In the NExTAME process, the feed is typically light fluid catalytic cracker (FCC) gasoline and/or a light pyrolysis gasoline fraction from which the diolefins are removed by selective

hydrogenation. It is an integrated process consisting of reactors and a distillation tower. Etherification is carried out close to the thermodynamic equilibrium in the reactors from where the reaction product is directed into a distillation tower. A side stream containing un­reacted tertiary olefins and alcohol is recycled back into the reactors. This proprietary tech­nology leads to a high conversion of a broad range of olefins and alcohol. The process does not require a separate alcohol recovery section.

Similar to NExTAME, the NExETHERS process consists of fixed bed reactors and two distil­lation towers. In fixed bed reactors, the etherification synthesis is completed to the thermo­dynamic equilibrium in the reactors. The reaction product is then directed into the main fractionator. The bottom product of the main fractionator comprises the ether product. A side stream containing unreacted tertiary olefins and unreacted alcohol is recycled back into the etherification section. The second tower is used to recycle the rest of the excess alcohol and to remove oxygenates from C4 cut which can be directed for alkylation without any fur­ther oxygenate removal or drying.

There are three common methods for producing molecular precursors for synthetic fuels from biomass, and several variants of each method, dependent on the specific feedstock characteristic. These three methods are gasification, pyrolysis, and liquefaction. Pyrolysis and liquefaction will both produce of form of bio-oil that can be processed along with petro­leum oil stocks and made into useful fuel products, while gasification will produce gaseous products such as carbon monoxide, methane, and hydrogen (commonly called syngas or synthesis gas in this process), and can be further refined directly to produce specific fuel molecules.

1.3. Gasification

Gasification is a process in which carbonaceous materials are exposed to heat and a sub-stoi­chiometric concentration of air to produce partially oxidized gaseous products that still have a high heating value with relatively lower concentrations of carbon dioxide due to limited oxygen [20]. Syngas can be catalytically reformed into a liquid fuel through the Fischer — Tropsch process, which converts carbon monoxide and hydrogen into long-chain hydrocar­bons. By-products of the process include ash (formed from alkali-metal promoters present in the original reaction), char and tars that are created due to inefficiencies in mixing and heat distribution. This can be problematic when using water based biomass as the feedstock, since there will either be very high costs (in both energy and cost) to dry, or numerous un­wanted products formed through side reactions. Three main types of gasification reactor are commonly used in industry: fixed bed, fluidized bed and moving bed. Each process has in-

herent advantages and drawbacks based on the complexity of the reactors, operating costs and product quality for use in the combustion of biomass. A more in-depth discussion of the design criteria and problems associated with using biomass as a fuel source for gasification reactors can be found in recent review articles [20, 21].

In liquid-liquid extraction, a water-insoluble organic extractant is mixed with the fermenta­tion broth [222]. Because solvents are more soluble in the organic phase than in the aqueous phase, they get selectively concentrated in the extractant. Although this technique does not re­move water or nutrients from the fermentation broth, some gaseous substrates might be re­moved because CO and H2 have much higher solubility in organic solvents than water [222, 223]. Oleyl alcohol has been the extractant of choice due to its relatively non-toxicity [224].

1.6. Perstraction

Liquid-liquid extraction is associated with several problems including toxicity to the mi­crobes, formation of emulsion, and the accumulation of microbes at the extractant and fer­mentation broth interphase [222]. In an attempt to remediate these problems, perstraction was developed and this technique employs membrane to separate the extractant from the fermentation broth. This physical barrier prevent direct contact between the microbe and the toxicity of extractant, but it can also limit the rate of solvent extraction and is susceptible to fouling [219, 221]

Pervaporation is an energy-efficient alternative to distillation for removing volatile organic compounds from water, especially ABE solvents from their dilute solutions in a fermentation broth. Pervaporation is able to enrich acetone, BuOH, and EtOH with respect to water. The se­lectivity of this process is based mainly on superposition of the thermodinamical liquid-vapor selectivity, the chemical affinity selectivity, and the kinetic diffusional selectivity of the materi­als used. The liquids to be separated are not stressed in any chemical, thermal, or mechanical way. Gudernatsch et al. demonstrated the technical feasibility of the pervaporation process in continuous fermentation runs. Composite hollow fiber membranes with transmembrane fluxes in the range of 2 kg m-2 h-1 and sufficient selectivity were prepared and characterized [182]. El — Zanati et al. designed a special cell to separate the butanol from butanol/water solutions of dif­ferent butanol concentrations between 6 and 50 g L-1. The temperature of the mixture feed to the cell was 33 °C while the pressure of permeation side was about ~0 bar. Results revealed that bu­tanol concentration changes non-linearly during the first 3 h, and then proceeds linearly. The percentage of butanol removal increases with increasing feed concentration [183]. A new type of pervaporation apparatus was designed and tested by Vrana et al. to develop an integrated fermentation and product recovery process for ABE fermentation. A cross-flow membrane module able to accommodate flat sheet hydrophobic membranes was used for the experiments. Permeate vapors were collected under vacuum and condensed in a dry ice/ethanol cold trap. The apparatus containing polytetrafluoroethylene membranes was tested using butanol-water and model solutions of ABE products. Parameters such as product concentration, component effect, temperature and permeate side pressure were examined [184].

Various kinds of polymeric, ceramic, and liquid membranes can be used for selective separa­tion of solvent vapors at the temperature of fermentation. Polymeric and ceramic membranes have rather poor solvent selectivity compared to liquid membranes even though they achieve reasonable solvent mass fluxes. Liquid membranes have stability problems due to various losses. Groot et al. used silicon tubing membrane technology in the BuOH/iso-PrOH batch fermentation and the substrate conversion could be increased by simultaneous product recovery [185,186]. Geng and Park carried out fermentation by using a low acid producing C. acetobutylicum B18 and a pervaporation module with 0.17 m2 of surface area was made of silicone membrane of 240 mm thickness. During batch and fed-batch fermentation, pervapo — ration at an air flow rate of 8 L min-1 removed butanol and acetone efficiently. Butanol concentration was maintained below 4.5 g L-1 even though C. acetobutylicum B18 produced butanol steadily. With pervaporation, glucose consumption rate increased as compared to that without pervaporation, and up to 160 g L-1 of glucose was consumed during 80 h [187]. Experiments using make-up solutions showed that BuOH and acetone fluxes increased linearly with their concentration in the aqueous phase. Fickian diffusion coefficients were constants for fixed air flow rates and increased at higher sweep air flow rates. During batch and fed-batch fermentation, pervaporation at an air flow rate of 8 L/min removed BuOH and acetone efficiently. BuOH concentration was maintained at <4.5 g/L even though C. acetobu­tylicum B18 produced BuOH steadily. Cell growth was not inhibited by possible salt accu­mulation or O2 diffusion through the silicone tubing. The culture volume was maintained relatively constant during fed-batch operation because of offsetting effects of water and product removal by pervaporation and addition of nutrient supplements [188]. Fadeev et al evaluated poly[1-(trimethylsilyl)-1-propyne] (PTMSP) dense films for n-butanol recovery from ABE fermentation broth. Flux decline of a PTMSP film during pervaporation of 20 g L-1 BuOH/water mixture was linear. PTMSP films change their geometry when exposed to alcohol and alcohol/water mixtures and then dried. As a result of the relaxation process, polymer film becomes thicker and denser, effecting membrane performance. PTMSP films that were treated with 70% iso-propanol/water show linear flux decline vs. pervaporation time. Strong lipid adsorption seems to occur on the membrane surface when fermentation broth is used as a feed causing flux decline within short period of time [189].

Oya and Matsumoto used a hydrophobic polypropylene porous hollow fiber membrane of surface area 0.3 m2, porosity 45%, and bubble point 12.5 kg/cm2, under reduced pressure [190]. By Knapp et al, a vinyl-type norbornene polymer with average molar weight ~5000 was found to be useful as pervaporation membranes with separation factor of ~10 for separation of n — butanol and isobutanol [191]. Various membranes like styrene Butadiene Rubber (SBR), ethylene propylene diene rubber (EPDM), plain poly di-Me Siloxane (PDMS) and silicalite filled PDMS were studied for the removal of ABE solvents from binary aqueous mixtures and from a quaternary mixture. It was found that the overall performance of PDMS filled with 15% wt./wt. of silicalite was the best for removal of butanol in binary mixture study. SBR perform­ance was best for the quaternary mixtures studied [192].

Composite membranes containing adsorbents such as silicalite or liquid extractants such as oleyl alcohol or other solvents proved to be effective materials in ABE solvent removal from fermentation broth. Thin-film silicalite-filled silicone composite membranes were fabricated by incorporating ultrafine silicalite-1 particles, 0.1-0.2 mm. It was found that with the increase of silicalite content in the top active layer, selectivity for n-butanol and n-butanol flux in­creased, while the total flux decreased. When the silicalite-1 content was over 60%, the active layer appeared to have defects, aggregation of silicalite-1 particles, which influenced the separation factor. By controlling the membrane thickness and silicalite-1 content, membranes with total flux of 600-700 g m-2 h-1 (n-butanol flux of 300 g m-2 h-1) and selectivity of 90-100 at 70 °C using 10 g L-1 of n-butanol as feed solution were obtained. The effects of operation temperature and feed solution concentration on membrane performances were studied [193]. A membrane with a silicalite-1 (its adsorption capacity for a mixture of acetone, butanol and ethanol were 8-12, 85-90 and <5 mg g-1, respectively, there was no apparent difference in absorption rate of butanol at 36° C and 79°C and desorption of butanol occurred efficiently at 78 °C and 1-3 Torr) to polymer ratio of 1.5:1 (g:g) (306 mm thick) had butanol selectivities of 100-108 and a flux of 89 g m-2 h-1 at feed butanol concnentrations ranging from 5 to 9 g L-1 and a retentive temperature of 78°C. A 170 mm silicone membrane under identical conditions had selectivity and flux of 30 and 84 g m-2 h-1, respectively. A thin silicalite membrane offered low selectivity and high flux, while a thick membrane offered high selectivity and low flux. The effect of butanol concentration (0.37-78 g L-1) on flux and selectivity was also studied [194].

Thongsukmak and Sirkar developed a new liquid membrane-based pervaporation technique to achieve high selectivity and avoid contamination of the fermentation broth. Trioctylamine as a liquid membrane was immobilized in the pores of a hydrophobic hollow fiber substrate having a nanoporous coating on the broth side. The coated hollow fibers demonstrated high selectivity and reasonable mass fluxes of solvents in pervaporation. The selectivities of butanol, acetone, and ethanol achieved were 275, 220, and 80, respectively, with 11.0, 5.0, and 1.2 g m-2 h-1 for the mass fluxes of butanol, acetone and ethanol, respectively, at a temperature of 54 °C for a feed solution containing 1.5 wt.% butanol, 0.8 wt.% acetone, and 0.5 wt.% ethanol. Mass fluxes were increased by as much as five times with similar selectivity of solvents when an ultrathin liquid membrane was used [195]. Other long-chain trialkylamines such as tri — laurylamine or tri-decylamine could also be used as liquid membranes [196]. Acetic acid in the feed solution reduced selectivity of the solvents without reducing the solvent fluxes due to coextraction of water which increases the rate of water permeation to the vacuum side. The liquid membrane present throughout the pores of the coated substrate demonstrated excellent stability over many hours of experiment and essentially prevented the loss of liquid membrane to the feed solution and the latter’s contamination by the liquid membrane [195].

In order to exclude toxic effect of the released liquid membrane ingredient, an oleyl alcohol based liquid membrane was developed. This liquid membrane was energy efficient and did not affect microorganism growth. Oleyl alcohol liquid membrane was proved to be useful for the separation of BuOH and isobutanol in a fermentation culture with immobilized Clostridi­um isopropylicum IAM 19239 [197].

An ionic liquid (IL)-polydimethylsiloxane (PDMS) ultrafiltration membrane (pore size 60 nm) guaranteed high stability and selectivity during ABE fermentation carried out at 37 °C. Overall solvent productivity of fermentation together with continuous product removal by pervapora — tion was 2.34 g L-1 h-1. The supported ionic liquid membrane (SILM) was impregnated with 15 wt.% of a novel ionic liquid (tetrapropylammonium tetracyano-borate) and 85 wt. % of polydi — methylsiloxane. Pervaporation, accomplished with the optimized SILM, led to stable and effi­cient removal of the solvents butan-1-ol and acetone out of a C. acetobutylicum culture [198].

Reverse osmosis for recovering water from broth can also be used to concentrate ABE fer­mentation products. Polyamide membranes exhibited BuOH rejection rates <85%. Optimum rejection of BuOH occurred at a pressure of 5.5-6.5 MPa and hydraulic recoveries of 50-70%. The flux range was 0.5-1.8 L m-1 [199]. Other membranes exhibited rejection rates as high as 98% and the optimal rejection of BuOH in the ferment liquor occurred at recoveries of 20-45% with flux ranging between 0.05-0.6 L m-2 min-1 [200].

Dialysis fermentation relieves BuOH toxicity with increased yield of product, and solvent extraction can be applied to the nongrowth side of the fermentor for concentration of the BuOH. C. acetobutylicum ATCC 824 and several other strains were studied for the fermenta­tion of corn, potato, and glucose [201].

The ability of cyclodextrins to form crystalline insoluble complexes with organic components was explored as a selective separation of dilute ABE products from Clostridium fermentation systems. A product or a product mixture at a concentration of 0.150 mM each was treated with a-cyclodextrine or p-cyclodextrine in aqueous solutions or nutrient broth. In the acetone — butanol-ethanol system and in the butanol-isopropanol system, a-cyclodextrine selectively precipitated 48% and 46% butanol after 1 h agitation at 30°. However, p-CD was superior for the butyric acid-acetic acid system because it selectively precipitated 100% butyric acid under the same conditions. Cooling the three-product system with a-CD to 4° for 24 h significantly increased the precipitates but decreased the selectivity for either butanol or butyric acid [202].

Hypercrosslinked microporous ion-exchanger resins proved to be suitable agents to adsorb butanol into solid phase from fermentation broth. This ensures fermenting with a microor­ganism capable of producing butanol in a suitable fermentation medium and recovering butanol from the fermentation medium [203].

Integration of the abovementioned (Chapter 6) methods ensures new possibilities in the economic ABE solvent recovery. Some representative examples without demand of complete­ness are discussed here.

Mawasaki et al performed continuous extractive butanol fermentation with the microbe immo­bilized in gel beads and presented the recovery system of butanol from the solvent by pervapo — ration with hollow fiber membrane. This system was expected to be advantageous to prevent the fouling of membrane because butanol-oleyl alcohol mixtures obtained from extractive fer­mentation do not include solid particles [204]. Pervaporation method could also be used for in situ alcohol recovery in continuous iso-PrOH-BuOH-EtOH fermentation with immobilized cells. Fermentation was performed in a stirred tank and in a fluidized bed reactor as well. In the integrated process, the substrate consumption could be increased by a factor of 4 if compared to continuous fermentation without pervaporation product recovery. Experiments with a pilot plant plate-and-frame pervaporation module were described for the separation and dehydra­tion of alcohols. This module was also coupled to continuous BuOH fermentation, however, sterilization of the module was troublesome, and it was frequently plugged by microbial cells [205]. ABE solvents were produced in an integrated fermentation-product recovery system us­ing C. acetobutylicum and a silicalite-silicone composite membrane. Cells of C. acetobutylicum were removed from the cell culture using a 500,000 molecular weight cut-off ultrafiltration membrane and returned to the fed-batch fermentor. The ABE solvents were removed from the ultrafiltration permeate using a silicalite-silicone composite pervaporation membrane. The sili- calite-silicone composite membrane (306 mm thick) flux was constant during pervaporation of fermentation broth at the same concentration of ABE solvents. Acetone butanol selectivity was also not affected by the fermentation broth, indicating that the membrane was not fouled by the ABE fermentation broth. The silicalite-silicone composite membrane was exposed to fermenta­tion broth for 120 h. Acetic acid and ethanol did not diffuse through the silicalite-silicone com­posite membrane at low concentrations. The fed-batch reactor was operated for 870 h. Totally 154.97 g L-1 solvents was produced at solvent yield of 0.31-0.35 [206].

Application of membrane-assisted extraction to butanol fermentation was investigated as a means of product separation and also as a way of alleviating the problems concerning the end — product inhibition. The coupled reactor-separator system was stable enough to sustain continu­ous operation lasting several weeks. The data on continuous run reaffirmed most of the advantages found in a previous study on fed-batch system in that the reactor separator system rendered high productivity and yields due primarily to reduced product inhibition. Improve­ment in productivity was particularly notable, as a fourfold increase over straight batch opera­tion was obtained. In normal continuous operation, spontaneous cell deactivation occurred after 200-400 h of operation despite the removal of inhibitory products. The presence of autoly — sin was one of the probable causes of cell deactivation. The cell viability, however, was pro­longed significantly when the bioreactor was operated under glucose-limited conditions [207].

A calcium alginate-immobilized continuous culture was used in a novel gas-sparged reactor to strip the solvents from the aqueous phase and reduce their toxicity. A dilution rate of 0.07 h-1 was found to give maximal solvent productivity at 0.58 g dm-3 h-1, although at 0.12 h-1 the productivity was slightly lower. In order to increase glucose uptake by the culture, feed glucose concentration was increased over time to attempt to acclimatize the culture. This resulted in productivity as high as 0.72 g dm-3 h-1 although this production rate was unstable [208].

An extractive acetone-BuOH fermentation process was developed by integrating bioproduc­tion, ultrafiltration, and distillation, providing simultaneous retention of biomass, selective removal of inhibitors from permeate and separation and purification of acetone, BuOH, and EtOH. Successive batch fermentations were performed with normal pressure distillation (98°) which permitted prolonging and enhancing (by a factor of 3) solvent production with very few volume exchanges of medium (average dilution rate was 0.002 h-1), and recovering the concentrated solvents online. Different operating conditions were also tested in order to study the presence of extracellular autolytic enzymes as inhibition factors. Extracellular autolytic activity was low most of the time, even without enzyme-inactivating heat treatment in the distillation boiler, and high-temperature distillation was deleterious to the culture medium. Improvements of the process were achieved, first, by managing continuous runs, providing a minimal renewal of the culture medium and, mainly, by decreasing the temperature and pressure of distillation. Solvent productivity reached 2.6 g L-1 h-1 for a 0.036 h-1 average dilution rate, corresponding to a feed concentration of 156 g L-1 glucose actually consumed [209].

Continuous extractive bioconversion processes were described for conversion of native starch granules to ABE solvent production using a selective adsorbent. In fermentation of carbohy­drates with C. acetobutylicum selective synthetic zeolite or crosslinked divinylbenzene — styrene copolymer sorbents are integrated in the process to adsorb the products from the medium continuously [210]. The conversion of glucose to ABE solvents by C. acetobutylicum employing extractive fermentation by using a combination of membrane technology and solid adsorbents integrated into the fermentation process was studied. The adsorbent used was a nitrated divinylbenzene-styrene copolymer. Its ability to adsorb fermentation broth constitu­ents was as follows: BuOH 82, EtOH 36, Me2CO 51, butyric acid 99, and AcOH 21 mg/g sorbent. The polymer was then heat treated to release the bound solvents. In a long term experiment using an adsorption column, 400 g glucose was added successively to the column and fermentation allowed to proceed for 320 h. A total amount of 67 g of solvent was recovered by heating 930 g polymer [211]. It was found that the in situ adsorption process using polyvinyl — pyridine as the adsorbent enhanced the fermentation rates and the reactor productivity by C. acetobutylicum. In typical traditional acetone-butanol fermentation process only about 60 g/L of glucose could be used in a batch operation mode and thus, at maximum only 21 g L-1 of the total final products concentration could be achieved. In the adsorption-coupled system an initial glucose concentration of 94 g L-1 was fermented when a weight ratio of the adsorbent to the fermentation broth of 3/10 was used. An overall product concentration of 29.8 g L-1 and a productivity of 0.92 g L-1 h-1 were achieved in the adsorptive batch fermentation system. Compared with the controlled traditional batch acetone butanol fermentation, the integrated process increased the final product concentration by 54% and the productivity by 130% [212]. Integration of a repeated fed-batch fermentation (C. acetobutylicum) with continuous product removal (poly(vinylpyridine) adsorption) and cell recycling resulted in inhibitory product concentration reduction. Because of the reduced inhibition effect, a higher specific cell growth rate and thus a higher product formation rate were achieved. The cell recycle using membrane separation increased the total cell mass density and, therefore, enhanced the reactor produc­tivity. The repeated fed-batch operation overcame the drawbacks typically associated with a batch operation such as down times, long lag period, and the limitation on the maximum initial substrate concentration allowed due to the substrate inhibition. Unlike a continuous operation, the repeated fed-batch operation could be maintained for a long time at a relatively higher substrate concentration without sacrificing the substrate loss in the effluent. As a result, the integrated process reached 47.2 g L-1 in the equivalent solvent concentration (including acetone, BuOH, and EtOH) and 1.69 g L-1 h-1 in the fermentor productivity, on average, over a 239.5-h period. Compared with controlled traditional batch acetone-BuOH fermentation, the equivalent solvent concentration and the fermentor productivity were increased by 140% and 320%, respectively [213].

Cells of C. acetobutylicum were immobilized by adsorption onto bonechar and used in a packed bed or fluidized bed reactor for the continuous production of ABE solvents from whey permeate. At dilution rates in of 0.35-1.0 h-1, ABE solvent productivities of 3.0 to 4.0 g L-1 h-1 were observed, but lactose utilization values were poor. When operated in an integrated system with product removal by liquid-liquid extraction, there was a decrease in productivity, but lactose utilization was increased markedly. Of the three extractants tested, oleyl alcohol proved to be superior to both benzyl benzoate and dibutyl phthalate [214].

Shah and Lee studied simultaneous saccharification and extractive fermentation (SSEF) to produce ABE solvents from aspen tree. In SSEF employing cellulase enzymes and C. aceto­butylicum, both glucan and xylan fractions of pretreated aspen are concurrently converted into acetone and butanol. Continuous removal of fermentation products from the bioreactor by extraction allowed long-term fed-batch operation. The use of membrane extraction prevented the problems of phase separation and extractant loss. Increase in substrate feeding as well as reduction of nutrient supply was found to be beneficial in suppressing the acid production, thereby improving the solvent yield. Because of prolonged low growth conditions prevalent in the fed-batch operation, the butanol-to-acetone ratio in the product was signifi­cantly higher at 2.6-2.8 compared to the typical value of two [215]. Integrated bioreactor — extractor was also tested in SSEF and production of ABE solvents from pretreated hardwood by C. acetobutylicum and cellulase enzymes. The SSEF system was constructed so that products of fermentation were extracted from the broth through a semipermeable membrane. In situ removal of inhibitory products was found to be beneficial in sustaining cell viability, thus allowing fed-batch operation of the bioreactor over a period of several weeks. Hardwood chips were pretreated by monoethanolamine in such a way that hemicellulose and cellulose were retained in high yield. The feed material thus prepd. was readily converted by SSEF. The ability of C. acetobutylicum to ferment both glucose and xylose was a major factor in simpli­fying the overall process into a single-stage operation [216].

HDO is a process closely related to hydrodesulphurization (HDS), which is highly devel­oped in the oil-refinery industry. In both processes, hydrogen is used to remove the heteroa­tom in the form of H2O and H2S respectively. This is the reason why several works on bio-oil HDO use catalytic systems already used in HDS processes, such as Co-Mo or Ni-Mo based catalysts. These catalysts are active in their sulphide form, so they need to be pretreated with H2S before operation to obtain Co-MoS2 or Ni-MoS2 active sites. Romero et al. [18] us­ing Co-MoS2 type catalysts for the HDO of 2-ethylphenol at 340°C and 7 MPa of hydrogen pressure proposed the reaction mechanism described in Figure 4. It is suggested that the oxygen from the molecule adsorbs on a vacancy of a MoS2 matrix. At the same time, the H2 from the feed dissociatively adsorbs on the catalyst surface forming S-H species. The addi­tion of a proton to the adsorbed oxygenated molecule leads to an adsorbed carbocation. This intermediate can directly undergo a C-O bond cleavage and the aromatic ring is regenerat­ed leading to ethylbenzene. The vacancy is afterwards recovered by elimination of water.

The problem of using MoS2 type catalysts for HDO of bio-oils is that during prolonged oper­ation sulfur stripping and oxidation of the surface of the catalyst occurs, causing deactiva­tion of the catalyst. The reason is that as compared to conventional oil, the sulfur content of bio-oil is very low (less than 0.1 wt % [19]). One alternative to avoid this problem is the co­feeding of H2S to the system, in order to regenerate the sulfide sites. For instance, in the HDO of alyphatic esters over a CoMoS2/Al2O3 and NiMoS2/Al2O3 catalysts a promoting ef­fect was observed in the activity of the catalyst when co-feeding H2S, however this co-feed­ing did not prevent from catalyst deactivation. This promoting effect was related to the increase in Bronsted acidity in the presence of H2S [20]. Nonetheless, the use of H2S has also some drawbacks. In the HDO of phenol over a Ni-MoS2-Al2O3 catalyst, it was observed an inhibitory effect of H2S, leading to a decrease in phenol conversion and not preventing cata­lyst deactivation. This was ascribed to the competitive adsorption between phenol and H2S [21]. Moreover, the formation of sulfur-containing compounds such as dimethyl sulfide, di — heptyl sulfide, hexanethiol and heptanethiol was observed in the HDO of aliphatic oxygen­ates over Co-MoS2 catalysts, even in the absence of sulfiding agents [22]. Therefore, the use of MoS2 type catalysts in bio-oil HDO seems challenging, becouse sulfur free bio-oil can be contaminated by sulfur, and because wood-based bio-oils contain high amounts of phenolic compounds that would compete with H2S for the active sites of the catalyst.

Another alternative is the use of bi-functional catalysts formed by the combination of transi­tion metals and oxophilic metals, such as MoO3, Cr2O3,WO3 or ZrO2. In this case, the oxo — philic metal acts as a Lewis acid site. The oxygen ion pair of the target molecule is attracted by the unsaturated oxophilic metal. The second step of the mechanism is hydrogen dona­tion. In this case, the hydrogen molecule is dissociatively adsorbed and activated on the transition metal. Finally, the activated hydrogen is transferred to the adsorbed molecule.

Regarding the support, y-Al2O3 is the most commonly used one. Nonetheless, it has to be taken into account the structural changes that y-Al2O3 might suffer under the typical oper­ating conditions in HDO. In contact with hot water (T > 350°C), y-Al2O3 is converted into a hydrated boehmite (AlOOH) phase with a significant decrease in the acidity and sur­face area [23]. Moreover, the relatively high surface acidity of Al2O3 is thought to pro­mote the formation of coke precursors. In fact, coke formation is one the main factors affecting the stability of the catalyst. Therefore, the use of less acidic or neutral support like active carbon or SiO2 is an interesting alternative [24]. For instance, Echeandia et al. [25] using Ni-WO3 on active carbon for the HDO of 1 wt% phenol in n-octane at 150-300°C and 15 bar observed lower coke formation on the surface of the active carbon with re­spect to alumina support. Based on product analysis, they also concluded that HDO of phenol occurs via two separate pathways: one leading to aromatics through a direct hy — drogenolysis route, and the other one to cyclohexane, through a hydrogenation-hydroge — nolysis route (see Figure 5). In terms of obtaining a final product with high octane number and reducing the consumption of hydrogen, direct hydrogenolysis reaction is preferred. Nonetheless, aromatics are harmful to human health and its content in transportation fuels is limited by legislation. Therefore, it is important to understand which sites are responsi­ble of each route, in order to obtain an upgraded product with the desired aromatic con­tent. CeO2 and ZrO2 supports have also shown to give good results in the HDO of different molecules. ZrO2-supported noble metal catalysts (Rh, Pd and Pt) [26] were com­pared with the conventional sulfided CoMo/Al2O3 catalyst in the HDO of Guaiacol in the presence of H2 at 300 °C. Sulfided CoMo/Al2O3 deactivated due to carbon deposition, and the products were contaminated with sulfur, however, neither problem was observed with the ZrO2-supported noble metal catalysts. As a conclusion, a good support for HDO should provide high affinity for the oxygen-containing molecule while presenting moder­ate acidity in order to minimize the formation of coke deposits.

Hydrogenoiysis /Nv

Benzene

Hydrogenation

Hvdrngenolvsis

Koine

Cvelohexanol

Figure 5. Scheme of phenol HDO. Adapted from [25].

1.1. Upgrading of real bio-oils

An important aspect in the HDO of bio-oils is the required degree of deoxygenation. It is assumed that the upgraded oil should contain less than 5 wt% oxygen so that the viscosi­ty is decreased to that required for fuel applications [17]. However, during the hydrotreat­ing, not only the oxygen is removed in the form of water, but also the saturation of double bounds occurs. This saturation has two significant negative effects. The first one is relat­ed to the quality of the upgraded oil, because the saturation of the aromatic components has a highly detrimental effect in the octane number. For instance, the octane number of toluene (119) decreases to 73 when the aromatic ring is hydrogenated [10]. The second negative effect is related to the consumption of hydrogen. According to Venderbosh et al. [27] in order to achieve 50% of deoxygenation 16 g H2/Kg of bio-oil is required, which is close to the expected stoichiometry value. Nonetheless, if the aim is to obtain the total re­moval of oxygen, the H2 consumption increases to 50 g H2/Kg of bio-oil; which means that the H2 consumption is 56% higher than the stoichiometry value. Some other studies sug­gest even higher H2 consumption requirements, 62 g H2/Kg of bio-oil [28]. This deviation of the H2 consumption from the stoichiometry value is explained on the basis of the differ­ent reactivity of the oxygenated compounds present in the bio-oil. High reactive com­pounds, such as ketones, are easily converted with low hydrogen consumption. However, more complex molecules, such as phenols, might suffer the hydrogenation/saturation of the molecule and therefore the hydrogen consumption exceeds the stoichiometric predic­tion at the high degree of deoxygenation.

In order to obtain high degrees of HDO but minimizing the hydrogenation of aromatics in bio-oil, two step hydrogenating processes have been developed. In the first stage, high reac­
tive and unstable compounds are transformed into more stable ones at low temperature (270°C, 136 atm H2) and without a catalyst. In the second step, a deeper HDO is carried out at higher temperatures (400°C, 136 atm H2) and using hydrotreating catalysts. The two-step hydrotreatment allows 13% reduction in hydrogen consumption for equivalent oil yield. Nonetheless, the reported octane number of the upgraded bio-oil, 72, is still lower than that of gasoline [17].

Environmental aspects should also be taken into account. Aromatic compounds have on one hand high octane number; however, they are also harmful to health. Indeed, environmental standards for aromatics in transportation fuels are becoming more restrictive. Thus, it seems challenging to achieve an agreement between obtaining oils with high octane number while fulfilling aromatic content policies.

Edilson Leon Moreno Cardenas,

Deisy Juliana Cano Quintero and Cortes Mann Elkin Alonso

Additional information is available at the end of the chapter http://dx. doi. org/10.5772/53351

1. Introduction

1.1. The trouble of organics solids wastes

In the protection of environment, the adequate handling of solids wastes occupy a main place, the integral handling of wastes is a term applied to all activities associated with the wastes management in the society. The main aim is the administration of wastes associated with the environment and public health. The handling of solid wastes is one of the main environmental problems in the cities due to its generations increase simultaneously with the growth of the cities, its industrialization and the increase of population. In addition, the actual life style carries out a high demand of consumption of goods that generally are thrown out in a short time; this generates more production of wastes and therefor having to search for solutions to the final disposition.

A solution for the trouble of the urban solids wastes is the implementation of process of reusing and giving value to the different materials that form what is known as "garbage", with the purpose of obtaining products or sub products that can be to introduce into new economic cycles. The maximization of reusing and giving value to solids wastes carry out benefits as: less consumption of natural sources, reduction of energy consumption, less environmental pollution, better use of the location where the garbage is placed and economic benefits from recovered materials. So the changes of consumption patron and the sustainable production are essential for the reduction of wastes production.

Is very difficult to stop the production of solids wastes, the idea is consider the solids wastes as a source of material reusable, raw matter, organics nutrients, biofuels and energetics fuel. The set of process to recover and treatment the wastes are known as valorization of solids wastes. This production of wastes is due to origin, social context and production activities [1]. During the valorization and reusing of wastes, is necessary take account aspect as recollection and transport, with this is possible to obtain highs benefices by the transformation. Addition­ally is necessary to include applications of new concepts related to the financial services, decentralized management, community contribution and the options of transformation, valorization and incorporation to economic cycles [2].

ABE fermentation can be conducted as batch, fed-batch, and continuous under anaerobic conditions. Batch fermentation is the simplest mode. The substrate is typical 40-80g/L and the efficiency decreased as substrate concentration upper than 80g/L (Shaheen et al, 2000). With optimized physiological and nutritional parameters, 20g/L n-butanol was obtained by C. beijerinckii ATCC 10132 in 72h (Isar and Rangaswamy, 2012). Fed-batch fermentation was adopted to avoid substrate inhibition. However, because of product inhibition, the substrate feeding seems ineffective. The solvent must be removed from the broth to decrease the product toxicity. The solvent can be removed by several ways such as liquid-liquid extrac­tion, perstraction, gas-stripping, and pervaporation etc. (Qureshi and Maddox, 1995; Qure- shi and Blaschek, 2001b). The whole systemic technique of high productivity was constructed by continuous feeding combined with product removal (Qureshi et al., 1992), such as using membrane reactor (Qureshi et al., 1999a). With these techniques, the fermenta­tion can be continuing for a long time and resulting in higher productivity. To improve the utilization efficiency of cells, the immobilization system is used (Huang et al., 2004; Qureshi et al., 2000; Lienhardt et al., 2002). Comparing with the free cell system, the immobilization system is easier to separate cells from product, can reach high cell concentration and pro­ductivity, and can decrease nutrient depletion and product inhibition.

Co-culture is another important way for butanol fermentation (Abd-Alla and El-Enany, 2012). C. beijerinckii NCIMB 8052 was entangled with ATCC 824 and thought as C. acetobuty — licum before the 16S rDNA based method was exploited (Johnson and Chen, 1995). These data implied that they could be cocultured before isolation. A microflora of four strain iso­lated from hydrogen-forming sludge of sewage performed a little high solvent yield (Cheng et al., 2012). Different strains possess various advantages, either with larger carbon sub­strate, higher butanol yield, or with high substrate and product tolerance. The co-culture should possess potential benefits and be harnessed fully after all the details are disclosed for each individual strain.

Algae are predicted to have first appeared approximately 1.5 billion years ago from an endosymbiotic event in which a eukaryotic cell engulfed a cyanobacterium [98]. The cyano­bacterium evolved into the modern day chloroplast, the algal organelle responsible for photosynthesis and carbon fixation. Today, algae can be found in a wide-range of environ­mental habitats from freshwater lakes and oceans to deserts and even the snow of the Antarctic

[99]. Along with this diversity of habitat, algae have evolved diverse cellular physiologies and genetics, resulting in a wealth of potential hosts and genetic sources for engineering fuel production. Many types of algae are currently under consideration for fuel production due to their natural TAG synthesis, including diatoms, green algae, eustigmatophytes, prymnesio — phytes, and red algae [100]. While many types of algae produce the fuel precursor TAG, few algal species have well-developed genetic tools available for engineering improved lipid production [101, 102]. Consequently, there are only a few reported examples of engineering algae for biofuel production.

To date, the only genetic mutation shown to improve lipid production in algae is the elimina­tion of starch biosynthesis, a competing carbon sink. The generation of mutants with impaired starch synthesis using random mutagenesis techniques resulted in up to a 10-fold increase in cellular lipid production in C. reinhardtii [56—58, 103]. Other targeted metabolic engineering attempts, such as overexpression of ACC in the diatoms Cyclotella cryptic and Navicula saprophila, failed to improve TAG biosynthesis [15, 96]. In addition to targeting overall TAG production, metabolic engineering strategies have been applied to influence the chemical composition of the fatty acid side chains. By expressing two heterologous TEs, the diatom Phaeodactylum tricomutum produced TAG with increased levels of lauric acid (C12:0) and myristic acid (C14:0) [104]. These shorter chain length fatty acids are more desirable for fuel production, and this demonstrates the potential to control the chemical composition of the fuel product and its associated properties with metabolic engineering. While examples of engi­neering algal TAG production are sparse, many engineering strategies have proven successful at improving the fatty acid content in plants. These strategies include expression of ACC and KASHI involved in fatty acid biosynthesis, expression of G3P dehydrogenase (GPD) for production of the glycerol backbone of TAG, expression of ATs such as DGAT, expression of TEs to release FFAs, and deletion of desaturases to alter the fatty acid composition [105]. Similar strategies may also be successful at improving TAG production in algae.

The metabolic engineering of algae is complicated by several factors. Most algae have a rigid cell wall structure that makes transformation difficult. A common transformation technique uses glass beads (or silicon carbide whiskers) along with a cell wall-deficient algal strain [106].

The cell wall can be removed using enzymatic techniques or through genetic mutation. Alternatively, a microparticle bombardment technique has been applied successfully to transform many different algal species [107]. In this technique, the recombinant DNA is coated onto a metal microparticle and ‘shot’ into the algal cell using a helium-powered ‘gun’. Other transformation methods include electroporation and the traditional plant transformation technique of Agrobacterium tumefaciens T-DNA-mediated transfer [107]. Once the recombinant DNA enters the cell, it must integrate into one of 3 algal genomes: nuclear, chloroplast, or mitochondrial (assuming the transformed DNA is not a stably maintained plasmid). DNA has been successfully integrated into the chloroplast genome via homologous recombination, whereby the recombinant gene and marker are flanked by homologous (i. e. matching) regions of the targeted chloroplast DNA, and the recombinant DNA replaces the matching region in the chloroplast. Unfortunately, homologous recombination does not occur in the nuclear genomes of many algae [108], and instead, the recombinant DNA is randomly integrated into the nuclear genome. This complicates metabolic engineering strategies due to the possibility of detrimental genetic effects resulting from the random integration and the lack of a technique for targeted gene knockout. Lastly, algal engineering attempts are often plagued by low gene expression. It has been discovered that many algae, like the model alga C. reinhardtii, employ RNA-mediated gene silencing [109]. Numerous strategies have been applied to combat the low gene expression brought about by gene silencing in algae, including codon optimization, the use of 5′ and 3′ untranslated regions which may participate in regulatory functions, and the inclusion of native intron sequences [108]. Knowledge of the gene silencing mechanisms in algae has led to the development of RNA interference (RNAi) technology for gene knock­down. RNAi exploits the native cellular machinery for gene silencing to reduce the expression of target genes [109]. As we continue to expand our knowledge of algal genetics, the list of engineered algae will rapidly increase. As evidence, the biofuel-relevant alga, Nannochlorop — sis sp., was recently shown to have a high efficiency of homologous recombination in the nuclear genome [110]. This will simplify future strategies for genetic engineering in Nanno — chloropsis sp. Another promising development is the construction of a plasmid for gene expression in C. reinhardtii that is now commercially available through Life Technologies [111]. The greater availability and standardization of tools for the genetic manipulation of algae will move algal engineering towards the advanced stages currently seen with other industrial organisms like E. coli and S. cerevisiae.

Cu has been extensively investigated in the glycerol hydrogenolysis reactions. Although its hydrogenation activity is generally lower than that of noble metals, its much lower price and its ability to catalyze C-O bond but not C-C bond hydrogenolysis make Cu catalysts at­tractive for this process. There are some works in the literature that report the use of other transition metals like Ni or Co, however, Cu based catalysts are predominant. Vapor phase glycerol dehydration reaction was studied by Sato et al. [69] over different copper catalysts at 513 K and atmospheric N2 pressure. They observed that basic MgO, CeO2 and ZnO sup­ports showed low acetol selectivity, while acidic supports, such as Al2O3, ZrO2, Fe2O3, and SiO2, effectively promoted acetol formation. The best results were obtained with Cu/Al2O3 catalyst. Increments in copper content lead to increments in acetol selectivity. Moreover, the activity of the Al2O3 support alone was rather low, which indicates that copper metal sites play a significant role in glycerol dehydration. Continuing with vapor phase processes, Akiyama et al. [64,70] studied glycerol hydrogenolysis in a fixed-bed down-flow glass reac­tor at temperatures between 340 and 473 K, atmospheric hydrogen pressure, and using Cu/Al2O3 catalysts. In the two step reaction they observed that glycerol dehydration to ace — tol was favored at relatively high temperatures. However, acetol hydrogenation to 1,2-PDO was favored at lower temperatures, because it is an exothermic reaction and the dehydro­genation of 1,2-PDO occurs preferentially at high temperatures. Based on these findings, they developed a reactor with gradient temperatures, at the top of the reactor glycerol dehy­dration reaction occurred at 453 K while at the bottom of the reactor acetol was hydrogenat­ed to 1,2-PDO at 418 K. Really high 1,2-PDO yields (94.9%) were reported.

Some of the best results in terms of glycerol conversion and 1,2-PDO selectivity were recent­ly reported using Cu on base supports. For instance, Yuan et al. [44] developed a Cu based solid catalyst (Cua4/Mg5.6Al2O86)via thermal decomposition of the as-synthesized Cua4Mg56Al2(OH)16CO3 layered double hydroxides. This bifunctional highly dispersed Cu — solid base catalyst was effective for hydrogenolysis of aqueous glycerol. The measured con­version of glycerol reached 80.0% with a 98.2% selectivity of 1,2-propanediol at 180 °C, 30 bar H2 and 20 h. The addition of Pd to the same catalytic system notably increased the activi­ty of the catalyst [71]. It was suggested that the hydrogen spill over from Pd to Cu favored glycerol hydrogenolysis to 1,2-PDO.

In the model of calculations was assumed that the discharge is homogeneous over the entire volume. It is justified at zero approximation, because the time of gas mixing in the radial direction is less than the times of characteristic chemical reactions. Also we neglect the processes in the transitive zone between the discharge to post-discharge. Thus, the time of gas
pumping through the transition region is too short for the chemical reactions to have a sufficient influence on the concentration of neutral components.

The total time of calculation is divided into two time intervals: the first one is the calculation of the kinetic processes of fast generation of active atoms and radicals in the discharge region. Those components accelerate the formation of molecular hydrogen, carbon oxides and production of other hydrocarbons. The second time interval is the oxidation of the gas mixture in the post-discharge region as a result of the high gas temperature and the presence of O and OH. These components remain in the mixture after the dissociation of water and oxygen molecules by electron impacts in the plasma. The oxidation of generated hydrocarbons has a noticeable influence on kinetics in the investigated mixture due to the high gas temperature.

Under the aforementioned conditions, the characteristic time of oxidation is approximately equal to the air pumping time through the discharge region (~10-3-10-2 s). The following system of kinetic equations is used in order to account for the constant air pumping through the system:

Nir Nj, Ni in the equation (3) are the concentrations of molecules and radicals; kj kM are the rate constants of the processes for the i-th component. The rates of electron-molecule reactions Sei are connected with discharge power and discharge volume. The last three terms in equation (1) describe the constant inflow and outflow of gas from the discharge region. The term Ki is the inflow of molecules of the primary components (nitrogen, oxygen, carbon dioxide, water and ethanol) into the plasma, G/Vy and kNt are the gas outflow as the result of air pumping and the pressure difference between the discharge region and the atmosphere. In order to define the initial conditions, the ethanol/water solution is assumed to be an ideal solution. Therefore, the vapor concentrations are linear functions of the ethanol-to-water ratio in the liquid. The evaporation rates K of C2H5OH and H2O are calculated from the measured liquids’ consumption. The inflow rates Ki of nitrogen and oxygen are calculated by the rate of air pumping through the discharge region:

G

Ki = — N° (4)

i V i

where N°i correspond to [N2] and [O2] in the atmospheric pressure air flow.

The gas temperature in the discharge region is taken to be constant in the model. In reality, the gas temperature T is dependent on the gas pumping rate and the heat exchange with the environment. Therefore, in order to take into account those influences, T is varied in the interval 800-2500K (similarly to the experimentally obtained temperature spread). After ~10-2 s, the balance between the generation and decomposition of the components leads to saturation

of concentrations of all species. This allows us to stop the calculations in the discharge region and to investigate the kinetics in the post-discharge region. System (3) is solved without accounting for the last three terms on the time interval without the plasma. The calculations are terminated when the molecular oxygen concentration reaches zero level.

The full mechanism developed for this experimental work is composed of 30 components and 130 chemical reactions between them and its closed to [11]. The charged particles (electrons and ions) are ignored in the mechanism, because of low degree of ionization of the gas (~ 10-6 — 10-5). Nitrogen acts as the third body in the recombination and thermal dissociation reactions.

In the non-equilibrium plasma almost the entire energy is deposited into the electron compo­nent. The active species, generated in the electron-molecular processes, lead to chain reactions with ethanol molecules.

Numerical simulation of kinetics showed that the main channels of H2 generation in the plasma were ethanol abstraction for the first 10-100^s, and hydrocarbon abstraction afterwards. Additionally, the conditions when the reaction between H2O and hydrogen atoms was the main channel of H2 production were found. A kinetic mechanism, which adequately described the chemistry of the main components, was proposed. The model did not account for nitrogen — containing species, and nitrogen was considered only as a third body in recombination and dissociation reactions. The comparison between experiments and calculations showed that the mechanism can adequately describe the concentrations of the main components (H2, CO, CO2,

CH4, C2H4, C2H6, and C2H2).

T=2023 К

t—— ‘—— 1—- 1—— 1—- ‘—— 1—— ‘—— 1—— ‘—— 1—- ‘—— r

6 8 10 12 14 16 18

Rate of C02 pamping, cm3/s

Figure 11. The dependence of the reaction main products of the flow rate of CO2 (inside discharge), T = 2023 K

However, it should be noted that with the increase in temperature to 2523 K leads to the fact that the output of the reactor is not observed almost no light hydrocarbons. They simply "fall apart” and burned. That leaves the most stable elements such as H2O, N2, CO2. This sug­gests that the increase in temperature up to these values is not advisable because of the de­crease in the yield of useful products (see Fig. 11 and Fig. 12a, b).

These calculations are based in good correspondence with the experimental data (see Fig. 8).

Other model hydrocarbon is bioglycerol (crude glycerol) which is a byproduct of the biodiesel manufacture. Biodiesel is a popular alternative fuel. It is carbon neutral, has emissions equivalent or below diesel, is biodegradable, non-toxic, and is significantly cheaper to manufacture than its petroleum equivalent. However there is one significant drawback: for every 10 gallons of biodiesel produced, roughly 1 gallon of bioglycerol is created as a byproduct.

Biodiesel is produced by mixing vegetable oil and potassium hydroxide KOH. Therefore, the large-scale production of environmentally friendly and renewable fuel may lead to possible bioglycerol accumulation in large quantities, which, in turn, can cause environmental prob­lems, as it is comparably bad fuel. In addition, it has a rather large viscosity of 1.49 Pa^s, which is larger for almost three orders of magnitude than ethanol and water viscosity. The solution to this problem would be "TORNADO-LE" usage for bioglycerol reforming. Pure glycerol chemical formula is СзНз^Н)з. However, bioglycerol contains various impurities (including a set of alkali).

Fig. 13 shows a photograph of burning discharge, where the working liquid is bioglycerol and working gas — air. Research is conducted by the SC polarity, because this mode has lowest liquid consumption.

Fig. 14 shows the typical emission spectrum of the plasma discharge in a "TORNADO-LE" where the working liquid is bioglycerol doped with alkali. It is registered at a current of 300 mA, voltage — 2 kV, air flow — 110 cm3/s. Optical fiber is oriented on the sight line, parallel to the liquid surface in the middle of the discharge gap. The distance from the liquid surface to the top flange equals 10 mm.

Emission spectrum (Fig. 14) is normalized to the maximum Na doublet (588.99 nm, 589.59 nm). It contains K (404.41 nm, 404.72 nm, 766.49 nm, 769.89 nm), Na (588.99 nm, 589.59 nm), Ca (422.6 nm) lines, and a part of continuous spectrum, which indicates that the there’s a soot in the discharge. Temperature, which is defined by the plasma continuous emission spectrum is 2700 ± 100 K.

The K, Na, Ca elements presence in the discharge gap complicates the plasma kinetics numeric modeling of the bioglycerol reform process. The gas flow rate at the system outlet is 190 cm3/s, i. e. by 80 cm3/s larger than the initial (110 cm3/s), which indicates bioglycerol reforming to the syn-gas. Liquid flow is 5 ml/min. Change of the CO2 share in the working gas weakly affects the spectrum appearance.

Based on the continuous nature of the plasma emission spectra, we compared the experimental results with the calculated spectra of the blackbody radiation. Calculations have been per­formed by using Planck’s formula.

Fig. 15 shows the computational grid with step of 200-300 K in the temperature range from 2500 K to 3500 K and the plasma emission spectrum in the case of bioglycerol, as a working fluid (air flow — 82.5 cm3/s, the flow of CO2 — 17 cm3/s, CO2/Air = 1/5, Id = 300 mA, U = 600 V). All spectra are normalized to the intensity, which is located at a wavelength of 710 nm.

The data in Fig. 15 show that the plasma emission spectrum coincides with the calculated by the Planck formula for the temperature T = 2800 ± 200 K. Since bioglycerol contains alkali metals, which represent an aggressive environment, the gas chromatography can’t be used. Therefore, in order to determine the gas composition, formed the bioglycerol reformation IR and mass spectrometry have been used.

With infrared transmission spectra one can see that the transition to bioglycerol increases the amount of such components as CO2 (2250-2400 cm-1), CO (2000-2250 cm-1), CH4 (3025-3200 cm-1), C2H2 (3200-3350 cm-1).