anaerobic fluidized bed reactor (AFBR) Activated sludge and digested sludge Glucose 29.04 L/L/d 1.8 mol/mol glucose Zhang ef a/., 2007b Polymethylmethacrylate (PMMA) immobilized cells Anaerobic sludge Sucrose 43.2 L/L/d 2.25 mol/mol sucrose Wu and Chang, 2007 Carrier-induced granular sludge bed (CIGSB) Sludge from wastewater treatment plant Sucrose 223.4 L/L/d 4.02 mol/mol Lee ef a/., 2006 Fluidized bed reactor (FBR) Sludge from wastewater treatment plant Sucrose 31.72 L/L/d Wu ef a/., 2007 Membrane bioreactor (MBR) Municipal sewage sludge Sucrose 40.08 L/L/d 1.51 mol/mol hexose Lee ef a/., 2007 Rhomboidal Enterobacter cloacae IIT-BT 08 Sucrose 1.855 mol/L/d Kumar and Das, 2001

Considering all the above-mentioned statements, making biofuels through bio­feedstocks refining can be an appealing alternative. However, the co-processing of triglyceride-based biomass in a refinery is necessary to enforce several previous studies. The stability of refining streams in the units and conditions of a refinery are well known, as well as the compatibility with the materials of the different systems. Nevertheless, this behaviour is unknown with pure vegetable oils or animal fats streams and their mixtures with petrol feedstocks. Stability problems during their storage might occur as a consequence of their low thermal and oxidative stability, and corrosion might arise from the free fatty acids that contain vegetable and residual oils and animal fats. Storage conditions can lead to density, viscosity or acidity changes, which might affect the processing of renewable materials (vegetable oils and animal fats) in the FCC unit (Geller et al., 2007). For example, in the FCC unit, viscosity of the sample is very important to control its vaporization. Likewise, potential polymers formed under storage conditions of vegetable oils and animal fats could lead to the deposition of gums in the tubes of the heat exchangers and the transfer lines prior to the FCC unit. Moreover, although acidity of vegetable oils or animal fats has a different origin compared to the acidity of oil products (the first one is referred to free fatty acids and the second one to naphtenic acids), acid limitation for refining streams (1.5 mg KOH/g approximately) (Humphries and Sorell, 1976; Piehl, 1988) might be a problem if free fatty acids cause corrosion. Corrosion problems associated with the mentioned free fatty acids of oils and fats are not important in the reaction section of a refinery unit, as the acids react rapidly because of the high temperatures reached. However, it cannot be said the same with the parts of the unit upstream the reactor as storage system where free fatty acids are intact. However, these issues have been poorly addressed in the literature.

We have recently studied the storage stability and corrosivity of a petrol feedstock and renewable materials mixtures under high temperature similar to that found in feed lines and heat exchangers prior to the FCC reactor (Melero et al., 2010a). Precisely, a low-saturated vegetable oil (soybean oil), a highly saturated vegetable oil (PO), animal fat unfit for human consumption and waste cooking oil were selected, whereas vacuum gasoil, hydrotreated vacuum gasoil and atmospheric residue were taken as petrol feedstocks. Storage stability studies were performed by means of an accelerated oxidation process in the presence of oxygen at 140°C according to the UOP 174-84 method (UOP, 1984). Physical properties as well as distillation curve of the samples studied were statistically unchanged after oxidation treatment. Likewise, water and/or sediment content in the samples were not evidenced after thermal treatment. Hence, according to the UOP 174-84 method, the different mixtures can be considered stable in storage at 77°C for periods of at least 180 days. Corrosion studies were also carried out following the UOP 174-84 method slightly modified by the presence of a carbon metal probe ASTM A 293 Gr C. The leaching of metallic species was monitored after thermal treatment. The results showed a negligible leaching of metallic species for pure petrol samples as well as for their mixtures with renewable materials. Hence, this preliminary study opens up good perspectives for the co-processing of triglyceride biomass feedstocks in the existing infrastructure of petroleum refineries, although further studies must be performed in the future.

The HTL process can be described by the following conceptual equation:

Biomass ^ Gas + water dissolved organics + solvent soluble [18 1]

hydrophobic organics + solvent insoluble hydrophobic organics + H2O

The gas consists primarily out of CO2 (and some CO and CH4), while the aqueous phase contains relatively small oxygenated hydrocarbons. The organics in the hydrophobic phase contain considerably less oxygen than the feedstock, in the order of 10-30 wt.% compared to more than 45% in the original oil. Without separation, the solvent-soluble (e. g. acetone) organics and solvent-insoluble organics define one product phase, sometimes referred to as bio-crude. This product has a thick paste-like appearance. The separated solvent-soluble organics form an oily substance: HLO. Table 18.1 lists selected properties of HLO. Solvent-insoluble organics are solid and char-like at room temperature.

Reported yields are in the following ranges: gas: 5-15 wt.%; water-dissolved organics: 10-25 wt.%; solvent-soluble hydrophobic organics: 30-60 wt.%; solvent — insoluble hydrophobic organics: 0-30 wt.%; and water: 10-30 wt.%. The residence time is a dominant process parameter with respect to the product distribution. After ca. ten minutes reaction time, the gas, water and water-dissolved organics yields become constant. After that time, a significant part of the oil (solvent-soluble organics) is transferred to solvent-insoluble organics (char).2 At low residence times (less than five minutes), it is possible to produce only very limited amounts of char-like product, but deoxygenation is then obviously limited. The effect of residence time on the product distribution is visualized in Fig. 18.2.

The chemistry of HTL is complex. Nevertheless, all HTL reactions can be classified according to their mechanism: ionic and free-radical reactions.14-16 Hydrolysis reactions, a class of decomposition reactions of organics involving breakdown by water, are typical ionic reactions catalyzed with bases or acids. Cellulose and hemicelluloses may be completely hydrolyzed under HTL conditions, while only partial hydrolysis of lignin is possible without a catalyst.17 It should be realized here that in the biomass structure, hemicellulose is bound partially to lignin as well, complicating the hydrolysis reactions and possibly promoting char formation as well. Complete dissolution (hydrolysis) of woody biomass was, however, recently demonstrated with Na2CO3.18 Ionic reactions are accompanied/followed by free radical decomposition reactions. These thermal reactions are favoured over ionic reactions at lower pressures, lower densities (gases) and higher temperatures.19,20 Susceptibility of biomass constituents towards thermal degradation decreases in the order: hemicelluloses, cellulose, lignin.21,22 In hot compressed water, thermal reactions tend to produce primary char, similar to the char-forming reactions in pyrolysis. During HTL, two types of char have been identified, viz. primary char and secondary char. Primary char is the residue that remains after conversion of solid biomass particles. Secondary char is the char produced via polymerization reactions of liquid decay products. Both phenol (lignin) and sugar derivatives have been identified as being susceptible
towards polymerization, but it may be reasonable to assume that primary char is caused primarily by the lignin and secondary char by the sugar derivatives.1,18,23 Anyway, these polymerization and polycondensation reactions lead to the increase of the average molecular weight of the oil and eventually lead to the formation of solvent-insoluble components (char). Due to the higher order in the reactants (~2) of these polymerization reactions, polymerization can be reduced by dilution. Na2CO3, as a catalyst, is known to prevent polymerization reactions and thus char formation, but its effect is more complex than just dilution.24-26 Appell et al.21 proposed the following biomass liquefaction mechanism with Na2CO3 and CO

Although the char yield (primary and secondary) is reduced by using alkalis, their use in HTC also has drawbacks. Alkalis react in the process, and the recovery of sodium or potassium from liquid and solid products could be a complicated and

costly procedure.26 In contrast, heterogeneous catalysts do not react away and are easy to recover. For that reason, Knezevic et al.2 studied the influence of a Ru/TiO2 catalyst on the HTL process. Compared to non-catalytic experiments, they found that the tested catalyst (1) increased the gas yield, (2) decreased the char yield and

(3) had a small effect on the oil yield. The lower char yields in catalytic tests suggest that the catalyst was able to convert the solvent-insoluble product (char) into gas. This was confirmed in an independent experiment using the solvent — insoluble product of glucose as a feedstock. In this test of 60 minutes at 350°C, ca. 30 wt.% of gas was produced that consists of CO2 and CH4. Without catalyst, under otherwise identical conditions, gas production was only 2-3 wt.%.1

Although the exact reaction pathway of HTL is not yet unravelled, at least four reactions have to be incorporated in a lumped (engineering) reaction path model: (1) depolymerization reactions, (2) decay reactions of the monomers (e. g. dehydration: glucose ^ HMF), (3) reactions causing the formation of gas (CO2 and/or CO by decarboxylation/decarbonylation), and (4) polymerization/ polycondensation reactions.

Knezevic et al.28 visualized the liquefaction of wood in closed quartz capillaries (see Fig. 18.3a). The formation of secondary char from glucose (Fig. 18.3) was also visualized in those capillaries.1

For a detailed overview of the chemical reactions in hot compressed water, the reader can refer to several reviews.14-16,29-31

Single reactor concepts for both catalytic biomass gasification and pyrolysis oil steam reforming have, up till now, not been able to produce a clean gas for a long period of time. This is mainly caused by the fact that no catalyst has yet been developed which is both mechanically strong and active for a full product gas conversion or is resistant against the ‘heavies’ formation (char and coke) which is accompanied with the initial biomass conversion step. To solve these problems, both in biomass gasification as in pyrolysis oil reforming staged systems have been proposed.

Downstream gasifiers upgrading/cleaning of the gas

Staged gasification was initiated at the University of Zaragoza by Corella et al.45,46 where two fluidized beds were used, one as the biomass gasifier and the second one as a (non-catalytic and catalytic) tar converter. When using a commercial catalyst, initially a clean gas was being produced. After a period of 1-2 hours of successful operation the catalyst started to lose its activity. The heavy tar content of the product gas was identified as being the cause of this deactivation. The introduction of a guard bed with calcined dolomite showed very promising results where no catalyst deactivation was found for a 48 hours on stream.20,47 Simell et al.48 were successful in a similar approach where calcined dolomite limestone was used as a guard material with a subsequent monolith catalytic Ni-alumina bed. Although the processes have shown its technical feasibility for hours (up to 100 h) of run time the processes were not developed commercially due to the unfavorable economics of that time.

The Biomass Technology Group B. V.49 is developing a different approach where the biomass is first pyrolysed (~500°C) after which the gas/vapor mixture is subsequently reformed autothermally with the addition of air. From a temperature of around ~1000°C, the product gas is essentially methane and tar free.

In recent years, supercritical fluid extraction using carbon dioxide (SC-CO2) has been intensively investigated to some traditional separation techniques, such as vacuum distillation or organic solvent extraction, as an alternative.

Supercritical carbon dioxide extraction is a process where carbon dioxide passes through a mixture of interest at a certain temperature and pressure until it reaches an extractor. This process is used because supercritical carbon dioxide has a low viscosity, a high diffusivity and a low surface tension that provides selective extraction, fractionation and purification, allowing its penetration in micro — and macro-porous materials (Dumont and Narine, 2007).

The major advantage of this method is the easy post-reaction separation of the components by depressurization. Another advantage is the low temperatures used for the majority of the experimentations because carbon dioxide has a critical temperature of 31°C. However, the use of high pressure conditions makes the system energetically expensive but can be economically viable at a rate of production superior to 25% using conditions of approximately 90 atm and 40°C (Mendes et al., 2002). At these specific conditions, only fatty acids are separated from tocopherol (Mendes et al., 2005).

The phase equilibrium data can provide fundamental and necessary information for designing a SC-CO2 separation process. A number of studies are available for this purpose (Stoldt and Brunner, 1998; Stoldt and Brunner, 1999; Chia-Cheng et al., 2000; Mojca et al., 2003; Pereira et al., 2004).

Simulation and thermodynamic modeling of the supercritical fluid extraction was reported by different authors (Vazquez et al., 2007; Fornari et al., 2009; Martinez-Correa et al., 2010).

Vazquez et al. (2007) described a process for the purification of squalene by using CC-SCCO2, a by-product obtained after the distillation and ethylation of olive oil deodorizer distillate (OODD), as raw material. The Group Contribution Equation of State was employed to simulate the separation process and to design the experimental extractions. As satisfactory agreement was found between the experimental and the calculated yields and phase compositions, a raffinate with a squalene concentration of up to 90% was obtained. Finally, the thermodynamic model was employed to develop optimal process conditions to enhance squalene recovery, including partial reflux of the extract product and recirculation of the supercritical solvent in a continuous countercurrent extraction column.

Several authors have studied the concentration of tocopherols directly from the DD, without carrying out any modification pretreatment of the raw material, namely the separation of tocopherols from FFA (Lee et al., 1991; King and Dunford, 2002).

However, the application of pretreatment like esterification leads to two advantageous results for the continuous process, one is that methyl esterified DD (ME-DD) has a higher solubility in SC-CO2 than DD. The other is that the viscosity is greatly reduced after removing most of the sterols.

The chemical modification of the DD combined with SC-CO2 has been reported by different authors (Bondioli et al., 1993; Nagesha et al., 2003; Liu et al., 2006; Vazquez et al., 2006; Fang et al., 2007; Vazquez et al., 2007; Torres et al., 2009). In this case, esterification and methanolysis of the DD produced a mixture containing tocopherols, phytosterol esters and FAME, with the process goal of the SC-CO2 process being the elimination of FAME to concentrate tocopherols and sterol esters in the raffinate.

Lee et al. (1991) studied the feasibility of tocopherols concentration from SODD by SC-CO2 at different temperatures and pressures. It was observed that by increasing the CO2 pressure, the SODD solubility also increases for all the studied temperature (45°C, 55°C and 70°C) and that esterified SODD has four to six times higher solubility in SC-CO2 than the sterol-removed SODD. The results showed that SODD initially containing about 13-14% tocopherols may require a countercurrent multistage column to be efficiently concentrated.

King and Dunford (2002) described a solid fluid fractionation method to recover sterol-enriched triglyceride fractions from vegetable oil DD (rice bran and soybean oil DD) using a pilot scale high pressure packed column.

It was possible to obtain oil fractions containing 20-31% sterols and 30-38% TAG, respectively. The method consists of two extraction steps, one carried out at 14 MPa and 45°C and the second extraction was performed at 20 MPa and 80°C. The described method does not leave any solvent or chemical residues in the final product, nor generates additional waste streams requiring subsequent disposal. However, another purification step should be applied in order to obtain a high-purity sterol fraction.

Bondioli et al. (1993) described a process to recover squalene from OODD after transformation of the FFA into TAG in order to increase the separation efficiency. OODD was converted into FFA by saponification and splitting. The mixture was further dried and esterified with glycerol in the presence of an acid catalyst into the corresponding TAG, the latter ones being extracted with SC-CO2. The process was carried out on a pilot-plant scale with a column operating in the countercurrent mode. Using this process, squalene was recovered in high purity and yields of about 90%.

Nagesha et al. (2003) described a process where SC-CO2 extraction of chemically modified SODD was studied at three levels of pressure (180-300 bar) and temperature (40-60°C) to optimize the conditions for enrichment of tocopherols in the raffinate. The modification process includes esterification, saponification, acid hydrolysis and cold crystallization to remove sterols, and again esterification of the FFA obtained from acid hydrolysis of the triglycerides (Fig. 22.8). After modification, SODD containing about 90% of FAME showed improved solubility in SC-CO2 and a better extraction rate. Since FAMEs are more volatile, they were extracted preferentially over tocopherols and other high molecular weight compounds. The extraction at higher pressures and temperatures resulted in a better yield of FAME along with tocopherols and this in turn decreased the degree of enrichment of tocopherols in the raffinate. However, a specific level of pressure and temperature of the extraction caused the increase in the solubility of FAME due to their volatility and results in the enhanced enrichment of tocopherols in the raffinate. It was observed that the enrichment of tocopherols (36%) to ten times the original concentration of the feed (4%) occurred at an extraction pressure of 180 bar and a temperature of 60°C.

The recovery of tocopherols and sterols from sunflower oil deodorizer distillates (SfODD) using countercurrent supercritical carbon dioxide extraction (CC-SCCO2) has been studied by Vazquez et al. (2006). The chemical transformation of the SfODD composition significantly enhances the concentration of minor lipids in the raffinate product. This pretreatment resulted in a mixture (ethylated SODD) which mainly consists of tocopherols, sterols and fatty acid ethyl ester (FAEE). Additionally, the reaction step produced a solid phase, mainly consisting of sterols, which was isolated from the liquid product.

After two consecutive extractions with hexane, the sterol purity in the new solid phase increased up to ca. 88%, which corresponds to 18% of recovery of the total sterols present in the original SODD. A similar procedure was accomplished replacing hexane by ethanol. In this case, the purity of the sterols obtained was similar, although the recovery was reduced to ca. 10%. This low value of recovery indicates a higher solubility of the sterol solid phase in ethanol compared to that in hexane.

The main drawback of the CC-SFE process described in the present study is related to the high amount of unidentified compounds present in the original SODD (20%). Around 50% of this unidentified material corresponds to non­volatile compounds which preferably accumulated in the raffinate.

CC-SCCO2 extractions of the ethylated and original SODD resulted in a 3.7- fold increase in the tocopherol + phytosterol concentration (ca. 80% recovery)

with the ethylated material, while only a 1.3-fold increase was obtained with the original SODD.

Additionally, during the formation of FAEE, partial crystallization of free sterols occurs, and around 20% of the sterols present in the original SODD can be recovered with high purity (88%) in the solid phase.

Liu et al. (2006) studied the vapor-liquid phase equilibrium data for SC-CO2 and methyl esterified DD (ME-DD) at 40°C and in the pressure range of 9.7-16.2 MPa in order to determine the feasibility of SC-CO2 to concentrate natural tocopherols from SODD. The results showed that the separation factor between tocopherols and FAME was from 2.5 to 3.8 at 40°C and 9.7-16.2 MPa, which is fundamental and necessary for future process designs. For this purpose a modification process of DD was applied that includes esterification, cold crystallization for removing sterols and alcoholysis. The FFAs obtained from TAGs by alcoholysis were further esterified to FAMEs. The detailed procedure is shown in Fig. 22.9. Through such pretreatment, the obtained methyl esters (ME-DD) contained 52% FAME and 8% of tocopherols and other compounds. After the reactions, most of the sterols were easily removed because of their low solubilities in FAMEs below 4°C.

22.9 Oil deodorizer distillates (DD) modification process (from Liu et al., 2006).

Fang et al. (2007) described a process where SC-CO2 fractionation was employed to concentrate tocopherols from ME-DD. The initial pressure, feed location, temperature gradient and ratio of CO2 to ME-DD were optimized for separating FAMEs. For the following tocopherol concentration step, a final pressure of 20 MPa resulted in the greatest average tocopherol content (>50%) and tocopherol recovery (about 80%).

ME-DD was prepared from DD through the pretreatment process that include two steps of esterification and methanolysis, which converted FFA and glycerides into FAMEs. The two reactions were conducted with the catalysts of sulfuric acid and sodium methoxide, respectively. After each reaction, the mixture was washed until neutral. Finally, the mixture was stored in a refrigerator for 12 h. As a result most of the sterols were crystallized and removed by filtering under reduced pressure. A fractionation column was required for the ME-DD separation. Low pressure (the initial pressure) was used in combination with a temperature gradient along the column to separate the FAMEs. Then, the pressure was increased to separate the tocopherols from other impurities.

Torres et al. (2009) reported a two-step enzymatic reaction to obtain phytosterol esters, where the SODD was initially modified by the addition of oleic acid in order to decrease the DD melting point. After esterification steps, the product obtained comprised mainly FAEEs, tocopherols and phytosterol esters, together with minor amounts of squalene, FFAs, free sterols and triacylglycerols. The FAEEs were eliminated by SC-CO2 and the phytosterol esters and tocopherols were concentrated in the raffinate. The separation between the last two compounds was carried out in an isothermal countercurrent column (without reflux), with pressures ranging from 200 bar to 280 bar, temperatures of 45-55°C and solvent- to-feed ratios from 15 kg/kg to 35 kg/kg. Using these extraction conditions, the fatty acid esters were completely extracted. The phytosterol esters were concentrated in the raffinate up to 82% with a yield of 72%.

To increase the amount of liquid product a technique known as fast pyrolysis (more properly defined as thermolysis since it is the thermal degradation of a chemical compound into a range of product chemicals) was developed in the early to mid-1980s.57-61 The yield of oil or liquid can be as high as 80% of the feedstock biomass but is normally around 65-75%. The corresponding amounts of char and gas are typically 10-25 and 10-20%, respectively. Fast pyrolysis processes occur in the time frame of a few seconds or less and the product range obtained is highly dependent on chemical kinetics as well as heat and mass transfer in the reaction chamber as discussed by Bridgwater.62 The residence time of solid in fast pyrolysis is of the order of a few seconds or less with a heating rate of tens or hundreds of °C/s and the temperature range used is normally above 650-1000°C. Bridgwater and Peacocke have reviewed various forms of fast pyrolysis reactors.63 It is clear, that to achieve such heating rates that these reactors will be considerably more complicated than for conventional pyrolysis. There are some obvious requirements for equipment used in fast pyrolysis and these are:

1 High heating rate to allow reactants to react in short periods of time and minimise char formation.

2 Very rapid heat transfer to allow feedstock to reach optimum temperature during their short residence time in the reactor.

3 Rapid cooling (condensation) of the pyrolysis products on exit from the reactor which prevents further reaction to char and tar-like materials.

It is difficult experimentally, as well as being expensive, to achieve the high heating rates required in static (batch) reactors. This is particularly true for the large scale reactors that would be required commercially. Whilst the heating rates may be achieved in small volume fixed bed reactors (fixed bed — the solid does not move) they can be more readily achieved if the feedstock is fed in and out of the reactor held at fixed temperature. This is achieved using a gas flow to suspend the feedstock solid. Probably the best design is a fluidised bed reactor (FBR).53,63 Here, a particulate solid, such as fine sand, together with a powder of the feedstock is supported on a high velocity gas flow forming a fluid that can be passed trough the reactor.64 The support solid is normally re-circulated through the reactor after separation of the products and this is known as the circulating FBR. The need to suspend the feedstock in the fluid requires that it is in the form of a fine particulate. This is also a necessity for rapid heat transport from the environment of the reactor to the solid. For these reasons the solid feedstock is normally ground or milled into particulates of the order of 1mm diameter or less. The FBR also facilitates the rapid cooling required to collect the bio-oil. Separation of char is usually achieved using cyclone technologies. Despite the apparent technical complexity, FBRs are well-established technology64 and probably represent the most cost-efficient form of pyrolysis.65

The first fluidized bed gasifier was developed by Fritz Winkler of Germany in 1921. It was later used for powering gas engines. From 1921, several companies were involved in making fluidized bed gasifiers which proved to be more efficient and competitive with other technology.

These types of gasifiers provide excellent gas-solid mixing. Fluidized bed gasifiers can be operated at lower temperatures — around 800-900°C — than fixed bed gasifiers. This directly affects NOx emission reduction. Also better fuel flexibility and efficiency in process carbon dioxide capture are some of the advantages of this type of gasifier.

A fluidized bed gasifier mainly consists of a bed of hot solid that is fluidized by the gasifying agent (air, oxygen, or steam). When the feedstock is fed into the hot bed, it undergoes gasification in the presence of a gasifying agent and the product gas leaves from the top of the gasifier. If the bed solid leaving the furnace is captured and again re-circulated into the gasifier, it is called a circulating fluidized bed (CFB) and if not then it is a bubbling fluidized bed (BFB). The bubbling bed gasifier is generally operated at a lower velocity (2-2.5 m/s) to ensure particles do not leave the reactor. The circulating fluidized bed is operated at higher velocity (3-5 m/s) and particles leaving the reactor are separated in a cyclone and fed back into reactor. CFB gasifiers are very suitable for large-scale syngas production (Tijmensen et al., 2002; Hamelinck et al., 2004; Wang et al., 2008; Zhang, 2009). CFB gasifiers are also considered to be rather fuel flexible and are most suitable for feedstocks with high volatile matter content and high char reactivity, such as biomass. Moreover, they offer short residence time, high productivity, low char/ tar contents, high cold gas energy efficiency and reduced ash-related problems (Wang et al, 2008). The gas produced by CFB gasifiers, operated at ~900°C contains, however, beside H2 and CO, considerable amounts of CO2, H2O, and

16.7

Temperature distribution along the height of the fluidized bed gasifier (Higman and Burgt, 2003).

hydrocarbons like CH4, C2H4, benzene, and tars. Thus, the product needs further treatment in a catalytic reformer to convert the hydrocarbons to H2 and CO.

Temperature distribution along the height of the fluidized bed gasifier is shown in Fig. 16.8. In the case of fluidized bed gasifiers, the temperature is more uniformly distributed.

Figure 16.9 shows different types of fluidized bed gasifiers that have been commercially developed. The Winkler gasifier, invented in 1920, was probably the first type of gasifier to use fluidization on an industrial scale to gasify pulverized coal and started with a capacity of 2000 m3/h of product gas. Foster Wheeler CFB is an air blown gasifier operating at atmospheric pressure. Depending on the fuel and the application needs, it operates at a temperature within the range of 800- 1000°C. The hot gas from the gasifier passes through a cyclone, which separates most of the solid particles associated with the gas and returns them to the bottom of the gasifier. In the twin reactor gasifier, the pyrolysis, gasification, and combustion take place in different reactors. In the combustion zone, the tar and gas produced during pyrolysis are combusted and heat the inert bed material. The bed material is then circulated into the gasifier and the pyrolysis reactor to supply heat. The char and heat carrier from the pyrolyser are taken into the gasifier. The gasification of char in the presence of steam produces the product gas. The residual char and the heat carriers from the gasifier are taken back into the combustor. This system was developed to overcome the problem of tar. The KBR transport gasifier is a hybrid gasifier having characteristics of both entrained bed gasifier and fluidized bed reactor. The KBR gasifier operates at considerably higher circulation rates, velocities (11-18 m/s), and densities than a conventional circulating fluidized bed. This results in higher throughput, better mixing, and higher mass and heat transfer rates. The solids that are transported are separated from the product gas in two stages and returned to the base of the riser. The gasifier operates at 900-1000°C and 11-18 MPa (Higman and Burgt, 2008). EBARA’s TwinRec

Process gasifier is used primarily to recover recyclable materials by removing their organic components through gasification and combustion (Steiner et al., 2002). Bharat Heavy Electrical Limited (BHEL) developed a pressurized fluid bed gasifier to take into account the higher ash conent of coal. Raw product gas from the cyclone is cycled and mixed with the feedstock in the drier zone. Again the feedstock is separated and cooled gas is taken for cleaning while the feedstock is supplied to the gasifier. BHEL is developing a 125 MWe IGCC demonstration plant at Auraiya in Uttar Pradesh, India (Higman and Burgt, 2008).

As discussed in the previous paragraphs, Fe and Co are the industrially relevant catalysts that are currently commercially used in FT, with the choice of catalyst depending primarily on the target product (waxes vs. gasoline and olefins) and the feedstock. Cobalt is the catalyst of choice for GTL processes, using natural gas as feedstock and a H2/CO syngas molar ratio of 2, while Fe is used for CTL processes with a low-hydrogen content syngas. Few studies have investigated in depth the type of catalysts suitable for the BTL-FT process, starting from biomass feedstock (Escalona et al, 2009; Jun et al., 2004; Lapidus et al., 1994; van Steen and Clayes,

2008) . It is of crucial importance to explore the differences between GTL and CTL on the one hand and BTL on the other, in order to successfully implement the FT reaction in the BTL process. Both configurations currently investigated for the BTL process (full conversion and once through FT, see Section 19.2) require high overall and per pass CO conversion and high C5+ selectivity. As cobalt is more active than iron, cobalt has been so far used as the catalyst of choice for economic and exergetic evaluations of the BTL process. However, as analysed in an excellent recent review by van Steen (van Steen and Clayes, 2008), it is debatable whether this is truly the optimal choice of catalyst for the BTL process. van Steen argues that although Fe catalysts can operate with a lower hydrogen content syngas such as that from biomass gasification, a WGS reactor after gasification might be required for both cobalt and iron catalysts in order to obtain a good productivity. Since cobalt yields a higher productivity at high conversion levels, it seems to be the catalyst of choice for BTL synthesis of linear, heavier hydrocarbons if clean syngas is available. However, given that biomass syngas contains several poisons for FT catalysts, such as sulphur-, chloride — and nitrogen — containing compounds, and keeping in view the fact that Fe catalysts are reported to be more resistant to sulphur (van Steen and Clayes, 2008) and ammonia poisoning (Koizumi et al., 2004), the financial risk of operating the FT reactor with an iron-based catalyst seems to be lower. In real operation, deviations from design conditions are inevitable and contamination of the syngas entering the FT reactor is possible. In such case, iron catalysts would be less severely affected than the cobalt ones. Even in the case that the catalyst should be replaced, the much lower cost of iron compared to cobalt offers obvious economical advantages.

Wrapping up, both cobalt and iron catalysts should be considered as options for the FT reactor in the BTL process. A number of scenarios for the BTL process should be developed with both type of catalysts, while the overall process design should be coupled with catalyst developments in both cases in order to clearly prove the superiority of the one catalyst system to the other for commercial application.

Most advanced biofuel production technologies today are focused towards converting lignocellulosic biomass into transportation fuels. Lignocellulosic biomass refers to plant biomass that is composed of cellulose and hemicellulose, which are natural polymers of carbohydrates and lignin. Cellulose and hemicellulose are tightly bound to the lignin by hydrogen and covalent bonds. Lignocellulose comes in many different types such as wood residues (e. g. sawmill residues), crop residues from agriculture (e. g. corn stover and cereal straws), industrial residues from agro-food processing operations (e. g. wheat bran and sugar beet pulp) and dedicated energy crops (primarily rapidly growing energy grasses such as Miscanthus and switchgrass, and wood species).

21.3.1 Biochemical routes — sugar platform

In order to distinguish biofuels derived from lignocellulose from those derived from existing agricultural commodities (see Section 21.2.1), often the term ‘cellulosic’ is added to the biofuel. This term indicates that these biofuels are based on converting the main carbohydrate fractions, cellulose and hemicellulose, of lignocellulosic biomass into fuels.