17.2.1 Introduction

As discussed in the introduction alcohols are produced via the synthesis gas route. The crucial step here is not the synthesis gas production via gasification but the conversion of synthesis gas into alcohols. For methanol this is a common practice although this is not straightforward: high pressure (50-100 bars), the right composition and purity of CO and H2 is necessary. For ethanol a number of synthesis routes have been proposed. Table 17.1 shows possible routes for ethanol production based on synthesis gas routes. Two routes can be distinguished: direct and indirect synthesis via intermediate products. The direct route would be the most interesting route from a molecular efficiency point of view, but the present catalytic systems (Cu-Zn-Co oxides/halides, Rh(CO)12 on La2O3) only produce ethanol with relatively low yields (28-49 wt%). This is why the indirect route for ethanol production is still the most used one. Of this last method there are again different kind of options among which are homologation (= chain extention) and a sequence of carbonylation, esterification and hydrogenation.

17.2.2 Hydrocarbonylation van methanol to ethanol

The hydrocarbonylation is a synthesis gas based route via which extention of the alcohol with a CH2 group is carried out; in this case the extension of methanol to ethanol:

CH3OH + CO + 2 H2 « C2H5OH + H2O

By carrying out the reaction in tetrahydrofuran as solvent, good yields can be achieved (89%). The catalysts used in this process are based on metals like Cu, Fe and Ni.

20.2.1 Reforming of fossil feedstock

Gasification (see next Section 20.2.2 for further information) is used for heavy feedstocks like heavy fuel oil and coal while reforming is used for lighter feedstocks, namely natural gas, associated gas, and naphtha. Steam (and dry) reforming are catalytic processes where commercially nickel on alumina-based catalysts are used. To reach high enough conversions to synthesis gas, the temperature of reactor outlet is approximately 850-1000°C at operating pressures up to 30 bar. The inlet temperature can be significantly lower (~500°C). Excess steam, when compared to the stoichiometric reactions, is used to steer the equilibrium conversion of methane and avoid kinetic carbon deposition. Four essentially different steam reforming processes can be identified, namely:

(1) One-step or single-step steam reforming: An externally fired tubular reactor is being used over which a temperature gradient is applied to reach full conversion.

(2) Two-step reforming: Initial pre-reforming of the heavier components in the feedstock towards methane and carbon oxides is done around 350-550°C,7 which is then followed by a consecutive (higher temperature) reformer.

(3) Auto-thermal reforming: Next to steam, oxygen is supplied internally to generate heat for the strongly endothermic reforming reactions.

(4) Partial (catalytic) oxidation: Part of the feedstock is combusted without adding additional steam which generates very low hydrogen over carbon monoxide ratios (<2).

For biomass, the pre-reform reaction looks like:

C6H10O5 + H2O ^ 3CO2 + 3CH4 [20.7]

Since steam reforming catalysts and catalysts used in downstream processes (like for methanol and Fischer-Tropsch production) are very sensitive towards sulfur poisoning, the feedstocks are desulfurized. For a complete and detailed review on steam (and dry) reforming of fossil feedstocks, the reader is referred to Rostrup-Nielsen et al.8

20.2.2 Gasification of fossil feedstock

Already around 1850 there was a considerable coal gasification industry. The Siemens gasifier (1861) and the Winkler gasifier (1926) were successful low — temperature (<900°C) air blown systems producing fuel gas. In 1938, the Koppers — Totzek entrained flow gasifier came into commercial use. This gasifier produced synthesis gas (CO + H2) on continuous basis containing no tars and methane at approximately 1850°C and atmospheric pressure from oxygen-entrained coal. At the end of the 1940s and the early 1950s, Texaco and Shell developed technologies for the production of the synthesis gas by oil gasification. These were entrained — flow reactors with top-mounted burners (atomizers) in the down-flow. Operating pressures and temperatures were up to 80 bar and in the range of 1250-1500°C, respectively. Apart from Texaco and Shell, Lurgi also developed oil gasification technology, known as multi-purpose gasification. Nowadays, most oil gasifiers are part of a refinery and are used for poly-generation of power, H2, synthesis gas and steam. As a result of the oil crisis of the early 1970s coal gasification was taken up again. It was again Texaco and Shell (together with Krupp-Koppers) who developed entrained-flow high pressure (20 bar to 70 bar) and high temperature (>1300°C) coal gasification. A good and complete review on gasification is given by Higman and Van der Burgt.1

DD represents a good source of valuable minor compounds such as sterols, tocopherols and squalene, which can be recovered and further used as food additives, in the pharmaceutical industry and cosmetics. Furthermore, the FFAs, one of the major compounds present in DD, are mostly used as additives for animal food, fluidizing agents, for lecithin or as medium-grade soaps. Such fatty acids can also be used as precursors in a wide variety of molecular synthesis schemes such as the production of dibasic acids of different chain lengths (Gangopadhyay et al., 2007). Alternatively, DD has non-food applications, such as their mixing with fuel oil (5-10%) to fire steam boilers (Svensson, 1976). A great interest was shown in DD for its possible application in the production of high-quality (biodiesel) or low-quality (biofuel) methyl esters.

Rough estimates of the quantity of DD are available. The amount mainly depends on the content of FFA, gums and impurities present in the oil and on the efficiency of refining. Using Mielke (2009) figures for the worldwide vegetable oils production in 2009 and assuming that (1) palm oil is entirely physically refined (100% RBD), (2) soybean oil is mainly chemically refined (100% NBD) and (3) rapeseed and sunflower oils are mainly physically refined (75% RBD/25% NBD), one can estimate the DD production. Considering that by physical refining a multiple of 1.2 times the FFA content is removed from crude oil as DD (Vries, 1984), the DD production can be estimated (Table 22.2). The FFA content before deodorization step for the chemically refined oil (NBD) is difficult to estimate. However, it is generally considered that the vegetable oil contains ca. 0.10% FFA before deodorization and 0.05% FFA after deodorization (final oil) in order to make it suitable for human consumption.

The palm oil production increased significantly in 2009 (46.5 mil. t/year) compared with 2007 (36.8 mil. t/year) that accordingly determined an increase in the DD (RBD).

Considering Mielke (2009) figures for the vegetable oil production in 2009, approximately 3.1 mil. t/year of DD should have been produced, where 3.02 mil. t/ year comes from physical refining (DD-RBD) and 0.03 mil. t/year comes from chemical refining (DD-NBD).

Owing to the high content of FFA present in the DD from physical refining (RBD), this side-product is suitable for the biodiesel production and the DD from chemical refining (NBD) is mainly valorized for the recuperation of minor compounds. However, there is no clear distinction in the literature of the origin of the feedstocks.

Even with the improvements noted above, hydrogen yields of fermentative hydrogen production processes are restricted by the existing metabolic pathways to 2 or 4 mol H2/mol glucose consumed, for butyrate or acetate fermentation, respectively. The techniques already discussed in Section 13.3.6 could not increase yields beyond these limits. In practice, typical yields in the range of 1 to 2 mol Hj/ mol of glucose result in 10-20% chemical oxygen demand (COD) removal, since the main part of the organic content of the wastewater remains in the liquid phase in the form of various VFAs and solvents. Even under optimum conditions of 4 mol H2/mol glucose, about 60-70% of the organic matter of the feed remains in solution (Venkata et al., 2008; 2009). Further utilization of the organic matter contained in the effluent of a fermentative hydrogen producing bioreactor, could increase the overall energy output of the process. The development of a two-stage process involves the fermentation of the substrate to hydrogen and organic acids in the first stage and additional energy extraction by feeding the effluent of the first stage reactor to a second stage.

One approach to utilize/reuse the remaining organic matter in producing a second useable form for energy (an energy carrier) is to produce CH4 in a second stage. Integration of an acidogenic process with a subsequent methanogenic process for combined hydrogen and methane generation, offers several advantages such as a higher performance of the process in terms of waste stabilization efficiency and net energy recovery (Ghosh et al., 1985). Such a two-stage system has been proposed so far for organic solid wastes rich in carbohydrates such as food wastes (Han and Shin, 2004), cheese whey (Antonopoulou et al., 2008a; Venetsaneas et al., 2009), olive mill wastewaters (Koutrouli et al., 2009), household solid waste (Liu et al., 2006b), a mixture of pulverized garbage and shredded paper wastes (Ueno et al., 2007a) and wastewater sludge (Ting and Lee, 2007). A combined hydrogen — and methane-generation process has already been scaled up to the pilot plant stage, for organic solid wastes (Ueno et al, 2007b). The hydrogen and methane production rates were 5.4 m3/m3/d and 6.1 m3/m3/d, respectively while the process COD removal efficiency was 80%. The overall efficiency of this combined process is demonstrated by the fact that methane yields were twofold higher than a comparable single-stage process (Ueno et al., 2007b).

Another approach to increase the overall energy extraction is to couple the fermentative hydrogen production with photofermentation with the aim to recover additional hydrogen. In such a two-stage process, the rich in organic acids effluent of fermentation which is produced in the first stage by anaerobic fermentative bacteria could be converted to hydrogen in the second step by non-sulfur purple photosynthetic bacteria which capture light energy, using a photobioreactor. This combination of both kinds of bacteria not only reduces the light energy demand of the photosynthetic bacteria but also enhances the hydrogen yield as well (Das and Veziroglu, 2001). Intensive research has been carried out in this area (Nath et al., 2008; Chen et al., 2008b) in the last few years. However, there are important factors limiting the practical application of such a process. One of them is that the involved hydrogen enzyme, nitrogenase, is potentially sensitive to the nitrogen content of the medium/substrate since nitrogen inhibits enzyme activity, as well as represses nitrogenase synthesis. However, this limitation can be potentially overcome either by genetic manipulation (Drepper et al., 2003) or selection (Rey et al., 2007) to remove nitrogenase regulation. In addition, one of the most severe constraints is that photosynthetic efficiencies are very low since at even moderate light intensities, the main part of captured light is dissipated as heat (Hoekema et al., 2006). This means that there will be a demand for large surface areas for the production of hydrogen contributing to the total cost and render the development of a two-stage process of fermentation-photofermentation, far from practical application.

Another approach to increase the overall energy recovery could be the coupling of fermentation with the additional hydrogen production, via a MEC. In this two — stage system, the organic acids which are typical by-products of hydrogen fermentation will be converted to hydrogen in a MEC (Liu et al., 2005; Rozendal et al., 2006). Specifically, the electrogenic bacteria, catabolize the substrates and use the anodic electrode as terminal electron acceptor while supplementary voltage (>200 mV) is added in order to drive hydrogen evolution at the cathode. Thus, a sequential second stage of a MEC after a fermentative hydrogen production first stage could completely convert the effluent of first step to hydrogen, achieving in principle, 12 mol H2/mol glucose with only a small electricity supply. However, the fact that the yields for MEC which have already been reported in the literature are quite lower than the respective yields of dark anaerobic fermentation process, in combination with the high cost of cathodic electrodes and the reduction of the electrical input, limit the practical applicability of this promising technology.

The use of cellulosic biomass in a petroleum refinery needs to overcome the recalcitrant nature of this material and convert it into a liquid product, which is done by fast pyrolysis or liquefaction to produce bio-oils or by hydrolysis routes to produce aqueous sugars and solid lignin. Catalytic cracking of bio-oils, sugars and lignin produces olefins and aromatics from biomass-derived feedstocks. Unfortunately, large amounts of coke are obtained under cracking of these compounds over acid solid catalysts, and hence, the improvement of reaction conditions must be addressed in the future. Likewise, the obtained hydrocarbon mixture usually contains a relevant presence of oxygenated compounds that limit its use as a transport fuel.

Triglyceride-based biomass has more appealing properties for its processing in FCC units (lower oxygen content, higher effective hydrogen index, close physical properties to conventional FCC streams . . .). Although the results described in literature are very promising, most of them are in laboratory scale and little work has been addressed in pilot plant under realistic FCC conditions, and hence, we are still far from a commercial stage. Likewise, another issue in mind is the compatibility of these biomass-based streams in the refinery framework upstream FCC unit (storage, transfer lines, heat exchangers. . .). This topic has been poorly addressed in the literature, but it is a crucial key for the utilization of biomass — derived feedstocks in a petroleum refinery.

We honestly think that the co-processing of biomass feedstocks in petrol refineries is an interesting approach to reach the integrated biomass conversion process in bio-refineries. Furthermore, FCC unit and using oleaginous raw material as feedstocks are shown as the most appealing alternatives.

Figure 18.4 presents a scheme of a typical HTL process. Prior to feeding into the process, biomass is pretreated to ensure that the feedstock has desired properties: rheological properties, water content, degree of fragmentation of biomass components, etc. Feeding biomass water slurries is a particular challenge due to the problems of biomass settling and filtering and blocking of the process lines, particularly for relatively high biomass/water ratios. Heating to the desired temperature in the range of 250-370°C is performed while water is kept in liquid phase by pressure regulation. HTL can convert very wet streams to a gas without paying a huge energetic penalty, if heat exchanging has taken place efficiently. For this, heat exchange between the reactor effluent and the feed stream is essential that requires operation at high pressures to avoid phase transition (see Chapter 20 of this book). The efficiency of the heat exchanger can be high leading to a feed stream outlet temperature of only 50-100°C below the reactor outlet stream. Make-up heat for the reactor has to be supplied externally. In most cases, tubular reactors have been used in continuous installations. Typically, residence times of 10-90 minutes have been applied. In most reported pilot HTC work, the residence time at an elevated temperature was significantly longer than the ca. five minutes for minimal char formation (see chemistry section), resulting in considerable secondary char formation. In pilot plants, the feed stream was heated externally or

by heat exchange with the reactor effluent. For both cases, it holds that the heating trajectory is already taking several minutes because of heat transfer limitations. The only way to heat biomass in less than five minutes is injecting the feed directly into a hot liquid that may have a negative influence on the energy efficiency in case of a very wet feedstock.

The water phase or oil phase can be recycled for use as a solvent and/or as a dilutant. Keeping the concentration of organics low will decrease the char yield. A low concentration of organics can be achieved by using a (very) diluted feedstock, back mixing or by recycling of water (or oil) over the reactor.

Upon cooling the reactor effluent of HTL, three different products, being also three different phases, are present: a hydrophobic organic phase, an aqueous phase with organic compounds dissolved in it and a gas phase, consisting mainly of CO2. Separating gases and the water phase is straightforward. The product gas is made available at a high pressure (> 200 bar), and thus, for its application, expensive gas compression can be avoided. No reported results were found on the separation of oil and char.

Non-catalytic conversion of biomass under these conditions (230-400°C) is very susceptible to the formation of carbonaceous deposits (see also Figure 20.2). In fact without a catalyst, the product distribution consists only for ca. 10 wt% of permanent gases (primarily CO2) and 90 wt% condensed products. This is mainly caused by sugars and their decay products as they can easily polymerize in hot compressed water.58 Pacific Northwest National Laboratory (US) developed a catalytic process for the destruction of organic waste at ca. 350°C while producing a methane rich gas.59-61 Tests were carried out at laboratory and pilot scale focusing on both catalyst and process development. Ruthenium on rutile titania, ruthenium on carbon and stabilized nickel catalysts showed the highest activity and the best stability. With these catalysts, nearly 100% gasification of model components (1-10 wt% organics in water) was achieved. The gas produced consisted of nearly only CH4 and CO2, as dictated by the overall thermodynamic equilibrium. The catalytic process was carried out in a series of fixed bed reactors. When using feedstock materials with the tendency to produce char/coke, a continuous stirred-tank reactor (CSTR) was required before the fixed bed to soften the feed and to prevent the buildup of solids. Pilot plant runs using complex feeds like potato waste and manure were carried out. The required liquid hourly space velocity (LHSV) was in the range of 1.5-3.5 Nm3feed/m3cat/h. For a waste disposal process these LHSVs are acceptable, but for the production of gaseous energy carriers from biomass the activity is rather low. Waldner62 reported high extents of gasification and equilibrium methane yield of concentrated (up to 30 wt%) wood sawdust slurries using Raney Nickel as catalyst at 400°C. For complete gasification, 90 minutes reaction time was required in their batch reactor.

The catalysts employed accelerate the rate of the gasification reaction relative to the rate of poly condensation/polymerization reactions, or they are able to gasify the formed polymers, or a combination of both. However, after comparing reaction rates it can be argued that the majority of the gas is produced via gasification of partially polymerized components: in non-catalytic experiments with monomer sugars as feed maximal oil (polymerized components) yields are obtained for reaction times of 2-5 minutes,58 whereas in catalytic test 30 up to 90 minutes reaction time62 are needed to achieve complete gasification. Van Rossum et al3 proposed a simplified lumped reaction path scheme for the conversion of small carbohydrates (< C6) in hot compressed water (see Figure 20.8). Savage63 and Kruze57 reported extensive reviews on catalysis and reactions in supercritical water.

Huber et al.15 and Cortright et al.6 reported interesting catalysis around 230°C for the production of hydrogen rich gas from small oxygenated hydrocarbons. They were able to decrease the methane formation rate via C-O bond cleavage and methanization (hydrogenation) while maintaining the high rates of C-C bond cleavage and shift for hydrogen production. Cortright used a Pt catalyst, Huber a Raney nickel catalyst promoted with tin. High hydrogen yields were obtained for methanol, ethylene glycol and glycerol. However, with sorbitol and glucose as feedstock already significant amount of methane were being produced next to hydrogen. Though in an embryonic stage, the methodology of decelerating methane producing reactions at catalytic sites while keeping a high rate of catalytic hydrogen production seems promising to produce hydrogen rich gas at conditions for which overall chemical equilibrium dictates a methane rich gas, viz. at sub critical temperature and at the combination of high temperature and high concentration of organics. In this concept, it will be important to decrease

homogeneous reactions to undesired by-products (oil/char/CH4) and to increase the reaction rate. This is quite a challenge for both catalyst and reactor design.

T. LE ANH, School of Transportation Engineering, Hanoi University of Science and Technology, Vietnam, I. K. REKSOWARDOJO, Institut Teknologi Bandung, Indonesia and K. WATTANAVICHIEN,

Chulalongkorn University, Thailand

Abstract: This chapter summarises findings on the use of biofuels in conventional diesel engines. A number of biofuels such as vegetable crude oil, pure plant oil and biodiesel in different forms, which are derived from many types of raw materials such as jatropha, coconut, palm, kapok nut and cat-fish, are investigated to find the impact of these biofuels on the engine’s combustion characteristics, performance, exhaust emissions and durability. The concept of using biofuels in engines is also mentioned to determine ways of utilisation of biofuels in engines that match both the demands of biofuels use and the design of the engines.

Key words: biofuels utilisation, engine performance, exhaust emission, durability.

23.1 Introduction

Biofuels are now recognised as the most suitable alternative fuels for engines which were originally designed to use fossil fuels. Although the process of formation of fossil fuels still continues through the effect of underground heat and pressure, the current rate of consumption is higher than the rate of formation. Consequently, fossil fuels are considered to be non-renewable, that is, they are not replenished as fast as they are consumed. Biofuels, including ethanol, biodiesel and several other liquid and gaseous fuels, constitute a very promising renewable energy resource with the potential to displace the consumption of a substantial amount of petroleum worldwide during the next few decades.1-4 A clear trend in that direction is already in process.

Research on the production and utilisation of biofuels in engines is therefore regarded as a priority not only for developed nations but also for developing countries. Although the use of biofuels is currently low, the amount is continuously increasing in every country. However due to the fact that biofuels are produced from many different sources, characteristics and quality also vary, so the utilisation of different biofuels in internal combustion engines must be carefully investigated to determine the effects on engine performance and material components.

In this chapter the utilisation of biofuels in conventional diesel engines is considered. The use of crude jatropha oil (CJO), degummed jatropha oil (DJO), pure plant oils (PPOs), and biodiesels produced from crude palm oil (CPO), Jatropha curcas, coconut oil, kapok nut oil and cat-fish fat in neat form (100% biodiesel) together with various blends of biodiesel with conventional diesel is described. In addition, the use of mixed biodiesel derived from different raw materials is also considered as a possible solution for improving the quality of biodiesels.

Findings regarding the utilisation of biofuels in diesel engines are presented from case studies conducted in ASEAN (Association of Southeast Asian Nations) countries, especially Indonesia, Thailand and Vietnam, where high priority has been given to the development and use of biofuels.

Catalytic pyrolysis has been an area of research aimed at developing viable methods to improve the quality of products compared to normal pyrolysis. The role of the catalyst is two-fold. Firstly, it lowers the temperature of the pyrolysis process and unstable hydrocarbons combine to form increased amounts of oil. Secondly, the catalyst adds a ‘cracking’ effect which deoxygenates the pyrolysis products by accelerating carbon dioxide and, more hopefully, hydrogen production.94,95 In this way, catalytic pyrolysis produces more hydrocarbon-like oil with lower tar and viscous content. Before detailing some of the current work in the area of catalytic pyrolysis, it is worthwhile reviewing very briefly the fundamentals of catalysis so that the challenges faced by this particular technology can be properly assessed.

Biomass properties such as higher volatile content, higher moisture contents, and complex reaction kinetics create challenges in predicting its performance as a fuel for gasification and make it equally difficult to design a gasifier to obtain the desired output. A number of methods have been proposed and used to predict the performance of fluidized bed biomass gasifier. They are zero dimensional, 1D, 2D, and 3D. Very limited work has been done on 2D and 3D modeling. The most frequently used method is equilibrium modeling, as it is easy and gives a quicker prediction of gasifier performance. Equilibrium modeling is the zero dimensional space independent modeling method and is helpful in identifying the maximum possible conversion of biomass and the theoretical efficiency (Huang and Ramaswamy, 2009). Ramayya et al. (2006) have used a stoichiometric equilibrium model to carry out a feasibility study of the coffee husk fluidized bed gasifier. Adhikari et al. (2007) have used a non-stoichiometric equilibrium model for studying the hydrogen production from steam reformation of glycerine and the optimum condition for higher hydrogen yields was at temperatures higher than 900 K, with a water to glycerine ratio of 9 at atmospheric pressure. Non­stoichiometric equilibrium models have been used for modeling steam gasification of coal and pure carbon fuel to predict production of hydrogen from ethanol in the presence of CaO (Florin and Harris, 2008). Jarungthammachote and Dutta (2007 and 2008) used the stoichiometric model to study a downdraft waste gasifier and non-stoichiometric model for both spout bed and spout fluid bed gasifiers. Thus, equilibrium modeling can be very helpful in modeling fluidized bed gasifiers for use with non-convectional biomass like coffee husks, glycerine, ethanol, etc., whose reaction kinetics are not identified correctly. The equilibrium model considers only the mass and energy balance and does not take into account the kinetics of the reaction, so the results obtained may differ a lot from the practical results.

Gasification consists of homogenous and heterogeneous reactions, whose reaction kinetics, as well as mass and energy transfer phenomenon, depends on the operating conditions, and accordingly the product gas composition and yield changes. To overcome this disadvantage and to predict the gasification performance more closely to reality, different simulators were developed. Nikoo and Mahinpey (2008) have used ASPEN PLUS for simulating an atmospheric bubbling fluidized bed biomass gasifier. The AS PEN PLUS simulator uses Gibbs free energy for simulation of product from homogenous reaction and reaction kinetics for char gasification. A shrinking core model has been used for kinetic modeling. Enden and Lora (2004) used a CSFB simulator to predict the performance of fluidized bed biomass gasifiers in terms of maximum char conversion obtained, as well as the amount of tar present in the gas produced, its hot and cold gas efficiency and the heating value of the gas produced, considering point to point mass and energy balance, chemical reactions kinetics and fluidization dynamics. So, once the preliminary sizing is done, one can use this simulator to evaluate whether the designed gasifier will give the desired performance or not.

Guo et al. (2001) have used a hybrid neural network model for simulating a steam fluidized bed gasifier to predicting the gas yield and composition of the gas. Because of its complex nature, it is rarely used for modeling the gasification process.

Other mathematical models were also developed to simulate the fluidized bed biomass gasifier. Raman et al. (1981) developed a one dimensional model to study the gasification of feedlot manure. Their work does not consider the devolatilization step and considers only the char gasification and water gas shift reaction. Ji et al. (2009) have used a 1D non-isothermal model to study steam gasification of biomass in the fluidized bed. Ergudenler et al. (1997) have developed a kinetic — free homogeneous equilibrium model for predicting the steady state performance of a fluidized bed straw gasifier. The Department of Energy has developed a design chart, but this is for the gasification of coal.

The first step in the design of a gasifier is to define the input and expected output. Depending upon the gasification process choice and its configuration, the input and output parameters can also vary, but in general the input and output could be listed as below:

Design input:

1 Fuel

(a) Proximate and ultimate analysis

(b) Feedstock temperature

2 Gasifying medium

(a) Choice of the medium steam, oxygen, air, or a mixture in suitable proportion.

(b) The gasifying medium may be chosen based on the following criteria:

(i) The desired heating value of the product gas.

(ii) Hydrogen content of the product gas can be maximized with steam, but if it is not a design priority, oxygen, or air could be a better option.

(iii) If an nexpensive source of external heat, such as waste recovery is available, steam is a good choice.

(iv) If N2 in product is not acceptable, steam or oxygen are to be chosen.

(v) Capital investment is lowest for air as the medium, followed by that for steam. Much larger investment is needed for an oxygen plant, which consumes a large amount of auxiliary power as well.

3 Product

(a) Desired composition of product gas

(b) Desired heating value

(c) Desired output of the gasifier (Nm3/s or MWth produced)

Design outputs:

1 Geometry

(a) Reactor configuration, its cross-section area and height

2 Operating parameters (a) Reactor temperature

(b) Input temperature of the gasifying medium

(c) Amount and relative proportion of the gasifying medium

3 Product

(a) Yield of product gas

(b) Composition of product gas

4 Performance parameters

(a) Carbon conversion efficiency

(b) Cold gas efficiency