The main advantage of biohydrogen is that it is a clean, CO2 neutral energy source, which can be used in fuel cells to produce electricity efficiently and water as a by-product, compared to other fossil fuels, the oxidation of which is accompanied by CO2, NOx, particulate and other emissions. Moreover, the high efficiency of electricity generation in fuel cells that utilize hydrogen is independent of the scale of the fuel cell. This feature allows the application of fuel cells (and consequently, the use of hydrogen) at both large scale (e. g. industrial plants) and small scale (e. g. vehicles) (Gosselink, 2002). In biohydrogen processes, the energy of sunlight is first captured in the plant biomass and then transferred to H2 as an energy carrier or is directly harvested in the form of H2. The chain of sunlight energy to hydrogen production and then hydrogen storage and distribution to the ultimate electricity generation comprises a sustainable energy scheme, which can be applied to replace gradually the fossil fuel economy. Technological advances reducing the limitations of biohydrogen processes could render the hydrogen economy easier to implement.

There are few economic analyses in the literature about hydrogen production. Most studies have been conducted at lab scale, while problems related with scaling up have not yet identified. Ressnick (2004) performed a comparative economic analysis applying a series of economic models to predict the capital and operating costs of the various approaches having been tested at a lab scale. The estimation has been based on a capacity of 50 million SCFD (standard cubic feet per day) and, based on the specific hydrogen production rate reported in the literature for the various biohydrogen processes considered, the size of the plant was assessed (Table 13.7).

The capital cost is mostly affected by the land area required in every approach. The annual sunshine limitations have not been taken into account in this analysis, and this has an impact on all photo-dependent processes. The specific rate of hydrogen production is also crucial in the above analysis. Any increase or decrease of these values would dramatically affect the economic analysis. As a result, more efficient bioreactor designs that could improve the rate and decrease the reactor volume, shrinking the capital cost further.

The operating costs also vary significantly depending on the biotechnology used. The high operating costs of direct biophotolysis are attributed to the labour costs associated with the large land area involved. In the case of the water-gas shift and the photo and dark fermentation technologies, the cost of the feedstock is crucial. If the cost of CO and glucose is excluded from the analysis, the operating costs of these technologies would decrease to 17.44, 5.60, 4.43 $/GJ respectively.

The gasification system in its commercial development may face the following technical and non-technical challenges (Basu, Acharya, and Dutta, 2009).

Technical barriers

1. Availability is the most important factor that prevents the wide scale use of gasifiers in the mainstream energy industries. Present day gasifiers have not reached the standard (>90%) expected in utility industries. The 140 MW high-temperature Winkler gasifier of Rheinbrau started with an availability of 45%, which rose to 89% in 1997 (Renzenbrink et al, 1997). This forced EPC to specify a stand-by gasifier in order to meet the overall unit availability matching the industry standard.

2. Complex operation due to a large amount of ancillary equipment, such as oxygen separation units (OSU), gas sweeteners, etc.

3. Gasification of high alkali biomass (agglomeration problems), RDF, and waste (mercury removal problem).

4. Tar control in order to avoid the problems in gas cooling and filtration, as well as its removal without producing toxic waste water.

5. Poor carbon conversion is a major problem. It has been less than 90.5%. High carbon levels in ash reduces the ash quality, and deprives the plant owners from the revenue expected from the sale of gasifier ash for their end-use. High carbon in ash imposes an additional burden of disposal cost.

Non-technical barriers

Beside the above technical challenges, several other challenges retarded the market penetration of gasification.

1. higher investment;

2. fuel availability and price level for non-conventional fuels, such as biomass;

3. subsidies available for biomass-only plants but not for co-firing;

4. economic competitiveness with steam cycles, co-firing, etc.;

5. complicated and costly means of power production;

6. small difference in efficiency compared with steam cycles.

The economics of gasification are heavily dependent, however, on the nature of the feedstock and on the location of the gasifier relative to both the source of the feedstock and to the ultimate user of the product (see Table 16.1).

Notwithstanding the complexity of the FT plants, all XTL (where X = C for coal, G for natural gas or B for biomass) processes consist of the three main sections illustrated in Fig. 19.1: gasification to syngas and gas cleaning/conditioning, FT synthesis reactor and product upgrading section. Variations and different available options for biomass gasification (pressure, use of oxygen, air medium, etc.), type of FT reactor and catalyst and target products lead to large number of possible process configurations to produce FT liquids from biomass (Tijmensen et al., 2002). All concepts, however, can be grouped into two main categories: full conversion FT, aimed at maximizing FT liquids production, and once through FT, with co-firing of the off gas with natural gas in a gas turbine for electricity production, aimed at maximizing energy efficiency.

Several studies have investigated the technical feasibility and economics of the different BTL-FT processes in order to identify the most promising system configurations (Hamelinck et al., 2004; Henrich et al., 2009; Tijmensen et al., 2002; van Vliet et al., 2009). The outcome of these studies is not conclusive as there are large uncertainties concerning technology status and economic values. Although both biomass gasification technologies and syngas conversion technologies are commercially available and have been demonstrated at a commercial scale, there is very limited commercial experience in integrating biomass gasification with downstream processes for the production of liquid transportation fuels. There is, in general, a common consensus that R&D efforts should focus on the following key issues: gasifier designs, syngas quality, product selectivity in chemical synthesis and process integration and scale (E4tech, 2008).

The following paragraphs consist of a description of the main processes, reactor types and catalytic materials employed in these three main sections of the BTL-FT process.

The production costs of advanced biofuels are too high to become market competitive without any governmental support. Technological breakthroughs are necessary to change this situation. Currently, the main focus within the production processes is on the production of the specific biofuels concerned. Primary residues (residues resulting from crop farming) and secondary residues (process residues) are used for feed applications in case they meet the quality requirements, and to produce heat and/or power, both for internal process use and to be fed to the national grind. A major problem is that in case the biofuel production capacity increases the amount of residues will overload existing markets for these products, resulting in decreased market prices. This was illustrated by oil-crop biodiesel derived glycerine production in recent years, resulting in the closure of a lot of biodiesel production facilities in Europe. The same situation is now occurring for conventional bioethanol derived DDGS (dried distiller’s grains with solubles). Another problem is that the production of heat and/or power from the process residues are low quality applications, resulting in relatively low market prices in case no governmental support is given (green power and/or heat support). The production of higher added-value Bio-based Products from these residues in integrated biorefinery facilities is necessary to maximise full biomass-to-products value chains, potentially making the production costs of the biofuels market competitive without any governmental support.

The combustion engine study aims to investigate comparative results of using crude palm diesel (blending 10% of CPO in diesel, 10% CPO diesel) on engine combustion of a CI IDI swirl chamber engine. The experiments, conducted on a Ford Ranger WL81 2.499 litre engine, were composed of two parts. First, measure and analyse in-cylinder pressure and fuel injection line pressure by using crude palm diesel and diesel fuel. Second, study combustion phenomena of both fuels in the swirl chamber by means of engine visioscope. Results show details of phenomena of spray, flame propagation. Two-colour method was also employed to evaluate flame temperature and distribution of soot in flame. And finally, to compare results of visualised combustion phenomena with heat release that estimated from in-cylinder pressure information.

Many properties of the 10% CPO diesel fuel can be attributed directly to the thickening effect of the CPO on the diesel fuel. In this study, blending 10% of CPO by volume in diesel can meet the Thai diesel fuel specification. The primary properties of both the baseline diesel and the crude palm diesel blend are shown in Table 23.1. The higher density and higher viscosity of CPO compared with diesel fuel resulted in slight increasing of these properties in the resulting crude palm diesel blends. The blend also has roughly 5% less energy per volume and less cetane value than diesel fuel. The 10% CPO diesel shows the slight reductions in T90 point that may affect the poor long-trip economy. The addition of CPO to diesel fuel will degrade the cetane number of the resulting 10% CPO diesel blend. The flash point of 10% CPO diesel is controlled by high flash point of the CPO. The flash point of 10% CPO diesel is higher than that of diesel fuel.

The engine under study is a commercial IDI, water cooled four cylinders, in-line, natural aspirated engine. The following chart displays the main dimensions:

The engine was connected to an AVL alpha 40 eddy-current dynamometer. In-cylinder pressure was taken by AVL piezoelectric pressure transducer model GU12P. Fuel line pressure was taken by a KISTLER 607C1 pressure transducer.

Indicating data were captured with Cussons P4503 shaft encoder and Cussons P4500 autoscan. Direct photography was taken with an AVL Engine Visioscope. The system consists of a PixelFly VGA Colour CCD camera (resolution 640 x 480 pixel), an AVL control unit, AVL 364C crank angle encoder, an optical linkage to the camera and the endoscope. The optical access for the endoscope to the swirl chamber of the fourth cylinder was prepared through the cooling system of the cylinder head. The visioscope software controls the triggering of the digital camera within a crank angle tolerance of 0.1°CA. The endoscope has a viewing angle of 30° forward view. To capture the spray images, the light source unit with fibre optic (40 mJ/flash with 20 ps duration at frequency of 10 Hz) was used.

The schematic arrangement of experimental set up is shown in Fig. 23.12.

The experiments were carried out at constant speed, steady state conditions at selected high probability operating points along ECE 15 driving cycle, as shown in Table 23.2.17,18 For the combustion analyses, images of simultaneous complex spray, inflammation and combustion processes in the swirl chamber were taken. Speed, torque, fuel consumption, engine operating pressure and temperature for both fuels were recorded during each test.

Comparison of in-cylinder pressure, fuel line pressure, fuel injection rate, heat release rate, net heat release and mass fraction burned is shown in Fig. 23.13.18 The measurement of in-cylinder pressure and fuel injection line pressure has

indicated that 10% CPO diesel has approximately 1° of early injection timing compared with diesel. The 10% CPO diesel also has longer ignition delay and higher amount of fuel injected mass (mf) due to its lower energy density. The maximum in-cylinder pressure of 10% CPO diesel is similar to diesel. Net heat release and mass fraction burned of 10% CPO diesel are also lower than diesel.

Comparison of maximum in-cylinder pressure (Pmax), SOI, ignition delay and fuel injected mass (mf) as engine operates with diesel and 10% CPO diesel are summarised in Table 23.3.

23.13 Comparison of in-cylinder pressure, fuel line pressure, fuel injection rate, heat release rate, net heat release and mass fraction burned as engine operates with diesel and 10% CPO diesel at 2000 rev/min, 30 Nm.18

The images of spray formation at selected operating points of reference diesel and 10% CPO diesel are shown in Fig. 23.14 (a) and (b), respectively.1718 The figures show that 10% CPO diesel has approximately 1°-2° of early injection timing compared with diesel. The early injection timing is probably due to the higher isentropic bulk modulus and higher viscosity of CPO compared with diesel, resulting in a slight increase in these properties in the resulting blends.19 The comparison of the observed spray formation between reference diesel and 10% CPO diesel are summarised in Table 23.4. It was found that, using OEM injection pump and standard injector in a pre-chamber, with 10% CPO diesel the observed sprays were wider than that of reference diesel. The difference in spray angle tends to reduce with increasing speed. The observed spray penetration with 10% CPO diesel is also longer than reference diesel in low to medium engine speed range. The higher the engine load, the longer the spray penetration was observed.

Summarising the results of these sections, as shown in Fig. 23.15, it can be noted that the visible combustion course in a swirl chamber occurs without any starting aids.17,18 The visible inflammation appears above the fuel jet. From there the flame engulfs the whole swirl chamber very quickly. This process needs some delay times. The comparison of the observed luminous spray combustion between reference diesel and 10% CPO diesel is shown in Table 23.5. It was found that 10% CPO diesel has shown a longer ignition delay period than diesel. The combustion for both fuels tends to start faster with increasing speed. After this ignition delay, the burning area rotates under the influence of the swirl. This motion can be observed for nearly the entire burn duration after complex luminous inflammation has occurred. In the low speed and load range, 10% CPO diesel

23.14 (a) and (b) Images of liquid fuel spray in the pre-chamber for reference diesel and 10% CPO diesel respectively. The crank angles at which the images were acquired are written on the left of the images.

combustion duration tends to have a slightly shorter period than diesel. This may be due to the benefit of oxygen content in the fuel.

Using the ‘Thermovision’ software from AVL List GmbH,20 the temperature of radiating soot particles was calculated from the three spectral intensities in the flame images using the two-colour method. In the temperature images, shown in Fig. 23.16, purple — blue — green — yellow — red — white in the original colour image denote the temperatures ranging from 1800 to 3000 K.

Table 23.5 Comparison of the first appearance of luminous flame, end of luminous flame and luminous flame duration between reference diesel and crude palm diesel in an IDI engine

First appearance End of luminous Luminous flame

of luminous flame (°CA) duration in

flame (°CA) pre-chamber (°CA)

23.16 Flame temperature images of spray combustion in the pre­chamber for reference diesel and 10% CPO diesel. The crank angles at which the images were acquired are written at the top of the images.

(Continued )

The difference in combustion is much more obvious when looking at the flame. The in-cylinder combustion temperature of 10% CPO diesel combustion is lower than diesel combustion. From Fig. 23.17, the flame areas of temperature above 2400 K for diesel and 10% CPO diesel at 2000 rev/min, 30 Nm are compared. It was found that diesel fuel showed greater amount of flame areas of temperature above 2400 K.

In the soot distribution images, the same colour scale denotes soot densities ranging from thin to dense soot. The appearance of luminous combustion flame comes from the radiation of soot particles occurred in the fuel mixture oxidation zone. Prediction of soot density distribution at selected operating points of diesel and 10% CPO diesel are shown in Fig. 23.18. It is noted that soot density in 10% CPO diesel combustion flame tends to be slightly lower than that in diesel.

Comparative studies of engine fuelled with reference diesel and 10% CPO diesel were investigated. Visualised images show the effects of CPO in 10% CPO diesel blend. The injection timing of 10% CPO diesel is approximately 1° earlier compared with the injection timing of reference diesel. Observed 10% CPO diesel fuel sprays have shown either longer spray tip penetration length or wider spray angle than the reference diesel.

23.17 Flame area with temperature above 2400 K for 10% CPO diesel and diesel at 2000 rev/min, 30 Nm.

23.18 Soot concentration distribution images of spray combustion in the pre-chamber for reference diesel and 10% CPO diesel. The crank angles at which the images were acquired are written at the top of the images.

(Continued )

Images of spray combustion indicate that the period of 10% CPO diesel combustion phenomena occurred more retardedly with respect to TDC than diesel. As its consequence, together with the lower heat of combustion, the predicted combustion flame temperature and soot density distribution, using the two-colour method, are lower than the reference diesel. The combustion for both fuels tends to start faster with increasing speed. The observed combustion duration of 10% CPO diesel is slightly shorter than that of diesel.

There are many forms of reactors used for the study, analysis and large scale semi­commercial testing in catalytic combustion. A full description of these is beyond the scope of this paper and the reader is referred to a number of papers. Samolada et al. have presented a typical laboratory set-up for gram quantities of biomass in a fixed bed reactor and this is very typical for small scale studies.133 Numerous pilot plant size (kg type quantities) reactors have also been studied and these are largely centred on fluidised bed type systems.141 These pilot stage systems will have hoppers and grinders for feeding real biomass samples (as particle size is critical in determining thermal transfer efficiencies), risers (for the fluidised bed generation), pyrolysis reaction chambers and systems for recovery and recycle of the catalyst. On the larger scale there are many variations of the methodology which allow product optimisation to different molecular weights as well as different products. The methods developed for commercial use and study have been summarised by Meng et al.142 These methods include, for example, catalytic steam pyrolysis where the addition of water promotes steam reforming reactions within the overall pyrolysis process. Quick contact cracking essentially involves a recirculating fluidised bed fast pyrolysis technique combined with a simple cracking catalyst which allows catalyst to be circulated in the fluidised bed and coke at the catalyst removed by oxidation during the recycle. The current development of large scale industrial plant is summarised by Dominov et al.143 The results of a commercial trial of catalytic pyrolysis technology with emphasis on the design and construction of plant were reported by Xie and Wang.144

Despite the complexity of these technologies and individual reactor designs, simple representation of the main methodologies can be made and these are represented in Fig. 14.1. In the simplest fixed bed system (Fig. 14.1A), catalyst and powdered biomass or other hydrocarbons are mixed and the composite placed in a tube (held in place by ceramic wool or sinter disks). The tube is externally heated to produce a high heating rate. The pyrolysis reaction and the catalytic cracking/reforming reaction can also be separated into two separate processes (Fig. 14.1B). This allows the cracking/reforming reactions to be run at different temperatures as well as secondary input of gas (water, hydrogen) to improve that catalytic process. Fluidised bed reactors can also be run with an in situ (i. e. in the pyrolysis reactor) catalytic process (Fig. 14.1c) or an ex situ process. One of the biggest advantages of an ex situ process is that the catalyst can be used in a fixed bed which prevents the impact damage that occurs in a fluidised bed which can severely limit catalyst lifetime. Catalysts can be periodically regenerated by oxidative treatment to remove coke using a twin reactor tube.

This approach will be illustrated using the example of steam gasification of char.

Let A denote the air supply in kg dry air per kg dry fuel, F the amount of dry fuel required to obtain one normal (at 0°C) cubic meter of the gas and Xc the carbon content of the fuel (kg carbon/kg dry fuel). This carbon is split between CO, CO2, and CH4 in the product gas. We know that 1 kmol gas occupies 22.4 Nm3. So, for 1 Nm3 of gas produced, one can write the carbon (moles) balance between inflow and outflow streams corresponding to:

where V represents the volumetric fraction of a constituent of the gas.

Similarly, one can develop three more equations balancing hydrogen moles, oxygen moles, and nitrogen moles.

rx, K^v-4 %,). [169]

2 22.4

cv [2V +V +V +2V )

_ I V W °if, [16.10]

2 22.4

We assume the product gas to be made of CO, CO2, CH4, H2, H2O, O2, and N2. The volume fractions of individual gases make up the total product gas, which is 1 Nm3 for F kg feed. So:

VCO2 + VCO + VH2 + VCH4 + VH2O + VO2 + VN2 = 1 [16.12]

Together one gets five equations. But it is necessary to find eight unknowns: VCO, VCO, VCH, VH, VHiq, VO2, V^, and F, the fuel feed for production of 1 Nm3 of

the product gas. To find the eight unknowns it is necessary to obtain three more equations. These are obtained from reaction kinetics. The following three overall gasification equations are relevant for steam gasification:

Water gas reaction: C + H2O CO + H2 + 131 kJ/mol [16.13]

Methanation reaction: C + 2H2 CH4 — 75 kJ/mol [16.14]

Shift reaction: CO + H2O H2 + CO2 — 41 kJ/mol [16.15]

For oxygen or air-gasification, other relevant sets of equations are to be considered. If allowed, these equations could reach equilibrium when forward and backward reaction rates are equal. Let PCO and PCO be the partial pressure of CO and CO2, respectively, under equilibrium.

The partial pressure of CO, PCO corresponds to the volume fraction of CO. So, PCO = VCO. P when the pressure of the reactor is P.

For the water gas reaction,

[16.16]

’ h, o

For the shift reaction,

P P V V

J, _’h,’co, _ H, CO,

I’pWS — p p ~ у у

COJ HjO CO HjO

For the Methanation reaction the equilibrium equation is:

[16.18]

Values of these equilibrium constants at different temperatures may be taken from Basu (2006, p. 68). Equations [16.8] to [16.12] are solved along with Eq. [16.16] to Eq. [16.18] to get eight unknowns. More details of this method are given in Basu (2006).

Although hydrocracking yields an appealing spectrum for the production of diesel, it is not an attractive option for gasoline. The relatively low extent of branching achieved in hydrocracking yields a product in the gasoline range with a low octane number. In addition, hydrocracking is considered an expensive process due to the high pressure operation and high hydrogen consumption. The fluid catalytic cracking (FCC) process has been investigated as an interesting option for the cracking of FT waxes aimed at the production of FT gasoline (Dupain et al., 2005, 2006; Lappas, 2007; Lappas and Vasalos, 2006; Lappas et al., 2007; Triantafyllidis et al., 2007).

The FCC process is the most important refinery process mainly for the production of gasoline from heavy petroleum fractions such as atmospheric and vacuum gas oil (VGO). In the FCC unit, the long hydrocarbons are cracked in the 480-540°C temperature range over zeolite catalysts to smaller n — and i-paraffins, n — and i-olefins and aromatics. Conventional FCC feedstocks are relatively aromatic, with a high sulphur and nitrogen content, in contrast to FT waxes that are highly paraffinic with extra low aromatics content (< 1 wt.%) and virtually zero sulphur (< 5 ppm) (see Table 19.3). Both the development, therefore, of new catalyst formulations and optimization of the overall process parameters are very critical to optimize the yield and quality of FCC products from FT waxes.

Lappas et al. (2007) compared the crackability of conventional VGO feed and FT wax provided by CHOREN over a typical refinery FCC E-cat. As can be seen in Fig. 19.10, the FT wax is much more crackable than VGO due to the highly paraffinic molecules of wax compared to VGO that contains a significant amount of aromatics. In fact, the cracking rate of the wax molecules was calculated as about 4.2 times faster than that of the VGO molecules. Moreover, coke formation was much less compared to VGO, again due to the paraffinic nature of the feed and the absence of aromatic compounds or coke precursors even at high conversion levels. Very high conversions, over 80 wt.%, can be achieved with conventional

FCC catalysts at very low catalyst/oil ratios and low temperatures. In Table 19.4, a comparison between the two feeds regarding the product distribution at 70 wt.% conversion is given. The table shows that gasoline (C5, 221°C) yield is about the same with both feeds. Gasoline from VGO has, as expected, a higher octane number; however, the research octane number (RON) of the wax gasoline is still acceptable. The RON of the wax gasoline was almost constant and independent of the conversion exactly due to the low aromaticity of this gasoline (Lappas et al., 2004). Dupain et al. (2006) also observed that the cracking of wax to gasoline is a primary reaction with a gasoline selectivity that is independent of conversion level or temperature. Despite the lower R ON number, gasoline from the cracking of FT waxes in an FCC unit is very promising due to the low content of aromatics in the product and the extremely low sulphur and nitrogen concentrations, leading to the production of a very clean gasoline. Moreover, it was found that the diesel range LCO product produced from the catalytic cracking of FT waxes is better than the respective produced from the cracking of conventional FCC feedstocks. The degree of branching in the diesel product is

lower than that of the gasoline, improving marginally the cetane number but acting very beneficially for the diesel cloud point and pour point, in addition to the very low sulphur and nitrogen content (Dupain et al, 2006).

The addition of ZSM-5 additive to a conventional E-cat was found to enhance the cracking rate of FT waxes, enhancing the cracking of gasoline range olefins to gas range olefins and especially propene and butene (Dupain et al., 2006). This was attributed to the diffusions of the initially formed smaller olefins in the ZSM-5 pores. The olefins are not able to leave the ZSM-5 pores rapidly enough, and they are thus easily activated and overcracked to gas range olefins (Dupain et al., 2006). Use of pure ZSM-5 resulted in an octane-enhancing effect of the produced gasoline due to the enhanced formation of olefins and aromatics. Triantafyllidis et al. (2007) investigated the potential utilization of various microporous (zeolites H-Y and H-ZSM-5) and mesoporous [amorphous silica — alumina (ASA) and Al-MCM-41] aluminosilicates as catalysts or active matrices in the cracking of FT waxes towards the production of liquid fuels. Focus was placed on the effect of porous and acidic characteristics of the materials on products yields and properties. According to the authors, the type of catalyst plays a significant role in the product selectivities. The percentage conversion of wax, the product yields [gasoline and liquefied petroleum gas (LPG)] and the RON of the produced gasoline are shown in Fig. 19.11 for different investigated microporous and mesoporous catalysts. The behaviour is typical for the two zeolitic catalysts when used in FCC of petroleum fractions, where H-Y zeolite is being utilized as the main active cracking component of the catalyst and ZSM-5 is being used as an additive in small amounts, leading to lower gasoline and higher LPG yields, and usually to higher RON. Similar trends are observed in Fig. 19.11

for the cracking of FT waxes. One of the main reaction pathways that ZSM-5 catalyzes with higher rates than H-Y is the cracking of paraffins, thus making it very active in the conversion of waxy feedstocks in agreement with the results of Dupain et al. (2006). The 3%-crystalline H-ZSM-5 sample, not diluted with ASA, showed high conversion activity (79 wt.%), very close to that of the diluted catalyst of the crystalline H-ZSM-5. It can thus be suggested that the acid sites present in this sample are much more active for the conversion of wax compared to those of Al-MCM-41 and ASA, although the very low crystallinity H-ZSM-5 sample consists mainly of X-ray diffraction (XRD) amorphous aluminosilicate phase. Figure 19.12 shows the yields (wt.% on feed) of various gasoline components. The data in Fig. 19.12 can also be used for a qualitative comparison of catalytic performance with regard to selectivity towards specific gasoline components, especially in the case of H-Y and H-ZSM-5-based catalysts, which showed a similar percentage conversion of wax (Fig. 19.11). The H-Y-st. catalyst presented a significant selectivity towards the production of branched paraffins (22 wt.% on feed) compared to much lower yields with the rest of the catalysts (3.5-4 wt.%). The increased formation of branched paraffins in gasoline is considered as a major target towards the production of environmentally friendly fuels in accordance with the EU regulations. Olefins were also higher with the H-Y-st. catalyst (15 wt.% on feed) compared to the rest of the catalysts (~12 wt.%), while naphthenes were 1-2 wt.% for all the catalysts. As far as aromatics are concerned, the H-ZSM-5 catalyst led to higher yields compared to the rest of the catalysts. The high RON values of gasoline with the H-ZSM-5 catalyst (~92, see Fig. 19.11) were mainly attributed to the high aromatics content, while in the case of H-Y-st. catalyst, the high RON (~87) was mainly attributed to the relatively high C5-C7 olefins and iso-alkanes yields. The 3%-

H-ZSM-5 sample showed similar trends with the fully crystalline H-ZSM-5 with regard to the yields of gasoline components, except for the case of aromatics, which are significantly lower with the former sample. Interestingly, the RON of the gasoline produced from the 3%-crystalline H-ZSM-5 sample remained considerably high (84). The yield of aromatics with the Al-MCM-41 sample was very low, but they cannot be compared with those of the rest of the catalysts due to the relatively low percentage conversion of wax with the mesoporous catalytic material.

In general, research has shown that the cracking of highly paraffinic FT waxes under FCC conditions can yield an interesting spectrum of renewable fuels, both in the gasoline and diesel range, by adapting the process parameters and catalyst formulations. Optimization of catalyst’s acidic and porosity properties as well as of process parameters is necessary in order to visualize a potential commercialization of the FCC-based upgrading of FT waxes.

Currently, secondary (process) residues of biochemical conversion processes, for example lignin, are used in thermo-chemical conversion processes (combustion)

Targeted at hemicellulose

Targeted at cellulose

21.6 Aquathermolysis-pyrolysis biorefinery concept.

to produce heat and/or power, partly used to energetically drive the biochemical conversion process. However, specifically for lignin a variety of downstream valorisation processes are being developed to produce more financial value out of this secondary residue stream, for example supercritical lignin conversion, enzymatic lignin conversion and catalytically supported fast pyrolysis (see Fig. 21.7).

Lignin, which consists of polyaromatic oxygenated compounds, represents a major fraction of biomass (10-30%) and is currently used as a low-grade fuel to provide heat in the pulp and paper industry, but it would be highly desirable to produce value-added products from lignin. Lignin can be converted into a transportation fuel by dehydroxygenation or zeolite upgrading. These are the same methods used to upgrade bio-oils, which contain a large fraction of lignin-derived products. Thring et al. (2000) studied zeolite upgrading of lignin with HZSM-5 zeolite as a catalyst in a fixed bed reactor operating at an atmospheric pressure, over a temperature range of 500-650°C and weight hourly space velocities of 2.5-7.5/ hour. The liquid product fraction, which consisted of mostly aromatic hydrocarbons (mainly benzene, toluene and xylene — with toluene dominating), was maximized at a temperature of 500°C and a space velocity of 5/hour. On the other hand, the gas product consisted of olefins, light hydrocarbon gases, CO and CO2 and was produced at the highest yield at a temperature of 650°C and a space velocity of 5/hour. Among the light hydrocarbon gases produced from the lignin, ethylene and propylene were the olefins produced in the highest quantities. Coke and char formation was particularly high at the low reaction temperatures employed in this work but decreased rather drastically with increasing temperature. For instance, at a space velocity of 5/hour, 50 wt.% of the lignin was converted to coke and char when a reaction temperature of 500°C was used compared to only 21 wt.% at 650°C. Small FCC pilot tests were run to determine the crackability of pyrolysis oil and pyrolytic lignin blended with VGO (Holmgren et al., 2007). In the blends, the VGO serves as a hydrogen donor. Compared to VGO, the pyrolysis oil and pyrolytic lignin tend to form high levels of coke. For the blends of VGO with pyrolysis oil or pyrolytic lignin, the acid bio-oils appeared to increase the crackability of the VGO and shift VGO yields towards increased light ends and lower LCO and clarified slurry oil (CSO), which is an economically attractive outcome. Nevertheless, the high levels of coke obtained with both blends (7% and 9%, respectively) would be unacceptable for most FCC units.

Glycerol is produced from biomass through fermentation of sugars and mainly by transesterification of vegetable oils during biodiesel production. The glycerol

market is currently undergoing radical changes, driven by very large supplies of glycerol arising from biodiesel production. Glycerol is currently too expensive to be used as a fuel; however, as biodiesel production increases, the price of glycerol will decrease. Corma et al. (2007) studied the catalytic cracking of aqueous glycerol and its mixture with VGO in a microactivity test (MAT) reactor at 500- 700°C with six different catalysts. Products from this reaction include olefins (ethylene, propylene and butanes), aromatics, light paraffins (methane, ethane, propane), CO, CO2, H2 and coke. The ZSM-5 catalyst had the highest level of olefins and aromatics and the lowest level of coke (< 20%) in the catalytic cracking of glycerol, whereas the other catalysts had high coke yields (30-50%). When glycerol is fed together with VGO, interactions between the hydrocarbon components and the glycerol reaction intermediates occur, resulting in final selectivities better than those calculated by considering a simple additive effect. These experiments showed that mixtures of VGO with biomass-derived feedstocks can help to transfer hydrogen from the VGO to the biomass molecules. One option for further improving the olefin and aromatic yields for co-feeding of glycerol and petroleum-derived feedstocks into an FCC reactor might involve adding ZSM-5 to the FCC catalyst because ZSM-5 produced more olefins and less coke than FCC catalyst.

Sugars can be used as feedstock for fuels production by different processes. Chen (1976) discussed the conversion of carbohydrate materials to petroleum — type hydrocarbons. The process is composed of microbial conversion of agricultural carbohydrate materials to alcohols followed by direct conversion of the oxygenated microbial reaction product to a hydrocarbon product comprising a substantial highly aromatic fraction. This latter conversion was carried out in the presence of a ZSM-5 zeolite at about 260-540°C. Later, Chen and co-workers (Chen and Koening, 1990; Chen et al., 1986) passed concentrated sugars, including glucose, xylose, starch and sucrose, over ZSM-5 at a temperature from 300°C to 650°C and observed hydrocarbon, CO, CO2, coke and water as products. The addition of methanol to the feed decreased the amount of coke and increased the hydrocarbon products. The hydrocarbon products consisted of gaseous alkanes (methane, ethane, propane), liquid alkenes and alkanes (butane, pentene, hexane) and aromatics (benzene, toluene, C8-C10 aromatics). One of the problems of this reaction is that when methanol is not used, 40-65% of the carbon is converted into coke.