Fixed bed gasifiers have a relatively low throughput and therefore for large-scale applications, as in the case of BTL, with very strict requirements concerning the purity of the syngas, are considered unsuitable (Wang et al., 2008). Based on throughput, complexity, cost, and efficiency issues, circulating fluidized bed (CFB) (Hamelinck et al., 2004; Wang et al., 2008; Tijmensen et al., 2002; Zhang, 2010) and entrained flow gasifiers (van der Drift et al., 2004) are very suitable for large-scale syngas production. Examples of CFB gasifiers employed for the gasification of biomass that have reached a certain degree of commercialization are the Lurgi circulating fluidized bed process, the Foster Wheeler gasifier, the VVBGC gasifier constructed under the EU-funded project Chrisgas, the UCG (Ultra Clean Gas) programmed by VTT, etc. (Higman and van der Burgt, 2008). Slagging entrained flow gasifier manufacturers are Shell, Texaco, Krupp-Uhde, Future-Energy (formerly Babcock Borsig Power and Noell), E-gas (formerly Destec and Dow), MHI (Mitsubishi Heavy Industries), Hitachi, and Choren (formerly UET) (van der Drift et al., 2004).

The biomass gasification technology that came closer to commercialization for syngas production in a BTL-FT plant was the Choren Carbo-V patented biomass gasification process (Fig. 18.3). The process is a good example of the application of entrained flow gasifiers in the BTL process and was used in the first demonstration 15,000 tons per year BTL plant in Freiberg Germany, coupled with the Shell SMDS (Shell Middle Distillate Synthesis) Fischer–Tropsch process (Rudloff, 2005). The Choren Carbo-V patented gasification process consists of three stages: low temperature, high temperature, and endothermic entrained-bed gasification (Rudloff, 2005). During the first stage, the biomass is continuously carbonized through partial oxidation with oxygen at temperatures between 400 and 500°C, ie, it is broken down to a tar-containing gas (volatile parts) and solid carbon (char). The tar-containing gas is then fed to the high-temperature gasifier, where it is partially oxidized using oxygen as the gasification agent. The heat, which is released as a result of the oxidation process, warms up the carbonization gas to temperatures that exceed the ash melting point of the fuels that have been used, ie, 1300–1500°C. At these temperatures any unwanted longer-chain hydrocarbons, eg, tar and even methane, are broken down. The gas that is produced primarily consists of carbon monoxide, hydrogen, carbon dioxide, and steam. The char from the low-temperature gasifier is cooled, ground down to pulverized fuel, and is then blown into the stream of hot gas coming from the combustion chamber in the entrained flow gasifier. A huge amount of heat is absorbed when gasifying the char and this allows lowering the temperature of the gas to 800–900°C in a matter of seconds. This “chemical quenching” process produces a tar-free gas with a low methane content and with high proportions of combustible carbon monoxide and hydrogen. Unfortunately, the company declared insolvency in 2011 as described in detail in Section 18.5. In 2012, Linde acquired the Carbo-V^® Technology of the insolvent Choren Industries GmbH and all related patents and trademarks. According to their press release (Linde, 2012) Linde plans to offer “the Carbo-V^® Technology as licensor and also as an engineering and contracting company for commercial projects on a strongly growing market”.

A more successful example is the GRE (Güssing Renewable Energy) multifuel gasification plant, located in Güssing, Austria. The plant is the world's first functioning fast internally circulating fluidized bed (FICFB) gasification plant. The gasification part of the plant consists of two interconnected fluidized bed systems, as shown in Fig. 18.4. Biomass is ground into small pieces and is fed to the reactor where it is gasified in anaerobic conditions at 850°C in the presence of steam. The bed material in the reactor is olivine which acts as a heat transfer medium and provides a stable temperature in the gasifier. After gasification, the synthesis gas is purified and cooled. Syngas cleaning includes dedusting with a bag filter and washing to reduce the concentration of tar, ammonia, and acid gas components. According to the company (Güssing Renewable Energy, 2015), the special syngas technology makes it possible to redirect all side products back into the process, as a result of which no waste or waste water is created during gas purification. Moreover, the product gas is completely free from nitrogen. The use of steam as a gasification agent in the FICFB concept results in a product gas with high quality (low nitrogen, optimal H₂/CO ratio, high heating value) without the need for pure oxygen. It also enables different syngas applications to be realized at smaller scale (10–100 MW fuel) (Rauch, 2014). A part of the remaining coke is taken to the burning section with the help of circulating channel material used as a carrier, where it is burnt (see Fig. 18.4). The produced heat from combustion raises the temperature of the carrier that is transferred back to the gasifier to maintain the temperature of the gasification reaction at the desired level. The flue gas is discharged separately, its heat content entering the district heating network. The Güssing plant has been operating at the commercial demonstration scale of 8 MW thermal input (50 dry tons per day) producing 2 MW electricity and 4.5 MW of district heat for more than the last 12 years with 70,000 h of operations through 2012. There are several further plants already built and in operation exploiting the developed gasification technology, as shown in Table 18.1 (Rauch, 2014). Besides the production of heat and power, a Fischer–Tropsch pilot plant has also been built in the CHP plant to demonstrate the use of the bio-derived syngas for fuel production. The specifics of the whole BTL process chain are discussed in Section 18.5 which presents the main advancements toward commercialization of the biomass-to-liquids process.

The presence of impurities in the syngas produced by the gasification step is inevitable. Syngas contains different kinds of contaminants, such as particulates, condensable tars, BTX (benzene, toluene, and xylenes), alkali compounds, H₂S, HCl, NH₃, and HCN. The catalysts employed in the FT reactor for the synthesis of the liquid fuels are notoriously sensitive to such impurities, and especially sulfur and nitrogen compounds, which irreversibly poison FT catalysts. Alkaline metals and tars deposit on catalysts and contaminate the products, while particles cause fouling of the reactor. Therefore, extensive cleaning of the syngas is required prior to entering the FT reactor. Moreover, the concentration of inert gases (ie, CO₂, N₂, CH₄, etc.) must be approximately less than 15 vol% (Boerrigter et al., 2004). Indicative syngas specifications for Fischer–Tropsch synthesis are shown in Table 18.2 (Boerrigter et al., 2004).

The first step in all syngas cleaning configurations considered so far is the removal of BTX and larger hydrocarbons, the tars. BTX should be removed upstream of the active carbon filters in the syngas cleaning train, as active carbon adsorbs BTX and would therefore require frequent regeneration, reducing process reliability. Tars normally condensate at the typical FT reactor conditions and foul downstream equipment, coat surfaces and enter pores in filter and sorbents. Therefore, tars should be removed to a concentration below condensation point at the operating pressure of the FT reactor (Hamelinck et al., 2004). Three processes can be used for tar removal. Thermal cracking of tars involves high temperatures, 1000–1200°C, and tars are cracked in the absence of catalyst with the use of steam or oxygen. However, thermal cracking has low thermal efficiency, requires expensive materials and results in the production of large amounts of soot. Catalytic cracking/reforming of tars in the presence of dolomite/olivine, nickel-based catalysts, or alkalis (Wang et al., 2008) overcomes these limitations. Still, this technology is not yet proven and costs are increased due to catalyst consumption (Milne et al., 1998). Alternatively, tars can be removed at low temperature by advanced scrubbing, using a special organic washing liquid (“oil”). Such a system has been developed by ECN that patented the OLGA tar removal technology (Boerrigter et al., 2004). It should be mentioned that the use of entrained flow gasifiers removes the need for a tar cracking/removal step as the high gasifier operating temperatures (1300–1500°C) yield a tar-free syngas.

After the removal of the tars, the other contaminants can be removed from the syngas by either the conventional “wet” low-temperature or the “dry” high-temperature cleaning. The wet gas cleaning technology is proven and has been well commercialized for large-scale coal gasification systems (Zhang et al., 2007). The general approach involves the quenching of the raw hot gas with water to cool the gas and remove solid particles (eg, dust, soot, ash) and the volatile alkaline metals (Boerrigter et al., 2004). NH₃ is then removed by a water washer along with halides and H₂S is removed either by absorption or the Claus process to elementary S. In the final step, the gas passes through ZnO and active carbon filters, which remove H₂S and remaining trace impurities and act as guard beds for the FT catalyst. Although proven, this technology has efficiency penalties and requires additional wastewater treatment. Many research efforts have been focused on the development of dry hot syngas cleaning processes, which appear to be potentially more efficient and cleaner than the proven conventional wet technology (Sharma et al., 2010). Hot gas cleaning consists of candle or ceramic filters for removing solid contaminants and sorbents for fluid contaminants, through which the high temperature of the syngas can be maintained, achieving efficiency benefits and lower operational costs. Dry gas cleaning can be especially advantageous when preceding a reformer or shift reactor, as these processes have high inlet temperatures. However, as aforementioned, the performance and reliability of the filters and sorbents has still to be proven at high temperatures, especially above 400°C, for a commercial implementation of the dry gas cleaning technology. Recent developments and critical review of the different syngas cleaning technologies have been published and can be found in Sharma et al. (2008, 2010).

After the gas cleaning train, the biomass-derived syngas has to be conditioned in order to adjust the H₂/CO ratio to that required for the FT reactor. Typical conditioning includes steam reforming of methane and light hydrocarbons to CO and H₂ over a nickel catalyst, following by a water gas shift reactor. Finally, as the concentration of inert gases must be kept below 15 vol% (Boerrigter et al., 2004), CO₂ is removed with amine treating. The purified and conditioned synthesis gas is then compressed to the required pressure and is fed to the FT reactor.

An extensive amount of research has been performed on several aspects of the Fe and Co catalysts, including fundamental, basic and applied research. These efforts include investigation of the effect of promoters, supports, additives, pretreatments, preparation, and generally all chemical and physical properties of the materials in order to increase catalyst activity, enhance selectivity to the desired products, inhibit formation of unwanted products, especially methane, and improve resistance to sulfur poisoning. A summary of improved modified Fe and Co catalysts employed in industry for the FT process is presented in Table 18.3 (Bartholomew, 1990).

Conventional hydrocracking takes places over a bifunctional catalyst with acid sites to provide isomerization/cracking function and metal sites with hydrogenation–dehydrogenation function. Platinum, palladium, or bimetallic systems (ie, NiMo, NiW, and CoMo in the sulfided form) supported on oxidic supports (eg, silica-aluminas and zeolites) are the most commonly used catalysts, operating at high pressures, typically over 10 MPa, and temperatures above 350°C.

In recent years, considerable research has been ongoing to investigate the effect of the operating conditions, both experimentally (Calemma et al., 2005, 2010; Rossetti et al., 2009; Rosyadi et al., 2011) and computationally (Pellegrini et al., 2004; Fernandes and Teles, 2007), and the catalytic material on the yield and quality of the FT-wax hydrocracking products. Concerning the operating conditions, it was found that wax hydrocracking requires milder pressure and temperature, as the paraffinic nature of the wax implies higher availability of hydrogen in the unit (little hydrogen consuming aromatics) and thus suppressed coke formation (de Klerk, 2008). FT-wax hydrocracking to middle distillates is favored at pressures ranging from 3–5 MPa and temperatures between 250–300°C (Calemma et al., 2010) and yields a product containing light paraffins up to C₂₄, as presented in a product sample chromatograph obtained from FT-wax hydrocracking experiments performed in CPERI (Fig. 18.8). At these conditions, middle distillate yield (C₁₀–C₂₂) reaches up to 80–85 wt% at intermediate conversion levels (∼60 wt%) (Calemma et al., 2010). At higher conversions, a small reduction in the middle distillate yield can be observed, indicating an increase of consecutive hydrocracking reactions leading to lighter products. Still the consecutive reactions are limited, allowing the reaction to be carried out at high conversions without lowering significantly the middle distillate selectivity (Calemma et al., 2010).

Extensive work has also been conducted by our group as part of the EU-funded IP RENEW project which explored technology routes for the production of BTL fuels (Lappas et al., 2004). More specifically, the operating conditions (temperature, pressure, H₂/oil ratio) were investigated in experiments with different commercial hydrocracking catalysts in a specially designed hydroprocessing pilot plant unit. The main conclusions were that with all catalysts hydrocracking temperature appears to play the most important role and influences significantly the product yields, as shown in Fig. 18.9. It was shown that the yields of naphtha and kerosene in the product increase as the temperature increases and so does the conversion. However, the diesel yield is maximized at a certain temperature and then decreases as a result of higher conversions achieved at higher temperatures (RENEW, 2008). Moreover, it was shown that the yield of gasoline and diesel in the product decreases as the H₂/oil ratio decreases and so does the conversion. The diesel selectivity is also slightly decreased as a result of the decreasing yield and conversion. Studies by Calemma et al. (2010) showed additionally that the composition of FT diesel, specifically the ratio of iso- and n-paraffins, is also influenced by the operating parameters.

The nature of the catalyst also affects significantly the product quality and yield. Experiments performed in CPERI with three different commercial hydrocracking catalysts showed measurable differences in diesel selectivity at isoconversion as a function of the catalytic material (Fig. 18.10) (RENEW, 2008). Catalysts loaded with a noble metal (particularly Pt) were reported to show better performances in terms of selectivity for hydroisomerization and products distribution in comparison with non-noble metal-based catalyst (Archibal et al., 1960; Gibson et al., 1960). Calemma et al. (2001) reported high diesel selectivities obtained over a Pt/SiO₂-Al₂O₃ catalyst during the hydroprocessing of FT waxes and attributed the observed results to the mild Brönsted acidity, high surface area, and pore size distribution of the support. Similar conclusions were also reached in a recent collaborative work between our group in CERTH and the University of Alicante in Spain (Iliopoulou et al., 2015). The work aimed at investigating the effect of acidity and mesoporosity in low loading (0.1 wt%) Pt/based catalysts supported on conventional microporous ZSM-5 zeolite or micro/mesoporous ZSM-5 prepared by controlled desilication/acid extraction, BETA zeolite and a commercial, amorphous, mesoporous silica/alumina support on the hydrocracking/hydroisomerization of FT waxes using n-hexadecane as the model compound. Hydroisomerization experiments performed at 30 bar, 275°C and WHSV 4 h⁻¹ demonstrated that the activity of the catalysts reduces in the following order Pt/BETA > Pt/ZSM-5 > Pt/ZSM-5 (meso) >> Pt/ASA. However, all zeolites led mainly to the production of a high amount of light gases and C₅–C₁₀ cracking products. This behavior was attributed to a combination of the acidic and textural properties of the zeolitic supports. Although the microporous/mesoporous ZSM-5 and BETA supports had a lower number of Brönsted acid sites than microporous ZSM-5 and were thus expected to have lower cracking activity, the opposite was observed. The increased mesoporosity and the much larger pore size of the two first zeolites seems to enhance the accessibility of hexadecane to the acid sites, leading to increased conversion and cracking reactions. It is interesting to note however that the ratio of iso/normal alkanes in the cracked products was much higher for the Pt/BETA catalyst. This could also be attributed to the mesoporosity of BETA which allows more space for skeletal rearrangement. The deposition of Pt on an amorphous support such as silica-alumina with negligible Brönsted and high Lewis acidity led to a catalyst with much lower activity than its zeolitic counterparts, but with very high selectivity toward iso-hexadecane, which was the only liquid product formed in the case of Pt/ASA.

Concerning the type of active metal, Zhang et al. (2001) showed that Pt performs better than Ni and Pd supported on tungstated zirconia for the hydroisomerization of the model compound n-hexadecane. The use of hybrid catalysts based on Pt/WO₃/ZrO₂ with addition of sulfated zirconia, tungstated zirconia, or mordenite zeolites was studied by Zhou et al. (2003). According to the authors, hybrid catalysts based on Pt/WO₃/ZrO₂ provide a promising way to obtain higher activity and selectivity for transportation fuels from FT products. Given the high cost of noble metals, hydroprocessing of FT waxes has also been studied over nickel catalysts (de Haan et al., 2007). De Haan et al. (2007) demonstrated the benefit of using non-sulfided nickel catalysts. In conventional hydroprocessing units, catalysts are sulfated to avoid poisoning by the sulfur species in crude oil. However in the case of the sulfur-free FT waxes, use of a sulfided catalyst implies the continuous addition of sulfur-containing compounds to avoid catalyst deactivation (de Klerk, 2008). Other advantages of developing a non-sulfided catalyst for the hydrocracking of FT waxes are a simplified, less costly and environmentally friendly process (no H₂S in the tail gas) (de Haan et al., 2007). Nickel supported on a commercial silicated alumina yielded results which compare favorably with those of a commercial sulfided NiMo catalyst, with diesel selectivities of 73–77% at a conversion of ∼52% (de Haan 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 sulfur and nitrogen content, in contrast to FT waxes that are highly paraffinic with extra-low aromatics content (<1 wt%) and virtually zero sulfur (<5 ppm) (see Table 18.4). The development therefore of new catalyst formulations, as well as optimization of the overall process parameters, are both 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. 18.11, 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 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 18.5, 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 higher octane number; however the RON number of the wax gasoline is still acceptable. The RON of the wax gasoline was almost constant and independent of the conversion due exactly 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 RON 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 sulfur and nitrogen concentrations, leading to the production of very clean gasoline. Moreover, it was found that the diesel-range LCO product produced from the catalytic cracking of FT waxes is better than that 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 sulfur 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 and Al-MCM-41) aluminosilicates as catalysts or active matrices in the cracking of Fischer–Tropsch waxes toward the production of liquid fuels. Focus was placed on the effect of porous and acidic characteristics of the materials on product yields and properties. According to the authors, the type of catalyst plays a significant role in the product selectivities. The percent conversion of wax, the product yields (gasoline and LPG), and the research octane number (RON) of the produced gasoline are shown in Fig. 18.12 for different investigated microporous and mesoporous catalysts. The behavior is typical for the two zeolitic catalysts when used in fluidized catalytic cracking (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. 18.12 for the cracking of F-T 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 amorphous silica-alumina (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 XRD amorphous aluminosilicate phase. Fig. 18.13 shows the yields (wt% on feed) of various gasoline components. The data in Fig. 18.13 can also be used for a qualitative comparison of catalyst performance with regard to their selectivity toward specific gasoline components, especially in the case of H-Y- and H-ZSM-5-based catalysts, which showed similar percent conversion of wax (Fig. 18.12). The H-Y-st. catalyst presented a significant selectivity toward 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 toward the production of environmentally friendly fuels in accordance with EU regulations. Olefins were also higher with the H-Y-st. catalyst (15 wt% on feed) compared to the rest 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. 18.12) were mainly attributed to the high aromatics content; while in the case of Η-Υ-st. catalyst the high RON (∼87) was mainly attributed to the relatively high C₅–C₇ olefin and iso-alkane yields. The 3%-crystalline 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 (81). The yield of aromatics with the Al-MCM-41 sample was very low but 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 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 F-T waxes.

In the frame of the European project RENEW, BtL-naphtha was also identified to have suitable properties for use in future power trains like homogeneous charge compression ignition engines. It was however found that further upgrading of the naphtha fraction is needed for optimized engine performance, targeted toward mild reduction of its cetane number via isomerization (RENEW, 2008). Through its participation in the EU projects RENEW and OPTFUEL, our group in CERTH established hydroisomerization as a viable upgrading option for BtL-naphtha on both laboratory (Heracleous et al., 2013; Iliopoulou et al., 2014) and pilot scales (Iliopoulou et al., 2012). Investigation of the effect of the zeolitic support (mordenite, ZSM-5, and beta zeolite) on the isomerization performance of a series of low-loading (0.1 wt%) Pt catalysts showed that the use of ZSM-5 leads to the synthesis of the most active material with satisfactory selectivity to isoparaffins. The superior performance of Pt/ZSM-5 was attributed to its high Brönsted acidity and the formation of homogeneously dispersed, cubic-shaped and highly crystalline Pt particles on the zeolite surface, as shown in the TEM image in Fig. 18.14 (Iliopoulou et al., 2014). In an effort to further increase selectivity to large isoparaffins, we prepared and tested a low-loading Pt catalyst supported on a hierarchical ZSM-5 support. Introduction of mesoporosity was achieved by treating the support with alkaline and acidic solutions (Heracleous et al., 2013). Fig. 18.15 shows selectivity to C₄–C₈ isoparaffins at 40% conversion in the hydroisomerization of surrogate naphtha—simulating the composition of BTL-naphtha—at 30 bar and temperature range 240–300°C. A clear increase in i-C₈ species at the expense of i-C₅ and i-C₆ is apparent for both the alkaline- and acid-treated Pt catalysts, showing clearly that the introduction of mesopores in the ZSM-5 structure improves selectivity, possibly due to the improved diffusivity within the zeolite pores and the reduced residence time of the intermediates within the pore channels.

As discussed above, the upgrading of the BTL naphtha byproduct into gasoline/HCCI fuel can greatly improve the economics of the overall BTL-FT process. However, this option has yet to be considered in commercial operations. Kreutz et al. (2008) argue that it is still uncertain whether the additional gasoline blending stock value can justify the great capital and operational costs that these upgrading units impose on the BTL-FT process.