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Transcript of 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
Dynamic Article Links
www.rsc.org/catalysis PERSPECTIVE
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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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