The research of gasification/steam reforming of pyrolysis oil was initiated by the National Renewable Energy Laboratory (NREL) in the USA. In the nineties of the twentieth century, NREL published the first results30 on steam reforming of acetic acid (HAc) and hydroxyacetaldehyde (HAA) with the aim to produce hydrogen. HAc and HAA were chosen as model compounds because they represent a part of the pyrolysis oil, which was identified as a possible renewable biomass chemical and energy carrier. A fixed bed microreactor was used to convert the model compounds using grounded commercial catalysts (G-90C and C18HC from United Catalysts Inc.). The thermal stability of the compounds was given as an indicator for coke formation. Both HAc and HAA were catalytically converted to hydrogen-rich gas at a reactor temperature of ~700°C (for HAA a lower inlet temperature was chosen) and a steam over carbon ratio (S/C) > 2. Further tests with model compound reforming,31,32 including the vapors of cellulose, xylan and lignin, spraying of glucose, xylose and sucrose onto a fixed catalytic bed in combination with catalyst screening ultimately led to the first actual reforming of the aqueous soluble phase of pyrolysis oil.33

Two commercial naphtha/C2-C3 steam reforming catalysts (UCI G90C and the ICI 46-series) showed very promising results in their ability to convert the aqueous soluble phase of pyrolysis oil with only minor coking at high steam over carbon ratios (20-30).33 However, an increase of methane concentration during a test could be observed. To feed the aqueous soluble phase of pyrolysis oil, adjustments had to be made to the atomizer system in order to directly add the reactant to the catalytic bed. With the improvement of the feeding system, fixed bed reforming of the aqueous soluble phase of pyrolysis oil was still limited to 3-4 hours of operation due to carbonaceous deposits on the catalyst and in the freeboard.34 To overcome this run time barrier, the reactor bed was changed from a fixed to a bubbling fluidized bed where the commercial catalyst was grounded to a particle size of 300-500 pm. A different catalyst than the ones used before, namely the naphtha reforming catalyst C11-NK from Sud-Chemie, was now being used. The liquid feed was added to the reactor via an externally water cooled atomizer system which was either vertically or horizontally placed.16,34 The aqueous pyrolysis oil fraction was reformed in the fluidized bed.

Besides catalyst attrition (5%/day), also some catalyst deactivation was observed leading to a rising methane concentration which leveled off at roughly

1.5 vol%. Additionally, methane co-reforming experiments were done where, at co-reforming conditions, two times less unconverted methane was observed than when only methane was being steam reformed.

Steam reforming of the whole pyrolysis oil was done by Van Rossum et al. in a bubbling fluidized bed using both a dedicated35 and commercial reforming catalyst.36 Initially, a methane free syngas was being produced but in time the methane content increased till it reached the production level of noncatalytic pyrolysis oil gasification. The catalysts activity was then limited to enhancing the water gas shift reaction, coke/char gasification and some pre-reforming activity for C2-C3 hydrocarbons. The catalyst showed, similar to the catalyst used by the NREL, high levels of attrition.

Because the chemical mechanism of steam reforming oxygenated compounds is different37 than methane and naphtha, many research groups have been trying to develop new catalyst formulations using model compounds of pyrolysis oil and the aqueous phase of pyrolysis oil to produce hydrogen and synthesis gas while minimizing coke formation. The details of these investigations are beyond the scope of this chapter. Interested readers are referred to the following publications: Refs. 38, 39, 40, 41, 42, 43 and 44.

Schwarzer and Gutsche (2005) described a process for preparing citrostadienol — free phytosterols. The patented invention relates to a process for the production of phytosterols by alkali-catalyzed transesterification of DD, neutralization of the catalyst and removal of the unreacted alcohol. The phytosterols are crystallized in hydrocarbon by lowering the temperature, optionally after the addition of an adequate quantity of aqueous methanol, and are then removed and purified by filtration, washing and drying. The resulting products have a citrostadienol content of less than 0.5%.

Pan et al. (2005) described a method to separate phytosterol and synthesize vitamin E succinate (VES) from rapeseed oil deodorizer distillate (RSDD). The RSDD was esterified with methanol in the presence of sulfuric acid and further cooled to room temperature. After storage at -3°C for 12 h, the esterified RSDD was centrifuged and the precipitate was separated as raw sterols (wet cake). The filtrate was used as the raw material to synthesize VES. The content and total recovery of phytosterol was above 85% and 80%, respectively, after the first crystallization and 95% and 45%, respectively, after the second crystallization.

Isolation of tocopherol succinates from sterol-removed, succinated DD mixture by crystallization was investigated by Lin et al. (2007). Membrane technology was also evaluated for its effectiveness to separate tocopherol succinates from mixtures containing sterols and tocopherols. Crystallization was conducted at low temperature with different solvents, including hexane, petroleum ether and a mixture of acetone and methanol (4:1, v/v). The crystallization results showed that recovery of tocopherol succinates from the cake fraction was poor with all solvents tested, with less than 10% of original tocopherol succinates in the raw material being crystallized under conditions employed. Among the solvents tested, hexane was better for the recovery of non-a-tocopherol succinates in the cake fraction. Furthermore, a high property of FFA was co-crystallized along with tocopherol succinates for all solvents used, leading to tocopherol succinates contents in the cake fractions lower than that in the raw material. Two nanofiltration membranes (DS-7 and AP01) were also examined using hexane or petroleum ether as a solvent. The recovery of tocopherol succinates was over 60%. However, their concentration was increased only by 6%. A combined process was then evaluated which include crystallization before succinylation, succinylation, first-stage membrane separation and second-stage membrane separation. The final tocopherols concentration derived from this combined process was as much as that of the original DD.

De Areal Rothes and Verhe (2005) developed a process where acid esterification was applied to corn oil deodorizer distillate (CoDD) in order to increase the ratio of free/esterified sterols that influenced the yield of crystallization. The sterol fraction was further isolated from the FAME by dry fractionation (crystallization). However, more research is required in order to obtain high-purity compounds.

Demirbas and Arin have summarised the characteristics of each form of pyrolysis commonly used.54 Slow or conventional pyrolysis is characterised by relatively low temperatures, slow heating rates and high residence times. The methodology has been used for the production of charcoal for centuries. The methodology is based on using large solid pieces of feedstock (since heating the solid can be performed slowly and there is no requirement for rapid heat transfer) and heating in situ to a set temperature for a period of time.55 Heating rates are of the order of 1°C/min and residence rates around a few seconds or longer. Higher temperatures are around 600-700°C. Typical yields are around an equal division as solid, gas and liquid.54-56

Various gasifier technologies have been developed over many decades and tailored to suit specific needs. These processes operate at pressures from atmospheric to >20 bars and at temperatures between ~700-1500°C. According to the way the feedstock is brought in contact with the gasifying agent, the gasifier is classified into the following types.

16.4.1 Fixed bed gasifier

Fixed bed gasifiers, which consist of a fixed bed of biomass through which the oxidation medium flows in updraft or downdraft configuration, are simple and reliable designs and can be used to gasify wet biomass economically on a small scale for CHP applications (Wang et al., 2008). However, they produce syngas with large quantities of tar or/and char, due to the low and non-uniform heat and mass transfer between the solid biomass and the gasifying agent
(Wang et al., 2008). The product gas must be extensively cleaned before use. Moreover, the throughput for this type of gasifier is relatively low and, therefore, for large-scale applications, as in the case of biomass to liquid (BTL), with very strict requirements concerning the purity of the syngas, fixed bed gasifiers are considered unsuitable.

There are three types of fixed bed gasifiers, as shown in Fig. 16.5: updraft (counter-current), downdraft (co-current), and cross draft gasifier. In case of the updraft gasifier, the fuel is supplied at the top and the air at the bottom so that fuel moves against the air flow, while in the case of the downdraft, air is introduced above the oxidation zone and the product gas is removed from the bottom. In the case of the cross draft, feedstock moves downward while the product gas leaves in a sideways direction. Figure 16.6 shows temperature distribution along the height of the gasifier (see Table 16.2).

In the case of the updraft gasifier, the tar content of the product gas is high, thus it cannot be used directly for the engine applications. As in an updraft gasifier, the pyrolysis zone lies above the combustion zone; the tar formed does not pass through the combustion zone, thus resulting in higher tar content in the product

Moving-bed gasifier (dry ash)

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

gas. This is the opposite of the downdraft gasifiers, where all the pyrolyzed product passes through the oxidation zone, thus product gas has lower tar content. As in the case of the updraft gasifier, the hot gases pass upward so their energy is available to vaporize the moisture. Due to this property, updraft gasifiers can gasify relatively higher moisture content fuel than downdraft gasifiers. Constriction in the oxidation zone of the downdraft gasifier, however, makes its design more complicated and difficult to scale up. Cross draft gasifiers are used mainly for charcoal gasification. However, during the process, the temperature could reach 1500°C, which could lead to material problems (Stassen and Knoef, 2001).

Figure 16.7 shows different designs of downdraft gasifiers. The Imbert type has a narrow constriction near the oxidization zone for efficient combustion of the fuel. Conversely, the stratified type does not have any narrow constriction, making it easier to design and scale up. Another design is the multi-stage downdraft gasifier, which was developed and tested at the Asian Institute of Technology biomass research laboratory. In this type, air is supplied at two stages. Similarly, in the case of the two-stage gasifier, the biomass is first pyrolyzed in a separate zone and then the tar formed is combusted in another gasification zone to supply the heat required for gasification. This type of gasifier can produce product gas with a tar content well below 50 mg/Nm3. In the case of the vortex gasifier, the air is supplied so as to create a vortex that causes the volatile pyrolysis product to move up and in the presence of air become combusted. Although gasification and pyrolysis in the vortex gasifier take place in a single reactor, the tar content in a product gas is similar to that of the multi-stage gasifier (Fock and Thompson, 2001). The multi-fuel downdraft gasifier commercialized in China can be operated with wood, corncobs, hard nut shells, sawdust, and hard coal (see Table 16.3).

Source: Chopra and Jain, 2007.

Cobalt-based catalysts are especially interesting from the commercial point of view due to their rather high activity and selectivity with respect to linear hydrocarbons. Furthermore, they exhibit higher stability, smaller negative effect of water on conversion and higher resistance to attrition in slurry bubble column reactors (Khodakov, 2009). Cobalt catalysts are only used for the LTFT process, as at higher temperatures, excess methane is produced (Dry, 2002). As the cost of cobalt is higher than that of Fe, it is desirable to increase the surface metal exposure, and therefore, Co-based catalysts are mostly supported on high-surface area stable supports such as Al2O3, TiO2 or SiO2 (Oukaci et al., 1999). Zeolites have also been studied as supports (Bessell, 1995). According to a review by Iglesia (1997), the use of support-precursor pairs with intermediate interaction strengths and the slow and controlled reduction of impregnated precursors appears to be the most promising route to the synthesis of supported Co catalysts with high Co concentrations and modest dispersions (0.10-0.15). SiO2 is considered the ideal support for Co FT catalysts, as its high surface area favours high Co dispersion at high Co loadings, while its surface chemistry enables high reduction of Co3+ or Co2+ to Co0 (Dalai and Davis, 2008). The latter is especially important, as metal Co is the active phase for FT and cobalt oxide is reduced at more than 300°C, temperature higher than the LTFT, implying that pre-reduction of the catalyst should take place prior to loading the reactor with consequent increase of cost and complexity. Promotion with small amounts of noble metals, for example Pt, Ru or Re, also enhances the reduction process (Iglesia et al., 1993). Although, in general, cobalt catalysts are less influenced by the presence of promoters than iron-based ones, the presence of noble metals is claimed to increase activity and selectivity to C5+ products via enhancement of the hydrogenolysis of the carbonaceous deposits and thus the cleaning of the catalytic surface (Iglesia et al., 1993).

Biodiesel is produced by reacting vegetable oils or animal fats with a low molecular weight mono alkyl alcohol (in most cases methanol). The so-called transesterification is typically performed at about 60°C in the presence of an alkaline catalyst such as sodium methoxide. For every ten tons of biodiesel produced about one ton of glycerol (or glycerine) is formed as a co-product.

In order to become more competitive and less dependable on politics, the biodiesel industry is looking for cheaper feedstocks, better control over feedstock supplies, improved conversion technology and new ways to increase the value of glycerine.

A significant improvement in conversion technology may come from heterogeneous catalyst systems which give easier catalyst separation, enable higher conversions, and yield a higher quality crude glycerine than current homogeneous alkaline catalysts. Finding new outlets for glycerine is also vital for the biodiesel industry to become more competitive. A promising approach is to convert glycerol to fuel additives, commodity chemicals and polymer building blocks such as 1,2-propanediol; 1,3-propanediol or epichlorohydrin (Pagliaro and Rossi, 2008).

The DTDL.2007G/19 project includes many testing objects such as engines, passenger cars, light duty vehicles,34,37-39 the findings from testing engines are presented in detail as follows.

The testing objects include 02 diesel engines D243, Belarus made. One used market diesel, another used biodiesel B5. These engines are usually used in tractors and fishing boats. Specifications of the testing engines are shown in Table 23.8.

Comparative tests were conducted on load curves and speed curves to investigate impacts of B5 fuel on engine’s performance. To assess exhaust emissions, R49 driving cycles (equivalent to Euro 2 emission standard — the one currently applied for heavy duty vehicle engines in Vietnam) were used for the testing engines.

Table 23.8 Specifications of the testing engines

Specifications

D243

In-line, diesel, four stroke Mechanical direct injection Four

110 mm x 125 mm 4.749 litres 16.4:1

80HP/2200rpm

The two testing engines were also operated within 300 hours durability tests to assess engine components, lubrication oil, as well as engine’s performance and exhaust emissions.

The test-cell used to conduct comparative tests and durability tests is high dynamic engine AVL test-cell for heavy duty vehicle’s engines at Laboratory of Internal Combustion Engine, Institute of Transportation Engineering, Hanoi University of Technology.

Emission bench CEBII was used for gaseous emissions analysis. Particulate matter was sampled by the AVL Smart sampler 743. The testing apparatuses are presented in Fig. 23.19.

638 Handbook of biofuels production

Zeolites are now common catalytic materials in widespread use in the petroleum industry and belong to a wider group of aluminosilicates known as molecular sieves.165 Zeolites are porous solids with pore sizes around 2-10 A and are described by IUPAC as microporous. A typical example is shown in Fig. 14.2. Unlike the aluminosilicates that are present in common clays that have a layered structure, zeolites have rigid honeycomb-like pore structures that can survive many reactions (ion-exchange, hydrogenation, hydroxylation and de-hydroxylation)

without swelling, contraction and collapse of the pore structure. Zeolites are usually hydrated and contain water which is only weakly held. Zeolites are strong solid acids. They are well known for their ion exchange capability which allows the ‘doping’ of the structure by hetero atoms and, thus, their chemical modification. They are crystalline materials with AlO4 and SiO4 units tetrahedrally linked through oxygen anions. In recent times, phosphate groups have been introduced in an effort to generate larger pore systems. Both the atomic and pore structure is periodic and almost 200 zeolite structures have been detailed.166 Zeolites can be synthesised via hydrothermal condensation of cation precursors in the presence of an organic template167 but also occur naturally, and the most important natural zeolite is clinoptilolite.168

The general catalytic chemistry of zeolites is well reviewed169 and, in particular, the hydrocarbon cracking properties detailed.170 Zeolites have high active site densities due to the in balance of anionic charge between the AlO4 and SiO4 units.171 The first found use in petroleum refineries almost 50 years ago where their high activity to hydrocarbon cracking coupled to shape selectivity provided by their pore structure, allowed industrial chemists to achieve greater yields of useful petroleum products.172 The year 1972 saw the advent of a synthetic high silica zeolite catalyst to the energy industry, HZSM-5, that was able to produce aromatic, petroleum-type products from a wide range of hydrocarbon sources.173 This material has become the standard against which most pyrolysis catalysts are assessed. The wealth of data available that demonstrates the effectiveness of zeolite catalysts in pyrolysis is now considerable and, therefore, only some of the more recent data is reviewed here.

Zeolites have been widely applied to the catalytic pyrolysis for waste polymer treatments. Despite the advances in polymer recycling for re-use as construction, consumer and retail materials this can be expensive (shipping to low-cost economies for sorting, energy intensive, use of solvents, environmental issues, etc.), result in low-grade products as well as being technically difficult for some polymer types.174 The opportunity to develop thermolysis methods including pyrolysis (as well as hydrogenation and gasification) for energy generation using waste polymer has been reviewed by Mastellone et a/.175,176 The mechanisms of polymer thermal decomposition are now well accepted.177 Whilst all aluminium containing zeolites have good cracking capability, it is generally found that the larger pore zeolites such as zeolite-Y (0.74 nm) produce less gas and low molecular weight (hydrocarbons with three to six carbons) products and a higher relative content of aromatics (which increase the product octane number and ensures the smooth running of internal combustion engines and use as a transport fuel178) compared to the smaller pore systems such as ZMS-5 (0.55 nm) and zeolite-A (0.4 nm).41,179-182 The reason for the behaviour is apparently due to the increased number of acid sites on zeolite-Y type systems183 as well as the small pore size of ZMS-5 which limits the diffusion of larger hydrocarbon moieties during reaction.184

Zeolite catalysts have also been used for the pyrolysis of biomass and related materials. In a study of the pyrolysis of a synthetic bio-oil showed that HZMS-5 was a better catalyst for aromatic production than a range of other zeolite, transition metal and mesoporous catalysts.133 The catalyst completely removed water and oxygen from the bio-oil in the conditions used but the authors also found that all the catalysts studied increase gas production at the cost of liquid generation. Uzun and Sarioglu also found that zeolite catalysts decreased liquid yield whilst increasing the proportion of gas and this seems to be a general finding from many studies.124 As above for polymeric materials, these authors also found that zeolite-Y increased the aromatic content but did maintain the highest liquid yield.124 Carlson et al. found that ZMS-5 was the catalyst that provided greater aromatic quantise for a range of biomass materials (cellulose, cellobiose, glucose and xylitol) when compared to other zeolites and mesoporous silica.127 Aho et al. studied a range of zeolite-b catalysts over a range of Si:Al ratios of 25-300 and found that the increasing acidity resulted in both greater gas yields as well as the amount of coking.185 These authors also demonstrated that the catalyst was a prerequisite in generation of polyaromatics from biomass. Zeolite catalysts have been widely used for the catalytic pyrolysis of vegetable and plant oils. One of the first studies of HZSM-5 was for the pyrolysis of corn and peanut oil in 1979.186 A high aromatic yield was found with the product mixture being akin to a high-grade petrol (gasoline).186 Similar findings were later reported by Milne and co-workers.173 It is generally thought that the HZMS-5 catalyst is the most effective type of zeolite catalyst for the conversion of vegetable oils to quality fuel materials.187,188 Similar effectiveness was found for the conversion of palm oil.189,190

The preliminary sizes and operating parameters of a fluidized bed gasifier may be estimated as follows:

Main flow rates

Using the desired power output (Q), the volumetric lower heating value (LHVv) of the producer gas and assumed gasifier efficiency, one can calculate the biomass flow rate (F) for production of 1 Nm3 of gas as:

і ■. ■■ [ p:. ‘.I;u I [ І6.35]

LHKng

One can find the required air flow rate, Mair for an air-gasification unit assuming a value of equivalence ratio, ER within the range of (0.2-0.3) (Basu, 2006)

Mair = mth. ER. F kg air/Nm3 product gas, [16.36]

where mth is the stoichiometric amount of air kg/kg and F is the fuel feed rate.

For steam gasification gasifiers, one can use published data on the steam/ biomass ratio for a similar fuel to find the amount of the gasification medium Mmed as the first guess.

As the gasifying agent will also be the fluidizing gas, it is necessary to check that it would be adequate for proper fluidization. The fluidizing velocity is chosen based on the bed particle characteristics taking into consideration the bed hydrodynamics and the process occurring in gasifier.

Reactor dimension

The volume flow rate of the gasifying medium (steam/air or oxygen) divided by the chosen fluidizing velocity will give the reactor cross-section area.

л. [16.37]

bed л r r

PmJJ

where pmed and U are the density and the fluidization velocity of the medium (air, oxygen, steam, or their mixture) at the operating bed temperature, respectively.

Bed height

The bed height should be chosen such that it provides the required residence time for better carbon conversion and would avoid slugging. Also, its selection is governed by operating cost considerations, as higher bed height means a higher pressure drop, taller reactor, and greater auxiliary power consumption.

Freeboard height and its diameter

Ideally, the freeboard height should exceed the transport disengaging height (TDH), so that the particle entrainment with upward flowing gases will be low, but in most cases, that is too expensive. A compromise between cost and performance is used.

Besides the diesel main product, naphtha, a gasoline fraction of less value is produced as a by-product. Straight-run FT naphtha has a low octane number, is olefinic and has high levels of oxygenates (Gregor and Fullerton, 1989). The chemical composition of two naphtha streams produced via LTFT and HTFT process is summarized in Table 19.7. Currently, the BTL-FT synthetic naphtha is rather sold as a low-cost chemical feedstock and cannot be used as a fuel. Untreated naphtha can also be used as an energy source for the production of heat and power or can be alternatively reformed on-site to synthesis gas and fed to the FT reactor to increase the process yield (Bienert, 2007). In the frame of the EU-funded NICE (New Integrated Combustion System for Future Passenger Car Engines) project, Renault/Regienov and Volkswagen tested naphtha fuels in experimental homogeneous charge compression ignition (HCCI) engines and found significant improvements compared to standard diesel fuel (RENEW, 2008). In this context, although BTL-FT naphtha is not a suitable fuel for conventional engines, it may be advantageous for future power trains like HCCI

and combined combustion system (CCS) being even more efficient and having less emissions. It should, however, be mentioned that the requirements for these future engines are not clear for the time being.

Even though the light FT by-product naphtha is not suitable for application as a fuel in its present form and in conventional gasoline engines, it could be upgraded by an additional isomerization or reforming unit to boost its octane number and fulfil the above, as discussed in Section 19.5.3. It should be noted that the production of finished gasoline blendstock is not yet considered because of the added cost and energy expenditures associated with upgrading naphtha to gasoline with the current technology.