gtl challenges

698 Catal. Sci. Technol., 2011, 1, 698–713 This journal is c The Royal Society of Chemistry 2011

Cite this: Catal. Sci. Technol., 2011, 1, 698–713

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

Eduardo Falabella Sousa-Aguiar,*ab

Fabio Bellot Noronhacand Arnaldo Faro, Jr.

d

Received 31st March 2011, Accepted 16th May 2011

DOI: 10.1039/c1cy00116g

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.

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 desulfurisa-

tion, these new units are very energy consuming and reduce the

overall thermal efficiency of the refinery. Still, refinery

a Petrobras Research Centre (CENPES), Ilha do Fundao, Q7, CidadeUniversitaria, CEP 21949-900, Rio de Janeiro, Brazil.E-mail: [email protected]; Fax: +55-21-38657484;Tel: +55-21-38656643

b Federal University of Rio de Janeiro (UFRJ), School of Chemistry,Department of Organic Processes, Centro de Tecnologia, Bloco E,Ilha do Fundao, Rio de Janeiro, Brazil

c National Technology Institute (INT/MCT), Av. Venezuela 82/518,CEP 21081-312, Rio de Janeiro, Brazil

d Federal University of Rio de Janeiro (UFRJ), Institute ofChemistry, Department of Physical Chemistry, Centro deTecnologia, Bloco A, Ilha do Fundao, Rio de Janeiro, Brazil

Eduardo Falabella

Sousa-Aguiar

Eduardo Falabella Sousa-Aguiar, Chemical Engineer,MSc, DSc, has 35 yearsexperience in Catalysis. Hehas been Professor in theFederal University of Rio deJaneiro for 30 years and aSenior Advisor in PetrobrasResearch Centre (CENPES),where he is currently the co-ordinator of XTL projects. Hehas authored over 300 scientificpapers and two books, havingadvised over 30 MSc andPhD theses. He has been theBrazilian focal point for theinternational program CYTED,

being also an adviser for ICS-UNIDO. He has received manyawards; deserving particular attention is the prestigious BrazilianNational Technology Award received in 2008.

Fabio Bellot Noronha

Fabio B. Noronha receivedhis B.S. degree from FederalUniversity of Rio de Janeiro in1987, M.Sc. degree in 1989,from COPPE/Federal Univer-sity of Rio de Janeiro andPh.D. degree in 1994 fromCOPPE/Federal University ofRio de Janeiro and Institut desRecherches sur la Catalyse—Lyon, France. In 1996, hejoined the Catalysis group ofNational Institute of Techno-logy (INT). He worked ina postdoctoral position withProf. Daniel Resasco at

Oklahoma University from 1999–2000. He has been involved instudies for conversion of natural gas and biomass to hydrogen,syngas and fuels.

CatalysisScience & Technology

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This journal is c The Royal Society of Chemistry 2011 Catal. Sci. Technol., 2011, 1, 698–713 699

processes are very energy efficient even with deep desulfurisa-

tion 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 via catalytic 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, how-

ever, may have an elegant alternative which is the activation of

methane via halogenation, 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 con-

tribution 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 lubri-

cants and food grade wax.

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.

Fig. 1 Main steps of a traditional GTL process.

Arnaldo Faro

Arnaldo C. Faro Jr. receivedhis B.S. degree in IndustrialChemistry from the FederalUniversity of Rio de Janeiroin 1968 and his Ph.D degreein 1984 from the University ofEdinburgh, Scotland. In 1969,he joined the PETROBRASR&D Centre, CENPES, wherehe worked at the CatalystDivision up to 1993, mainly inthe development of hydro-processing catalysts. In 1994he joined the Institute ofChemistry of the FederalUniversity of Rio de Janeiro,

where he works up to the present date as an Associate Professor inthe Physical Chemistry Department and is head of the Hetero-geneous Catalysis laboratory.

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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).7 Fig. 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

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 a-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

POX is based on the exothermic non-catalytic reactions of

methane and oxygen inside a combustion chamber. This tech-

nology 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 1C) for obtaining high methane con-

version 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.

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 DH2980/kJ mol�1

Steam Reforming (SMR)CH4 + H2O - CO + 3H2 206CO + H2O - CO2 + H2 �41Catalytic partial oxidation (CPO)CH4 + 1/2O2 - CO + 2H2 �38Autothermal reforming (ATR)CH4 + 1.5O2 - CO + 2H2O �520CH4 + H2O - CO + 3H2 206CO + H2O - CO2 + H2 �41

Fig. 3 H2/CO ratio from different syngas technologies. (Reprinted

with permission from ref. 10. Copyright 2004.)

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Another combination of steam reforming and partial oxida-

tion 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 second-

ary 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 differ-

ences 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

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 combus-

tion 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. There-

fore, 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 cata-

lyst.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.

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

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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 litera-

ture.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 compre-

hensive 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 commer-

cialization 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/ZrO223,24 and

ceria-based catalysts25–27 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 develop-

ment, there are still many issues to be addressed before CPO

technology can achieve commercialization.20 One of the dis-

advantages of this route is the highly flammable mixture that

may ignite at temperatures above 250 1C. 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 1C.31 Oxygen 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 compres-

sion (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

In the past decades, two industrial consortia have been

actively working on the development of ceramic-based mem-

brane 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 (La1�xCax)yFeO3�d.32 The second consortia

led by Praxair and comprising Amoco, BP, Statoil, Phillips

Petroleum and Sasol developed the oxygen transport mem-

brane (OTM) technology for the production of oxygen from

air separation.7,36

Despite the efforts carried out to develop the ceramic mem-

brane 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 substitu-

tion 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.

Fig. 5 An oxygen transport mechanism through a ceramic mem-

brane based on a perovskite material. (Reprinted with permission

from ref. 31. Copyright 2000 Elsevier.)

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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 technol-

ogies. 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.

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 technol-

ogy; (vi) oxygen is not required; (vii) lower NOx 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

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

Fig. 6 Cross section of a plate-type reformer combining catalytic combustion/steam reforming reactions. (Reprinted with permission from

ref. 38. Copyright 2003 Elsevier.)

Fig. 7 Haldor Topsøe convection reformer. (Reprinted with permission

from ref. 40. Copyright 2001 Elsevier.)

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convert natural gas into liquids via GTL 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 conven-

tional steam reformers. The SR reaction was also carried out

at contact times below 1 ms (90 and 900 ms) 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 ms,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 con-

trolled, 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 - CnH(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:

Wn/n = (1 � a)2an�1

where Wn is the weight fraction of hydrocarbon molecules

containing n carbon atoms, a is the chain growth probability

or the probability that a molecule will continue reacting to

form a longer chain. Generally, the value of a 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.

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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 1C,

employing an iron-based catalyst and Low-Temperature

Fischer–Tropsch (LTFT), operated at lower temperatures

(200–240 1C), 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

(B2), 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 con-

version 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 B10 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.66 CoO

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 crys-

talline 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-methyl-

imidazolium 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 1C 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 me-

chanisms 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 deactiva-

tion 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

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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-5,72 FeCuMg + ZSM-5,73 Co/MCM-22,74

Co-Ni/TiO2 + ZSM-5,75,76 Co/SiO2 + Pd/b-zeolite,77 alkali

promoted Fe + ZSM-578,79 have been tried. As expected, the

use of zeolite containing FT-catalysts80–82 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.

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 func-

tion 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 profit-

able 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 limita-

tion (reactants migrating to the active sites) will reduce CO

concentration inside the pellet, thereby favouring the forma-

tion of lighter products. On the other hand, the second

mechanism impacts the re-adsorption of a-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, g-Al2O3 nano-

fibers91 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 selecti-

vity 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 (TOS4 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 distribu-

tion 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

Fig. 8 Influence of zeolite on the yield of branched C5–C8 products.

(Reprinted with permission from ref. 82. Copyright 2007 Elsevier.)

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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/a-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 block4micro monoliths4monoliths4 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.

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. How-

ever, 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 298 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.

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 con-

tains 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

Fig. 9 Activity (CO conversion) and selectivity (C5+) for different

catalyst layer thicknesses at 250 1C, 10 bar and H2/CO = 2. (Rep-

rinted with permission from ref. 97. Copyright 2011 Elsevier.)

Table 2 FT product spectra (at 2 MPa)96

Catalyst type:FT temperature/1C

Fe: fused340

Fe: precip.235

Co: supported220

Selectivity (C atom basis)CH4 8 3 4C2–C4 30 8.5 8C5–C6 16 7 8C7–160 1C (bp) 20 9 11160–350 1C (bp) 16 17.5 22+350 1C (bp) 5 51 46Water-soluble oxygenates 5 4 1a value 0.7 0.95 0.92C3 + C4: %alkenes 87 50 30C5 to C12 cut:%Alkenes 70 64 40%Oxygenates 12 7 1%Aromatics 5 0 0C13 to C18 cut:%Alkenes 60 50 5%Oxygenates 10 6 o1%Aromatics 15 0 0

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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 (o�47 1C) is needed to ensure that

the fuel remains pumpable under the low temperature condi-

tions 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 chemi-

cally 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 techno-

logies 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 1C 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 proper-

ties, due to the high degree of branching of the paraffinic

hydrocarbons. Cetane number is very high (470),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 selecti-

vities 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 func-

tion 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 hydro-

cracking 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 deter-

mining105 and the hydrogenation–dehydrogenation steps are

consequently in quasi-equilibrium, so that olefin concentration

is determined by thermodynamic factors.100 Isomerisation occurs

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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 a,a,g type are especially prone

to C–C bond rupture by b-scission,106 because they involve both

reactant and product tertiary carbenium ions, as illustrated in

Scheme 1.

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 proto-

nated 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/dehydro-

genation 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 platinum105,112–116 and palla-

dium,117 and mixed NiMo and NiW sulfides101,104,113,118 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.100

Pt-promoted HY affords a high yield of gasoline-range hydro-

carbons (490%) 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 (470%) 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 b-zeolites are relativelywide-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 pro-

ducts 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 hydro-

carbon chain inside narrow straight pores, while the middle part

would be located within wide cavities, where branching reac-

tions 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 micro-

kinetic 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

Scheme 1 C–C bond rupture by b-scission from an a,a,g-tribranchedintermediate.

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content, pour-point, cloud-point, oxidation stability and visco-

sity 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 impor-

tance 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 detri-

mental 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 hydroisomerisa-

tion, 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 produc-

tion, 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. Hydroisomerisa-

tion 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 HIDWhas

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, respec-

tively, 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

second4 third4 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 hydro-

isomerization 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 hydroisomerisa-

tion 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 cata-

lysis,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 con-

centration 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 condi-

tions (300 psig and 230 1C). 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

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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 hydro-

isomerisation. Although GTL technologies are well estab-

lished, many catalytic challenges still exist in the three steps

described above.

Regarding the generation of synthesis gas, the main technol-

ogies 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 E 2) at low steam-to-

carbon (S/C) ratio, without causing high carbon formation.

Nevertheless, fundamental studies have led to a greater under-

standing 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 profit-

able 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 out-

standing 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 technol-

ogy 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 hydro-

isomerisation, 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

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platinum supported on sulfated zirconia and on polyoxo-

cation ([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 hydroisomerisa-

tion 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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