The main catalytic challenges in GTL (gas-to-liquids) processes

Abstract

In the present review the main catalytic challenges for GTL processes are discussed. It is considered that GTL comprises three main catalytic areas, namely synthesis gas generation, Fischer–Tropsch synthesis and upgrade. Each one is analysed and the main characteristics of traditional and innovative catalysts are presented. For syngas generation, steam methane reforming, non-catalytic partial oxidation, two-step reforming, autothermal reforming and catalytic partial oxidation of methane are discussed. For Fischer–Tropsch, we highlight the role of nanocatalysis, hybrid zeolite-containing catalysts, diffusion limitations and selectivity to high molecular weight hydrocarbons. Also, new reactors technologies such as micro reactors are presented. Finally, special attention is paid to the main upgrade steps (Hydrocracking and Hydroisomerisation/Dewaxing), the new mechanisms of isomerisation being discussed for bifunctional zeolitic catalysts.

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Eduardo Falabella Sousa-Aguiar

Eduardo Falabella Sousa-Aguiar, Chemical Engineer, MSc, DSc, has 35 years experience in Catalysis. He has been Professor in the Federal University of Rio de Janeiro for 30 years and a Senior Advisor in Petrobras Research Centre (CENPES), where he is currently the coordinator of XTL projects. He has authored over 300 scientific papers and two books, having advised over 30 MSc and PhD theses. He has been the Brazilian focal point for the international program CYTED, being also an adviser for ICS-UNIDO. He has received many awards; deserving particular attention is the prestigious Brazilian National Technology Award received in 2008.

Fabio Bellot Noronha

Fabio B. Noronha received his B.S. degree from Federal University of Rio de Janeiro in 1987, M.Sc. degree in 1989, from COPPE/Federal University of Rio de Janeiro and Ph.D. degree in 1994 from COPPE/Federal University of Rio de Janeiro and Institut des Recherches sur la Catalyse—Lyon, France. In 1996, he joined the Catalysis group of National Institute of Technology (INT). He worked in a postdoctoral position with Prof. Daniel Resasco at Oklahoma University from 1999–2000. He has been involved in studies for conversion of natural gas and biomass to hydrogen, syngas and fuels.

Arnaldo Faro

Arnaldo C. Faro Jr. received his B.S. degree in Industrial Chemistry from the Federal University of Rio de Janeiro in 1968 and his Ph.D degree in 1984 from the University of Edinburgh, Scotland. In 1969, he joined the PETROBRAS R&D Centre, CENPES, where he worked at the Catalyst Division up to 1993, mainly in the development of hydroprocessing catalysts. In 1994 he joined the Institute of Chemistry of the Federal University of Rio de Janeiro, where he works up to the present date as an Associate Professor in the Physical Chemistry Department and is head of the Heterogeneous Catalysis laboratory.

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1. Introduction: the GTL process

The increasing necessity for clean-burning fuels—with very low or even no sulfur, with the minimum content of aromatics and with minimum formation of nitrogen oxides, soot and unburned hydrocarbons—is changing in a rather drastic way the traditional goals of the refining industry. Although the generation of cleaner fuels may be achieved by introducing further processing capacity in refineries, such as desulfurisation, these new units are very energy consuming and reduce the overall thermal efficiency of the refinery. Still, refinery processes are very energy efficient even with deep desulfurisation of diesel. Nevertheless, desulfurisation processes usually require hydrogen, whose production via the shift reaction also produces CO₂. Hence, an improvement in air quality during the use of cleaner diesel may come at the expense of higher greenhouse-gas emissions during such diesel production. Therefore, the search for alternative feedstock such as natural gas or biomass has become a must in order to cope with more stringent regulations. In this new peculiar scenario, catalysts play an outstanding role.

Regarding potential feedstock, natural gas seems to be the most attractive one. After convenient processing, natural gas is essentially sulfur-free. Thus, one may expect a long term change from a predominantly oil-based refining industry towards an increasing dependence on natural gas. Since natural gas main component is methane, a rather inert molecule, one may also expect a great interest in the so called C1 chemistry, which has been traditionally carried out viacatalytic routes.

Two main catalytic areas of interest may be identified when natural gas is the main raw material. The first one concerns the transformation of natural gas or any methane-rich feedstock into syngas, a predetermined mixture of hydrogen and carbon monoxide. Such syngas may then (a) undergo the Fischer–Tropsch synthesis to produce a range of hydrocarbons in the form of a synthetic version of crude oil, a route known as “traditional GTL”; (b) be transformed into other gases (GTG, gas-to-gas), of which dimethyl ether (DME) is surely the main representative. Such routes, however, may have an elegant alternative which is the activation of methaneviahalogenation, aiming at generating either DME or olefins for petro chemistry. Finally, there have been many efforts to make viable the direct transformation of methane (from natural gas) into higher molecular weight hydrocarbons, particularly aromatics (non-traditional GTL). This last route would prevent the installation of the somewhat expensive syngas generation unit, which still represents the major contribution to the total cost of a GTL traditional process. Although some of these routes have been used for several years, it must be borne in mind that they still have many catalytic challenges.

The GTL process is surely the most important commercial route to produce higher molecular weight derivatives from natural gas. As depicted in Fig. 1, the GTL process normally comprises three steps,^1–3 namely reforming, Fischer–Tropsch synthesis and upgrading. The reforming step aims at generating syngas through the reaction of methane with water, although the reaction with CO₂ (dry reforming) may also take place. More recently, a new process to produce syngas denominated “autothermal reforming” has arisen. Such a process includes partial oxidation of methane, a very exothermic reaction, as a way of improving the thermal efficiency of the reforming step. Fischer–Tropsch synthesis is the second step and promotes the polymerisation of syngas into diesel, naphtha, paraffin and others. Finally, the upgrading step,⁴ which may include hydrocracking, hydrotreating and hydroisomerisation, whose intention is to either maximise diesel and naphtha production from paraffinic compounds, or to generate high quality lubricants and food grade wax.

Fig. 1 Main steps of a traditional GTL process.

Eventually, it must be borne in mind that GTL products are synthetic; therefore they present very high quality. GTL diesel is essentially sulfur-free, has very high cetane number (over 70) and very low aromatics content. Lubricants also display outstanding properties and may be compared to Type 4 lubricants,^1–3 the best there are.

2. Synthesis gas production

2.1 Large-scale syngas production technologies for GTL

Nowadays, large-scale syngas production technologies find widespread use in the manufacture of hydrogen in refineries, in the production of gasoline and diesel containing very low sulfur levels as well as in the synthesis of chemicals such as ammonia, methanol, formaldehyde and acetic acid.^5,6 Syngas may be generated from different feedstocks, including natural gas, shale gas, naphtha, residual oil, petroleum coke, coal and biomass.

Considering GTL applications, natural gas is the preferred choice of feedstock due to the high capital costs of GTL technology that determines the economics of the process.^7,8 In particular, natural gas with low value is desired such as the associated gas, the so-called stranded or remotely located gas reserves as well as the large gas reserves.⁷ The use of the associated gas may also contribute to a reduction in the amount of flared gas.

The main technologies for producing syngas from natural gas for GTL applications are: steam methane reforming (SMR), non-catalytic partial oxidation (POX), two-step reforming and autothermal reforming (ATR).⁷Fig. 2 shows a scheme representing SMR, two-step reforming and ATR⁹ while Table 1 summarizes the main reactions occurring in the different syngas production processes.¹⁰

Fig. 2 Technologies for syngas productions: (a) steam reforming of methane; (b) two-step reforming; (c) autothermal reforming. (Reprinted with permission from ref. 9. Copyright 2000 Elsevier.)

Table 1 Synthesis gas reactions¹⁰

Reactions	ΔH298^0/kJ mol^−1
Steam Reforming ( SMR )
CH4 + H2O → CO + 3H2	206
CO + H2O → CO2 + H2	−41
Catalytic partial oxidation ( CPO )
CH4 + 1/2O2 → CO + 2H2	−38
Autothermal reforming (ATR)
CH4 + 1.5O2 → CO + 2H2O	−520
CH4 + H2O → CO + 3H2	206
CO + H2O → CO2 + H2	−41

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SMR is the most widely used commercial technology for syngas and hydrogen generation (Fig. 2a). SMR involves the endothermic reaction between methane and steam, which requires high temperatures to achieve maximum conversion. The reaction is performed in a fired tubular reactor filled with a Ni based catalyst supported on α-alumina containing a variety of promoters. High steam-to-carbon (S/C) ratios are used in the feed to inhibit carbon formation in the catalyst and thus, SMR produces a syngas with a high H₂/CO ratio (Fig. 3), making it well suited to H₂ production applications.¹⁰

Fig. 3 H₂/CO ratio from different syngas technologies. (Reprinted with permission from ref. 10. Copyright 2004.)

POX is based on the exothermic non-catalytic reactions of methane and oxygen inside a combustion chamber. This technology is very flexible, operating with different feedstocks besides natural gas.⁸ Since a catalyst is not used, carbon formation is not a problem and thus, steam is not required, reducing the CO₂ content in the syngas. However, it requires high operating temperatures (1300–1400 °C) for obtaining high methane conversion and for reducing soot formation. Another disadvantage of POX is high oxygen consumption, which significantly impacts the costs of a syngas plant since an air separation unit (ASU) is required. In addition, the H₂/CO ratio obtained is below 2 (in the range of 1.7–1.8) (Fig. 3), which may not be suitable for some industrial applications such as GTL.^8,10,11

However, these technologies are not capable of producing syngas with the desired H₂/CO ratio for GTL applications (Fig. 3).¹⁰ For the low-temperature Fischer–Tropsch process (LTFT), the H₂/CO ratio required is about 2. One approach to achieve the H₂/CO ratio suitable for FT synthesis is to use both technologies (SMR and POX) in parallel. The two streams containing syngas with different compositions are thus mixed in order to achieve the desired H₂/CO ratio. For example, the Shell plant in operation in Bintulu, Malaysia, operates in this manner.

Another combination of steam reforming and partial oxidation is the so-called two-step reforming (Fig. 2b). It comprises the primary steam reforming that takes place in a fired tubular reformer (as previously described for SMR), and the secondary steam reforming, which occurs in an adiabatic reactor. In the secondary reformer, the unreacted methane in the exit gas from the primary reformer (10–13%) is further converted to hydrogen and CO. The size of the steam reformer is reduced but usually requires oxygen. In the case of a GTL plant, the thermal energy necessary for the endothermic reaction in the secondary reformer is provided by the addition of pure oxygen, which reacts with methane, hydrogen and CO.¹¹ The Mossgas plant in South Africa uses the two-step reforming process to produce syngas.

Another strategy is ATR, which combines an endothermic reaction (SMR) and an exothermic reaction (POX) in the same reactor (Fig. 2c). In fact, the concept of the autothermal reactor is very similar in many aspects to the secondary reformer of the two-step reforming process. The main differences concern the burner and reactor design due to the different feed composition in each reactor.⁷

Currently, ATR or a combination of ATR and steam reforming (pre-reformer and/or heat exchange reformer) is the preferred technology for large-scale GTL plants due to the economies of scale.⁷ This is the technology of the Oryx plant joint venture between Qatar Petroleum and Sasol in Qatar. Therefore, ATR will be described in greater detail. A typical process flow diagram for ATR is shown in Fig. 4 and is comprised of various steps: adiabatic pre-reforming, ATR and heat recovery.^8,12,13

Fig. 4 Process diagram flow for ATR. (Reprinted with permission from ref. 13. Copyright 2009 Elsevier.)

In the adiabatic pre-reformer, the steam reforming of higher hydrocarbons present in the natural gas produces a mixture of methane, hydrogen, CO and CO₂.¹² The presence of a pre-reformer in the process reduces the consumption of oxygen since a higher preheat temperature to the ATR may be used.

The autothermal reformer consists of a burner, a combustion chamber and a catalytic bed in a refractory lined steel vessel.¹² In the ATR reactor, methane is partially burned with oxygen in the burner and in the combustion chamber. Therefore, ATR also requires an ASU albeit smaller than for POX. All oxygen is consumed in these regions. Steam and CO₂ reforming of the unreacted methane and the shift reaction occur in the reactor bed, which contains a Ni/MgAl₂O₄ catalyst.

Before introducing syngas from the ATR into the FT reactor, it must first be cooled, for example, by producing saturated high-pressure steam in boilers. An alternative is to use the process heat of the exit gas from the ATR reformer for steam reforming in a heat exchange type reformer (gas heated reforming—HTER) and for preheating the feed gas to the ATR reformer.¹³ In this reformer, heat exchange is mainly by convection, resulting in lower heat fluxes than in tubular reformers. The HTER may be combined with ATR in series or in parallel.¹³ The main problem of this technology is metal-dusting corrosion.

The composition of the syngas produced by ATR may be controlled through judicious selection of process conditions. The optimal H₂/CO ratio for GTL plants can only be achieved through recirculation of CO₂ or a CO₂ rich off-gas, which reduces the amount of steam in the feed. Operation at low steam-to-carbon (S/C) ratios not only improves the syngas composition but also reduces CO₂ recycle, which decreases the capital investment and energy consumption.^11,12 However, a corresponding reduction in the S/C ratio favors carbon formation in the pre-reformer and soot formation in the ATR reactor.¹²

Carbon formation on the pre-reformer at low S/C ratios occurs through reaction pathways involving the dissociation of methane and hydrocarbons and depends on feed gas composition, operating temperature and nature of the catalyst.¹² Soot formation in the ATR reactor is also dictated by operating conditions as well as both burner design and the catalyst used, the latter of which has to be able to convert soot precursors formed.

In recent years, significant progress has been made in the optimization of catalysts for SR and ATR. Fundamental studies have led to a greater understanding of the mechanism of carbon formation, even allowing for carbon free-operation at very low S/C ratios.¹¹ In fact, the catalyst is not the limiting factor for the operation of a tubular reformer and thus, further catalyst development should be very limited. The foremost challenge to significantly impact the GTL technologies is the development of alternative technologies.

2.2 Alternative technologies

The syngas production step may account for 60–70% of the total capital cost of a GTL plant.^7,8 Therefore, there is a great interest to develop new lower-cost syngas production technologies.

The catalytic partial oxidation of methane (CPO) is an alternative technology that involves the reaction between methane and oxygen (Table 1) in a reactor containing a catalyst without a burner.¹² Since the 1990's, CPO has been extensively studied and several reviews can be found in the open literature.^14–20 The reaction conditions (temperature, pressure) and reaction mechanism have been thoroughly investigated. The performance of different noble and non-noble transition metals for partial oxidation of methane was evaluated and a comprehensive review about catalyst screening carried out to date was published recently.²⁰ In particular, catalyst deactivation mainly due to carbon formation is an important issue for the commercialization of CPO.^19,20 The nature of the support plays an important role in the stability of catalysts for this reaction route.^21,22 Several studies have shown that Pt/ZrO₂^23,24 and ceria-based catalysts^25–27 are stable on CPO. More recently, we have reported that use of promoters such as cerium oxide improves the stability of Pt/ZrO₂ catalysts.^28–30 The improved performance was attributed to the higher reducibility and oxygen storage/release capacity of Pt/CeZrO₂ catalysts, which allowed a continuous removal of carbonaceous deposits from the active sites, favoring the stability of the catalysts. Further catalyst developments are still necessary such as: (i) stable catalysts for operating at high pressures (20 bar), typical of the Fischer–Tropsch process using Co catalysts; (ii) the control of metal particle size through stabilization by the support, and maintaining the ensemble size below the critical value required for carbon formation.²⁰

In spite of all the progress achieved on catalyst development, there are still many issues to be addressed before CPO technology can achieve commercialization.²⁰ One of the disadvantages of this route is the highly flammable mixture that may ignite at temperatures above 250 °C. Therefore, the reactants may not be pre-heated at high temperatures, resulting in high natural gas and oxygen consumption since part of the feed has to be burned to generate the heat required to achieve the reaction temperature. Taking into account that 40% of the capital costs of a GTL plant corresponds to the ASU, CPO is unlikely to be economically competitive with ATR technology.

However, if CPO is carried out in a ceramic membrane reactor, the costs associated with a conventional oxygen plant are eliminated and this technology becomes economically viable.^7,10 In the ceramic membrane reactor, air separation and the partial oxidation reaction take place in the same device. This technology is based on a dense ceramic membrane that exhibits both oxygen ionic and electronic conductivity at high temperatures, typically 800–900 °C.³¹Oxygen ions flow through these membranes by sequentially occupying the oxygen vacancies when they are heated to high temperatures. The driving force for oxygen permeation is established across the membrane by depleting the oxygen partial pressure on one side of the membrane through chemical reaction. Therefore, oxygen is transported from low pressure air feed to a high pressure fuel stream without the need of mechanical compression (Fig. 5).^31,32 For GTL applications, the membrane materials must be chemically and mechanically stable in the high-pressure, reducing natural gas feed side as well as in the low-pressure, oxidizing air feed side. They must have sufficient mixed electronic and oxygen ion conductivity to achieve high oxygen fluxes. Perovskite-type oxides with general formula ABO₃ with dopants on the A and/or B sites have been extensively used in ceramic membrane reactors for selectively transporting oxygen, generally showing the highest oxygen flux at high temperature.^33,34 These materials exhibit high oxygen permeability due to the presence of oxygen vacancies as well as thermal stability.³⁵

Fig. 5 An oxygen transport mechanism through a ceramic membrane based on a perovskite material. (Reprinted with permission from ref. 31. Copyright 2000 Elsevier.)

In the past decades, two industrial consortia have been actively working on the development of ceramic-based membrane technology for fuel production.^7,36 Air Products headed one consortium including ARCO, Ceramatec, Chevron, Norsk Hydro and others that developed the ion transport membrane (ITM) system based on a perovskite-type oxide with the formula (La_1−xCa_x)yFeO_3−δ.³² The second consortia led by Praxair and comprising Amoco, BP, Statoil, Phillips Petroleum and Sasol developed the oxygen transport membrane (OTM) technology for the production of oxygen from air separation.^7,36

Despite the efforts carried out to develop the ceramic membrane for the production of oxygen from high temperature air separation, there are still critical issues in the implementation of membrane separation technology for oxygen production. Further research is still necessary to improve the chemical, thermal and mechanical stability of the membrane materials while maintaining high ionic and electronic conductivities.^34,36 Perovskite materials exhibit loss of stability during long-term operation due to reaction with CO₂ or water.³⁴ Partial substitution of cation A or B may induce significant changes in chemical and thermal stability.³⁴ However, there is not a systematic and fundamental study that correlates the composition of the perovskite structure with oxygen permeability or chemical and structural stability, which is important to define the metal elements in the A and B sites.³⁴ Perovskites exhibiting higher oxygen flow rates are more easily reduced, which may result in the formation of cracks.³⁵ Therefore, the development of new materials for membranes has to take into account an appropriate balance between the oxygen permeation rate and chemical/thermal stability.

Concerning catalysts for membrane reactors, there are also important challenges. In particular, the significant catalyst deactivation due to carbon formation emphasises the fact that catalysts for CPO have to be resistant to coke as previously described.³⁷ Furthermore, most of the studies investigating CPO were carried out in a fixed-bed reactor with powder catalysts. In this case, thermal gradients may occur in the catalyst bed and the effect of heat transfer could be significant because total oxidation reactions are highly exothermic. This may lead to the formation of hot spots and consequently catalyst deactivation.¹⁵ However, the extremely high reaction rates of CPO due to its high exothermicity allow residence times in the millisecond range, which is characteristic of the compact reactors, including monoliths, foams and plate-type reactors.

The development of compact reformers may significantly contribute to reducing the cost of syngas generation technologies. One type of compact reformer is the so-called plate-type reformer. The concept of the plate reactor is based on coupling an endothermic with an exothermic reaction by means of indirect heat transfer. Then, catalytic combustion or another highly exothermic reaction is used to generate the heat required for the endothermic reaction. Fig. 6 shows a schematic section of a plate-type reactor based on catalytic combustion/steam reforming.³⁸ The reformer plates are arranged in a stack. One side of each plate is coated with a steam reforming catalyst, where the syngas reaction takes place. On the other side of the plate, catalytic combustion of the natural gas occurs, providing the heat to the endothermic steam reforming reaction.

Fig. 6 Cross section of a plate-type reformer combining catalytic combustion/steam reforming reactions. (Reprinted with permission from ref. 38. Copyright 2003 Elsevier.)

The advantages of the plate reformers are:³⁹ (i) a significant reduction in size and weight in comparison to conventional fired tubular reformers. The compact design enables the use of this technology in offshore platforms or remote sites for the conversion of associated gas to liquid fuels; (ii) standardized design and consequently, lower capital cost; (iii) increased thermal efficiency due to the better heat and mass transfer (and then the rate of the reactions that are limited by heat or mass transfer in a conventional reactor can be improved); (iv) faster start up (since each plate has a lower thermal inertia); (v) modular nature, which facilitates scaling up by using a large number of small units and in turn, makes it a flexible technology; (vi) oxygen is not required; (vii) lower NO_x emissions because the catalytic combustion used to provide the heat proceeds at lower temperature than homogeneous combustion. Consequently, lower operation temperature means less costly materials in reactor construction.

Therefore, a challenge is to downscale the mature steam reforming technology. Several companies have developed more compact steam methane reformers for onshore stand-alone syngas production. Haldor Topsøe has developed a convection reformer as shown in Fig. 7.⁴⁰ The reactor is similar to a reformer of the heat exchange type, where the process gas is heated in a counter-current configuration by the flue gas on the outside. BP and Davy Process Technology have also developed a compact reformer that was tested in the GTL facility in Alaska.⁴¹

Fig. 7 Haldor Topsøe convection reformer. (Reprinted with permission from ref. 40. Copyright 2001 Elsevier.)

Compact reactors may also address the challenge of dealing with the associated gas produced in oil fields located at remote and offshore sites in very deep water. This technology is an alternative to resolve the issue of flaring the associated natural gas or to the high costs associated with the high pressure reinjection of the gas into wells. Ongoing research aims at the installation of microchannel reactors (plate-type reactors with channel dimensions in the micro range) in offshore platforms and floating production storage offloading vessels (FPSO) to convert natural gas into liquids viaGTL technology. Compact GTL and Velocys are, among others, companies that are leading efforts in this way.

Microchannel reactors have been investigated in recent years with the aim of downsizing the SR technology.^42–50 The conventional SR of methane using fired tubular reactors filled with pellets of the Ni catalyst operates at contact times around 1 s. Some reports in the open literature indicate that the SR reaction using microchannel reactors has operated at less than 10 ms contact times.^42–45

Velocys developed a microchannel reactor for SMR with integrated catalytic partial oxidation followed by catalytic combustion of natural gas that provides the required heat for the endothermic SR of methane reaction in adjacent channels.⁴³ The reforming catalyst was 10% Rh supported on alumina doped with MgO, which was deposited on a FeCrAlloy plate. Higher rates of heat transfer were achieved for SR in the microchannel reactors at low reforming and combustion contact times (at around 4 ms) than in conventional steam reformers. The SR reaction was also carried out at contact times below 1 ms (90 and 900 μs) in a microreactor containing a Rh/MgO/Al₂O₃ catalyst.⁴⁵ The experimental and theoretical results showed that methane conversion of 88 and 17% was achieved for contact times of 900 and 90 μs, respectively. The model predicted that catalyst thickness has an important impact on the microchannel performance as long as the reaction is not heat and mass transfer limited. This result indicates that the techniques for catalyst coating and deposition are a key challenge in developing microreactors. The choice of metal alloy to be used as a catalytic substrate as well as surface treatments and catalytic coating techniques will directly affect the performance of the microchannel reactor.³⁷

The feasibility of the SR of methane in microreactors was also demonstrated theoretically.^38,46–50 The SR of methane was simulated using a parallel plate microreactor, where the propane combustion on Pt catalysts and SR of methane on Rh catalysts took place on opposite sides of the wall. These papers investigated the effect of operating conditions (flow rates, inlet composition and catalyst loading) and design parameters (wall material, channel size) on methane conversion and power output.

In contrast to the conventional SR technologies whose commercial catalysts have already been optimized, new catalyst formulations are necessary for compact reactors. Sufficiently active catalysts are required to carry out the SR of methane at very low contact times and thus, noble metals are the preferred choice. For example, nickel, the metal selected as a catalyst used in fired tubular reformers of large scale syngas technologies, has very low intrinsic activity. Since heat and mass transfer rates in microreactors are very fast and the process is kinetically controlled, the slow SR rate leads to slow removal of heat and thus to undesirably high temperatures.^48,50 In order to overcome this limitation, a reduction in the combustion rate or a larger reactor could allow the use of a Ni catalyst, suggesting that a better catalyst than Ni is required. Rh exhibits a higher intrinsic activity than Ni and thus, SR on Rh is approximately one order of magnitude faster than on Ni.⁴⁸

Another issue for developing new catalysts for compact reactors is catalyst long term stability. In this case, catalyst deactivation due to carbon formation is critical since it may lead to channel blockage. Different approaches may be adopted to reduce carbon formation based on either the prevention of carbon formation reactions in the first place, or on the rapid conversion of carbon, once formed, to gaseous products for ease of removal. The support may play a major role in SR of methane in assisting to remove carbon or suppress its formation. In general, alumina is the support used since FeCrAlloy is the selected metallic substrate for catalysts in microstructured reactors.⁵¹ The addition of dopants to an alumina support or the use of ceria and ceria-containing mixed oxides supported on alumina as support for the SR of methane are possible strategies to improve catalyst stability. Redox supports like ceria and ceria-containing mixed oxides improve catalyst resistance to carbon formation due to their high oxygen storage capacity (OSC) and oxygen mobility. This highly mobile oxygen may react with carbon species as soon as it forms and thus keeps the metal surface free of carbon, inhibiting deactivation. This approach has been successfully applied for the production of syngas by partial oxidation and autothermal reforming of methane on fixed-bed reactors using Pt or Pd supported on Ce/Al₂O₃ or CeZr/Al₂O₃.^52–57

Therefore, the development of novel catalysts especially designed for compact reactors offers great opportunities and challenges. The search for new metals, less expensive than Rh, the reduction of the metal loading of the catalysts on the microreactors may be outlined as examples of future research.

3. Fischer–Tropsch synthesis

The famous Fischer–Tropsch synthesis is probably the most important step in the GTL process. The original process was developed by Franz Fischer and Hans Tropsch, working at the Kaiser Wilhelm Institute in the 1920s. The synthesis involves several reactions leading to a variety of hydrocarbons, but the overall reaction may be described as follows:

(2n + 1)H₂ + nCO → C_nH_(2n+2) + nH₂O

Hence, FT synthesis may be regarded as a polymerisation reaction that uses syngas as a reactant, producing hydrocarbons of several molecular weights. It is well established that the product distribution of hydrocarbons formed during the Fischer–Tropsch process follows the so-called Anderson–Schulz–Flory (ASF) distribution, which can be expressed as:

W_n/n = (1 − α)²αⁿ⁻¹

where W_n is the weight fraction of hydrocarbon molecules containing ncarbon atoms, α is the chain growth probability or the probability that a molecule will continue reacting to form a longer chain. Generally, the value of α is determined by the characteristics of the catalyst and the specific process conditions.

The FT-synthesis is traditionally catalysed by transition metals; cobalt, iron, and ruthenium are the most common metals used in the literature. Although nickel may also be used, methane formation (“methanation”) is favoured when this metal is employed, therefore in commercial catalysts nickel is discarded.

Normally, the FT step of the GTL processes is divided into two main areas: High-Temperature Fischer–Tropsch (or HTFT), which is operated at temperatures of 330–350 °C, employing an iron-based catalyst and Low-Temperature Fischer–Tropsch (LTFT), operated at lower temperatures (200–240 °C), using a cobalt-based catalyst (although iron-based catalysts can also be used). The type of catalyst used also determines the syngas composition. In fact, cobalt-based catalysts are highly active for Fischer–Tropsch but display almost no activity for water-gas-shift reaction.⁵⁸ For that reason, cobalt-based catalysts require a higher H₂/CO ratio (∼2), whereas iron-based catalysts are more suitable for low-hydrogen-content synthesis gases such as those derived from coal due to its promotion of the water-gas-shift reaction. Pressure has also an influence on Fischer–Tropsch selectivity; increasing the pressure leads to higher conversion rates and also favours formation of long-chained alkanes.

Despite the great technological knowledge acquired after so many years of existence of the Fischer–Tropsch process, FT-catalysts still face several challenges. Among them, it is worth mentioning the following, which will be detailed in the next sections:

(a) lower costs of production;

(b) selectivity to high octane gasoline;

(c) increased selectivity to high molecular weight products;

(d) new reactor systems.

3.1 Lower cost

As previously mentioned, Fe and Co are the main active components of most commercial FT-catalysts. Cobalt intrinsic activity is higher than iron, however site density is higher in iron catalysts, resulting in a higher overall activity. As conversion increases (therefore increasing the partial pressure of water) this advantage of iron catalysts tends to disappear. Hence, depending on the process conditions, either Co or Fe catalysts may have better productivity.^59,60 Modern FT industrial units are using cobalt-based catalysts. Traditional cobalt-based FT-catalysts usually present rather low cobalt dispersion, with average cobalt particle sizes of about 20 nm. Nevertheless, the existence of smaller particles would mean a more efficient use of cobalt, implying lower costs of the catalyst, since cobalt is indeed a somewhat expensive element. In fact, cobalt is about 1000 times more expensive than iron.⁶¹

The preparation of catalysts with smaller cobalt particle sizes is well known. However, these catalysts are often not as active as expected. Indeed, turn-over frequency seems to decrease for particles smaller than ∼10 nm. The phenomenon of lower TOF values for smaller particles has been referred to as the cobalt particle size effect.⁶²

For cobalt particles larger than 10 nm the particle size effect seems to be absent. In fact, it has been shown that in the range of 9–200 nm the TOF was not influenced by the cobalt particle size.⁶³ Also,⁶⁴ it has been claimed that selectivity to higher molecules (C₅₊) is insensitive to Co dispersion (0.5–10%). However, concerning catalysts with even smaller particle sizes the results reported in the literature are controversial. Some groups observed lower activities for smaller cobalt particles,⁶⁵ whereas others reported the opposite.

This controversy is apparently caused by problems to obtain fully reduced small cobalt particles on oxide supports.⁶⁶CoO can react with these supports both during synthesis and during the reduction treatment resulting in compounds like CoAl₂O₄, CoSiO₃ or CoTiO₃. This explanation for the lower activity of small cobalt particles supported on oxide supports is referred to as a secondary particle size effect.⁶²

In order to study the influence of the cobalt particle size on the FT reaction without the interference of the effects caused by the support material, inert carbon supports have been tried (carbon nanofibers—CNF).⁶⁷ Particles smaller than 6 nm presented much lower activities, suggesting that an optimal particle size should be in the range of 6–8 nm. Such results have been corroborated by other publication.⁶⁸ Nevertheless, it has been observed that selectivity to C₅₊ increases with the increasing particle size. Therefore, the search for lower cost may be a compromise between smaller particles (less cobalt) and higher selectivity to liquid fraction (bigger particles).

More recently,⁶⁹ ionic liquids have been proposed as an alternative way of stabilising nanoparticles of cobalt. Ionic liquids are liquid compounds that present ionic-covalent crystalline structures or electrolytes entirely composed of ions which are liquid at ambient temperature. Indeed, cobalt nanoparticles with a size of around 7.7 nm prepared in 1-alkyl-3-methylimidazolium bis(trifluoromethanesulfonyl)imidate ionic liquids are effective catalysts for the Fischer–Tropsch synthesis, yielding olefins, oxygenates, and paraffins (C₇–C₃₀). The nanoparticles may be easily prepared by the decomposition of Co(CO)₈ in the ionic liquid at 150 °C and can be reused at least three times provided they are not exposed to air. It must be borne in mind that the use of ionic liquid stabilised Co-particles may open a new horizon in the field of three-phase (slurry) reactors for Fischer–Tropsch.

Also, calcination in the presence of NO has been proposed as an alternative way of controlling particle size distribution.⁷⁰ In a very recent publication, a highly active Co/SiO₂ catalyst has been prepared with a narrow particle size distribution with a surface-average size of 4.6 ± 0.8 nm. Such catalysts displayed an unprecedented high FT activity and the narrow particle size distribution led to an activity enhancement of approximately 40% compared to Co/CNF, which had a wider particle size distribution (5.7 ± 1.4 nm).

Particles may also undergo sintering, which is probably the major cause for catalyst deactivation. Actually, three mechanisms have been proposed for Co catalyst deactivation:⁷¹ (1) sintering of the Co active phase, (2) carbon deposition and (3) surface reconstruction. The understanding of these deactivation mechanisms is certainly fundamental for the design of new FT catalysts, representing an interesting area of study. Also, they are a major issue for the regeneration of such catalysts. Indeed, attempts to introduce a three-step regeneration process based on the previous mechanisms have been successful, restoring the FTS performance of the spent catalyst to that of the fresh catalyst.

3.2 Fischer–Tropsch for gasoline production

It is well established that FT reaction is one of the best ways to produce high cetane diesel with practically no sulfur and a very low aromatic content. However, although Fischer–Tropsch has been studied for many years, one question still remains: is it capable of producing high octane gasoline?

In order to achieve such a goal, one must:

(a) increase the gasoline yield;

(b) enhance gasoline isomerisation.

Since the products slate in Fischer–Tropsch is a function of the ASF distribution, the first step would be an alternative mechanism leading to a new products distribution, or rather, the search for a new catalytic system. Such a bi-functional system must present acidic sites to promote isomerisation along with the traditional FT-sites. Aiming at breaking the ASF distribution, many authors have tried mixtures of traditional FT-catalysts and some acidic oxides, which would promote the isomerisation of the alkane chain formed via Fischer–Tropsch. As a matter of fact, systems such as FeCoK + Pd/ZSM-5,⁷²FeCuMg + ZSM-5,⁷³Co/MCM-22,⁷⁴Co-Ni/TiO₂ + ZSM-5,^75,76Co/SiO₂ + Pd/β-zeolite,⁷⁷ alkali promoted Fe + ZSM-5^78,79 have been tried. As expected, the use of zeolite containing FT-catalysts^80–82 seems to be promising for yielding branched products (Fig. 8). Recently, an excellent review on this subject has been published.⁸³ Nevertheless, all hybrid systems containing different zeolites did present a significant deactivation. Understanding the mechanism of such steep deactivation is surely the main key for the development of a new FT catalytic system that would allow the generation of high octane gasoline.

Fig. 8 Influence of zeolite on the yield of branched C5–C8 products. (Reprinted with permission from ref. 82. Copyright 2007 Elsevier.)

An interesting hybrid system⁸⁴ has also been proposed to produce gasoline via Fischer–Tropsch. Such a system comprises (a) an iron-based FT-catalyst; (b) a traditional methanol synthesis catalyst (Cu/ZnO/Al₂O₃); and (c) an acidic oxide (HZSM-5). The authors claim that adding Cu/ZnO/Al₂O₃ to a physical mixture of a FTS catalyst and HZSM-5 increases conversion and selectivity to LPG and gasoline. This hybrid system may be regarded as a mixture of a dimethyl ether (DME) synthesis catalyst (methanol catalyst + zeolite) and a FT-catalyst, which reinforces the oxygenate mechanism proposed for FT-synthesis in the presence of iron-based catalysts.^85–87

3.3 Increased selectivity to high molecular weight products

The product distribution in Fischer–Tropsch is surely a function of the ASF distribution. According to ASF distribution, products ranging from methane to heavy solid paraffin may be generated. Since naphtha, diesel and paraffin are more profitable products than, for instance, LPG, most GTL plants aim at producing these fractions, which are normally called the “C₅₊”. Although still a matter of dispute, metallic particle size seems to influence the selectivity in FT-synthesis, as previously mentioned. More recently, another factor has called the attention of the scientific community. Diffusion limitations in porous supports are somehow related with selectivity in FT.

Two steps in the diffusion mechanism play an important role under FT-synthesis conditions.^88,89 The first one concerns the migration of the reactants to the active sites, whereas the second one regards the diffusion of the products towards the outer surface of the catalyst pellet. It is clear that the size of the pellet will impact both steps. The first diffusion limitation (reactants migrating to the active sites) will reduce CO concentration inside the pellet, thereby favouring the formation of lighter products. On the other hand, the second mechanism impacts the re-adsorption of α-olefin, thereby provoking an increase in the selectivity to paraffin and higher molecular weight products as the size of the pellet increases. “Eggshell” type catalysts have been proposed⁹⁰ to optimise diffusion limitations. In these catalysts, Co is deposited on the outer part of the pellet, forming a thin external layer.

In a very relevant scientific contribution, γ-Al₂O₃ nanofibers⁹¹ presenting simultaneously a very high surface area (321 m² g⁻¹) and a hierarchical macro–mesoporous structure have been used to prepare supported CoRu catalysts at two loading levels (20 wt% Co–0.5 wt% Ru and 30 wt% Co–1.0 wt% Ru). Such unique catalysts have been used to elucidate the relative significance of diffusion and dispersion effects during FT synthesis. It has been shown that in the absence of diffusion limitations, both FT activity and selectivity are mostly determined by Co⁰ dispersion. Thus, particle size effects (lower TOF and higher CH₄ selectivity for Co⁰ nanoparticles below 8–10 nm in size) previously mentioned are indeed observed. Nevertheless, catalyst porosity governs catalyst performance in the pseudo-steady state (TOS > 7–8 h), when diffusion issues start to be determining. In this case, for high metal loadings (30 wt% Co), the nanofibrous alumina support presented the highest specific activity and productivity to diesel when compared to other supports such as wide pore commercial aluminas. In contrast, wide pore supports produced more waxy hydrocarbons (C₂₃₊).

3.4 New reactor systems

As previously mentioned, Fischer–Tropsch products distribution is governed by the chain growth probability parameter (alpha). This parameter, which determines the selectivity in FT-synthesis, strongly depends on the reaction temperature, since the activation energy of the termination step is higher than that of the growing step.^92,93 High temperatures favour the formation of undesirable light products, mainly methane.

Considering that Fischer–Tropsch is highly exothermic, it is worth noticing that the ideal reactor for FT would be the one in which an excellent control of temperature could be achieved. Although conventional reactors such as fixed bed and slurry have been used, new concepts of the reaction system are being claimed as the best systems for Fischer–Tropsch, mainly those related to micro channel reactors. Micro reactors (or micro channel reactors) represent a new area of knowledge in the field of Chemical Engineering, known as “Process Intensification”. Process Intensification (PI) was defined as a “reduction in plant size by at least a factor 100”⁹⁴ and represents a paradigm shift in Process Design. Undoubtedly, PI leads to a substantially smaller, cleaner, and more energy efficient technology.^95,96

It must be borne in mind that several offshore natural gas occurrences have been recently found, which has drawn the attention of the scientific community to offshore technologies. Such natural gas requires convenient processing either in a platform or a FPSO (Floating Production, Storage and Offloading) unit. It is obvious that offshore processing units are to be located on a confined area, therefore the use of technologies employing smaller equipment is a must.

In a recent publication,⁹⁷ Fischer–Tropsch synthesis in micro channels has been extensively studied. Different metallic supports (aluminium foams of 40 ppi, honeycomb monolith and micro monolith of 350 and 1180 cpsi, respectively) have been loaded with a 20% Co–0.5% Re/α-Al₂O₃ catalyst by the washcoating method, generating layers of different thicknesses deposited onto the metallic supports. The study evidences the viability of the use of structured supports for the Fischer–Tropsch synthesis. Indeed, the results show that both the supported catalysts and the micro channels block present better performance than the powder catalyst. Regarding selectivity, some very interesting conclusions have been drawn. As depicted in Fig. 9, the selectivity to C₅₊ depends on the type of support and mainly on the amount of catalyst deposited and its effect on the catalytic layer thickness; however, it does decrease as the CO conversion increases. Concerning structured systems, selectivity decreases in the following order: micro channels block > micro monoliths > monoliths > foams. The results may be related to the catalytic layer thickness in the case of the structured supports, and to the better temperature control in the case of the micro channels block. Hence, the main challenges regarding micro reactors applied to Fischer–Tropsch synthesis seem to be the control of the coating process for a given configuration.

Fig. 9 Activity (CO conversion) and selectivity (C₅₊) for different catalyst layer thicknesses at 250 °C, 10 bar and H₂/CO = 2. (Reprinted with permission from ref. 97. Copyright 2011 Elsevier.)

4. Upgrade

4.1 Overview

The conventional technology for GTL involves the reforming of natural gas to produce essentially a mixture of carbon monoxide and hydrogen (synthesis gas) and the conversion of this mixture by FT synthesis to a mixture of hydrocarbons with varying chain lengths. The selectivity for hydrocarbons in a given molecular weight range may be controlled, to a certain extent, by the choice of catalyst and process conditions. However, due to the characteristic kinetics of the Fischer–Tropsch process (AFS distribution), the production of a wide molecular weight distribution is unavoidable. With such kinetics, the selective synthesis of a product with a narrow range of chain lengths is theoretically impossible, except for methane or for an infinite chain length.

Table 2 ⁹⁸ shows typical product distributions of FT synthesis for the two main established FT technologies, namely high-temperature Fischer–Tropsch (HTFT) and low-temperature Fischer–Tropsch (LTFT) and the catalysts used industrially, namely unsupported iron and supported cobalt catalysts.

Table 2 FT product spectra (at 2 MPa)⁹⁶

Catalyst type: FT temperature/°C	Fe: fused 340	Fe: precip. 235	Co: supported 220
Selectivity (C atom basis)
CH4	8	3	4
C2–C4	30	8.5	8
C5–C6	16	7	8
C7–160 °C (bp)	20	9	11
160–350 °C (bp)	16	17.5	22
+350 °C (bp)	5	51	46
Water-soluble oxygenates	5	4	1
α value	0.7	0.95	0.92
C3 + C4: %alkenes	87	50	30
C5 to C12 cut:
%Alkenes	70	64	40
%Oxygenates	12	7	1
%Aromatics	5	0	0
C13 to C18 cut:
%Alkenes	60	50	5
%Oxygenates	10	6	<1
%Aromatics	15	0	0

---

The syncrude from HTFT synthesis is more olefinic, rich in oxygenates (mainly alcohols, carboxylic acids and ketones) and contains aromatics, while syncrude from LTFT synthesis contains mainly n-alkanes, n-olefins and alcohols. Product saturation increases with carbon number and although straight-run LTFT naphtha and distillate contain a fair amount of olefins and oxygenates, the heavier products are mostly n-paraffin waxes.^98,99 Some branched compounds can also be obtained in LTFT.¹⁰⁰

As in the case of straight-run petroleum refining, straight-run FT product distribution does not match market demands in terms of quantity and quality. Thus, a fairly large proportion of C₁–C₄ gases is produced, especially in HTFT and, in parallel with the naphtha, kerosene and diesel fractions, a heavy residue is always obtained, basically comprised of high molecular weight n-alkanes, especially in LTFT. Assuming ideal AFS kinetics, the maximum straight run middle distillates yield (C₁₀–C₂₀ cut) achievable is about 40 wt%.¹⁰⁰

Despite the essentially nil nitrogen and sulfur contents, the linearity of the hydrocarbons obtained and the almost absence of aromatic compounds are detrimental to the octane rating of the naphtha fraction obtained from FTS.

The same characteristics are extremely favorable for the cetane number (CN) and particulate emission of FT diesel. In LTFT with cobalt catalysts, a diesel fraction with CN in the 75 range can be obtained, while in HTFT, hydrogenation of the diesel fraction for olefin removal leads to a CN rating of about 50, close to the 45–50 required by the market.⁹⁸ The absence of heteroatoms, aromatics and naphthenes make this fuel ideal for the environmental impact reduction. However, the same linearity and absence of aromatics are detrimental to the cold-flow properties, lubricity and density of straight-run FT diesel.^98,99

Essentially the same problem exists with the kerosene fraction of straight-run FT products. The freezing point specification of jet fuel (<−47 °C) is needed to ensure that the fuel remains pumpable under the low temperature conditions experienced during high altitude flight and places a limit on the amount of linear hydrocarbons in the fuel.⁹⁹

Due to their paraffinic character, FT heavy fractions are adequate as a raw material for base oils with high viscosity index for lube oil production. However, they have to be chemically transformed, since their large proportion of n-paraffins is highly detrimental to pour- and cloud-points of the product. A typical value for a LTFT paraffin wax is 92 wt% n-alkanes.¹⁰¹

de Klerk has recently critically analyzed the refining technologies available for the processing of straight-run FT products, including gases, naphtha, middle distillates and residues,⁹⁹ to produce useful fuels and products. The present review is focused on the chemical transformation of FT waxes to produce middle distillates and base oils for lube production, especially the catalytic challenges involved in these processes. As can be seen from Table 2, the heavy fraction of LTFT products, boiling above 350 °C contains about 50% of the total carbon.

It should be emphasized that LTFT residue transformation by fluid catalytic cracking (FCC) to produce gasoline has been considered by several groups.⁹⁹ Recently, for example, Dupain et al.¹⁰² reported that this highly paraffinic feedstock has a high reactivity and can be more than 90% converted by FCC to produce a gasoline fraction (70 wt%) with a very low aromatics concentration. As a result of the formation of i-alkanes, n-olefins and i-olefins the gasoline is expected to have an acceptable octane number. de Klerk, however, argues that the high hydrogen content of LTFT material does not favor coke formation and additional fuel would have to be burned in the catalyst regenerator to keep the FCC heat balance.⁹⁹

4.2 Middle distillate production by hydrocracking of FT waxes

Although investigated by Sasol since the 70's, middle distillate production by hydrocracking (HCC) of FTS waxes was only implemented in the 90's at the Shell plant in Malaysia.⁹⁸ Since then, it has been implemented in the Oryx GTL plant in Qatar, using a ChevronTexaco isocracking technology with a Chevron proprietary HCC catalyst.⁴

The diesel produced by wax HCC has good cold-flow properties, due to the high degree of branching of the paraffinic hydrocarbons. Cetane number is very high (>70),⁴ and the amount of heteroatomic contaminants (nitrogen, sulfur and oxygen compounds) is virtually nil. Results of engine tests show that this diesel leads to significant reduction of CO, hydrocarbons, particulate matter and polyaromatic hydrocarbon emissions in comparison with petroleum derived ones.¹⁰³ The absence of aromatics and sulfur has a negative impact on lubricity which is, according to Calemma et al.,¹⁰³ generally well below the accepted standards. More importantly, the almost total absence of aromatic and naphthenic compounds causes the density of the diesel produced from either LTFT or its hydrocracking product to be well below specification.^4,98,99,103 For that reason, and due to its very high cetane rating and low pollutant emission, diesel from FT wax hydrocracking is more suitable for blending with other diesel fractions in order to adjust their properties to specification than to direct use as a fuel.

The main goal of FT wax hydrocracking to produce middle distillates should be to keep a high selectivity to the desired product range at the highest possible conversion. It is possible to adjust process conditions in order to optimize the middle distillate yield, since selectivity to these products usually decreases with increasing conversion, so that a maximum yield exists at a certain conversion level. With current technology, by combining FT synthesis and hydrocracking, diesel selectivities above 80% are achievable.¹⁰⁴ However, the maximum yield depends strongly on the catalyst.

An ideal catalyst for FT wax hydrocracking should be selective for the rupture of central bonds in long chain n-paraffin molecules, should minimize successive cracking of the primary cracked products and the cracking of molecules in the middle distillate range already present in the feed.¹⁰⁰ Furthermore, the cracked products should be branched in order to improve cold-flow properties of the product.⁹⁸

4.2.1 Hydrocracking mechanism. It has been known for a long time that hydrocracking involves a bifunctional mechanism, where a “metal” function is responsible for hydrogenation–dehydrogenation reactions and an acidic function is responsible for isomerisation and cracking reactions. The word “metal” appears between quotes because in many hydrocracking catalysts the metal function is actually given by a transition metal sulfide phase. The so-called ideal hydrocracking mechanism involves the fast formation of an olefin by dehydrogenation of a paraffin molecule. The olefin then migrates to an acidic site, where it is protonated to form secondary carbenium ions. These carbenium ions undergo isomerisation and cracking reactions resulting in product carbenium ions which are transformed into saturated products through the reverse elementary steps. The acid catalyzed steps in this mechanism are generally assumed to be rate determining¹⁰⁵ and the hydrogenation–dehydrogenation steps are consequently in quasi-equilibrium, so that olefin concentration is determined by thermodynamic factors.¹⁰⁰Isomerisation occurs through sequential monobranched, dibranched, and tribranched isomer formation. The cracking reaction requires the formation of dibranched and especially tribranched intermediates.¹⁰⁵ Tribranched intermediates of the α,α,γ type are especially prone to C–C bond rupture by β-scission,¹⁰⁶ because they involve both reactant and product tertiary carbenium ions, as illustrated in Scheme 1.

Scheme 1 C–C bond rupture by β-scission from an α,α,γ-tribranched intermediate.

It follows that the requirement that hydrocracking products be branched¹⁰⁰ is not a serious challenge, as through the bifunctional mechanism they are necessarily so.

It is generally accepted that equilibrium is quickly reached between products of non-branching isomerisation steps (hydride, methyl and ethyl shift), so that equilibrium exists between the alkanes with the same carbon number and degree of branching.¹⁰⁷ The limiting step is considered to be the increase in branching degree that involves the rearrangement of a secondary to a tertiary carbenium ion through a protonated cyclopropane intermediate.^107,108

4.2.2 Catalysts for FT residue hydrocracking. A proper balance between “metal” and acidic functions must exist in the ideal hydrocracking catalyst. The hydrogenation/dehydrogenation function must be strong enough to adequately supply the acidic sites with olefin molecules for carbenium ion production and quickly hydrogenate the product olefin to avoid secondary cracking.¹⁰⁰ A too strong metallic function may lead to a shift in selectivity to isomerisation, rather than cracking,¹⁰⁹ and the appearance of hydrogenolysis reactions that produce undesirable light gases.^110,111

Both precious metals, mainly platinum^{105,112–116} and palladium,¹¹⁷ and mixed NiMo and NiW sulfides^{101,104,113,118} are used as active phases for hydrogenation/dehydrogenation. There is a large experience in the petroleum refining industry in the use of mixed sulfide-based residue hydrocracking catalysts, since petroleum-derived residues contain sulfur and nitrogen, which are strong poisons for precious metal catalysts. On the other hand, FT waxes are free of these contaminants. This allows the use of the much more strongly hydrogenating precious metal catalysts and therefore lower operating temperatures. Mixed sulfide catalysts are cheaper, but have to be maintained in the sulfided state during process operation by addition of an organosulfur compound.¹⁰¹ The lower hydrogenating power of the mixed sulfide catalyst as compared to precious metals renders the former more selective for hydrocracking and the latter to hydroisomerisation.¹⁰⁹

Bouchy et al. have remarked¹⁰⁰ that the possibility of secondary cracking increases with an increased average residence time of olefinic intermediates in the vicinity of acid sites. Therefore, any diffusional limitation or confinement effect resulting in a too strong adsorption of the intermediates should be minimized. For this reason, amorphous mesoporous supports, like silica-aluminas,^{101,103,104,114,116,119} have been more frequently used than zeolite supports and, when these are used, a zeolite with little shape selectivity, such as USY is the usual choice.¹⁰⁰

It is generally found that high middle distillate yields are obtained with solids with weak to medium acidic strength.¹⁰⁰Pt-promoted HY affords a high yield of gasoline-range hydrocarbons (>90%) while Pt-promoted HZSM-5 affords a larger amount of gas products due to its strong acid sites.¹¹⁰

Recently investigated systems using acidic supports other than silica-aluminas or zeolites include platinum supported on sulfated zirconia¹²⁰ and on polyoxocation ([AlO₄Al₁₂(OH)₂₄(H₂O)₁₂]⁷⁺ and [Zr₄(OH)₁₄(H₂O)₁₀]²⁺)-pillared montmorillonite.^110,115 The latter were reported to afford higher yield of diesel-ranged hydrocarbons (>70%) than HZSM-5, HY, WO₃/ZrO₂, and SiO₂–Al₂O₃ supported catalysts, due to the appropriately weak acid strength, high thermal stability, large BET surface area, and large pore size. A Pd–Al₂O₃ catalyst has been prepared by an anionic surfactant templating method and was found to be more active than alumina-supported palladium (Pd/Al₂O₃) for the production of middle distillates due to its higher palladium dispersion and high medium strength acidity.¹¹⁷

Ultra-stable Y zeolite (USY) and also β-zeolites are relatively wide-pore zeolites that do not display shape selectivity, as far as hydrocracking or hydroisomerisation of n-alkanes is concerned. Thybaut et al. have remarked that such behavior leads to a very wide product distribution in hydrocracking, ranging from products as light as LPG over the more valuable fractions naphtha, kerosene, diesel, and lube oil base stocks, to products that are barely lighter than the original feedstock.¹⁰⁵ They also remarked that the use of zeolites with straight parallel narrow pores, such as ZSM-22, leads to the phenomenon known as pore mouth catalysis,^121–127 whereby only linear hydrocarbons or the linear part of branched hydrocarbons can penetrate the pores and branching reactions can only occur at the pore mouth, involving the portion of the molecule that remains outside the pore. This leads to a high selectivity for isomerisation near the extremity of the hydrocarbon chain, since the multibranched intermediates involved in hydrocracking cannot be formed. And if cracking occurs under appropriately severe conditions, undesirable light hydrocarbons are produced. Thybaut et al. then speculated whether there could be some zeolite pore structure that would allow the adsorption of both extremities of the linear hydrocarbon chain inside narrow straight pores, while the middle part would be located within wide cavities, where branching reactions could occur, eventually leading to cracking at the desired middle part of the chain to maximize the production of valuable hydrocarbons. They proposed that a hypothetical structure consisting of Y zeolite supercages joined by ZSM-22 segments could have this property. By simulating the reaction of n-dodecane in this type of structure using a single-event microkinetic model (SEMK), they estimated that with this type of structure the percentage of C₆ products obtained by central cracking in the chain can be increased from 25% with non-shape-selective Y zeolite up to 93%. They proposed that this is a promising approach for the development of zeolite catalysts for the selective hydrocracking of Fischer–Tropsch waxes into middle distillates.¹⁰⁵

4.3 Isomerisation dewaxing for base oil production

4.3.1 Molecular structure and properties of lube oils. Some of the most important properties of lubricating oils are sulfur content, pour-point, cloud-point, oxidation stability and viscosity index. Viscosity index (VI) is a standard empirical measure, widely accepted by the lubrication industry, inversely related to the change in viscosity of the oil with temperature. Its importance lies in the fact that the lube oil has to perform its duty both at low cold start temperature and at high temperature under heavy duty operation. The VI is disfavoured by the presence of naphthenic and aromatic hydrocarbons in the lube. Aromatic hydrocarbons, besides lowering the VI, are detrimental to oxidation stability, decreasing the useful life of the lubricant. On the other hand, linear paraffins impair the oil's cold flow properties.¹²⁸

Normal grade base oils (groups I and II, according to API classification) have VI's in the 80 to 119 range. High grade group III oils are produced by modern hydroprocessing technology in petroleum refineries, including hydroisomerisation, and have VI's above 120.¹²⁸ One way to further improve the quality of lube oils is to use feeds for hydroprocessing with a molecular composition already close to the ideal one for a high grade lubricant.¹²⁸ The heavy fraction produced in FT synthesis, especially the one produced in HTFT, has all the features required for a premium feedstock for base oil production, due to its very low sulfur, naphthene and aromatic hydrocarbon content. However, it cannot be used directly as a base oil, due to its high n-paraffin content. Hydroisomerisation dewaxing (HIDW) is the most adequate process for adjusting the cold flow properties of the oil by conversion, rather than removal of the n-paraffins.

Proper design of catalysts and process conditions for HIDW has to take into account molecular characteristics desired for obtaining proper cold-flow properties without compromising VI. It is generally accepted that increasing the degree of branching of alkanes contributes to decrease in the VI of the oil.¹²⁹ Miller et al.¹³⁰ suggested that minimizing the overall branching while maximizing the branching towards the middle of the lubricant base oil molecules provides fluids with a high VI and low pour points.

Kobayashi et al.,¹³¹ however, using NMR data on lube base oils prepared by hydrocracking/isomerisation of Fischer–Tropsch waxes, showed that the VI could be correlated to a single parameter, (ACN)²/ABN, where ABN and ACN are, respectively, the average branching number and the average carbon number of the oil, implying that the position of the branching is not important to determine the VI. Later¹³² they found that the position and the degree of branching in hydroisomerised FT residues are correlated with each other, which would explain why ABN and ACN alone were able to correlate VI data. In the order of decreasing probability, the carbon branch location is second > third > fourth position in the chain, and so on, and the probability of the seventh and eighth or inner carbon atoms was almost equal. A trend of increasing proportion of branches located at the second carbon was observed with increasing degree of branching.¹³²

Verdier et al.,¹²⁹ also in an NMR-based study, found that the presence of methylenes in non-branched alkyl chains contributed to an increase of the VI, while branching and aromaticity negatively affected the VI. Methyl branching seemed to have a much smaller detrimental effect on the VI than aromaticity, and the position of the methyl branches did not seem to be important.

4.3.2 Catalysts for HIDW. From the previous discussion, it seems that process conditions and catalyst design should be aimed at branching the largest possible fraction of the n-paraffin molecules present in the feed, but limiting as much as possible the number of the branches, and avoiding the occurrence of hydrocracking reactions.

Thus, catalysts that have a high hydrogenation activity and a low degree of acidity are best for maximizing hydroisomerization versus hydrocracking¹⁰⁹ since a strong hydrogenating power limits the degree of branching by hydrogenating primary isomerization products. Platinum or palladium are generally found to be the most appropriate metallic phase for HIDW catalysts, rather than mixed-sulfides or a base metal such as nickel.^109,120

The support should be selective for adsorption of linear alkanes and the pores should be small enough to limit the occurrence of branching reactions inside them that lead to multibranched hydrocarbons which are deleterious for the VI and are precursors of hydrocracking reactions. Medium pore zeolites and, especially, those with parallel straight pores and ten-membered ring pore openings, such as ZSM-22, ZSM-23, ZSM-48 and SAPO-11 (a silica-alumino phosphate), have been shown to be excellent acidic components for hydroisomerisation catalysts for long-chain n-paraffins, as recently reviewed by Bouchy et al.¹⁰⁰

These zeolites have pores with the appropriate geometry for the occurrence of the pore mouth catalysis effect alluded to above. For this reason, they are very selective for 2-methyl branching of short chain linear alkanes.¹²⁶ With long chain hydrocarbon, as the ones relevant for HIDW of FT waxes, a second effect appears, which has been named key–lock catalysis,^100,121,126 whereby both extremities of the hydrocarbon chain penetrate neighbouring pores emerging at the zeolite crystal surface and the branching occurs at the central part of the chain by reaction on acidic sites at the external surface of the zeolite between pore openings. The position of the central branching relative to that of chain-end branching depends on the distance between the openings of the neighbouring pores.

Bouchy et al. have studied the hydroisomerization of n-octadecane in a series of closely related zeolites of the ZSM-48 family.¹⁰⁰ Maximum isomer yields of up to 77% at conversions approaching 100% were obtained in some cases and significant selectivity differences were observed between the different but related zeolites. This shows that subtle differences in the arrangement of pore openings at the crystal surfaces, detailed topology of the zeolite channels and concentration and position of aluminium atoms strongly influence catalyst activity and selectivity. This provides interesting opportunities for fine tuning of catalyst performance to suit specific ends.

Apart from zeolitic catalysts, some reports have appeared in the literature concerning HIDW with platinum deposited on amorphous supports, such as zirconia-supported tungsten oxide.^120,133 The largest n-hexadecane isomerisation yield reported was 71% at a 86% conversion level, with a catalyst containing 0.5% Pt and 6.5 wt% W, under very mild conditions (300 psig and 230 °C). About 72% of the hexadecane isomers were mono- or dimethyl branched. Reduced tungsten oxide species have been proposed to be the active sites and the role of the hydrogenating metal is not completely clear. The fact that no correlation was found between activity and platinum dispersion in Pt/WO₃/ZrO₂ plus the observation that olefin addition was detrimental to instead of promoting alkane conversion suggest that the conventional bifunctional mechanism does not operate in this case.

5. Conclusions

GTL technologies, based on the traditional Fischer–Tropsch synthesis, have been known for many years and faced some ups and downs over the years. However, due to more stringent environmental regulations, new interest has arisen regarding technologies capable of producing clean-burning fuels and high cetane diesel, among them the GTL technologies.

The GTL technologies comprise three main steps, namely, the syngas generation, the Fischer–Tropsch synthesis and the Upgrading, which encompasses hydrocracking and hydroisomerisation. Although GTL technologies are well established, many catalytic challenges still exist in the three steps described above.

Regarding the generation of synthesis gas, the main technologies are steam methane reforming (SMR), non-catalytic partial oxidation (POX), two-step reforming and autothermal reforming (ATR), which combines an endothermic reaction (SMR) and an exothermic reaction (POX) in the same reactor. The main challenges in this area seem to be related to the generation of a correct H₂/CO ratio for GTL (H₂/CO ≈ 2) at low steam-to-carbon (S/C) ratio, without causing high carbon formation. Nevertheless, fundamental studies have led to a greater understanding of the mechanism of carbon formation; hence little has been made in terms of new catalysts development.

Since syngas production step may account for 60–70% of the total capital cost of a GTL plant, alternative technologies have been proposed, deserving attention is the catalytic partial oxidation of methane (CPO). Although several catalysts may be found in the literature, such as Pt/ZrO₂ and ceria-based catalysts, in all cases catalyst deactivation mainly due to carbon formation is an important issue. Also, stable catalysts for operating at high pressures and the control of metal particle size are yet to be developed.

More recently, ceramic membrane reactors, in which both air separation and the partial oxidation reaction take place, have been applied to CPO. Indeed, should CPO be carried out in a ceramic membrane reactor, the costs associated with a conventional oxygen plant would be eliminated. Much effort has been made to make CPO membrane reactors commercial; nevertheless, further research is still necessary to improve the chemical, thermal and mechanical stability of the membrane materials while maintaining high ionic and electronic conductivities.

Compact reformers using micro channel technology are also the future of syngas generation. However, in contrast to the conventional SR technologies whose commercial catalysts have already been optimized, new catalyst formulations are necessary for compact reactors. Sufficiently active catalysts are required to carry out the SR of methane at very low contact times and thus, noble metals are the preferred choice. Moreover, life of such catalysts must be rather long, therefore catalyst deactivation due to carbon formation becomes critical since it may lead to channel blockage.

Concerning Fischer–Tropsch synthesis, it must be borne in mind that, despite such synthesis exists for more than 80 years, some important catalytic challenges deserve special attention. Lower costs of production, selectivity to high octane gasoline, increased selectivity to high molecular weight products and new reactor systems are related to a more efficient use of the active metal in the catalysts, that is to say, smaller particles. Modern FT industrial units are using cobalt-based catalysts and traditional cobalt-based FT-catalysts usually present rather low cobalt dispersion. The preparation of catalysts with smaller cobalt particle sizes is well known, however, as recently proven, particles smaller than 6 nm cause a steep drop in activity. Hence, there is a search for methods of preparation that furnish a narrow particle size distribution. The use of ionic liquids seems to be an interesting option as well as calcination in the presence of NO. Furthermore, more studies regarding stability and mechanisms of catalyst deactivation have to be performed.

Fischer–Tropsch is well known as an excellent chemical route to produce diesel; nevertheless, the naphtha produced via this route is not suitable for the gasoline pool. Recently, new hybrid catalytic systems containing a zeolite component have been proposed as an alternative route to promote isomerisation as FT-synthesis takes place. Such systems use the FT mechanism via oxygenates but present a considerable deactivation in some cases. Reducing deactivation is certainly an important challenge.

Another interesting issue of FT-synthesis concerns the products slate. Naphtha, diesel and paraffin are more profitable products than, for instance, LPG, therefore most GTL plants aim at producing these fractions, which are normally called the “C₅₊”. In recent publications, diffusion limitations in porous supports have been associated with selectivity in FT. Several mesoporous systems have been studied and outstanding results have been obtained with a hierarchical macro–mesoporous structure. Commercial production of such supports is surely a hurdle to overcome.

Again, as previously mentioned for syngas generation, micro reactors have been proposed as a cutting edge technology for GTL and Fischer–Tropsch. Microreactors would allow an excellent temperature control, thereby controlling alpha, the degree of polymerisation. Furthermore, the use of compact technologies would allow exploitation of off-shore gas reserves, since the plant could be accommodated on a FPSO. In terms of catalysts, the control of the coating process for a given configuration, generating a convenient layer of catalyst, deserves more attention.

The last step in GTL processes is the upgrading step, which may include hydrocracking, hydrotreating and hydroisomerisation, aiming at either to maximise diesel and naphtha production from paraffinic compounds, or to generate high quality lubricants and food grade wax.

Hydrocracking (HCC) is already a commercial process, yet some interesting features are required for the ideal catalyst. Indeed, a proper balance between “metal” and acidic functions must exist in the ideal hydrocracking catalyst. For that reason, acidic supports other than silica-aluminas or zeolites include platinum supported on sulfated zirconia and on polyoxocation ([AlO₄Al₁₂(OH)₂₄(H₂O)₁₂]⁷⁺ and [Zr₄(OH)₁₄(H₂O)₁₀]²⁺)-pillared montmorillonite have been studied with good results.

Finally, one must look at hydroisomerisation and dewaxing processes (HIDW). In contrast to what has been described for HCC, HIDW is not yet a commercial reality. Proper design of catalysts and process conditions for HIDW has to take into account molecular characteristics desired for obtaining proper cold-flow properties without compromising the viscosity index. Catalysts presenting a high hydrogenation activity and a low degree of acidity are best for maximizing hydroisomerisation versus hydrocracking. Recently, alternative acidic zeolites have been tested with good results. New mechanisms, namely the pore-mouth and the key–lock ones have been suggested to explain the performance of such catalysts. The linear carbon chain penetrates with one end into a pore opening (pore mouth) or with both ends each into a different pore opening (key–lock). Those new mechanisms provide interesting opportunities for fine tuning of catalyst performance to suit specific ends.

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