The release of C5 and C6 sugars from lignocellulosic biomass deconstruction supports the growth of heterotrophic microorganisms and the metabolic conversion of sugars into biofuel. Representative hydrocarbon-based fuel titers produced by engineered, heterotrophic hosts are listed in Table 1. The most common heterotrophic hosts for biofuel production are the model organisms Escherichia coli and Saccharomyces cerevisiae. These hosts are attractive candidates for fuel production due to their fast growth rates, well-known genetics and regulation, advanced molecular tools for genetic engineering, and established use in the industrial setting. Neither E. coli nor S. cerevisiae naturally produce significant amounts of hydrocarbon-based fuels, necessitating the application of metabolic engineering techniques. Heterotrophic organisms that naturally produce hydrocarbon-based fuels are also potential hosts for large-scale biofuel production. For example, Bacillus subtilis naturally produces higher concentrations of isoprene than other commonly known bacteria like E. coli [55]. B. subtilis is also a model organism for Gram-positive bacteria with established tools for genetic modification, advancing its appeal as a host for isoprene production. Similarly, heterotrophic algae can produce significant quantities of TAG. This has motivated some preliminary investigation into engineering the model green alga, Chlamydomonas reinhardtii, for TAG production [56—58]. While most meta­bolic engineering efforts have focused on these model heterotrophic hosts, genetic tools can be developed for other organisms with desirable fuel production traits.

Most heterotrophic hosts for biofuel production utilize the Embden-Meyerhof-Parnas (EMP) pathway for sugar catabolism (Figure 4). The EMP pathway has evolved for efficient carbon utilization and is typically not rate-limiting for fuel production. As such, EMP pathway enzymes are not often targeted for genetic manipulation. However, the organic feedstock from lignocellulose deconstruction is comprised of a range of sugars, including hexoses: glucose, mannose, and galactose, and pentoses: xylose and arabinose [79]. A major concern in convert­ing these sugars into fuel is the efficient utilization of all available hexoses and pentoses. While some organisms like E. coli can naturally metabolize these different forms of sugar, others, like S. cerevisiae, can only utilize specific forms [80]. S. cerevisiae does not naturally express path­ways for catabolizing pentoses. There are two known pathways for xylose catabolism, both of which have been expressed in S. cerevisiae [81—83]. Xylose can be converted into xylulose-5- phosphate (Xu5P), an intermediate in the pentose phosphate pathway (PPP), through expres­sion of a xylose isomerase (XI) and xylulose kinase (XK) [82]. Alternatively, the XI can be replaced by a xylose reductase (XR) and xylitol dehydrogenase (XDH) [81, 82]. Complications

in these two xylose utilization pathways include the inhibition of XI by xylitol (Xol) and the reducing equivalents required by XR and XDH [80]. Successful strategies for engineering xylose utilization in S. cerevisiae include expression of a fungal XI from Piromyces sp. E2 along with overexpression of the non-oxidative PPP pathway [84] and expression of XR and XDH from the xylose-fermenting yeast Pichia stipitis [85]. Two pathways have also been expressed in S. cerevisiae for arabinose utilization [86, 87]. The bacterial pathway for arabinose catabolism consists of 3 enzymatic steps, while the fungal pathway involves 5 enzymatic steps, 4 of which require cofactors of NADPH or NAD+ (Figure 4). Efficient arabinose utilization in S. cerevi — siae has been achieved through heterologous expression of a bacterial arabinose catabolism pathway along with overexpression of the non-oxidative PPP and evolutionary engineering [88]. While most of these metabolic engineering examples focus on utilizing sugars for fermentation to ethanol, the strategies for engineering carbon utilization can also be applied for hydrocarbon-based fuel production.

Unlike S. cerevisiae, E. coli can utilize the hexoses and pentoses derived from lignocellulose; however, the carbon catabolite repression (CCR) system in E. coli leads to inefficient, diauxic growth [89]. Through CCR, E. coli sequentially consumes different sources of organic carbon based on substrate preference, leading to delayed and often incomplete utilization of unpre­ferred sugars like xylose and arabinose. This translates into lower productivities and yields along with downstream complications due to the presence of unmetabolized sugars [80]. As a result, CCR is often targeted by metabolic engineering to alleviate these undesired effects. A common engineering strategy is to use mutants of the transcriptional activator CRP (cyclic AMP receptor protein) which have been modified to eliminate the allosteric requirement for cAMP, thereby leading to expression of the pentose catabolizing pathways in the presence of the preferred substrate, glucose [90]. The phosphotransferase system (PTS), responsible for the preferential uptake of glucose, has also been deleted to encourage simultaneous utilization of mixed sugars [91]. Lastly, deletion of methylglyoxyal synthase was shown to improve the co­metabolism of sugars, ostensibly due to elimination of methylglyoxyal, an inhibitor of sugar metabolism [92]. Through modifying the components of CCR, E. coli can be engineered to efficiently utilize the organic carbon mixture resulting from lignocellulose degradation.

In addition to the hexoses and pentoses derived from lignocellulosic biomass, glycerol may soon become an inexpensive organic carbon source for fuel production. Glycerol is a byproduct of the conversion of TAG into biodiesel during algal biofuel processing, and thus, large quantities of glycerol may be available for use as an organic carbon source. The main pathway for aerobic glycerol utilization involves a two-step conversion to produce the glycolytic metabolite DHAP [93]. The glycerol utilization pathway is not a common target for metabolic engineering, yet glycerol has been reported as a supplementary carbon source for the produc­tion of isoprenoid-based fuels, farnesol and a-farnesene [47, 48]. Future metabolic engineering efforts may focus more on glycerol utilization as the availability of glycerol increases.

Second generation biofuel production still remains to be demonstrated at large scales, yet the overall process is easily integrated with current technologies. Equipment and practices used for agricultural harvesting can be directly applied to harvesting lignocellulosic biomass. In fact, some agricultural processes already produce biomass waste streams that can be utilized for feedstock, such as corn stover. Moreover, commercial fermenters can be employed as bioreactors for the microbial fuel conversion. The main technical difficulties in large-scale lignocellulosic fuel production center on provision of the carbon source. The quantities of biomass needed to support industrial-scale fuel production will require a significant invest­ment of land and nutrient resources, and the supply will be subject to varying climate conditions. A supply chain infrastructure must also be constructed to harvest the biomass and transport it to the production facilities. A primary technical focus of current research on lignocellulosic-derived fuels is the deconstruction of biomass into useable sugars. The thermal, chemical, and enzymatic processes for biomass deconstruction have been a limiting factor for economical second generation biofuel production [94, 95]. As the cost of biomass deconstruc­tion is reduced with new technology, the large-scale production of second generation biofuels will begin to contribute to the world’s supply of renewable energy.

A direct glycerol hydrogenolysis mechanism was recently proposed by Yoshinao et al. [50]. The experiments were carried out using Rh-ReOx/SiO2 and Ir-ReOx/SiO2 catalysts at 393 K and 80 bar H2 pressure. The low reaction temperature implies that the dehydration-hydroge­nation route was not further possible, due to the endothermic character of glycerol dehydra­tion and the required activation energy, and suggests the energetically more favored direct hydrogenolysis reaction [51]. They suggested a direct hydride ■ proton mechanism. The se­lected catalysts are able to activate hydrogen easily and to form hydride species. It is pro­posed that glycerol is adsorbed on the surface of ReOx clusters to form alkoxide species. Glycerol can form two adsorbed alkoxides: 2,3-dihydroxypropoxide and 1,3-dihydroxyiso-
propoxide; it is suggested that the formation of 2,3-dihydroxypropoxide is preferred as it re­quires a smaller adsorption cross-section than 1,3-dihydroxyisopropoxide [52]. Next, the hydride attack to the 2-position of 2,3-dihydroxypropoxide gives 1,3-PDO, while the hy­dride attack to the 3-position of 2,3-dihydroxyisopropoxide yields 1,2-PDO. The higher se­lectivity to 1,3-PDO obtained (1,3-PDO/1,2-PDO ratio = 2.7) is explained on the basis of the higher stability of the six membered-ring transition state that leads to the formation of 1,3- PDO as compared to the stability of the seven membered-ring transition state that leads to the formation of 1,2-PDO (Figure 11).

A different direct glycerol hydrogenolysis mechanism was established by Chia et al. [53] try­ing to explain the hydrogenolysis of different polyols and cyclic ethers over a Rh-ReOx/C catalyst. They concluded from DFT calculations that the — OH groups on Re associated with Rh are acidic. The acidic nature of ReOx was also reported before [54]. Such acidic Re sites can donate a proton to the reactant molecule and form carbenium ion transition states. In the case of glycerol hydrogenolysis, the first step involves the formation of a carbocation by pro­tonation-dehydration reaction. This carbocation is stabilized by the formation of a more sta­ble oxocarbenium ion intermediate resulting from the hydride transfer from the primary — CH2OH group. Final hydride transfer step leads to 1,2-PDO or 1,3-PDO [53]. The authors also reported that the secondary carbocation is more stable than the primary carbocation. Nevertheless, higher selectivity to 1,2-PDO was obtained (1,3-PDO/1,2-PDO ratio = 0.65).

DC and Impulse Plasma-Liquid Systems

Valeriy Chernyak, Oleg Nedybaliuk, Sergei Sidoruk, Vitalij Yukhymenko, Eugen Martysh,

Olena Solomenko, Yulia Veremij, Dmitry Levko, Alexandr Tsimbaliuk, Leonid Simonchik,

Andrej Kirilov, Oleg Fedorovich, Anatolij Liptuga, Valentina Demchina and Semen Dragnev

Additional information is available at the end of the chapter http://dx. doi. org/10.5772/53764

1. Introduction

It is well known [1] that hydrogen (H2) as the environmentally friendly fuel is considered to be one of the future most promising energy sources. Recently, interest in hydrogen energy has increased significantly, mainly due to the energy consumption increase in the world, and recent advances in the fuel cell technology. According to the prognosis, in the next decades, global energy consumption will be increased by 59%, and still most of this energy will be extracted from the fossil fuels. Because of the traditional fossil fuels depletion, today there’s a growing interest in renewable energy sources (f. e. — bioethanol, biodiesel). Bioethanol can be obtained from the renewable biomass, also it can be easily and safely transported due to its low toxicity, but it’s not a very good fuel. Modern biodiesel production technologies are characterized by a high percentage of waste (bioglycerol) which is hard to recycle.

It is common knowledge [2] that addition of the syn-gas to the fuel (H2 and CO) improves the combustion efficiency: less burning time, rapid propagation of the combustion wave, burning stabilization, more complete mixture combustion and reduction of dangerous emissions (NOx). Besides, the synthesis gas is an important stuff raw for the various materials and synthetic fuels synthesizing. There are many methods of synthesis gas (including hydrogen) production, for example — steam reforming and partial liquid hydrocarbons oxidation. Also,

there is an alternative approach — biomass reforming with low-temperature plasma assistance. Plasma is a very powerful source of active particles (electrons, ions, radicals, etc.), and therewith it can be catalyst for the various chemical processes activation. However, a major disadvantage of chemical processes plasma catalysis is weak processes control.

There is a bundle of electrical discharges that generate both equilibrium and non equilibrium plasma. For plasma conversion — arc, corona, spark, microwave, radio frequency, barrier and other discharges are used. One of the most effective discharges for the liquid hydrocarbons plasma treatment is the "tornado" type reverse vortical gas flow plasma-liquid system with a liquid electrode ("TORNADO-LE") [3]. The main advantages of plasma-liquid systems are — high chemical plasma activity and good plasma-chemical conversions selectivity. It may guarantee high performance and conversion efficiency at the relatively low power consump­tion. Moreover, those are systems of atmospheric pressure and above, and this increases their technological advantages.

Also, syn-gas ratio — hydrogen and carbon monoxide concentration ratio should be mentioned. As well, it should be taken into consideration that for efficient combustion (in terms of energy) of the synthesis gas it should contain more hydrogen, and in the case of the synthesis materials — they should contain more CO.

Relatively new possible solution to this problem — carbon dioxide recycling. Many modern energy projects have difficulties with the large amount of CO2 storing and disposing. And it is also known that the addition of CO2 to plasma during the hydrocarbons reforming may help to control plasma-chemical processes [4]. That is why the objective of the research is to study the influence of different amounts of CO2 in the working gas on the plasma-chemical processes during the hydrocarbons conversion.

This research deals with hydrocarbons (bioethanol, bioglycerol) reforming by means of the combined system, which includes plasma processing and pyrolysis chamber. As a plasma source the "tornado" type reverse vortical gas flow plasma-liquid system with liquid electrode has been used [5].

Qualitatively new challenge is connected with a selectivity of the plasma chemistry strength­ening by the transition of the chemical industry to "green chemistry". The last is a transition from the traditional concept of evaluating the effectiveness of the chemical yield to the concept that evaluates the cost-effectiveness as the exclusion of hazardous waste and non-toxic and/or hazardous substances [6].

A quantitative measure of the environmental acceptability of chemical technology is the ecology factor, which is defined as the ratio of the mass of waste (waste) to the mass of principal product. Waste is all that is not the principal product.

By the way, the most promising approaches in green chemistry is the implementation of processes in supercritical liquids (water, carbon dioxide) [7].

Water in supercritical condition unlimitedly mixes with oxygen, hydrogen and hydrocarbons, facilitating their interaction with each other — oxidation reactions are very fast in scH2O (supercritical water). One particularly interesting application of this water — efficient destruc­tion of chemical warfare agents. When mixed with other substances scH2O can be used not only for oxidation but also in the reactions of hydrolysis, hydration, the formation and destruction of carbon-carbon bonds, hydrogenation, and others.

Besides, the use of pulsed electrical discharges in the liquid brings up new related factors: strong ultraviolet emission and acoustic or shock waves. In literature it can be found that systems with energies more than 1 kJ/pulse, that have negative influence on the lifetime of such systems. Reasonable from this perspective is the usage of pulsed systems with relatively low pulse energy and focusing of acoustic waves. In addition, the acoustic oscillations in such systems can be used as an additional mechanism of influence on chemical transformations.

In using of acoustic oscillations for chemical reactions the most attention is paid to systems with strong convergent waves. However, the processes during the collapse of the powerful convergent waves are studied unsufficiently. In the literature the systems of cylindrical, spherical or parabolic surfaces used in the focusing of shock waves for technological needs are known [8]. However, among their disadvantages should be noted that partial usage of the energy of acoustic wave and the problem of it’s peripheral sources synchronization, which leads to distortion of the shock wave front ideality and reduces the focusing effectiveness.

Probably, more perspective method of using acoustic waves is their generation by single axial pulse electric discharge with further reflection from an ideal cylindrical surface. This approach can provide better symmetry of compression by convergent acoustic wave both in the gas and in the liquid. Probably, such mechanism can be exploited for scH2O production

In addition, the re-ignition of electrical discharge at the moment of collapse convergent acoustic waves can lead to the plasma temperature increasing due to compression of the discharge channel, as well as the appropriate amplification of acoustic waves after the collapse.

It’s clear that plasma-liquid systems (PLS) mentioned above have some sharp differences. Therefore, the first section of this article presents the results of our research on the addition of CO2 to the "TORNADO-LE". And the second section of the article is devoted to investigation of double-impulse system in underwater electric discharge.

The appropriate temperature for optimal fermentation ability of C. Acetobutylicum strains strongly depends not only on the type of strain, but on the composition of the medium and raw materials as well, and is strongly influenced by a series of factors such as presence or absence of additives, sugar concentration, pH, and others. McNeil and Christiahsen studied the effect of temperature on the solvent production by C. acetobutylicum in the range 25 to 40°C [49]. It was found that the solvent yield decreased with increasing temperature. Consid­ering total solvent yield and productivity only, the optimum fermentation temperature was found to be 35°C [49]. Comparison of the solvent production by using strains of C. acetobu — tylicum and C. butylicum from whey showed that higher yields of solvents were observed at 37 °C or 30 °C, respectively [50]. Oda and Yamaguchi [51] concluded that temperature control played important role in the solvent yield and the optimal temperatures were not found to be the same during different stages of the process. Harada [52] concluded that the yield of BuOH was increased from 18.4-18.7% to 19.1-21.2% by lowering the temperature from 30 °C to 28 °C when the growth of the bacteria reached a maximal rate.

Fouad et al. [53] studied fourteen different media in the fermentative production of acetone and butanol. The highest total yields were achieved in medium containing potato starch and soluble starch as C sources. Compositon and pH of the medium have important influence on ABE fer­mentation. The contaminants in the media have decisive effect on the ABE fermentation. For ex­ample, hydrolysates obtained by enzymatic saccharification of wheat straw or cornstover pretreated by steam explosion in classical or acidic conditions, were found non-fermentable in­to acetone-butanol. A simple treatment involving heating the hydrolysates in presence of calci­um or magnesium compounds such as Ca(OH)2 or MgCO3 at neutral pH values restored normal fermentability to these hydrolysates [54]. Sugar concentration of the media also influences the ABE fermentation. Fond et al. [55] studied growing of C. acetobutylicum in fed-batch cultures at different feeding rates of glucose. The sugar conversion to BuOH and Me2CO increased with in­creasing the glucose flow whereas, on the contrary, conversion to butyric acid was highest at slow glucose feeding rate. The AcOH concentration was constant at different flows of glucose and the solventogenesis was not inhibited at high flow of sugar [55].

The amount and chemical form of inorganic and organic nitrogen sources basically affect on the ABE process. They influence also strongly depends on presence or absence of other important additives. Among studied inorganic nitrogen compounds, ammonium nitrate and urea could stop the fermentation in the middle, (NHJHSO^ NH4Cl, and (NH4)2HPO4 resulted acetone-rich fermentation, while (NH4)2CO3 and NH4OH gave BuOH-rich fermentation [56]. Baghlaf et al [57] studied the effect of different concentrations of corn steep liquor, fodder yeast, soybean meal, corn bran, rice bran, and KH2PO4 in the ABE fermenatation, and the organism preferred utilization of natural organic sources. The best concentration of KH2PO4, favouring the ABE production was found to be 2 g/L. Oda [58,59] occurred a little effect of adding (NH4)2SO4 to EtOH-extracted soybean meal in the yield of solvents, however, cane molasses and dried yeasts were good supplements to the same soybean meal. Addition of asparagine retarded the fermentation. When used as the sole N source, soybean press cake and egg white were good; the others tested were, in the order of decreasing suitability, EtOH-extracted soybean meal, casein, fish protein, zein, gluten, yeast protein, and gelatine. With peanut cake as the N source, Ca salts were not desirable. The stimulants tested were mostly effective: they were, in the order of decreasing effect, liver (best), rice bran-clay, a-alanine, a-methylphene — thylamine-H2SO4, p-alanine, p-aminobenzoic acid, naphthaleneacetic acid, and cane molasses — clay (the last two were slightly worse than the control without stimulant). Doi et al. [60] could occur that growth-promoting amino acids in the casein acid-hydrolyzate can be divided into three groups: the bacteria required isoleucine, valine, and glutamic acid; asparagine, serine, threonine, alanine, and glycine accelerated fermentation. Leucine, phenylalanine, methionine, tryptophan, proline, lysine, histidine, and arginine were not required for growth and cystine and tyrosine inhibited fermentation.

As it was mentioned earlier, the choice of catalyst and operating parameters affect the reac­tions that take place within the hydroprocessing reactor. The key operating parameters of hydroprocessing include the reactor temperature, hydrogen partial pressure, liquid hourly space velocity and hydrogen feed-rate.

1.1.4. Temperature

Most catalytic hydrotreating and hydrocracking reactors operate between 290-450°C. The temperature range is selected according the type of catalyst and feedstock type to be proc­essed. In the first stages of the catalyst life (after its loading in the reactor) the temperature is normally kept low as the catalyst activity is already very high. However as time progresses and the catalyst deactivates and cokes, the temperature is gradually increased to overcome the loss of catalyst activity and to maintain the desired product yield and quality.

1.1.5. Hydrogen partial pressure

Hydrogen partial pressure affects significantly the hydrotreating reactions as well as the cat­alyst deactivation. The catalyst deactivation rate is inverse proportional to the hydrogen partial pressure and to hydrogen feed-rate. However high hydrogen partial pressures corre­spond to high operational costs, which rise even higher for high olefinic feedstocks that ex­hibit higher hydrogen consumption due to the saturation reactions. Therefore hydrogen partial pressure should be balanced with the catalyst activity and catalyst life expectancy in order to optimize the overall process.

This process which uses reactive distillation technology is developed by combined expertise of Huls AG and UOP. The feed consists of MeOH or EtOH and hydrocarbon streams con­taining reactive tertiary olefins such as isoamylene and IB. Reaction takes place over an acid­ic ion exchange resin at mild temperature and moderate pressure.

In the MTBE case, feed first passes through an optional water wash system to remove the resin contaminants. The majority of the reaction is carried out in a simple fixed-bed reactor. The reactor effluent feeds the reactive distillation column containing a proprietary packing where simultaneous reaction of the remaining IB and distillation occur.

Overhead from the reactive distillation column is routed to MeOH recovery, a simple coun­tercurrent extraction column using H2O, and a MeOH-H2O distillation column. The recov­ered MeOH is recycled to the reactor section. Hydrocarbon raffinate is typically sent to a downstream alkylation or oligomerization unit.

1.2. Synthetic fuels

Unlike biofuels, which transform biological molecules into petroleum substitutes, synthetic fuels take a raw biological material, and through chemical processing, create compounds identical to petroleum fuel. This has a very distinct advantage over common biofuels in that there are no compatibility issues between the traditional fuels nor is there a need for any en­gine or fuel line modifications required. Synthetic fuels are usually made by utilizing a com­plex biological molecule and through thermal processing, break down the material into simple chemical building blocks (i. e., methane, carbon monoxide, hydrogen, etc.) and re­form them into target chemicals. There are limitations with synthetic fuels production, espe­cially when pertaining to production from aquatic and marine biomass where the water content is naturally higher than 99% by weight in its natural state, since initial breaking down of the products is usually through thermal processing that require dry or near dry conditions. However, since algae and aquatic biomass has such diverse characteristics and high cellular energy density, there is benefit for using either algae where the lipids have been extracted or whole algal cells as feedstock for these thermal synthetic fuel processes and thus can be considered as an option for production of synthetic fuels.

To retain high cell densities in reactor, microbes can be grown as biofilm attached to carrier ma­terial. Planktonic cells can be retained in the fermentation broth by installing solid/liquid sepa­rators such as membranous ultra-filtration units, spiral wound filtration systems, hollow fibres, cell-recycling membranes and centrifuges [214—216]. The concentrations of solvents from gas fermentation rarely exceed 6% [w/v] so a cost — and energy — efficient product recovery process is required. Furthermore, acetogens also exhibit low resistance towards solvents like ethanol [217, 218] and butanol [219, 220] so an in situ/online product recovery system can en­hance solvent productivity by decreasing solvent concentrations (and hence toxicity) in the fer­mentation broth. Distillation has been the traditional method of product recovery but the associated high energy costs have led to the development of alternative methods such as liq­uid-liquid extraction, pervaporation, perstraction, and gas stripping [24, 221].

End product inhibition can be reduced by in situ removal of inhibitory fermentation products as they form. The first experiments were performed by Bekhtereva who studied the effect of BuOH on the ABE fermenting process and on the development of Clostridium acetobutylicum in concentrated mash. Experimental removal of neutral products from the substrate during fermentation was tested by continuous extraction with castor oil. This oil could extract acetone 13-60, EtOH 5-20 and BuOH 50-88% from the wort. By adding the oil to the medium in varying amounts depending on the carbohydrate content, it was possible to ferment corn mash of 3-5 times the usual concentrations. The yield of acetone was 20-37 g L-1 of wort, that of all neutral products 60-100 g L-1. Their concentrations in the wort under the oil layer was usually lower than in control vessels, e. g., total products 1.4-2.3%, BuOH 0.4% against 1.2-1.3% in usual fermentation. The extraction was beneficial to the development of the bacteria [151]. Other starch-containing materials could also be fermented in the usual manner, and BuOH and Me2CO were continuously removed by means of a solvent immiscible with H2O, e. g. castor oil [152]. The extraction processes were coupled to batch, fed-batch, and continuous BuOH fermentation to affirm the applicability of recovery techniques in the actual process. In batch and fed-batch fermentation, a 3-fold increase in the substrate consumption, in continuous fermentation ~30% increase could be achieved. [142].

Toxicity and selectivity of 13 organic compounds were tested in extractive batch fermentations performed with C. acetobutylicum. Among them, oleyl alcohol and mixed alcohol (the mixture of oleyl alcohol and C18 alcohol) were the best for acetone-BuOH fermentation. The orthogonal — cross-test method with 3 elements and 3 levels was used to evaluate effects of fermentation temperature, inital glucose concentration, and solvent/water ratio on extractive batch ABE fermentation. Extractive batch ABE fermentation in a stirred fermentor was studied at different initial glucose concentrations at 41/35° and at solvent/water ratio 1:2. When initial glucose concentration was 110 g L-1, at the end of extractive fermentation the BuOH concentration in the broth and in the solvent was 5.12 and 22.3 g L-1, respectively. The total BuOH and ABE concentrations based on the broth volume were ~16.27 and 33.63 g L-1, the conversion ratio of glucose was 98% and the total ABE yield was 0.312. In situ extractive fermentation could eliminate the inhibition of BuOH on microbial growth, increased the initial glucose concen­tration and reduced the wastewater amount, thus the consumption of energy could be reduced for the separation and purification of the products [153]. Roefler et al [154] studied the effect of six extractants in batch extractive fermentation: kerosene, 30 wt.% tetradecanol in kerosene, 50 wt.% dodecanol in kerosene, oleyl alcohol, 50 wt.% oleyl alcohol in a decane fraction, and 50 wt.% oleyl alcohol in benzyl benzoate. Best results were obtained with oleyl alcohol or a mixture of oleyl alcohol and benzyl benzoate. In normal batch fermentation of C. acetobutyli- cum, glucose consumption is limited to ~80 kg m-3 due to accumulation of BuOH in the broth. In extractive fermentation using oleyl alcohol or a mixture of oleyl alcohol and benzyl benzoate, >100 kg m-3 of glucose can be fermented. Maximal volumetric BuOH productivity was increased by ~60% in extractive fermentation compared to batch fermentation. BuOH produc­tivities obtained in extractive fermentation compare favorably with other in situ product removal fermentations [154].

A medium for ABE fermentation by C. acetobutylicum was mixed with 0.2-5.0% 1-octanol or 2- ethylhexanol and various parameters of fermentation were studied. Glucose consumption, cell growth, ABE formation, and acetate and butyrate formation were inhibited, especially at higher solvent concentrations. Octanol was more toxic than 2-ethylhexanol [155]. A mathematical model for simultaneous fermentation and extraction of the products was derived for ABE pro­duction by immobilized C. acetobutylicum cells in a microporous hollow fiber based tubular fermentor-extractor. The solvent, 2-ethyl-1-hexanol, is used for in situ dispersion-free extrac­tion of products. Both predicted and experimental data follow the same trend. The experimen­tally observed value of total solvent productivity increased by >40% as a result of in situ solvent extraction [156]. Unfortunately, good extractants for BuOH, such as decanol, are toxic to C. ace — tobutylicum. The use of mixed extractants, namely, mixtures of toxic and nontoxic coextrac­tants, was tested to circumvent this toxicity. Decanol appeared to inhibit BuOH formation by C. acetobutylicum when present in a mixed extractant that also contained oleyl alcohol, however, maintenance of the pH at 4.5 alleviated the inhibition of BuOH production and the consump­tion of butyrate during solventogenesis. A mixed extractant that contained 20% decanol in oleyl alcohol enhanced BuOH formation by 72% under pH-controlled conditions. A mechanism for the effects of decanol on product formation is proposed [157]. The same mixed extractant that contained 20% decanol in oleyl alcohol were combined by Wang et al. to carry out in-situ extrac­tive acetone-butanol fermentation, resulting 19.21 g/L of butanol concentration. Butanol pro­ductivity could be 62.8% higher than that of control; meanwhile, total organic solvent productivity increases by 42.3% as compared to the control [145].

BuOH fermentation was carried out by contact with solvent containing C10-14 alcohols, as well. A seed culture of C. acetobutylicum IAM19013 was inoculated and mixed with tridecanol. The broth was anaerobically fermented, with stirring, at 37 °C for 60 h. The solvent layer at the top of the fermentor was circulated to the bottom. The concentrations of BuOH in the solvent and vapor were 41.6 g L-1 and 66%, respectively [158]. Higher alcohols, e. g. C16-18 unsaturated alcohols and C16-20 branched alcohols were also tested for continuous extraction of BuOH from the medium during the fermentation period. Extraction of the BuOH from the medium by using unsaturated or branched alcohols innoxious to the microorganism markedly increased BuOH yield. Thus, C. acetobutylicum was anaerobically cultivated at 37°C on a medium containing 10% glucose and, after 30 h, 40% oleyl alcohol was added to the broth to remove the BuOH from the aqueous phase and thereby reactivate the fermentation. This increased the total BuOH concentration to 2.5-fold in an additional 70 h [159]. Oleyl alcohol was found to be one of the best solvents for in-situ extractive ABE fermentation. Its butanol partition coefficents value was varied between 3.0 and 3.7 depending on the composition of the broth, nontoxic, nonmiscible and its boiling point is high as compared to ABE solvents. Batch and fed-batch extractive fermentation by C. acetobutylicum was studied with oleyl alcohol as extractant. Extractive fermentation could reduce the product inhibition, increase the initial glucose concentration and increase the fermentation rate. A mathematical model was suggested to describe batch fermentation processes. The proposed model could simulate the experimental data fairly well [160].

In situ removal of inhibitory products from C. acetobutylicum resulted in increased reactor productivity; volumetric butanol productivity increased from 0.58 kg m-3 h-1 in batch fermen­tation to 1.5 kg m-3 h-1 in fed-batch extractive fermentation using oleyl alcohol as the extraction solvent. The use of fed-batch operation allowed glucose solutions of up to 500 kg m-3 to be fermented, resulting in a 3.5-5-fold decrease in waste water vol. Butanol reached a concentra­tion of 30-35 kg m-3 in the oleyl alcohol extractant at the end of fermentation, a concentration that is 2-3 times higher than is possible in regular batch or fed-batch fermentation. Butanol productivity and glucose conversion in fed-batch extractive fermentation was compared with continuous fermentation and in situ product removal fermentation [161].

In ABE fermentation using C. acetobutylicum IAM 19012, it was necessary not only to keep BuOH concentration below the toxic level (2 g L-1), but also to control glucose concentration at <80 g L-1 and pH between 4.5 and 5.5. The amount of glucose consumed could approximatly be estimated as 4 times the volume of gas evolved, and BuOH was produced from glucose with an average yield of 0.173. It was thus possible to estimate the concentration of glucose and BuOH at any fermentation time using the volume of gas evolved as an indicator. As oleyl alcohol was an excellent extracting solvent for BuOH, a fed-batch culture system for the microorganism was developed, where withdrawing and feeding operations of the solvent were done automatically based on gas evolution [162]. Ohno combined the fermentation by C. beijerinckii ATCC 25752 which perfectly inhibited the process at the BuOH concentration of 12 kg m-3 with the extraction with oleyl alcohol and removing the butanol from its mixture with oleyl alcohol which was car­ried out by prevaporation with hollow fiber membrane. When the BuOH concentration in oleyl alcohol was 22 kg m-3, the BuOH flux was 3.6’10-4 kg m-2 h-1 at 35 °C [163].

Extraction with the non-toxic immiscible solvent, oleyl alcohol was combined with fermenta­tion performed with immobilized C. acetobutylicum to ferment glucose to ABE solvents in a fluidized-bed bioreactor. The extracting solvent had a distribution coefficient of near 3 for butanol. Nonfermenting system tests indicated that equilibrium between the phases could be reached in one pass through the column. Steady-state results are presented for the fermentation with and without extractive solvent addn. One run, with a continuous aqueous feedstream containing 40 g L-1 glucose, was operated for 23 d. Steady state was established with just the aqueous feedstream. About half of the glucose was consumed, and the pH fell from 6.5 to 4.5. Then, during multiple intervals, the flow of organic extractive solvent (oleyl alc.) began into the fermenting columnar reactor. A new apparent steady state was reached in about 4 h. The final aqueous butanol concentration was lowered by more than half. The total butanol production rate increased by 50-90% during the solvent extraction as the organic-to-aqueous ratio increased from 1 to 4, respectively. A maximal volumetric productivity of 1.8 g butanol h-1 L-1 was observed in this nonoptimized system. The butanol yield apparently improved because of the removal of the inhibition. More substrate is going to the desired product, butanol, and less to maintenance or acid production, resulting in a 10-20% increase in the ratio of butanol relative to all products [164].

Whole broth containing viable cells of C. acetobutylicum was cycled to a Karr reciprocating plate extraction column in which acetone and butanol were extracted into oleyl alcohol flowing counter-currently through the column. A concentrated solution containing 300 g L-1 glucose was fermented at an overall butanol productivity of 1.0 g L-1 h-1, 70% higher than productivity of normal batch fermentations. The continuous extraction process provides flexible operation and lends itself to process scale-up [165].

A new type of bioreactor containing a porous permeable wall to recover the biobutanol pro­duced in anaerobic ABE fermentation processes was developed [166, 176]. The ferment liquor is contacted with a non-toxic organic solvent as oleyl alcohol and the butanol in the fermentation liquor distributes between the organic phase and the ferment liquor. The butanol containing solvent located at one side of the permeable wall is in diffusion equilibrium with a same kind of auxiliary solvent with lower butanol concentration located at the other side of the permeable wall. Due to concentration difference, butanol diffuses from one side of the wall to the other side. The concentration difference is kept to be constant by continuous removal of the butanol form the auxiliary solvent phase in which the butanol concentration is always lower than in the extractant phase but much higher than the butanol concentration in the ferment liquor phase. In this way, the primary extractant solvent contacting the ferment liquor is only a transmitting me­dia between the ferment liquor and a small volume of the auxiliary solvent separated with the permeable wall. Energy demand of the distillation to remove the butanol from the auxiliary sol­
vent is less than energy demand of the direct butanol recovery from the ferment liquor or from the extractant phase [166]. The porous composite membranes used as permeable walls for ABE production can be prepared by the method of Tamics et al. [167].

Not only simple alcohols but polyols can also be used in extractive fermentation systems for ABE production. Mattiasson et al. [168] produced acetone and BuOH by C. acetobutylicum in an aqueous two-phase system using 25 % polyethylene glycol 8000. Bacteria remained in the lower phase, and the partition coefficients of acetone and BuOH favoring the upper phase were 2.0 and 1.9, resp. Mean productivity was estimated at 0.24 g BuOH L-1 h-1, producing 13 g BuOH L-1 in 50 h. Poly(propylene)glycol 1200 is the highest partition coefficient reported to date for a biocompatible ABE extracting solvents. Extractive fermentations using concentrated feeds produced ~58.6 g L-1 acetone and BuOH in 202 h, the equivalent of 3 control fermentations in a single run. Product yields (based on total solvent products and glucose consumed) of 0.234-0.311 g g-1 and within-run solvent productivities of 0.174-0.290 g L-1 h-1 were consistent with conventional fermentation reported in the literature. The extended duration of fermen­tation resulted in an overall improvement in productivity by reducing the fraction of between — run down-time for fermentor cleaning and sterilization [141].

Two aqueous two-phased systems involving polyol-type extractants were investigated to determine their ability to reduce product inhibition in the acetone-BuOH-EtOH fermentation. An industrial-grade dextran (DEX) and a hydroxylpropyl starch polymer (Aquaphase PPT (APPT)) were tested as a copolymer with polyethylene glycol (PEG) to form a two-phased fermentation broth. Two-phase fermentation performances in the DEX-PEG and APPT-PEG 2-phase systems were compared to a single-phase conventional fermentation through a series of batch runs. Effects of the phase-forming polymers on C. acetobutylicum also were investi­gated. With a BuOH partition coefficient of 1.3, the BuOH yield with the two-phase system was increased by 27% over conventional fermentation [169].

Dibutyl phthalate is one of the ester-type extractants used in extractive fermentation of glucose, glucose-xylose mixtures and hydrolyzates of lignocellulosics to acetone-butanol solvents. Dibutyl phthalate has satisfactory physical properties, nontoxic and mildly stimulates the growth of the organism used, C. acetobutylicum P262. Sugar concentrations mainly in the range of 80-100 g L-1 resulted in solvent concentrations of 28-30 g L-1 in 24 h extractive fermen­tation compared to 18-20 g L-1 for nonextractive control fermentation. Conversion factors of 0.33-0.37 g solvents g-1 sugar consumed were obtained. Rapid fermentation was achieved by high cell concentrations and cell recycling from every 24 h fermentation to succeeding similar 24 h fermentation. Somewhat higher nutrients were also helpful. By this means, 255 L of acetone-butanol solvents were obtained per ton of aspen wood, 298 L per ton of pine, and 283 L per ton of corn stover. Such high product yields from inexpensive substrates offer the prospect of economic viability for the process [170].

Induction of flocculation of Clostridia led to a reduction of the specific solvent production rate. Cells adhering to sintered glass are better than flocculating cells for continuous BuOH-acetone fermentation. Due to low toxicity, in-situ application of paraffin, oleic alcohol or stearic acid butyl ester with the cells in the fermenter is possible. Solvent production by Clostridia can be considerably enhanced by the extractive process. Extraction may be directly integrated into a continuous fermentation. Separation of BuOH from oleic acid is easy due to the high boiling point of the extractant (260 °C) being far above the boiling point of BuOH (117 °C). Thus, BuOH can be obtained by normal distillation and the extractant can be recycled [171].

BuOH could be manufactured by cultivating BuOH-producing microorganisms such as C. ace — tobutylicum in a medium containing a fluorocarbon extractant. The generation time, the mean BuOH production rate, and the mean final BuOH concentration in the C. acetobutylicum cul­ture medium containing Freon-11 (1 g L-1) were increased by 29, 19, and 12%, respectively. Pro­duction of acetone and EtOH was not affected [172]. Continuous fermentation of a carbohydrate substrate with continuous extraction of the product by CFCl3 took place in a cylin­drical fermentor, with an inlet at the center and a filter membrane concentric with the outer wall, allowing the medium to diffuse outward and to retain microorganisms. The collected medium is pumped to an extractor, where it contacts CFCl3 or another material with a high solvency for BuOH and a low solvency for H2O and then separated into two phases. The extracted medium is recycled to a feed tank. The solvent is removed from BuOH in an evaporator, where BuOH is collected and the solvent pumped to a compressor and re-utilized [173].

Organic solvents having relatively high distribution coefficients for BuOH against water, often higher alcohols, esters, and organic acids, are very toxic to the microorganisms for BuOH fermentation. Most fermentation inhibition caused by solvent toxicity was eliminated by re­extracting the primary extractant solvent from the residual phase, to be recycled from the product extraction column to the fermentor by paraffin as an extractive fermentation process applied externally to product extraction. After selecting 2-octanol as the extractant from the standpoint of energy consumption in BuOH recovery, a two-stage-extraction BuOH extractive fermentation process having the possibility of reducing the production cost of BuOH was proposed [174]. Heptanal shows strong toxic effect towards C. Acetobutylicum R1 and T5 strains [175] but it has extremely high distribution coefficent (11.5) for butanol [175,176]. Ex — situ extraction with heptanal and recycling the residual broth into a new fermentation cycle proved to be unsuccesfull because the broth contained approx. two times higher heptanal concentration than the toxic limit. Diluting the recycled broth or extracting it with a secondary non-toxic apolar solvent such as hexane to remove the residual dissolved heptanal, inoculate the recycled broth with fresh bacteria in each cycle showed that 4-5 cycles of fermentation could be obtained without important decreasing in the ABE yields and productivity [175]. A multiple solvent extraction is described by Shi et al [177].

Mathematical formulation was made for the performance evaluation considering two types of solvent-supplying strategies. One is to add multiple solvents simultaneously and the product is removed at one time. Another is to add them one by one consecutively. Computer simula­tions were made for batch, fed-batch, and repeated fed-batch operation of acetone-BuOH fermentation to show the power of the approach. Significant improvement in terms of productivity and product concentration is expected when two extractants such as oleyl alcohol and benzyl benzoate are used, as compared to using only one solvent [177]. A two-stage — extraction butanol extractive fermentation process was developed and studied using a bench — scale extractive fermentation plant with a butanol production capacity of ~10 g h-1. The production rate equation for extractive fermentation was simply expressed by a previously reported equation multiplied by an equation for the extraction raffinate recycling effect. A butanol production-cost calculation program for the two-stage-extraction process determined the optimum operational conditions to be when butanol concentration, residual sugar concentration and recycling ratio were 6 kg m-3, 15 kg m-3 and 3, respectively. These optimal conditions were achieved in the bench-scale plant when it was operated with total sugar concentration, dilution rate and recycling ratio of 113 kg m-3, 0.158 h-1 and 3, respectively [178].

A special kind of in-situ extractive fermentation is the so-called perstraction, where a selective membrane is located between the broth and the extractant phase. Both sides of the membrane contact with each phase and ensures a medium betwen two immiscible phases to exchange butanol content. Due to lack of direct contact betwen two phases, toxicity or other problems can be eliminated and a dispersion-free extraction is possible, leading to an easy operation of the equipment, but the mass transfer in the membrane becomes important. This extraction processes were coupled to batch, fed-batch, and continuous BuOH fermentation to affirm the applicability of the recovery techniques in the actual process. In batch and fed batch fermen­tation a 3-fold increase in the substrate consumption could be achieved, while in the continuous fermentation it increases by~30% [142]. Jeon and Lee [179] described a fed-batch operation for enhanced separation with a semipermeable silicon membrane which showed high specific permeability to BuOH and acetone. Among various solvents examined, oleyl alcohol and polypropylene glycol were the most suitable as extractants. In fed-batch operation of the membrane-assisted extractive BuOH fermentation system, significant improvements were found in comparison to a straight batch fermn. The total glucose uptake per run was raised to 10 times of the value normally found in batch fermentation. The solvent productivity increased by a factor of 2. The total solvent yield increased by 23% due to reduction of acid production and reuse of cells in the fed-batch operation [179]. A continuously operated membrane bioreactor was connected to a 4-stage mixer-settler cascade and Clostridium acetobutylicum was cultivated in this reactor. BuOH was selectively extracted with butyric acid-saturated decanol from the cell-free cultivation medium, and the BuOH-free medium was refed into the reactor. Due to high boiling point of decanol, recovery of BuOH from the decanol solution is easy. Both partition coefficient and selectivity of BuOH in the cultivation medium-decanol system are sufficiently high for removing it from the medium. Direct contact of cells with the decanol phase causes cell damage. However, decanol is practically insoluble in the fermenta­tion medium, thus the contact of the cell-free medium with the solvent phase does not influence cell growth neither product formation. At a dilution rate of D=0.1 h-1, BuOH productivity was increased by a factor of 4 by removing BuOH from the medium [180].

As it has been stated in the introduction, a general characteristic of bio-oils coming from the pyrolysis of biomass is their high oxygen content (35-40 wt%). More than 300 compounds have been identified in bio-oil, most of them containing oxygen atoms. The exact composi­tion of the bio-oil depends on the type of biomass fed. These compounds can be classified in five broad categories: (i) hydroxyaldehydes, (ii) hydroxyketones, (iii) sugars and dehydrosu­gars, (iv) carboxylic acids, and (v) phenolic compounds [16]. Hydroprocessing of biomass — derived oils differs from processing petroleum because of the importance of deoxygenation as compared to nitrogen or sulfur removal. Bio-oil hydrodeoxygenation (HDO) process im­plies complex reaction networks that includes cracking, decarbonylation, decarboxylation, hydrocracking, hydrogenolysis, hydrogenation and polymerization. The upgrading process should yield a product with lower amount of water and oxygen, decreased acidity and vis­cosity, and higher HV. The complexity of the reactions and the high variety of oxygenated compounds make the evaluation of bio-oil upgrading difficult and has brought the use of model compounds such as phenol, guaicol, 2-ethylphenol, methyl heptanoate or benzofuran to test different catalysts and to understand the main characteristics of the HDO process. El­liot [17] has reported the HDO reactivity of different organic compounds that are typically present in bio-oils (see Figure 3). Olefins, aldehydes and ketones can easily be reduced by H2 at temperatures as low as 150-200 °C. Alcohols react at 250-300 °C by hydrogenation and thermal dehydration to form olefins. Carboxylic and phenolic ethers react at around 300 °C. Regarding the operating pressures, due to the low solubility of hydrogen in organic and aqueous solutions, high pressures are required to guarantee high availability of hydrogen in the vicinity of the catalyst (80-300 bar of H2 pressure) [15].