deactivation of cobalt and nickel catalysts in fischer ...952256/fulltext01.pdf · gas and biomass)...

Deactivation of cobalt and nickel

catalysts in Fischer-Tropsch

synthesis and methanation

Javier Barrientos

Doctoral Thesis in Chemical Engineering

KTH Royal Institute of Technology

School of Chemical Science and Engineering

Department of Chemical Engineering and Technology

Stockholm, Sweden 2016

Deactivation of cobalt and nickel catalysts in Fischer-Tropsch synthesis and

methanation

JAVIER BARRIENTOS

TRITA-CHE Report 2016:31

ISSN 1654-1081

ISBN 978-91-7729-060-5

Akademisk avhandling som med tillstånd av KTH i Stockholm framlägges till offentlig

granskning för avläggande av teknologie doktorsexamen fredagen den 23

september kl. 10:00 i sal F3, KTH, Lindstedtsvägen 26, Stockholm. Fakultetsopponent: Claude Mirodatos, Université Lyon 1, CNRS, Lyon, France.

© Javier Barrientos 2016

Tryck: Universitetsservice US-AB

“True wisdom comes to each of us when we realize how little we understand

about life, ourselves, and the world around us.”

Socrates

Abstract

i

Abstract

A potential route for converting different carbon sources (coal, natural

gas and biomass) into synthetic fuels is the transformation of these raw

materials into synthesis gas (CO and H2), followed by a catalytic step which

converts this gas into the desired fuels. The present thesis has focused on two

catalytic steps: Fischer-Tropsch synthesis (FTS) and methanation. The

Fischer-Tropsch synthesis serves to convert synthesis gas into liquid

hydrocarbon-based fuels. Methanation serves instead to produce synthetic

natural gas (SNG). Cobalt catalysts have been used in FTS while nickel

catalysts have been used in methanation.

The catalyst lifetime is a parameter of critical importance both in FTS

and methanation. The aim of this thesis was to investigate the deactivation

causes of the cobalt and nickel catalysts in their respective reactions.

The resistance to carbonyl-induced sintering of nickel catalysts

supported on different carriers (γ-Al2O3, SiO2, TiO2 and α-Al2O3) was studied.

TiO2-supported nickel catalysts exhibited lower sintering rates than the other

catalysts. The effect of the catalyst pellet size was also evaluated on γ-Al2O3-

supported nickel catalysts. The use of large catalyst pellets gave considerably

lower sintering rates. The resistance to carbon formation on the above-

mentioned supported nickel catalysts was also evaluated. Once again, TiO2-

supported nickel catalysts exhibited the lowest carbon formation rates.

Finally, the effect of operating conditions on carbon formation and

deactivation was studied using Ni/TiO2 catalysts. The use of higher H2/CO

ratios and higher pressures reduced the carbon formation rate. Increasing the

temperature from 280 °C to 340 °C favored carbon deposition. The addition

of steam also reduced the carbon formation rate but accelerated catalyst

deactivation.

The decline in activity of cobalt catalysts with increasing sulfur

concentration was also assessed by ex situ poisoning of a cobalt catalyst. A

deactivation model was proposed to predict the decline in activity as function

of the sulfur coverage and the sulfur-to-cobalt active site ratio. The results also

indicate that sulfur decreases the selectivity to long-chain hydrocarbons and

olefins.

Keywords: cobalt, nickel, Fischer-Tropsch synthesis, methanation,

deactivation, carbonyl, sintering, carbon formation, sulfur, poisoning.

Sammanfattning

iii

Sammanfattning

En potentiell väg för att omvandla olika kolkällor (kol, naturgas och

biomassa) till syntetiska bränslen är omvandlingen av dessa råvaror till

syntesgas (CO och H2), följt av ett katalytiskt steg som omvandlar denna gas

till de önskade bränslena. Denna avhandling har fokuserat på två katalytiska

steg: Fischer-Tropsch syntes (FTS) och metanisering. Fischer-Tropsch-

syntesen tjänar till att omvandla syntesgas till flytande kolvätebränslen.

Metanisering syftar istället till att producera syntetisk naturgas (SNG).

Koboltkatalysatorer har använts i FTS medan nickelkatalysatorer har använts

i metanisering.

Katalysatorns livslängd är en parameter med avgörande betydelse både

för FTS och metanisering. Syftet med denna avhandling var att undersöka

orsakerna till deaktivering av kobolt- och nickelkatalysatorerna i sina

respektive reaktioner.

Motståndet mot karbonyl-inducerad sintring av nickelkatalysatorer på

olika bärare (γ-Al2O3, SiO2, TiO2 och α-Al2O3) studerades. Nickelkatalysatorer

med TiO2 som bärare uppvisade lägre sintringshastigheter än andra

katalysatorer. Effekten av katalysatorpelletstorlek utvärderades också för

nickelkatalysatorer med γ-Al2O3 som bärare. Stora katalysatorpellets

uppvisade betydligt lägre sintringshastigheter. Beständigheten mot

kolbildning för de ovannämnda bärarbelagda nickelkatalysatorerna

utvärderades också. Åter uppvisade nickelkatalysatorer med TiO2 som

bärarmaterial de lägsta kolbildningshastigheterna. Slutligen undersöktes

effekten av driftsbetingelser på kolbildning och deaktivering vid användning

av Ni/TiO2 katalysatorer. Användningen av högre H2/CO-förhållanden och

högre tryck minskade kolbildningshastigheten. Ökning av temperaturen från

280 °C till 340 °C gynnade kolbildingen. Tillsats av ånga minskade också

koldbildningshastigheten men accelererade katalysatordeaktiveringen.

Nedgången i aktivitet hos koboltkatalysatorer med ökande

koncentration svavel bedömdes också med ex situ förgiftning av en

koboltkatalysator. En deaktiveringsmodell föreslogs för att förutsäga den

minskade aktiviteten som funktion av svavelbeläggning och förhållandet

svavel-kobolt på de aktiva sätena. Resultaten tyder på att svavel minskar

selektiviteten till långa kolväten och olefiner.

Nyckelord: kobolt, nickel, Fischer-Tropsch-syntes, metanisering,

deaktivering, karbonyl, sintring, kolbildning, svavel, förgiftning.

List of appended papers

v

List of appended papers

Paper I

Deactivation of supported nickel catalysts during CO methanation.

J. Barrientos, M. Lualdi, M. Boutonnet and S. Järås.

Applied Catalysis A: General 486 (2014) 143–149.

Paper II

CO methanation over TiO2-supported nickel catalysts: A carbon formation

study.

J. Barrientos, M. Lualdi, R. Suárez París, V. Montes, M. Boutonnet and S.

Järås.

Applied Catalysis A: General 502 (2015) 276–286.

Paper III

Further insights into the effect of sulfur on the activity and selectivity of

cobalt-based Fischer-Tropsch catalysts.

J. Barrientos, V. Montes, M. Boutonnet and S. Järås.

Catalysis Today (2015), http://dx.doi.org/10.1016/j.cattod.2015.10.039.

Paper IV

The effect of catalyst pellet size on nickel carbonyl-induced particle sintering

under low temperature CO methanation.

J. Barrientos, N. González, M. Lualdi, M. Boutonnet and S. Järås.

Applied Catalysis A: General 514 (2016) 91-102.

NOTE: All papers are reproduced with permission from the copyright holders.

http://dx.doi.org/10.1016/j.cattod.2015.10.039

Contribution to appended papers

vii

Contribution to appended papers

The author of this thesis was the main responsible for these publications. All

co-authors had substantial contributions to the conception of the work,

planning, acquisition, analysis or interpretation of the data and results.

Paper I

I had the main responsibility for writing this paper. I performed most of the

experimental work.

Paper II


experimental work. V. Montes performed the XPS, ICP-MS and TEM

measurements.

Paper III


experimental work and modelling. The ICP-MS analyses were performed at

ALS Scandinavia.

Paper IV

I had the main responsibility for writing this paper. N. González performed

most of the experimental work under my supervision. I performed all the mass

and heat transfer calculations and modelling.

Other publications and conference presentations

ix


Peer-reviewed publications

R. Suárez París, L. Lopez, J. Barrientos, F. Pardo, M. Boutonnet, S. Järås,

Catalytic conversion of biomass-derived synthesis gas to fuels, Catalysis:

Volume 27, The Royal Society of Chemistry (2015) 62-143.

R. Suárez París, M. E. L' Abbate, L. F. Liotta, V. Montes, J. Barrientos, F.

Regali, M. Boutonnet, S. Järås, Hydroconversion of paraffinic wax over

platinum and palladium catalysts supported on silica-alumina, Catalysis

Today (2015), http://dx.doi.org/10.1016/j.cattod.2015.11.026.

Presentations in conferences (presenting author in bold)

J. Barrientos, N. González, M. Boutonnet, S. Järås – The effect of Zr, Mg,

Ba and Ca oxide promoters on the stability of Ni/γ-Al2O3 methanation

catalysts. Poster presentation at the 13th Nordic Symposium on Catalysis,

Lund, Sweden (2016).

J. Barrientos, B. Venezia, V. Garcilaso de la Vega, M. Boutonnet, S. Järås –

The effect of Zr on the catalytic performance of Co/γ-Al2O3 Fischer-Tropsch

catalysts. Oral and poster presentation at the 13th Nordic Symposium on

Catalysis, Lund, Sweden (2016).

J. Barrientos, N. González, M. Lualdi, M. Boutonnet, S. Järås – The effect

of catalyst pellet size on nickel carbonyl-induced particle sintering under low

temperature CO methanation. Oral presentation at the 11th Natural Gas

Conversion Symposium, Tromsø, Norway (2016).

J. Barrientos, B. Venezia, V. Garcilaso de la Vega, M. Boutonnet, S. Järås –

The effect of Zr on the catalytic performance of Co/γ-Al2O3 Fischer-Tropsch

catalysts. Poster presentation at the 11th Natural Gas Conversion Symposium,

Tromsø, Norway (2016).



x

J. Barrientos, M. Boutonnet, S. Järås – The effect of sulfur on the activity

and selectivity of cobalt-based Fischer-Tropsch catalysts. Oral presentation at

Syngas Convention 2. Cape Town, South Africa (2015).

J. Barrientos, M. Boutonnet, S. Järås – The effect of sulfur loading on the

catalytic performance of cobalt-based Fischer-Tropsch catalysts. Poster

presentation at the 12th Nordic Symposium on Catalysis. Oslo, Norway (2014).

J. Barrientos, M. Boutonnet, S. Järås - Cobalt- and iron-based FT catalysts

for biomass-derived diesel production. Poster presentation at the 4th

International Workshop of COST Action CM903 (UBIOCHEM). Valencia,

Spain (2013).

J. Barrientos, M. Lualdi, M. Boutonnet, S. Järås – Deactivation of

Methanation Catalysts: A Study of Carbon Formation on Titania-Supported

Nickel Catalysts. Poster presentation at the 11th European Congress on

Catalysis – Europacat XI. Lyon, France (2013).

J. Barrientos, M. Boutonnet, S. Järås - A comparison of cobalt- and iron-

based Fischer-Tropsch catalysts for biomass-derived diesel and SNG

coproduction. Poster presentation at the 2nd International Conference in

Catalysis for Renewable sources: Fuel, Energy and Chemicals, CR2. Lund,

Sweden (2013).

M. Lualdi, J. Barrientos, M. Boutonnet, S. Järås – Deactivation of

methanation catalysts. Poster presentation at the 15th Nordic Symposium on

Catalysis. Mariehamn, Finland (2012).

Table of contents

xi

Table of contents

PART I INTRODUCTION ................................................................................................................................. 1

CHAPTER 1 SETTING THE SCENE ...............................................................................................................................3 1.1. Scope of the work ............................................................................................................................................ 4

CHAPTER 2 THE FUEL SYNTHESIS PROCESS .................................................................................................................5 2.1. Syngas production ............................................................................................................................................ 5

2.1.1. Coal and biomass gasification .................................................................................................................. 5 2.1.2. Reforming of natural gas .......................................................................................................................... 7

2.2. Syngas conditioning and purification ............................................................................................................... 9 2.2.1. Adjustment of the H2/CO ratio and removal of CO2 ................................................................................. 9 2.2.2. Removal of syngas impurities ................................................................................................................. 10

2.3. The Fischer-Tropsch synthesis ........................................................................................................................ 11 2.3.1. Cobalt-based FT catalysts ....................................................................................................................... 13 2.3.2. Low-temperature Fischer-Tropsch reactors ........................................................................................... 17 2.3.3. Upgrading of Low-temperature Fischer-Tropsch products .................................................................... 20

2.4. Methanation .................................................................................................................................................. 21 2.4.1. Nickel-based methanation catalysts ....................................................................................................... 22 2.4.2. Methanation reactors ............................................................................................................................ 24

2.5. Co-production of liquid fuels and SNG via Fischer-Tropsch synthesis and methanation ............................... 28 CHAPTER 3 DEACTIVATION OF COBALT- AND NICKEL-BASED CATALYSTS .........................................................................29

3.1. Carbonyl-induced metal particle sintering ..................................................................................................... 29 3.1.1. Nickel carbonyl-induced sintering .......................................................................................................... 29 3.1.2. Cobalt carbonyl-induced sintering ......................................................................................................... 31

3.2. Hydrothermal sintering .................................................................................................................................. 31 3.2.1. Hydrothermal sintering during methanation ......................................................................................... 33 3.2.2. Hydrothermal sintering during FTS......................................................................................................... 33

3.3. Re-oxidation ................................................................................................................................................... 34 3.3.1. Re-oxidation of cobalt nanoparticles...................................................................................................... 34 3.3.2. Re-oxidation of nickel nanoparticles ...................................................................................................... 35

3.4. Formation of inactive metal-support compounds.......................................................................................... 36 3.5. Carbon formation ........................................................................................................................................... 37

3.5.1. Carbon formation during FTS ................................................................................................................. 38 3.5.2. Carbon formation during methanation .................................................................................................. 39

3.6. Poisoning by syngas impurities ...................................................................................................................... 42 3.6.1. Poisoning of cobalt FT catalysts .............................................................................................................. 42 3.6.2. Poisoning of nickel methanation catalysts ............................................................................................. 43

3.7. Other deactivation causes .............................................................................................................................. 44 3.7.1. Attrition and crushing ............................................................................................................................. 44 3.7.2. Weak anchoring of cobalt particles to the carrier .................................................................................. 45 3.7.3. Surface reconstruction ........................................................................................................................... 45 3.7.4. Decoration of metal particles by support species .................................................................................. 45

PART II EXPERIMENTAL .............................................................................................................................. 47

CHAPTER 4 CATALYST PREPARATION AND CHARACTERIZATION .....................................................................................49 4.1. Catalyst preparation ....................................................................................................................................... 49

4.1.1. Incipient wetness impregnation ............................................................................................................. 49 4.1.2. Wet impregnation .................................................................................................................................. 50

4.2. Characterization of supports and catalysts .................................................................................................... 50 4.2.1. N2 adsorption ......................................................................................................................................... 50 4.2.2. Hg intrusion ............................................................................................................................................ 51 4.2.3. Temperature-programmed reduction .................................................................................................... 52 4.2.4. Temperature-programmed hydrogenation ............................................................................................ 52 4.2.5. H2 static chemisorption .......................................................................................................................... 53

Table of contents

xii

4.2.6. X-ray diffraction ...................................................................................................................................... 53 4.2.7. X-ray photoelectron spectroscopy .......................................................................................................... 54 4.2.8. Inductively coupled plasma mass spectrometry ..................................................................................... 54 4.2.9. Scanning and transmission electron microscopy .................................................................................... 55

CHAPTER 5 CATALYTIC TESTING ............................................................................................................................ 57 5.1. Reactor setup ................................................................................................................................................. 57 5.2. Gas analysis and data treatment .................................................................................................................... 60

PART III RESULTS AND DISCUSSION ............................................................................................................ 63

CHAPTER 6 CARBONYL-INDUCED NICKEL PARTICLE SINTERING ..................................................................................... 65 6.1. Support and nickel particle size effects on carbonyl-induced sintering (Paper I) ........................................... 65 6.2. Catalyst pellet size effects on carbonyl-induced sintering (Paper IV) ............................................................. 68

CHAPTER 7 CARBON FORMATION IN METHANATION ................................................................................................. 75 7.1. Support effects on carbon formation (Paper I) ............................................................................................... 75 7.2. Nickel particle size effects on carbon formation (Paper I) .............................................................................. 78 7.3. Effect of the H2/CO ratio on carbon formation (Paper II) ............................................................................... 79 7.4. Effect of pressure on carbon formation (Paper II) .......................................................................................... 81 7.5. Effect of temperature on carbon formation (Paper II) ................................................................................... 83 7.6. Effect of steam addition on carbon formation (Paper II) ................................................................................ 85

CHAPTER 8 SULFUR POISONING OF COBALT FISCHER-TROPSCH CATALYSTS ..................................................................... 91 8.1. Effect of sulfur on activity (Paper III) .............................................................................................................. 91 8.2. Effect of sulfur on selectivity (Paper III) .......................................................................................................... 94

CHAPTER 9 CONCLUSIONS AND OUTLOOK ............................................................................................................... 99 9.1. Carbonyl-induced nickel particle sintering ..................................................................................................... 99 9.2. Carbon formation in methanation ................................................................................................................ 100 9.3. Sulfur poisoning of cobalt Fischer-Tropsch catalysts .................................................................................... 101

Acknowledgements ................................................................................................................................ 103 List of abbreviations ............................................................................................................................... 105 References .............................................................................................................................................. 109

1

Part I

Introduction

Chapter 1 – Setting the scene

3

Chapter 1

Setting the scene

Despite great advances in alternative and renewable energies during the

last decades, humankind is still highly dependent on fossil energy sources. In

2013 oil, coal and natural gas constituted 31.1 %, 28.9 % and 21.4 %,

respectively, of the world’s total primary energy supply [1]. This dependency

is even more pronounced in the transportation sector, which is totally

dominated by oil-derived fuels.

This extreme reliance on these non-renewable sources is an important

concern, mainly, for three reasons:

1) Depletion of fossil fuels will occur in the nearby future. The extinction

of coal, gas and oil reserves is estimated to occur in 110, 54 and 52 years,

respectively [2].

2) Oil and gas reserves are heterogeneously distributed around the world.

Oil is primarily located in the Middle East, Venezuela and Canada [3].

Gas reserves are mainly concentrated in the Middle East and Russia.

The energy security in importing countries is constantly subjected to the

exporting countries’ political stability.

3) The use of fossil fuels and the release of greenhouse gas emissions

(mainly CO2) contribute to the climate change. Moreover, their

combustion results in the emission of environmentally harmful

compounds (e.g. SO2 and NOx) responsible for local disasters such as

acid rain and smog.

These issues have encouraged governments to move towards a more

independent, sustainable and environmentally friendly energy system. Efforts

and resources are placed in electricity production via renewable energies

(solar, wind, hydroelectric, geothermal), development of hybrid and electric

vehicles, CO2 capture and storage, CO2 reutilization and replacement of fossil

fuels by biofuels.

Chapter 1 – Setting the scene

4

A potential route for the production of biofuels is the thermochemical

conversion of biomass into synthesis gas (CO and H2) followed by the catalytic

conversion of this gas into the desired fuels. The present thesis has focused on

two catalytic steps known as Fischer-Tropsch synthesis (FTS) and

methanation. The FTS serves to transform synthesis gas into liquid

hydrocarbon-based fuels (e.g. bio-diesel and bio-jet fuel). Methanation serves

instead to produce synthetic natural gas (i.e. bio-SNG).

It must be mentioned that synthesis gas (usually called “syngas”) can

also be produced from non-renewable sources such as coal and natural gas.

Both FT and methanation processes are already being applied at commercial

scale using these fossil raw materials. Unfortunately, no commercial FT or

methanation processes exist using biomass as feedstock, yet.

1.1. Scope of the work

The catalyst lifetime is a crucial parameter which affects the

profitability, or even feasibility, of a process. Fundamental understanding of

catalyst deactivation causes, safe operating conditions, and use of stable

catalyst materials in different situations is therefore, valuable.

The objective of the present work was to investigate the deactivation of

Fischer-Tropsch and methanation catalysts during the conversion of CO and

H2 into hydrocarbons. Cobalt-based catalysts were used in the Fischer-

Tropsch synthesis while nickel-based catalysts were used in methanation.

Chapter 2 – The fuel synthesis process

5

Chapter 2

The fuel synthesis process

The production process of Fischer-Tropsch fuels and synthetic natural

gas (SNG) from carbon sources (coal, natural gas and biomass) can be divided

into four steps:

1) Syngas production

2) Syngas conditioning and purification

3) Fuel synthesis (Fischer-Tropsch synthesis and methanation)

4) Fuel upgrading

The present chapter offers a brief overview of the entire process. Special

focus is given, nonetheless, to the Fischer-Tropsch and methanation steps.

The catalysts, reactors and other aspects concerning these two reactions are

presented in this chapter.

2.1. Syngas production

Syngas (H2 and CO) can be produced via gasification of coal and

biomass or via reforming of natural gas. The Fischer-Tropsch process is

named according to the raw material: coal-to-liquids (CTL), biomass-to-

liquids (BTL) and gas-to-liquids (GTL). If the desired product is synthetic

natural gas (SNG), the process can be named: coal-to-gas (CTG) and biomass-

to-gas (BTG).

2.1.1. Coal and biomass gasification

Coal gasification is usually carried out in entrained-flow (EF) gasifiers.

In these reactors, the gasifying agent (usually oxygen and steam) is co-fed with

dry pulverized coal particles. The gasifier usually operates at high pressure (>

25 bar), high temperatures (> 1200 °C) and very high turbulent flows. The coal

residence time in the gasifiers is in the range of 0.5-10 seconds [4]. Due to the


6

high temperatures, the concentration of CH4 in the outlet gas is rather low [5].

The size of EF gasifiers varies between 50 and 1000 MWth [6].

Biomass gasification can rarely be carried out in entrained-flow

gasifiers. Entrained-flow gasifiers require a fuel material such as a liquid,

slurry or solid which can be easily pulverized [7]. This technology is therefore

limited to biomass feedstocks such as black liquor, which is the spent cooking

liquid from the pulp and paper industry.

The most common biomass gasification technologies are fixed-bed and

fluidized bed gasifiers. Fixed bed technologies can be either downdraft (co-

current) or updraft (counter-current). In downdraft gasifiers the biomass is

fed from the top and the oxidizing agent flows downwards. In updraft gasifiers

the oxidizing agent flows upstream while biomass is introduced from the top

of the gasifier [8]. These two technologies are however limited to small scales

(< 10 MWth) [6].

Fluidized-bed technologies can be used for mid-large scales (between 5

and 100 MWth) [6]. There are two main modalities: bubbling fluidized bed

(BFB) or circulating fluidized bed (CFB) gasifiers. In BFB’s the solid particles

are suspended by the flowing gas. The gas velocity is approximately 1 m/s. In

CFB’s the solid particles are fluidized and so transported upwards with the exit

gas. The gas velocity in these gasifiers is in the range of 3 to 10 m/s. The

particles must therefore be separated by cyclones and recirculated [8].

The syngas composition depends on various factors such as the

feedstock composition, temperature, pressure and oxidizing agent (air,

oxygen and/or steam) [9]. Table 1 presents the outlet gas composition of three

gasifiers to offer an idea of the syngas composition resulting from biomass and

coal gasification. These examples have been chosen because these three

gasifiers operate with oxygen (instead of air) as oxidizing agent. Gasification

with O2 is convenient for this application because the presence of nitrogen in

the process results in some technical challenges [10-12]. Among others,

nitrogen acts as an inert gas and its presence results in larger downstream

equipment [13]. Furthermore, removal of N2 is required if syngas is aimed to

be converted into H2 or SNG.

Gasification with pure O2 therefore requires the installation of an air

separation unit. The investment costs of such a unit are significant [13].

Gasification with air may be, however, a more feasible option in some

situations such as in small-scale BTL applications [14, 15].


7

Table 1. Outlet gas composition and operating conditions of different types of gasifiers in the absence of N2.

Gasifier type Updraft fixed-beda CFBb EFc

Feedstock Wood chips Wood chips Coal Capacity (ton/day) 1 96 2155 Gasifying agent O2/steam O2 O2/steam Feed H2O/O2 (vol/vol) 2 0 Low (~0)

Temperature (°C) - 950 1400

Pressure (MPa) Atmospheric 1.8 3.0 H2 (dry vol%) 37 19 25 CO (dry vol%) 23 19 69 CO2 (dry vol%) 34 45 4 CH4 (dry vol%) 5 13 0

a Adapted from Ref. [16]. b Adapted from Ref. [17]. c Adapted from Refs. [6, 18].

2.1.2. Reforming of natural gas

There are different routes for converting natural gas and shale gas into

synthesis gas. These routes and their corresponding reactions are listed in

Table 2.

Table 2. Gas reforming routes (adapted from Refs. [19, 20]).

Gas reforming routes Reactions ΔH° (kJ/mol) Steam reforming (SR) 𝐶𝐻4 + 𝐻2𝑂 ⇄ 𝐶𝑂 + 3𝐻2

𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐶𝑂2 + 𝐻2

206 -41

Autothermal reforming (ATR)

𝐶𝐻4 + 𝐻2𝑂 ⇄ 𝐶𝑂 + 3𝐻2 𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐶𝑂2 + 𝐻2 𝐶𝐻4 + 1.5𝑂2 ⇄ 𝐶𝑂2 + 2𝐻2𝑂

206 -41 -520

Partial oxidation (POX) 𝐶𝐻4 + 0.5𝑂2 ⇄ 𝐶𝑂 + 2𝐻2 -38 Dry reforming (DR) 𝐶𝐻4 + 𝐶𝑂2 ⇄ 2𝐶𝑂 + 2𝐻2 247

Steam reforming (SR) of natural gas is usually carried out in

multitubular fixed bed reactors [19, 21]. The endothermic reforming reaction

occurs inside the reactor tubes which are heated by burning gas in the outer

shell. This technology is suitable for producing H2-rich synthesis gas. The

Fischer-Tropsch reaction, when using cobalt-based catalysts, requires a

H2/CO ratio of ca. 2. The ratio obtained in SR is higher but can be adjusted by

co-feeding CO2 in the reformer or by downstream separation of H2 [20].

The preferred route for GTL applications is usually oxygen-blown

autothermal reforming [20, 22]. Even though this technology requires the


8

production of O2, the ATR reactor is relatively simple (compared to SR

reactors) and the syngas production process becomes less expensive than SR

at large scales [20]. Moreover, a H2/CO ratio of 2 in autothermal reformers

can be achieved [23] (see Table 3). The scenario may be different in the case

of off-shore GTL applications. In this case, the technology of choice may be

steam reforming [24, 25] or air-blown autothermal reforming [26].

Table 3. Outlet gas composition from a pilot autothermal reformer (adapted from Ref. [20]).

Component Concentration in product gas (vol%)a H2 56.8 N2 0.2 CO 29 CO2 2.9 CH4 1 H2O 10.1

a Inlet H2O/C = 0.21, inlet O2/C = 0.59, 24.5 bar, exit temperature=1057 ◦C.

Partial oxidation is theoretically a very convenient choice for the

Fischer-Tropsch process since it results in a H2/CO ratio of 2. Nevertheless,

the safety problems related to premixing oxygen and gas make this route not

competitive [21, 27]. Also, dry reforming is hardly feasible. The reaction

results in a non-complete conversion of methane due to the thermodynamics

and the process economy depends on the cost of the CO2 available [22].

Moreover, the resulting H2/CO ratio is low if syngas is intended to be used in

the Fischer-Tropsch reaction.

The reforming of methane is catalyzed by nickel-based catalysts which

are very sensitive to sulfur impurities in the feed gas [28]. Therefore, the gas

needs to be purified prior to entering the reactor. Desulfurization of the feed

gas is carried out by hydrodesulfurization of organosulfur compounds (i.e.

conversion of these to H2S) followed by capture of H2S in a ZnO guard bed [21,

29].

Longer hydrocarbons (C2+) present in the feed gas are likely to cause

thermal cracking and carbon formation on nickel catalysts at high

temperatures. Therefore, these are usually pre-reformed at lower

temperatures (350-550 °C) before entering the main reforming unit [30].


9

2.2. Syngas conditioning and purification

The syngas obtained from reforming or gasification has a composition

which is not suitable for fuel synthesis. As can be seen in Table 1, the syngas

obtained from biomass or coal gasification usually has a low H2/CO ratio

(between 0.4 and 1.6). This ratio needs to be adjusted to the stoichiometry of

the Fischer-Tropsch and methanation reactions. Moreover, the gas can

contain significant amounts of CO2. Carbon dioxide usually behaves as an

inert gas but in some situations can have negative effects on the FT product

distribution [31]. Finally, the gas contains different kind of impurities (e.g.

sulfur, nitrogen compounds, alkali…) which dramatically affect the lifetime of

FT and methanation catalysts.

2.2.1. Adjustment of the H2/CO ratio and removal of CO2

The H2/CO ratio can be increased by means of the so-called water gas

shift (WGS) reaction. The reaction is described as follows:

𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐶𝑂2 + 𝐻2 (𝛥𝐻° = −41𝑘𝐽/𝑚𝑜𝑙)

The WGS reaction is exothermic and therefore the CO conversion at

equilibrium is favored at low temperatures. The WGS reaction is usually

carried out in adiabatic fixed bed reactors in a wide temperature range (200-

500 °C). The operating temperature depends on the catalyst employed and the

desired H2 content in the gas [19]. Different types of catalyst can be used. The

most common catalysts, used in the production of pure H2, are Fe3O4-based

catalysts and CuO-based catalysts. Nevertheless, CoMo-based catalysts can be

used as well (as in Haldor Topsoe’s coal-to-SNG process [32]).

The resulting H2/CO ratio can be increased in different ways. In some

cases, it may be convenient to use a small part of the syngas to produce pure

hydrogen (or hydrogen rich-gas) via WGS and CO2 removal. In this way,

hydrogen is not only used to increase the H2/CO ratio of the process gas but

also used for other process needs (e.g. catalyst reduction, regeneration or

hydrocracking of Fischer-Tropsch waxes). Hydrogen can also be produced via

reforming and subsequent WGS of the undesired short-chain hydrocarbons

produced in the Fischer-Tropsch reaction [33].

It must be mentioned that the Fischer-Tropsch reaction can also be

carried out using Fe-based catalysts which are active for both FT and WGS. In

this case, the H2/CO ratio of the gas resulting from coal or biomass gasification


10

does not need to be adjusted. Indeed, Fe-based catalysts are the choice in CTL

applications [34, 35]. Nevertheless, according to Botes et al. [36], there is no

evident benefit in choosing iron over cobalt catalysts in this situation.

Production of H2 is in any case required for other process needs, as explained

previously. Actually, the pilot BTL plant started by CHOREN employed cobalt-

based catalysts [37, 38].

Removal of CO2 is commonly performed using sorbents. The most

common technologies make use of physical sorbents such as methanol, in the

RectisolTM process [39], or a mixture of dimethyl ether and polyethylene

glycol, in the SelexolTM process [40-42]. These absorption processes, often

called acid gas removal processes, also serve to reduce the H2S in the process

gas. The removal of acid gases can also be performed by means of chemical

sorbents such as amines [40]. Pressure-swing-adsorption technologies using

activated carbon are also being considered to replace the conventional

absorption processes [43, 44].

2.2.2. Removal of syngas impurities

There are several compounds present in syngas which act as poisons for

cobalt- and nickel-based catalysts. Table 4 presents a list of typical impurities

and the recommended maximum concentration of these in the syngas feed.

Table 4. Impurities in syngas and FT catalyst tolerance levels (adapted from Refs. [45, 46]).

Impurities Maximum concentration (μg/Nm3) S compounds (e.g. H2S and COS) 10-20 N compounds (e.g. NH3 and HCN) 10-20 Halides (e.g. HCl, HF) 10 Alkali and alkaline metals 10 Tars 100-1000

Compounds such as COS and HCN are converted via hydrolysis to H2S

and NH3, respectively [47-49]. H2S, NH3, HCl, alkali and alkaline metals are

removed by means of sorbents at low temperatures [50]. Tars are usually

removed via steam reforming [51-53] and/or in wet scrubbers at low

temperatures. More information regarding syngas cleaning technologies can

be found in recent reviews [45, 46].


11

2.3. The Fischer-Tropsch synthesis

The Fischer-Tropsch synthesis is an exothermic reaction between H2

and CO which leads to water and a wide variety of hydrocarbons (gas, liquid

and waxes). The FT reaction leads mainly to n-paraffins and α-olefins and, to

a minor extent, branched hydrocarbons and oxygenates. The reaction is

usually described as follows [54, 55]:

𝐶𝑂 + 2𝐻2 → −𝐶𝐻2 − + 𝐻2𝑂 (∆𝐻° = −165 𝑘𝐽/𝑚𝑜𝑙)

The selectivity to different hydrocarbons depends on the reaction

temperature, pressure, feed composition and the catalyst employed. Long-

chain hydrocarbons are favored at low temperatures, high pressures and low

H2/CO ratios. The main metals that catalyze this reaction are iron, cobalt,

nickel and ruthenium. Nevertheless, only the two former metals have found

industrial application. Nickel is very selective to methane, which makes it

unattractive if the aim is to produce long chain hydrocarbons (i.e. liquid fuels)

[56]. Ruthenium is a highly active and selective FT catalyst [57-59]. However,

its high price has limited its use to scientific studies only.

Due to the high price of cobalt, compared to iron, there is a need to

maximize its specific metallic surface area. For that purpose, the active phase

(cobalt) is deposited as nanoparticles on a support (commonly Al2O3, TiO2 or

SiO2). Iron-based catalysts are commonly used in CTL due to their intrinsic

WGS activity and the possibility of using H2-poor syngas. Cobalt-based

catalysts, which are almost inactive for WGS, are used in GTL applications.

Nevertheless, as explained previously, there are no hard rules for such a

catalyst choice. Both catalysts can be used independently of the feedstock

material [36].

There are two main Fischer-Tropsch process modalities [36]: high-

temperature Fischer-Tropsch (HTFT) and low-temperature Fischer-Tropsch

(LTFT). The HTFT process operates in the range of 300-350 °C using Fe-based

catalysts. The main desired products are short-chain olefins, oxygenates and

hydrocarbons in the gasoline range. The LTFT process operates in the range

of 200-240 °C. Both Fe- and Co-based catalysts can be employed. The main

desired products are linear long-chain paraffins (middle distillates and

waxes). The FT waxes are later hydrocracked to maximize the yield to middle

distillates (jet fuel and diesel cut) [60].


12

The FT reaction is generally described as a polymerization reaction in

which a hydrocarbon chain grows by insertion of monomers containing one C

atom. The mechanism is usually divided into three steps: chain initiation

(monomer formation), chain growth and termination. The FT mechanisms

can be divided into two types: mechanisms in which the chain growth

proceeds via incorporation of CH2 monomers and mechanisms in which it

occurs via insertion of other monomers (e.g. CO and/or enols). Detailed

illustrations of the different FT mechanisms can be found in one of the present

author’s publications (not included in the present thesis) [61].

It is generally accepted that parallel mechanisms can occur on the

catalyst surface during FTS [62, 63]. Nevertheless, there is no full consensus

on the mechanistic details of formation of CH2 monomers [64-66]. Two

monomer formation pathways are proposed. In one pathway (known as the

“Satchler-Biloen mechanism” or direct CO dissociation), CO adsorbs and

directly dissociates into C and O adatoms. Afterwards, adsorbed C atoms are

hydrogenated and converted into CH2 [67]. In the other pathway (often called

H-assisted CO dissociation), H is bound to CO before it dissociates [68].

A useful tool for roughly estimating the FT product distribution,

independently of the reaction mechanism, is the so-called Anderson-Schulz-

Flory (ASF) model. This model has only one assumption: the probability of

chain growth (α) is independent of the hydrocarbon chain length (n) [69].

Based on this assumption, it is possible to derive an equation (see equation

2.1) relating the probability of chain growth (α) and the molar fraction (xn) of

hydrocarbons with the same carbon number (n) [70]. An analogous

expression (see equation 2.2) can also be derived in terms of mass fraction

(wn) [19, 71].

𝑥𝑛 = 𝛼𝑛−1 ∙ (1 − 𝛼) (2.1)

𝑤𝑛

𝑛= 𝛼𝑛−1 ∙ (1 − 𝛼)2 (2.2)

In Figure 1, the mass fraction of different hydrocarbon groups is plotted

against “α” according to equation 2.2. The blue and orange colored regions

indicate the typical HTFT and LTFT operational regimes, respectively [36]. It

is also interesting to mention that a new Fischer-Tropsch operating mode is

initiating in China using an intermediate temperature (240-290 °C) [72]. The

main desired products are hydrocarbons in the diesel range. This modality is


13

known as “Light fraction process technology” (LFPT) [73] or middle-

temperature Fischer-Tropsch (MTFT) [74].

The present thesis has focused on the utilization of cobalt-based

catalysts in the LTFT process. The following sections will therefore only

provide information regarding these catalysts and LTFT reactor technologies.

Excellent publications and reviews can be found on Fe-based FT catalysts [75-

77] and HTFT reactors [78-80].

Figure 1. Anderson-Schulz-Flory FT product distribution as function of the chain growth probability.

2.3.1. Cobalt-based FT catalysts

Cobalt catalysts are used by several companies (e.g. Sasol, Shell, Qatar

Petroleum, Statoil, Chevron) due to their suitability in the production of

middle distillates with high cetane number [81]. Cobalt catalysts are highly

active and selective to linear paraffins [55]. The most common carriers used

commercially are γ-Al2O3, SiO2 and TiO2 (mainly anatase).


14

There are different methods for depositing cobalt on the surface of these

carriers. The most common method is the impregnation of the carrier with a

liquid solution containing a cobalt salt, followed by drying and calcination

[82]. The resulting material is a carrier containing Co3O4 nanoparticles. Since

the active phase for FT is Co0, the catalyst needs to be reduced prior to

reaction.

Small nanoparticles can strongly interact with the support and form

spinel compounds that are difficult to reduce. This phenomenon is

particularly important in the case of Co/Al2O3 catalysts. Temperatures as high

as 700 °C are required to reduce CoAl2O4 species [83]. In order to facilitate the

catalyst reducibility, small amounts of a second metal (e.g. Pt, Re, Ru) are

added [84-86]. These metals are usually called reduction promoters.

Alternatively, small amounts of metal oxides (structural promoters) can be

added to the carrier (e.g. Zr, Ti, La) [87-93]. Other structural promoters can

be used to modify the FT product selectivity and other properties. For

instance, promotion with Mn favors the formation of long chain hydrocarbons

and olefins [94-96]. Yet, it should be noticed that excessive amounts of

promoters can have adverse effects. Table 5 presents a list of commercial-type

cobalt catalysts and their composition.

Table 5. Cobalt FT catalysts used and/or patented by FT synthesis companies*.

Company Support Reduction promoter

Structural promoter

References

Sasol γ-Al2O3 Pt Si [97]

Shell TiO2 Mn, V [98] GTL.F1 (Statoil) NiAl2O4 Re [99] ENI/IFP/Axens γ-Al2O3 Si [100]

Nippon Oil SiO2 Ru Zr [101] Syntroleum γ-Al2O3 Ru Si, La [102, 103]

BP ZnO [104, 105] Exxon Mobil TiO2 Re γ-Al2O3 [58, 106]

ConocoPhilips γ-Al2O3 Ru, Pt, Re B [107]

Compact GTL Al2O3 Ru, Pt [108] Oxford Catalysts/Velocys

SiO2 Pt, Re Ti [109]

*Adapted from Refs. [110, 111]. Deduced from literature and patents. The actual components may

differ from those used in pilot and commercial reactors.

The activity and selectivity of a cobalt catalyst is dependent on several

parameters such as the metal dispersion, degree of reduction, carrier and

promoters used. Probably, the most important factor which should be


15

considered when preparing a cobalt-based catalyst is the size of the cobalt

crystallites. Catalysts consisting of small cobalt particles (< 6 nm) present low

activity and selectivity to long chain hydrocarbons [112-114]. This particle size

effect is illustrated in Figure 2, where the turnover frequency (TOF) is plotted

as function of the cobalt particle size. The TOF is a measure of intrinsic activity

which is defined, in this context, as the number of CO molecules converted per

second per cobalt active site. As can be observed in the figure, the TOF

increases with increasing cobalt particle size up to ca. 6 nm. The TOF is

nevertheless constant for larger particles, i.e. the reaction is structure

insensitive for particles larger than 6 nm.

Figure 2. The influence of cobalt particle size on the TOF (220 °C, H2/CO = 2, 1 bar). Reproduced from Bezemer

et al. [113], with permission from ACS (American Chemical Society).

The scientific explanation for this unusual particle size effect is still a

matter of debate. Den Breejen et al. [115] propose that the lower activity found

on small particles is explained by a longer residence time of reaction

intermediates (CHx) and a lower coverage of CHx on the cobalt surface. The

authors also ascribe the higher selectivity to short chain hydrocarbons to a

higher H2 coverage. Yet, Prieto et al. [112] propose a different explanation

based on the fact that cobalt particles flatten during FTS. According to their

work, partially oxidized interfacial Co-support species form during FTS. The

relative concentration of these species is higher in flattened small

nanoparticles than in flattened large particles.

Another typical property of cobalt FT catalysts is their deviation from

the ideal ASF product distribution. This deviation is illustrated in Figure 3,

which presents the logarithm of “xn” (molar fraction) as function of “n”

(carbon number) for an actual cobalt-FT catalyst and for an ideal catalyst


16

following the ASF model. As can be seen in the figure, the ASF selectivity

follows a linear function when plotted in this manner. This statement can also

be easily understood by expressing equation 2.1 in logarithmic form (see

equation 2.3).

ln(𝑥𝑛) = 𝑛 ∙ ln(𝛼) + ln (1 − 𝛼

𝛼) (2.3)

Three main discrepancies are usually found between the actual

selectivity and the ASF model. Firstly, the actual selectivity to C1 is higher.

Secondly, the selectivity to C2 is lower. Finally, the actual FT selectivity

presents a non-constant “α” and apparently increases with increasing carbon

number.

0 5 10 15 20 25 30

-8

-7

-6

-5

-4

-3

-2

-1

0

ln (

xn)

Carbon number (n)

Ideal ASF distribution ( = 0.75)

Real FT product distribution

Figure 3. Deviation of cobalt FT catalysts from the ideal ASF product distribution (author’s own elaboration). The

real FT product distribution points correspond to a cobalt-based catalyst tested at 220 °C, 2 bar H2 and 2 bar CO (adapted from Ref. [116]).

The higher selectivity to methane has been explained by the presence of

different active sites, some sites active for methanation and others for

polymerization [117]. This explanation is, however, conflictive with other


17

studies which have shown that there is a relationship between the selectivity

to methane and the selectivity to long chain hydrocarbons [118]. This relation

suggests that both methane and long chain hydrocarbons originate from the

same sites. Another explanation for this high C1 selectivity is the increased

mobility of the methane precursor [119]. The low selectivity to C2 has been

explained by a very high re-adsorption rate of ethylene compared to other

olefins [120, 121].

Different explanations exist for this higher selectivity to long-chain

hydrocarbons. Iglesia et al. [120, 122] proposed that α-olefins re-adsorb on the

catalyst surface and are further converted to longer chain hydrocarbons. This

explanation is in line with the fact that the olefin-to-paraffin ratio decreases

with increasing carbon number. Another more recent explanation is based on

the hypothesis that long hydrocarbons present lower desorption rates, higher

residence times and thus lead to further polymerization [123, 124]. Other

explanations are based on the fact that the actual FT selectivity resembles a

combination of two ideal ASF models with different “α”. One explanation is

the presence of two different kinds of active site with different chain growth

probabilities [75]. Another explanation is the presence of two chain growth

mechanisms with two different monomers [116] or two hydrocarbon chain

termination pathways [119].

2.3.2. Low-temperature Fischer-Tropsch reactors

The main concern in the design of a LTFT reactor is the removal of the

reaction heat. Isothermal (or nearly isothermal) operation and low

temperatures are required to maximize the yield to long chain hydrocarbons.

At present, only two commercial reactor technologies exist: slurry bubble

column reactors (SBCR’s) and multitubular fixed bed reactors (MTFB’s).

However, a third technology, known as microchannel reactors, is under

commercialization.

Another important aspect in the design of a FT reactor is to limit the

conversion per pass and recycle the tail-gas. The conversion per pass is

dependent on the ease of temperature control and heat removal. Another

reason to limit the conversion per pass is the intolerance of FT catalysts to

high steam partial pressures. Partial pressures of steam below 5-6 bar or CO

conversions below 50-60 % are recommended in order to avoid excessive

catalyst deactivation rates [19]. Further information about the detrimental

effects of H2O on catalyst deactivation can be found in Chapter 3.


18

In slurry bubble column reactors (SBCR’s), the catalyst is suspended in

the wax product and the syngas flows upstream in the form of bubbles.

Isothermal operation is achieved by the natural motion of the catalyst-wax

mixture (i.e. the slurry) and by means of internal heat exchanger coils where

water or steam flows. The FT waxes are continuously separated from the

catalyst using sieves. The catalyst reduction is usually performed ex situ. In

order to protect the catalyst from oxidation in contact with air, the reduced

catalyst is impregnated with paraffinic waxes or FT waxes. Then, the reactor

is loaded with batches of solid waxes containing the catalyst [125, 126].

In SBCR’s the gas reactants (CO and H2) diffuse from the bubbles to the

interior of the catalyst particles, which are filled with waxes. Since cobalt

catalysts are highly active, the reaction may be mass and heat transfer limited

if insufficient attention is given to the SBCR fluid dynamics and the catalyst

properties. Mass and heat transfer limitations lead not only to a decrease in

the observed catalyst activity but also to a decrease in the selectivity to long

chain hydrocarbons [127]. For typical cobalt FT catalysts, this issue can be

overcome by using catalyst particles smaller than 100 μm [128-130]. Yet, even

smaller catalyst particles may be required if novel FT catalysts with superior

activity and/or different pore structure are used. The bubble regime and the

size of the bubbles are also parameters of great importance. The bubble regime

is a function of various operating conditions and variables (e.g. pressure,

temperature, conversion, wax composition, surface tension…). SBCR’s

operate in a heterogeneous bubble regime, known as churn-turbulent regime,

which is achieved at superficial velocities in the range of 0.1 to 0.4 m/s.

Further details about the bubble regime and the design of a SBCR can be found

in the work of Steynberg [78].

Multitubular fixed bed reactors (MTFB’s) are similar to shell and tube

heat exchangers. The syngas flows downstream inside the tubes, where the

catalyst is placed, and the cooling agent (steam or water) flows in the shell. In

order to minimize the pressure drop the tubes are filled with catalyst pellets

with a diameter of 1-3 mm. Homogeneous catalyst pellets of 1-3 mm would

result in severe mass and heat transfer limitations. Therefore, MTFB’s make

use of “egg-shell” type catalyst pellets in which the active phase (cobalt) is only

deposited in a thin layer located on the outer surface of the pellet [71, 131, 132].

Yet, careful attention has to be placed to the gas mass flux to avoid

interparticle and gas film temperature gradients. The diameter of the tubes is

limited to ca. 5 cm in order to minimize radial temperature gradients [78]. As

a result, these reactors contain a very large number of tubes. For instance,


19

Shell’s commercial MTFB’s using cobalt-based FT catalysts consist of 8000

tubes [133].

Microchannel reactors are similar to plate heat exchangers. The reactor

consists of parallel metallic plates with a separation of 0.1 to 10 mm [134]. The

catalyst is coated on these plates. The temperature is controlled by flowing

syngas and steam in alternating channels. This technology is gaining

considerable attention since it provides high yields to hydrocarbons in

relatively small reactor volumes (see Table 6). Microchannel reactors are

therefore being considered in small BTL and off-shore GTL projects [135-137].

The current capacity, dimensions, properties and catalyst

considerations of these three reactors are summarized in Table 6. Examples

of companies using SBCR’s are: Sasol, GTL F1, ENI, Nippon Oil, Syntroleum,

Exxon Mobil and ConocoPhillips. MTFB’s are used by Shell and BP.

Microchannel reactors are used by Velocys and Compact GTL [111]. At present,

only Sasol and Shell are using their technologies in commercial plants (see

Table 7). Shell used its technology in the demo-pilot Choren BTL plant with a

capacity of 330 bbl/day in Freiberg (Germany). Velocys used their

microchannel technology in a small BTL pilot plant with a capacity of 1

bbl/day in Güssing (Austria) [38].

Table 6. LTFT reactor properties and catalyst consideration (author’s own elaboration).

SBCRa MTFBa Microchannelb

Capacity (bbl/day) 14 000 6 700 125-200

Reactor height/length (m) 30 20 3.96

Reactor diameter (m) 7.8 6.2 1.37 Reactor productivity (bbl/m3-day) 10 11 21-34 CO conversion per pass (%) 55-65 30-35 65-75

Temperature control Good Poor Excellent Pressure drop Low Moderate Moderate Risk for mass and heat transfer limitations

Low High Medium

Catalyst mechanical strength requirement

High Low Low

Catalyst wax separation Difficult Easy Easy Catalyst replacement On-line Off-line Off-line Catalyst reduction Ex situ In situ In situ Catalyst regeneration Ex situ In situ In situ Poisoning Global Local Local

aAdapted from Refs. [19, 111]. bAdapted from Refs. [134, 138].


20

Table 7. Commercial GTL plants using cobalt catalysts.

Location (country)

Owner (Plant name)

FT reactor (technology provider)c

Plant capacity (bbl/day)

Startup Ref.

Bintulu (Malaysia)

Shell MTFB (Shell)

14 700 1993 [139]

Ras Laffan (Qatar)

Sasol/QPa (Oryx-GTL)

SBCR (Sasol)

34 000 2007 [140]

Ras Laffan (Qatar)

QP/Shell (Pearl-GTL)

MTFB (Shell)

140 000 2011 [139]

Escravos (Nigeria)

NNPCb/Chevron (Escravos-GTL)

SBCR (Sasol)

33 000 2014 [141]

aQP: Qatar Petroleum. bNNPC: Nigerian National Petroleum Corporation. cAdapted from Ref. [142].

2.3.3. Upgrading of Low-temperature Fischer-Tropsch products

As explained previously, the main LTFT products are middle distillates

(diesel and jet fuel cut) and waxes. If the middle distillates are to be

maximized, it is necessary to convert the FT waxes into shorter hydrocarbons.

This is usually performed in a hydrocracking and isomerization unit. In this

manner, yields to middle distillates of 60-70 % can be achieved [143].

It should be noted that the middle distillates produced via Fischer-

Tropsch synthesis have an excellent cetane number but a very low density. The

minimum density of diesel, according to the European Normative

(EN590:2009), is 820 kg/m3. The diesel fuel that can be obtained from FT has

a density of approximately 780 kg/m3 [143-146]. Unfortunately, the FT middle

distillates cannot be used for production of transportation fuels in countries

where this low density specification is applied. Nevertheless these are perfect

for blending with low quality crude oils with high density and low cetane

number. According to Klerk [147], a compliant fuel could be synthetized by

using part of the wax feed for production of aromatics (alkyl benzene) in a fluid

catalytic cracker. However, that would lead to a yield of less than 25 %.

Hydrocracking of FT waxes is slightly different from hydrocracking of

conventional crude oil. FT waxes are mainly linear paraffins and contain very

low amounts of oxygenates and aromatics. The main objective is selective

cracking to middle distillates and isomerization. Isomerization (i.e. increasing

the content of branched hydrocarbons) is performed in order to improve the

cold-flow properties of the final distillate [148].


21

FT wax hydrocracking is carried out using catalysts with a

hydrogenation-dehydrogenation function (provided by a metal) and an

isomerization-cracking function (provided by acid sites). Examples of metals

with hydrogenation-dehydrogenation function are noble metals (Pt and Pd)

[149] and metal sulfides (NiMoS, CoMoS, NiWS) [150, 151]. The

isomerization-cracking function is carried out by acidic carriers such as

zeolites, silicas, aluminas or amorphous silica-aluminas [150].

The reaction is performed in fixed bed reactors operating in trickle-flow

mode. The operating conditions are: 30-70 bar, 300-375 °C and very high

H2/wax ratios (500-8000 vol/vol) [150]. Due to this excess of hydrogen, the

rate of coke formation is very low, resulting in catalyst lifetimes of several

years [152].

2.4. Methanation

Methanation is simply the FT reaction between CO and H2 which leads

to methane. Nevertheless, the reaction between CO2 and H2 leading to

methane is also denoted methanation. The latter is actually a combination

between the former reaction and the reverse WGS reaction. The two reactions

are described as follows [153, 154]:

𝐶𝑂 + 3𝐻2 ⇄ 𝐶𝐻4 + 𝐻2𝑂 (∆𝐻° = −206 𝑘𝐽/𝑚𝑜𝑙)

𝐶𝑂2 + 4𝐻2 ⇄ 𝐶𝐻4 + 2𝐻2𝑂 (∆𝐻° = −165 𝑘𝐽/𝑚𝑜𝑙)

As discussed previously in the Fischer-Tropsch synthesis section, the

selectivity to methane and short chain hydrocarbons increases with increasing

temperature and decreasing pressure. The reader may therefore think that

high quality SNG (i.e. rich in methane) can be obtained by operating at very

high temperatures. Yet, the reaction is exothermic and involves a decrease in

the number of moles. Therefore, methane formation is thermodynamically

favored at low temperatures and high pressures.

Thermodynamic equilibrium calculations of the CO methanation

reaction have been performed by the present author in order to properly

understand how pressure and temperature influence the composition of the

equilibrated gas. Figure 4 presents the equilibrium CO conversion, selectivity

to CH4 and CO2, and final CO2 content in the dry gas for an ideal initial syngas

with a H2/CO ratio equal to 3. As can be observed, the CO conversion is

practically not limited by thermodynamics, especially if the reaction is carried


22

out at high pressures. Furthermore, the influence of the WGS reaction and so,

the formation of CO2, is significant, in particular at low pressures and high

temperatures.

Figure 4. Thermodynamic equilibrium considerations in methanation as function of temperature and pressure a)

CO conversion b) Selectivity to CH4 and CO2 and c) Outlet CO2 molar content in dry equilibrated gas. The colored area indicates the typical maximum CO2 concentration in natural gas quality specifications. The equilibrium curves have been calculated using the software Aspen HYSYS v7.1 using an initial gas composition of 75 % H2 and 25 % CO.

The amount of CO2 in the resulting SNG is a parameter of importance if

the gas is intended to be commercialized. The natural gas quality

specifications are different in different countries. In European countries, the

maximum CO2 molar content in natural gas varies between 1 and 3 % [155].

This issue is nevertheless solved either by removing CO2 from the final SNG

[154, 156] or by controlling and adjusting different process parameters such

as feed gas composition, temperature and pressure [157]. A very attractive

aspect of methanation is that SNG is completely replaceable with natural gas

and can be used in existing infrastructures [154].

2.4.1. Nickel-based methanation catalysts

Nickel-based catalysts are the most employed materials due to their

high activity, selectivity to methane and relatively low price [158, 159].

Nevertheless, MoS2-based catalysts are recently gaining attention due to their

resistance to sulfur and their intrinsic WGS activity, which allows for the use

of H2-poor syngas feeds [160]. However, the activity of MoS2 catalysts is low,


23

a characteristic that must be improved if they are to replace Ni catalysts in the

future [161].

The major concern in the use of nickel-based catalysts is their stability.

These catalysts are threatened by different deactivation phenomena at low and

high temperatures. At high temperatures and steam partial pressures there is

a risk for thermal sintering [162]. At low temperatures and high CO partial

pressures, there is a high risk for nickel carbonyl formation [163] and carbon

formation [164]. These deactivation causes are discussed in detail in Chapter

3.

Nickel catalysts have been used in the form of Raney Ni [165-167] and

as supported nickel catalysts. The most common catalyst is Ni/Al2O3 due to its

mechanical strength, thermal stability and high yield per unit cost [168-172].

Nevertheless, other supports have shown interesting properties. For instance,

Ni/TiO2 exhibits very high CO consumption rates and can apparently reduce

the rate of nickel carbonyl formation [173, 174]. Ni/SiC catalysts have also

been considered due to their high thermal resistance [175].

Different promoters can be used to enhance the stability of Ni/Al2O3

catalysts. Addition of ZrO2 to the support enhances the resistance towards

carbon formation and sintering [176, 177]. Promotion with small amounts of

MgO has also shown an improved carbon formation resistance [178-180].

Another important aspect of methanation with nickel catalysts is the

reaction’s structure sensitivity. Early surface science studies were suggesting

that the turnover frequency (TOF) was independent of the type of active site

[181]. In consequence, it was believed that the reaction was structure

insensitive. More recently, Andersson et al. [182] found evidence of the

methanation reaction actually being highly structure sensitive. In this work

they proved with experiments that the TOF increases with decreasing nickel

particle size. This observation was explained by means of density functional

theory (DFT) calculations which showed that the energy barrier for both direct

and indirect CO dissociation is lower on less coordinated sites (e.g. kinks and

steps) than on terraces. This study also showed by means of DFT calculations

that the H-assisted CO dissociation (indirect CO dissociation) path is more

favorable under methanation conditions.


24

2.4.2. Methanation reactors

The design of methanation reactors presents an important advantage

compared to FT reactors: the reaction does not need to be carried out

isothermally. Nevertheless, heat removal is still a major consideration in the

overall methanation process design since, as discussed previously, the

methane yield is thermodynamically maximized at low temperatures [183].

Moreover, extremely high temperatures can severely deactivate the catalyst.

Two main technologies have been proven suitable for methanation of

synthesis gas [154, 184]: fluidized bed reactors and series of adiabatic fixed

bed reactors with intercooling. The latter has already been commercialized by

Haldor Topsoe, Davy Technologies (Johnson Matthey) and Lurgi. Haldor

Topsoe’s process is known as TREMP (Topsoe Recycle Energy Efficient

Methanation Process). Johnson Matthey’s process is known as HICOM (High

Combined Shift Methanation). Other less developed technologies are:

internally cooled fixed bed reactors [183, 185], microchannel reactors [186,

187] and liquid phase methanation [188, 189].

Fluidized bed reactors aim to convert syngas into methane using a

single reactor [190-192]. This type of reactors is very suitable for that purpose

since the fluidization of solids facilitates the operation at isothermal

conditions. Moreover, fluidization allows for continuous removal/

replacement of the catalyst and on-line regeneration. Nevertheless, the

catalyst requires a high attrition resistance [193, 194]. This technology has

been demonstrated at industrial scale (2000 Nm3/h SNG, up to 20 MW) [154].

Adiabatic fixed bed methanation (also called high temperature

methanation [162]) is an attractive technology due to its simple reactor design

and the possibility of obtaining high quality superheated steam by cooling

down the gas effluents leaving the methanation reactors at high temperatures

[195]. A high methane yield can be obtained by means of several reactors in

series with intercooling in order to move the equilibrium towards further

methane production. A recycle is also needed in the first adiabatic reactor in

order to limit the adiabatic temperature increase, minimize catalyst

deactivation and reduce the number of reactors in series.

A process diagram of a high-temperature methanation process is

presented in Figure 5 for a better understanding of how the adiabatic fixed bed

concept can be used for production of methane-rich SNG. It consists of 4


25

methanation reactors (steps) as in the TREMP process [157]. The methane

concentration in the gas after each methanation step is presented in Figure 6.

The temperature span in the first adiabatic fixed bed reactor is dictated

by the catalyst limitations [196]. The minimum inlet temperature depends on

the catalyst resistance to carbon formation and nickel carbonyl formation

[197]. The maximum exit temperature depends on the catalyst resistance to

hydrothermal sintering [162]. A recycle of the exit gas (using either a

compressor or an ejector [198]) is required in order to operate inside these

temperature limits. As can be observed in Figure 6, the absence of recycling

could result in temperatures as high as 900 °C. Catalysts with high resistance

to the above-mentioned deactivation causes allow for a larger temperature

span and thus result in lower capital and operating costs associated with gas

recycling.

Early high temperature methanation processes were operated using a

rather narrow temperature span. In the old Lurgi process, the inlet and outlet

temperatures in the first reactor were approximately 300 °C and 450 °C [199].

Nowadays, the TREMP process can operate between 250 °C and 700 °C due to

catalyst improvements [184, 196, 197]. The catalyst employed by Haldor

Topsoe (named MCR-2X) consists of nickel supported on a pre-stabilized

alumina carrier [200]. A dual bed concept comprising a non-nickel catalyst

prior to the MCR-2X catalyst is used in order to operate at such low inlet

temperatures. Recently, the nature of this non-nickel catalyst has been

published on Haldor Topsoe’s website; its formulation is “Cu/Zn/Al” and it is

named GCC-2 [201].

The catalysts in the subsequent reactors are not subjected to such

hostile conditions. The CO concentration in the inlet gas is very low and the

outlet temperatures are lower as well. These reactors are mainly converting

the CO2 produced in the first reactor into methane. In order to produce a

methane-rich SNG it is suitable to separate the product water before the last

reactor in order to push the equilibrium toward further methane conversion

(see the change in the equilibrium curve in Figure 6). The actual SNG

composition resulting from the TREMP process contains ca. 98 % methane

[157].


26

Figure 5. Schematic diagram of a high temperature methanation process (adapted from Refs. [157, 198]). The

temperatures differ from the actual TREMP process. The outlet temperatures have been calculated using the software Aspen HYSYS v7.1 using an initial syngas composition consisting of 75 % H2 and 25 % CO, 20 bar and using the SRK equation of state.

Figure 6. Reaction equilibrium diagram of the high temperature methanation process presented in Figure 5

(adapted from Refs. [157, 185]). All the equilibrium curves and data have been calculated using the software

Aspen HYSYS v7.1.


27

A list of current coal-to-SNG commercial plants in operation is presented in Table 8. All the plants make use of adiabatic fixed bed methanators in series with intercooling and recycling. The largest single-train methanation technology has a capacity of 1.4 billion Nm3/year and is installed in the Qinghua plant [202]. Several coal-to-SNG commercial projects are under construction and will probably be operating in the coming years [184]. China, in particular, has approved several projects which will result in a total annual production of 37 billion Nm3 SNG, a fact that has awakened environmental concerns and criticism [203]. Finally, it should be mentioned that Göteborg Energi AB started up the first demonstration biomass-to-SNG plant in 2014 [204]. The project, known as GoBiGas, makes use of the TREMP technology and has a capacity of 20 MW SNG.

Table 8. Commercial coal-to-SNG plants in operation.

Location (country)

Owner (Plant/project name)

Methanation technologya

Plant capacity (billion Nm3 SNG/year)

Start-up

Ref.

Beulah, ND (USA)

BEPCb

(Great Plains Synfuels Plant)

Lurgi 1.5 1984 [205]

Yining (China)

Qinghua Group (Qinghua)

TREMP 5.5 2013 [202]

Chifeng (China)

Datang (Keqi) HICOM 4 2013c [206]

Ordos (China)

Huineng Coal Electricity Group (Huineng)

TREMP 1.6 2014 [207]

a Adapted from Ref. [184]. b BEPC: Basin Electric Power Cooperative. The resulting CO2 from the plant is sent to Saskatchewan

in Canada for capture and storage. c Stopped in 2014 due to an industrial accident and technical difficulties.

Finally, it may be noted that methanation has also found an application

in the conversion of coke-oven gas (gas resulting from the production of

metallurgical coke) into SNG [208-210]. The TREMP process is already

operating on coke-oven gas in two plants in China since 2013 [211, 212]. One

plant is located in Wuhai (client: Petrochina) and produces 900 million Nm3

SNG/year. The second plant is located in Shandong (client: China National

Offshore Oil Corporation) and produces 160 million Nm3 SNG/year.


28

2.5. Co-production of liquid fuels and SNG via Fischer-Tropsch

synthesis and methanation

A potential opportunity for increasing the technical and economic

feasibility of a BTL plant is the combination of the Fischer-Tropsch synthesis

and methanation. The idea of integrating these two processes was firstly

reported by Zwart and Boerrigter [213] when they recognized that the typical

composition of the FT off-gas (containing C1 to C4 hydrocarbons) resembled

that of the Dutch Groningen natural gas. A simple block diagram of this co-

production concept is presented in Figure 7.

Figure 7. Block diagram of biomass to SNG and FT fuels co-production concept (adapted from Ref. [213]).

This co-production concept presents some technical advantages in

comparison to the production of only FT liquid fuels. For instance, the FT

reactor no longer needs gas recycling. The unconverted syngas is sent to

methanation. Moreover, methanation can be operated at much higher

temperatures than FT, a fact that allows for the production of superheated

steam by heat exchanging, and thus increasing the energy efficiency of the

process.

Chapter 3 – Deactivation of cobalt- and nickel- based catalysts

29

Chapter 3

Deactivation of cobalt- and nickel-based catalysts

This chapter aims to summarize the main causes of deactivation of

cobalt- and nickel-based catalysts during Fischer-Tropsch synthesis (FTS) and

methanation. It is worth mentioning that these two types of catalyst suffer

from very similar deactivation phenomena. For that reason, this chapter is

divided into sections that describe different deactivation causes valid for both

types of catalyst.

3.1. Carbonyl-induced metal particle sintering

Carbonyl-induced metal particle sintering proceeds via the so-called

Ostwald ripening mechanism. In other words, sintering occurs via migration

of small metal clusters or monoatomic species from small to large crystallites.

The carbonyl-induced mechanism consists of three steps [163, 214]: formation

of a volatile metal carbonyl (i.e. Ni(CO)4 and Co2(CO)8) or mobile sub-

carbonyl species (i.e. Ni(CO)x and Co(CO)x, x<4) [215-217], diffusion of these

species inside the catalyst pores and decomposition over another metal

particle.

3.1.1. Nickel carbonyl-induced sintering

Nickel catalysts are very sensitive to this type of deactivation in

comparison with cobalt catalysts. This is somehow expected since the vapor

pressure of Ni(CO)4 is higher than that of Co2(CO)8 [218]. Carbonyl-induced

sintering can lead to an excessive growth of the nickel crystals within days and

thus, to a significant loss of metallic surface area [197, 217, 219]. At very

extreme conditions, Ni(CO)4 can also diffuse outside the catalyst pellet, move

across the catalyst bed and eventually, leave with the gas products [163].

The formation of volatile metal carbonyls is thermodynamically favored

at low temperatures and high CO partial pressures [220, 221]. It is generally

assumed that, under typical methanation conditions, the rates of formation


30

and decomposition of Ni(CO)4 are fast enough [222, 223] so that chemical

equilibrium with respect to Ni(CO)4 formation is always achieved at the nickel

surface [163].

The loss of methanation activity of nickel alumina catalysts caused by

carbonyl formation has been investigated in several studies [217, 219, 224].

Among these, the work performed by Shen et al. [163] in which they developed

a criterion for identifying “safe” and “unsafe” operating conditions (i.e.

conditions at which this sintering mechanism occurs) is worth presenting.

According to that work, stable catalytic activity can be achieved if the

equilibrium partial pressure of Ni(CO)4 is kept below 1 μPa. A diagram was

also developed in that work for an easier identification of safe and unsafe

operating conditions. That diagram was adapted for the present thesis and is

presented in Figure 8.

0 20 40 60 80 100 120

200

250

300

350

400

450

Unsafe operating

region

Te

mp

era

ture

(°C

)

CO partial pressure (kPa)

pNi(CO)4,eq

= 5 Pa

pNi(CO)4,eq

= 2.5 Pa

pNi(CO)4,eq

= 1 Pa

pNi(CO)4,eq

= 0.5 Pa

Safe operating

region

Figure 8. Diagram of the safe operating regions for alumina-supported nickel catalysts (adapted from Ref. [163]).

The equilibrium curves for different equilibrium partial pressures of Ni(CO)4 (pNi(CO)4,eq) were calculated using

the thermodynamic equilibrium data available in Ref. [225].

This type of deactivation can be minimized not only by exposing the

catalyst to a suitable temperature and CO partial pressure but also by

modifying the catalyst composition and properties. For instance, Vannice and


31

Garten [173] found that the concentration of nickel carbonyls in the gas phase

was orders of magnitude lower on a TiO2-supported catalyst than on a SiO2-

supported catalyst. This observation was attributed to a strong metal-support

interaction (SMSI) between nickel and titania.

Recently, Munnik et al. [224] studied how the spatial distribution of

nickel particles and their size affected the stability of the catalyst. It was found

that the distance between nickel crystals had a minor effect on the deactivation

rate. Therefore, they concluded that the sintering rate is not controlled by

diffusion of Ni(CO)4 from one crystal to another. In addition, it was found that

catalysts with small nickel particles deactivate to a larger extent than catalysts

with initially larger particles. This observation was explained by a higher

nickel carbonyl supersaturation which is sufficient to break the catalyst pore

walls, widen the pore diameter, and thus, allow for a further growth of the

crystals.

3.1.2. Cobalt carbonyl-induced sintering

Deactivation of cobalt-based FT catalysts via carbonyl-induced Ostwald

ripening has been, perhaps, underestimated. Indeed, this type of sintering has

not been considered in the more recent reviews dealing with deactivation of

cobalt-based FT catalysts [111, 226, 227]. It was only last year, in the work of

Kistamurthy et al. [228], that evidence of this type of deactivation has been

found. Their arguments were based on the following observations. Firstly,

sintering was only observed in the presence of CO. Secondly, they could

observe how small particles were reduced in size with time on stream while

large neighboring particles increased in size. Finally, cobalt sub-carbonyl

species have been previously considered to be mobile under FT conditions

[216].

This argument is in line with the work of Claeys et al. [229] in which

they studied sintering of cobalt particles by means of an in situ magnetometer.

No sintering was observed in water-hydrogen atmospheres at typical FT

temperatures. Significant changes in crystallite size were only observed in the

presence of CO.

3.2. Hydrothermal sintering

As its name indicates, hydrothermal sintering is caused by high

temperatures and high steam partial pressures. Sintering can occur due to two

mechanisms [230, 231]: Ostwald ripening, and particle migration and


32

coalescence. The latter, in contrast with Ostwald ripening, proceeds via

collision and coalescence of entire crystallites. The thermal stability of a

material is usually related to its solid-state diffusion coefficient and thus, to its

melting point. Two temperatures are used to indicate the occurrence of

thermal sintering: the Hüttig and the Tammann temperatures. Moulijn et al.

[230] recommend the following semi-empirical relations to calculate these

temperatures:

𝑇𝐻ü𝑡𝑡𝑖𝑔(𝐾) = 0.3 ∙ 𝑇𝑚𝑒𝑙𝑡𝑖𝑛𝑔 (𝐾) (3.1)

𝑇𝑇𝑎𝑚𝑚𝑎𝑛𝑛(𝐾) = 0.5 ∙ 𝑇𝑚𝑒𝑙𝑡𝑖𝑛𝑔 (𝐾) (3.2)

The Hüttig temperature represents the limit at which atoms at

crystallite defects become mobile. The Tammann temperature represents the

limit at which bulk atoms within the crystal exhibit mobility. Table 9 presents

the melting point, the Tammann and the Hüttig temperatures of metals and

compounds relevant for this thesis.

Table 9. Melting point, Tammann temperature and Hüttig temperature of metals and compounds relevant for the

present work (adapted from Moulijn et al. [230]).

Tmelting (°C) TTammann (°C) THüttig (°C)

Co 1480 604 253 Ni 1452 590 245 Pt 1755 741 335 Cu 1083 405 134 Al2O3 2045 886 422 SiO2 1713 720 323 TiO2 1857a 792 366

a Obtained from Ref. [232].

Sintering, however, can occur at lower temperatures than those presented in Table 9. Specially, for small particles and in the presence of steam [171, 231]. The morphology of the material also plays an important role. For instance, the thermal resistance of α-Al2O3 is higher than that of γ-Al2O3 [230, 233]. Another possible explanation for the occurrence of sintering at lower temperatures is that the melting point of a compound is not always well determined [230].


33

3.2.1. Hydrothermal sintering during methanation

The hydrothermal resistance of a nickel-based methanation catalyst is

of paramount importance. Nickel particle growth leads to a reduction of the

number of steps and kinks which are the most active sites for CO methanation

[162]. As a result, sintering leads not only to a loss of metallic surface area but

also to a reduction of the catalyst intrinsic activity (TOF). The thermal stability

of the catalyst represents, thereby, a constraint in the design of adiabatic fixed

bed reactors. Catalysts with higher stability allow for operation at higher

temperatures and thus, reduce costs associated with gas recycling [196].

Sintering of nickel catalysts has been extensively studied. Mathematical

models describing the crystallite growth as function of temperature, steam

content and sintering mechanism are available in the literature [231].

According to Rostrup-Nielsen [162], sintering under high-temperature

methanation conditions (600 °C, 30 bar) proceeds via atom migration.

Support sintering and pore collapse can also occur at these conditions [197].

Catalyst supports are therefore pre-stabilized in order to overcome this issue

[196, 200].

3.2.2. Hydrothermal sintering during FTS

Hydrothermal sintering of cobalt crystallites during FTS is considered by many researchers as one of the main causes of deactivation [234-236]. This may be surprising considering the low temperatures employed in this reaction (200-240 °C). The Hüttig temperature for cobalt is 253 °C. Nevertheless, the mobility of the crystals is affected by their size and their nature. Therefore sufficiently small particles could undergo sintering, especially at high CO conversions, conditions at which the steam content in the gas is significant.

Many scientific studies have provided convincing results which support the occurrence of this deactivation mechanism [237-240]. Bian et al. [238], tested cobalt-silica catalysts containing crystallites in the range of 6-10 nm at 10 bar, 200-240 °C for 60 h on stream. They found that increasing the

temperature up to 240 °C accelerated sintering of the cobalt crystallites. In addition, Shi et al. [240], found that the growth of cobalt crystallites (as determined by ex situ X-ray diffraction) increased with addition of water to the syngas feed. A deeper insight into the sintering of cobalt crystallites was offered in the work performed at ExxonMobil [237]. They analyzed catalyst samples by means of different ex situ techniques. Transmission electron microscopy (TEM) images of the spent samples suggested that crystallite migration is probably the predominant sintering mechanism under FTS.


34

Support sintering under FTS conditions is rarely observed. However it has been suggested as a possible deactivation mechanism when using high surface area silica supports. Huber et al. [241] found that steam treatments of cobalt-silica catalysts at 220 °C lead to a significant decrease in surface area as well as the formation of cobalt-silicate inactive compounds.

Sintering of cobalt crystallites is, to some extent, reversible [227]. Re-dispersion of cobalt can be achieved by an oxidation-reduction sequence at certain conditions. During oxidation, cobalt forms hollow particles by means of the so-called Kirkendall effect [242]. The inner void has the size of the original reduced crystallites. Then, the reduction results in the breaking of these hollow structure into smaller particles and thus in a re-dispersion.

3.3. Re-oxidation

Another concern in these two CO hydrogenation reactions is the risk for

re-oxidation of the metal active sites. The risk exists because steam, the major

reaction product, is an oxidizing agent. Operation at high CO conversion leads

to high H2O/H2 ratios and thus to harmful atmospheres. Furthermore, the

potential for oxidation is higher for nanoparticles than for bulk cobalt or

nickel.

3.3.1. Re-oxidation of cobalt nanoparticles

Re-oxidation of cobalt nanoparticles during FTS is quite a debated

topic. Van Steen et al. [243] performed thermodynamic calculations in order

to evaluate if re-oxidation was feasible under typical FTS conditions. The

study showed that only cobalt particles with a diameter smaller than 4.4 nm

may oxidize under relevant conditions (220 °C and H2O/H2<1.5). A

reproduction of the phase diagram developed in that work is presented in

Figure 9.


35

2 3 4 5 6 7 8

0

2

4

6

8

10

Co(II)O

pH

2O/p

H2 (

ba

r/ba

r)

dCo

(nm)

Co (fcc)

Figure 9. Stability region of spherical β-Co (fcc) and Co(II)O crystals in H2O/H2 atmospheres at 220 °C as

function of the diameter of a spherical metal Co crystallite “dCo” (adapted from Ref. [243]). The dotted lines

represent the calculated limits for an error in the value of the surface energy of β-Co of ±15%. The equilibrium

curves were calculated using the mathematical expressions presented in Ref. [243]. The chemical potentials of

the different components were obtained from the Aspen Plus V8.6 database.

Some researchers have still pointed out the occurrence of surface re-

oxidation [244-246]. For instance, Hilmen et al. [244] exposed cobalt-

alumina catalysts to 250 °C, 10 bar and a H2O/H2=1.5 atmosphere

(corresponding to approximately 75 % CO conversion), and found indications

of re-oxidation and a significant loss of cobalt surface area. Other studies have

observed contradictory results. Saib et al. [247] studied the change in the

degree of reduction of a cobalt-alumina catalyst (promoted with platinum)

with an average particle size of 6 nm in a 100 bbl/day slurry reactor during

operation at 230 °C, 20 bar and CO conversions up to 70 %. For that purpose,

catalysts samples were unloaded at different times and analyzed. It was found

that the degree of reduction of the catalyst increased from 53 to 89 % over a

period of 140 days. In other words, the catalyst underwent further reduction

rather than re-oxidation.

3.3.2. Re-oxidation of nickel nanoparticles

Re-oxidation of nickel is apparently not an important deactivation

cause in methanation [196]. Metallic nickel is stable at quite high

temperatures and H2O/H2 ratios. According to Rostrup-Nielsen [248], re-

oxidation takes place at H2O/H2 ratios in the range of 10 to 50 depending on

the nickel crystal size and the support interaction.


36

3.4. Formation of inactive metal-support compounds

Solid-state reactions between the metal and the support have also been

suggested as a potential deactivation mechanism. Strongly interacting carriers

such as alumina and silica are prone to react with the metal sites in the

presence of steam and form inactive spinel phases. A particularly studied case

is the reaction between cobalt and alumina which is described as follows:

𝐶𝑜 + 𝐴𝑙2𝑂3 + 𝐻2𝑂 ⇄ 𝐶𝑜𝐴𝑙2𝑂4 + 𝐻2

Alumina-supported cobalt catalysts have been widely studied in FT due

to their suitability in continuously stirred tank reactors and slurry bubble

column reactors. Thermodynamic calculations of this catalytic system have

shown that bulk oxidation of cobalt to a CoAl2O4-spinel phase is spontaneous

under relevant FT reaction conditions [249, 250]. However, it is claimed that

this reaction is kinetically restricted [249].

Yet, some studies have found results indicative of this type of

deactivation. For instance, Holmen et al. [244] investigated the formation of

these species by exposing reduced catalysts to a H2O/H2 treatment.

Subsequent temperature-programmed reduction analyses revealed the

formation of hardly reducible cobalt species (reducing at around 800 °C)

typical of cobalt-aluminate compounds. The results are in agreement with the

work of Goodwin and coworkers [251] who performed a similar study on

Co/SiO2 catalysts.

A possible explanation is that the formed species are not strictly

identical to CoAl2O4. Some researchers have suggested that the surface

compound is probably deficient in Co [252]. Others propose that steam

significantly accelerates the rate of metal-aluminate formation [253].

The formation of nickel-support compounds during methanation has,

unfortunately, been less investigated. Bartholomew et al. [171], investigated

the loss of activity of Ni/Al2O3 catalysts at high temperatures and in the

presence of steam (only 3% H2O in H2). The study concluded that the major

deactivation causes were sintering of the carrier and the nickel particles. No

indications of nickel aluminate formation were found.


37

3.5. Carbon formation

The formation of non-reactive carbon species is generally recognized as

one of the major deactivation mechanisms in FT and methanation [227].

These carbon species progressively cover the metal particles leading to a loss

of active surface area. Excessive formation of carbon can also obstruct the

carrier pores and restrict the access of reactants to metal particles.

There are different routes which can potentially lead to the formation of

carbonaceous inactive species. These routes are summarized in Figure 10. One

of the most considered routes is that initiating with the dissociation of CO into

adsorbed C and O atoms. Adsorbed C, usually called atomic carbon or Cα, can

then be hydrogenated and transformed into hydrocarbons. Nevertheless,

atomic carbon can also react with the metallic surface and form bulk carbides.

Alternatively, atomic carbon can polymerize and form amorphous species or

graphite. It may be mentioned that species such as polymeric carbon (Cβ)

could be partially hydrogenated especially if the H-assisted CO dissociation

path is dominant.

Figure 10. Possible carbon formation routes on cobalt- and nickel-based catalysts during Fischer-Tropsch

synthesis and methanation (adapted from [214, 226]).

Another particular carbon species typically formed when using nickel

catalysts is whisker carbon. This type of carbon diffuses through the nickel

crystal and grows under it. Its growth, in the form of fibers, changes the shape

of the crystal and displaces it. Moreover, the catalyst particles are destroyed


38

when the whisker impacts the pore walls, resulting in an increased pressure

drop across the reactor [254, 255].

Carbon can also form by means of the Boudouard reaction or via

dehydrogenation of hydrocarbons produced in the reaction. The latter is

usually called “coke” rather than carbon [19].

3.5.1. Carbon formation during FTS

One of the challenges in studying carbon formation on spent cobalt FT

catalysts is the fact that the surface is covered by the product waxes.

Nevertheless, characterization methods such as X-ray diffraction can be used

to detect crystalline phases, such as Co2C. Several researchers [234, 235, 256]

have identified this phase on spent samples after relevant FTS conditions.

A common technique for studying other carbon species present on spent

samples is temperature-programmed hydrogenation (TPH) and temperature-

programmed oxidation (TPO). The method consists in flowing H2 or O2

through the spent samples while increasing the temperature and monitoring

the amount of CH4 or CO2 (resulting from the reaction with carbon). For that

purpose, the wax needs to be previously extracted with an organic solvent (e.g.

tetrahydrofuran).

Among various studies, the work of Moodley et al. [257], in which they

identified three main carbon species by means of TPH of wax extracted

catalysts is worth presenting. The TPH profile from that work is presented in

Figure 11. As can be observed, the profile reveals the existence of three peaks.

The first peak was assigned to surface carbide species and residual

hydrocarbons. The second peak was ascribed to residual wax contained in

small pores. The last peak was assigned to polymeric carbon. Similar analyses

were performed on samples removed from the reactor at different times on

stream during a period of 6 months. The results indicated that polymeric

amorphous carbon was the only species accumulating during the run. Finally,

it was observed by means of transmission electron microscopy that the carbon

species were located both on the cobalt surface and on the alumina carrier. It

is believed therefore that carbon migrates from the cobalt surface to the

support.


39

Figure 11. Peak deconvolution of a methane profile for TPH of a wax-extracted Co/Pt/Al2O3 catalyst from a FTS

run in a slurry bubble column (reproduced from [257], with permission from Elsevier).

Different solutions have been proposed to improve the resistance of FT

catalysts to carbon formation. Addition of Ru is considered to retard the

formation of carbon as well as improve the activity and selectivity [57, 122,

258, 259]. Addition of potassium increases the resistance towards graphite

formation [260]. Boron promotion has been found to reduce the deposition of

carbon without affecting the catalyst activity and selectivity [261]. Hybrid

catalysts possessing cracking ability (e.g. cobalt supported on zeolites) have

also been proposed to minimize the accumulation of heavy hydrocarbons

[262, 263].

Finally, it should be mentioned that the build-up of very hard product

waxes (hydrocarbons with more than 100 carbon atoms) could also contribute

to the loss of activity [226]. Accumulation of hard waxes can reduce the

diffusion rate of reactants inside the catalyst and, consequently, slow down the

reaction rate. In this sense, the loss of activity is not strictly assigned to

deactivation but to an increasing influence of mass transfer on the reaction

rate.

3.5.2. Carbon formation during methanation

The formation of carbon species on spent nickel catalysts can be easily

studied by means of temperature-programmed techniques. The first

identification of carbon species formed during methanation was developed by

McCarty and Wise [264]. A summary of the identified species in different

studies and their temperatures of formation and hydrogenation is presented

in Table 10.


40

Table 10. Carbon species formed on nickel catalysts under methanation conditions (adapted from [265]).

Carbon species Designation Temperature of formation (°C)

Peak temperature for reaction with H2 (°C)

Adsorbed, atomic Cα 200-400 200 Polymeric, amorphous

Cβ 250-500 400

Whisker carbon Cv 300-1000 400-600

Nickel carbide Cϒ 150-250 275

Graphitic (crystalline)

Cc 500-550 550-850

The formation of whisker carbon during CO methanation has been

observed in very few studies [266, 267]. However, according to Rostrup-

Nielsen [164], carbon formation is not a problem in an adiabatic methanation

reactor at typical industrial conditions, independently of the recycle ratio and

the outlet temperature.

A useful tool for predicting the formation of whisker carbon is the so-

called principle of equilibrated gas [268]. By means of this principle, carbon

can be predicted above a certain temperature depending on the inlet gas

composition, temperature and recycle ratio [197]. The whisker carbon limits

for typical methanation conditions are presented in Figure 12. Carbon

formation is predicted at a temperature above ca. 600 °C if ideal graphite data

is used for the thermodynamic calculations [164]. This implies a minimum

recycle ratio of about 2.5 for carbon-free operation. Nevertheless, the carbon-

limits are moved to significantly higher temperatures if these are corrected

using whisker carbon equilibrium data and its dependence on nickel crystallite

size (see Figure 12 b)) [268]. The validity of this principle has been confirmed

by experience in pilot and commercial scale reactors. Methanation reactors

have been operating successfully at temperatures up to ca. 700 °C [196, 197].

The formation of bulk nickel carbide (Ni3C) during methanation, most

probably does not occur. Mirodatos et al. [269] showed that the formation of

bulk carbide occurs in the absence of H2. Nevertheless, if the catalyst is

exposed to H2/CO=3, only surface carbide (Cα) is formed. It is generally

believed that Cα and Cβ are the carbon species governing the nickel surface

during low-temperature CO methanation. Nevertheless, the presence of Cα on

spent catalysts cannot be related to deactivation since it is also the reaction

intermediate for methane formation.


41

Figure 12. a) Recycle ratio and outlet temperature of adiabatic equilibration of gas (adapted from [197]) (Data

calculated with the software Aspen HYSYS v7.1 using the inlet temperature and gas composition specified in [197]). b) Whisker carbon limits and Ni crystallite size in methanation (adapted from [164]). The whisker carbon limits have been calculated using the principle of equilibrated gas and using the equilibrium data presented in [268].

There are several factors among the catalyst properties which can affect the formation of polymeric carbon. The type of carrier and the nickel particle size apparently have an important effect [219, 267, 270]. Some studies [176, 177] have also shown that the addition of ZrO2 as a structural promoter in Ni/Al2O3 catalysts reduces the carbon formation rate.

A very interesting solution for diminishing the formation of carbon is the addition of CO2 to the feed. Mirodatos et al. [219], compared the deactivation of Ni/SiO2 catalyst in the absence and in the presence of CO2 in the feed gas (ca. 3% vol. CO2). They showed that the addition of CO2 leads to a significant decrease of the deactivation rate which could be assigned, by means of TPH analyses, to a lower amount of carbon formed. In other words, CO2 co-feeding enhances the carbon gasification rate via the reverse Boudouard reaction (CO2 + C → 2CO) rather than the carbon formation rate. This observation is in line with the work of Pedersen et al. [197] in which they showed that no deactivation occurs at low temperatures if CO is totally replaced by CO2.


42

As explained previously, Haldor Topsoe’s TREMP methanation technology uses a Cu-based catalyst (named GCC-2) prior to the nickel-methanation catalyst (named MCR-2X) to protect it from risks such as carbon formation [201]. This catalyst also serves as guard bed to capture sulfur impurities in the gas phase. A possible aspect which should be considered when applying this dual bed concept is the poor thermal resistance of Cu (as can be deduced from Table 9). Cu-based WGS catalysts are usually reduced at temperatures not higher than 230 °C in order to avoid excessive sintering [271]. Unfortunately, very little information is published regarding the use of this GCC-2 catalyst and the temperatures to which it is exposed.

3.6. Poisoning by syngas impurities

Poisoning is a general term which can be used to describe many types

of deactivation but, in this context, poisoning is defined as the blockage of

metal active sites by syngas impurities. Among the different impurities, sulfur

compounds (H2S, COS and organosulfur compounds) are the most studied

due to their harmful effects and their presence in coal, natural gas and

biomass-derived syngas. Sulfur compounds are known for their irreversible

chemisorption on metal sites and their possible influence on neighboring

unsulfided atoms.

3.6.1. Poisoning of cobalt FT catalysts

Sulfur poisoning of cobalt-based catalysts has been studied by many

researchers [272-276]. All studies observe a decrease in activity with

increasing sulfur content in the catalyst at any concentration. On the other

hand, there is still no agreement as concerns the effect of sulfur on the FT

product selectivity [226, 272, 277].

Regeneration of sulfur-poisoned cobalt catalysts is practically

impossible [214]. Sulfur could be removed by means of steam or H2 treatments

at very high temperatures [278] but it would dramatically sinter the catalyst.

Therefore, the syngas feed is desulfurized to levels below 20 μg/Nm3. This is

usually achieved by placing a ZnO guard bed prior to the FT reactor.

Poisoning by N compounds is less severe due to its reversible

adsorption. The catalyst can be easily regenerated by periodic in situ H2

treatments. In any case, N compounds are removed to levels below 10 μg/Nm3

prior to the FT reactor in order to avoid excessive periodic regenerations. This


43

is achieved by hydrolysis of HCN to NH3, followed by a water wash and a final

guard bed containing an acidic solid adsorbent [49].

Unfortunately, due to the lack of BTL industrial experience, there are

very few studies dealing with poisoning by typical biomass-derived impurities

such as tars and inorganic ashes (containing alkali, Mn, Fe, P or Cl

compounds). An intensive study of most of these impurities is summarized in

the work of Borg et al. [277]. In that work, they observed that Na, K and Ca

had the most adverse effects on catalyst activity. In addition, they found that

Cl did not have any detrimental effect but actually a positive effect on both

activity and selectivity to long-chain hydrocarbons. This observation is

somewhat surprising since others have recommended to limit HCl, HBr and

HF compounds in syngas to levels below 10 μg/Nm3 [279].

3.6.2. Poisoning of nickel methanation catalysts

Excellent studies can be found regarding the effect of sulfur on the CO methanation activity of nickel catalysts [280-286]. Just as for cobalt FT catalysts, the activity of nickel catalysts decreases with increasing sulfur coverage of the metal surface. Sulfur concentration levels as low as 10 μg/Nm3 in syngas are recommended to guarantee a decent catalyst life [287]. It is also known that the decline in activity is not linearly correlated with the decline of available nickel area. The decline in activity is more severe [281]. Moreover, it is known that the activation energy does not change with increasing sulfur coverage. This indicates that the methanation reaction requires a large ensemble of nickel atoms, which hardly can exist if the metal surface is partially blocked by sulfur [281, 287].

In situ H2S poisoning studies indicate that the normalized methanation activity is proportional to the square of the unsulfided nickel surface fraction [281]. However, studies with presulfided catalysts indicate a more severe deactivation [282]. This contradiction between in situ and ex situ poisoning studies is in line with the recent work of Khodakov et al. [288] in which they studied the effect of S poisoning before and during methanation. In that work, they concluded that, during methanation, CO dominates on the most active methanation sites (steps and kinks [182]) whereas S chemisorbs on terraces. However, in the absence of CO, S chemisorbs on steps and kinks as well, which leads to a more dramatic decrease in activity. This observation is also supported by the fact that the TOF, based on the unsulfided metallic surface area, remains constant.

Regeneration of sulfur-poisoned nickel catalysts is possible at high temperatures and in H2O/H2 atmospheres. This solution may be applied in


44

the case of catalysts used for adiabatic fixed bed methanation which are designed to resist higher temperatures. Regeneration with steam is more effective than regeneration with pure H2 [287]. A degree of regeneration (i.e. percentage of sulfur removed) of about 40 % can be achieved after exposure to an atmosphere of H2O/H2=2~3, 700 °C for a period of 24 h [289]. A

regeneration degree of ca. 80 % can be achieved in pure steam at 600 °C in only 3 hours [289]. Total sulfur removal can be achieved at higher temperatures. These steam treatments imply, though, a total oxidation of the catalyst [287].

3.7. Other deactivation causes

This section aims to give a short overview of other deactivation causes

which are, to the best of the author’s knowledge, studied less extensively or

not so well documented. Yet, it does not imply that these are less important.

3.7.1. Attrition and crushing

Just as in other catalytic applications, catalysts in fluidized bed reactors

and slurry bubble column reactors require high attrition resistance. This

property is vital in slurry bubble columns since the product wax cannot be

contaminated with catalyst fines and the filters separating the waxes from the

catalyst should never be clogged.

Wei et al. [290] studied the attrition resistance of typical commercial

cobalt catalysts. They found that the attrition strength decreases in the order

Co/Al2O3>Co/SiO2>Co/TiO2 (rutile)>Co/TiO2 (anatase). They also found that

the addition of small amounts of metals (Ru, Cu) or oxides (Zr, La, K, Cr) had

minor effects. The addition of Zr to Co/Al2O3 catalysts had a negative impact.

Moreover, the resistance of alumina carriers presents a maximum when

calcined at 300 °C [291]. Finally, the attrition strength of Co/Al2O3 catalysts

can also be improved by doping the carrier with a compound able to form a

spinel phase (e.g. Ni) followed by calcination at 550 °C [99, 226].

The mechanical strength (tensile and compressive strength [214]) of

catalyst pellets is also an important parameter, particularly in fixed bed

reactors where the catalyst have to withstand a high load and its breakage

results in a higher pressure drop across the bed. The strength of heat-treated

carriers with low porosity and surface area is usually larger than that of high

surface area materials [214]. Moreover, alumina can react with steam

becoming partially hydrated and fragile [292]. Nevertheless, the mechanical


45

strength is apparently not a major concern. This problem is rarely reported.

Current Ni/Al2O3 catalysts apparently overcome this issue even after exposure

to high temperatures and steam [196].

3.7.2. Weak anchoring of cobalt particles to the carrier

In slurry bubble column reactors using Co/Al2O3 catalysts, cobalt may

be washed out from the support if the particles are not sufficiently anchored.

It is hypothesized that the reason for this phenomenon is the partial

dissolution of alumina during catalyst preparation and formation of an

amorphous alumina layer which weakens the anchoring of cobalt [293, 294].

The problem is nevertheless solved by modifying the support with tetra-ethyl-

ortho-silicate followed by calcination to give surface concentration of ca. 2.5

Si atoms/nm2 [111].

3.7.3. Surface reconstruction

It has been observed that the cobalt surface changes significantly during

Fischer-Tropsch synthesis. Planar surfaces reconstruct into spike-shaped

cobalt islands which leads to a change in the nature of the active sites [216].

This phenomenon is believed to be responsible for the activation of the

catalyst [295]. Nevertheless, it has also been proposed as a deactivation

mechanism [227, 296]. The importance of this mechanism and its

contribution to catalyst deactivation is still not well understood and lacks

experimental evidence [226].

3.7.4. Decoration of metal particles by support species

Carriers such as TiO2, can partially be reduced during catalyst activation

with hydrogen, leading to mobile species (TiOx, x<2) that migrate onto the

metal particles [270, 297, 298]. This decoration of metal particles has also

been observed when using acidic oxides (e.g. Ta and W oxides) as coatings on

alumina carriers [299].

Vannice et al. [173, 174] found that the presence of these TiOx species

leads to an increase in the TOF of Ni/TiO2 catalysts. Iglesia et al. [71, 122] also

observed that the TOF of Co/TiO2 catalysts was higher than other catalysts

when testing these at atmospheric conditions. However, they observed that

the TOF was independent on the support when testing the catalysts at high

pressures. This observation was ascribed to a destruction of this metal support

interaction by exposure to the water produced in the FT reaction. Recently,


46

Prieto et al. [299], found that these species do increase the TOF of the

neighboring available Co atoms at relevant FT conditions but there seems to

be a compromise between the increase in the TOF and the extent of

decoration, i.e. the amount of exposed metal atoms.

47

Part II

Experimental

Chapter 4 – Catalyst preparation and characterization

49

Chapter 4

Catalyst preparation and characterization

4.1. Catalyst preparation

This section describes the preparation methods used to synthetize the

cobalt- and nickel-based catalysts employed in this work. Basically, all the

catalysts were prepared by impregnation of commercial carriers with an

aqueous solution containing the salt of a catalytic element (cobalt, nickel or

platinum).

The carriers were pretreated in air prior to impregnation. The γ-Al2O3

and SiO2 carriers were dried at 120 °C for 6 h and then calcined for 10 h at 500

°C (ramp = 1 °C/min). The TiO2 carrier was dried at the same conditions but

calcined at 700 °C for 10 h (ramp = 1 °C/min). The α-Al2O3 carrier was

prepared by calcining the γ-Al2O3 in flowing air at 1100 °C for 10 h.

The carriers were impregnated with aqueous solutions of

Co(NO3)2·6H2O, Ni(NO3)·6H2O or Pt(NO3)2·4NH3. Two different

impregnation techniques were used: incipient wetness impregnation and wet

impregnation. The former technique was employed when the solid material

was a fine powder. The latter was used when the carrier consisted of large

pellets.

4.1.1. Incipient wetness impregnation

The incipient wetness impregnation technique is probably the most

common catalyst preparation procedure due to its simplicity. The technique

consists in impregnating a pre-dried carrier with a liquid solution containing

the salt (precursor) of the catalytic element. For that purpose, the precursor

salt is usually dissolved in a volume of solvent equal to the pore volume of the

solid material. This aqueous solution is then added slowly (dropwise) to the


50

solid and drawn into the pores by capillary forces. The solution is added

continuously while stirring the solid powder until the pores of the carrier are

filled.

For most of the prepared catalysts, the amount of water required to

dissolve the salt was larger than the pore volume of the carrier. In these cases,

the aqueous solution had to be deposited batch-wise, drying the impregnated

carriers at 120 °C for 2 h after every deposition.

Once the impregnation of the carriers with the desired amount of salt

had been completed, the solids were dried for 3 h at 120 °C and calcined in air

in order to convert the precursors into oxides of the catalytic elements. The

cobalt-based catalyst was calcined at 300 °C for 16 h (ramp = 1 °C/min). The

nickel-based catalysts were calcined at 500 °C for 3 h (ramp = 2.5 °C/min).

4.1.2. Wet impregnation

The wet impregnation technique is defined as an impregnation in which

the volume of liquid is larger than the pore volume of the solid material. In

this case, the solids are usually immersed in a vessel containing the precursor

solution. After a specified immersion time, the solution is heated up until its

evaporation. The resulting impregnated pellets are finally dried and calcined.

This procedure was employed to impregnate large cylindrical γ-Al2O3

pellets (see Paper IV). In this case, the wet impregnation technique was more

convenient than the incipient wetness impregnation method. A dropwise

deposition of the liquid solution into such large pellets would have led to a

heterogeneous distribution of nickel in the catalyst batch.

4.2. Characterization of supports and catalysts

This section describes briefly the different characterization techniques

used to study the physicochemical properties of the catalysts and supports.

Some catalyst characterization results are presented here as well.

4.2.1. N2 adsorption

In order to study the physical properties of the catalysts and supports,

nitrogen adsorption measurements at the boiling temperature of liquid

nitrogen (-196 °C) were carried out in a Micromeritics ASAP 2000 unit. The


51

method consists in obtaining adsorption and desorption isotherms, by means

of a static vacuum procedure; these are used for estimation of the surface area,

pore volume and pore size distribution.

The Brunnauer, Emmett and Teller (BET) surface area was estimated

using the adsorption data obtained at relative pressures of between 0.06 and

0.2. The pore volume was estimated from a single point of adsorption at a

relative pressure of 0.998. The average pore diameter was estimated from the

BET surface area and the pore volume assuming that the pores are cylindrical.

4.2.2. Hg intrusion

The nitrogen adsorption technique is not useful when analyzing

materials in which the macropore contribution to surface area and pore

volume is significant [19]. In these cases, changes in macropore size with

varying pressure cannot be determined. The mercury intrusion method is then

preferred for studying the physical properties of macroporous materials (pore

size larger than 50 nm).

Mercury intrusion analyses were performed in a Micromeritics Mercury

Porosimeter, AutoPore IV. The method consists in measuring the volume of

mercury penetrating the pores as a function of the applied pressure (pressure

range: 2-60000 psia).

In the present work, the γ-Al2O3- and SiO2-based materials were

analyzed by nitrogen adsorption. The TiO2- and α-Al2O3-based materials were

analyzed by mercury intrusion. A summary of the physical properties of these

carriers is presented in Table 11.

Table 11: Physical properties of the calcined supports as determined by N2 adsorption and Hg intrusion.

Support BET surface area (m2/g)

Pore volume (cm3/g)

Average Pore Diameter (nm)

γ–Al2O3 a 193 0.50 10 γ–Al2O3 b 223 0.64 11 SiO2 360 0.64 5 TiO2 14 1.39 409 α-Al2O3 7 0.47 255

a Puralox SCCa-5/200 from Sasol b SA 6173 from Saint Gobain NorPro


52

4.2.3. Temperature-programmed reduction

Temperature-programmed reduction (TPR) analyses were carried out

on calcined catalyst samples in order to study the reducibility of cobalt and

nickel oxides. These analyses were performed in a Micromeritics AutoChem

2910 unit. The procedure consists in flowing a reducing gas (5 vol% H2 in Ar)

through the sample tube containing the catalyst while increasing the

temperature from ambient to 930 °C (ramp = 10 °C/min). Meanwhile, the

difference in thermal conductivity between the gas at the inlet and at the outlet

of the sample tube is monitored by means of a thermal conductivity detector

(TCD). The TCD signal was calibrated by performing a TPR analysis with Ag2O

standard, thus estimating the H2 consumption.

The degree of reduction (DOR) of the catalyst after different reduction

treatments was also estimated by means of TPR analyses. For that purpose,

the samples were reduced in situ prior to the previously described TPR

procedure. The H2 consumption determined in these analyses can then be

ascribed to the unreduced metal oxides still present after each reduction

treatment. The DOR was estimated by assuming that the unreduced cobalt

and nickel oxides after the reduction treatment were in the form of CoO and

NiO. As a result, the DOR could be calculated according to equation 4.1.

𝐷𝑂𝑅 = 1 −𝐴𝑇𝐶𝐷 ∙ 𝑓

𝑥𝑚𝑒𝑡𝑎𝑙 𝑀𝑚𝑒𝑡𝑎𝑙⁄ (4.1)

where “ATCD [a.u./g]” is the area resulting from integration of the TCD

signal (normalized per mass of catalyst, “f [moles of H2/a.u.]” is the

calibration factor correlating the TCD signal and the consumption of H2,

“xmetal” is the mass fraction of nickel or cobalt in the catalyst sample and

“Mmetal” is the atomic weight of nickel or cobalt.

4.2.4. Temperature-programmed hydrogenation

Temperature-programmed hydrogenation (TPH) analyses were

performed on spent nickel-based catalysts in order to study the carbon species

present after reaction. The TPH of the samples was performed in the reaction

setup by passing a flow of pure hydrogen through the catalyst bed while

increasing the temperature from room temperature to 690 °C (ramp=5

°C/min). The evolution of methane in the outlet gas, resulting from carbon

hydrogenation, was monitored during the treatment with a HiQuadTM

QMG700 mass spectrometer.


53

4.2.5. H2 static chemisorption

Hydrogen static chemisorption was performed on reduced catalyst

samples in order to estimate the metal dispersion and the average crystallite

size of cobalt or nickel. The analyses were conducted on a Micromeritics

ASAP2020C unit at 35 °C. These were performed on in situ reduced samples

in a pressure interval between 20 and 510 mmHg. The linear part of the

obtained H2 chemisorption isotherms was extrapolated to zero pressure in

order to estimate the H2 monolayer coverage. The metal dispersion was then

calculated assuming that the adsorption stoichiometry between an atom of H

and the metal (cobalt or nickel) was 1:1 [300].

The metal dispersion (DH) and the DOR can then be used to estimate

the average crystallite size of the metal, assuming that the crystals are

spherical. Equations 4.2 and 4.3 were used to estimate the crystallite size of

cobalt (d(Co0)H) and nickel (d(Ni0)H), respectively [301].

𝑑(𝐶𝑜0)𝐻 =96

𝐷𝐻

∙ 𝐷𝑂𝑅 (4.2)

𝑑(𝑁𝑖0)𝐻 =97

𝐷𝐻

∙ 𝐷𝑂𝑅 (4.3)

4.2.6. X-ray diffraction

X-ray diffraction (XRD) was used to identify the crystalline phases

present in solid powders (i.e. the catalysts and supports). The principle of this

analysis is that crystalline structures consist of repetitive arrangements of

atoms capable of diffracting X-rays. Every crystalline element or compound

presents diffraction angles for various planes and, thus, presents a unique

diffraction pattern. The patterns allow for identification of different crystalline

phases in solid samples.

The XRD measurements were conducted in a Siemens D5000

instrument with Cu-Kα radiation (2θ=10-90°, step size=0.02°) equipped with

a Ni filter, operated at 40 kV and 30 mA. The XRD analyses were also used to

estimate the diameter of the cobalt and nickel crystallites making use of the

Scherrer equation and assuming the particles to be spherical [302].

X-ray diffraction (XRD) was conducted on calcined, reduced (and

passivated) as well as on spent catalysts allowing for the determination of the


54

mean crystallite size of NiO (d(NiO)XRD) and Ni (d(Ni0)XRD). Unfortunately,

reduction and passivation in air of cobalt-based catalysts usually leads to a

significant or even complete oxidation of the Co0 crystallites. Therefore, the

cobalt crystallite size (d(Co0)XRD) was estimated from that of Co3O4 particles

(d(Co3O4)XRD) using equation 4.4 [303].

𝑑(𝐶𝑜0)𝑋𝑅𝐷 = 0.75 · 𝑑(𝐶𝑜3𝑂4)𝑋𝑅𝐷 (4.4)

4.2.7. X-ray photoelectron spectroscopy

X-ray photoelectron spectroscopy (XPS) is a technique which consists

in bombarding the catalyst surface with X-ray photons while monitoring

emitted core photoelectrons as function of electron energy. Since every

compound has its unique emission pattern, XPS allows for quantitative

chemical analysis of a thin surface layer. In the present work, XPS was used

to estimate the weight percentage of nickel on the catalyst surface (see Paper

II).

X-ray photoelectron spectroscopy (XPS) data were recorded on 4 mm ×

4 mm pellets with a thickness of 0.5 mm which were obtained by gently

pressing the powdered materials. The sample was later degassed to a pressure

below 2·10−8 Torr at 150 °C in the instrument pre-chamber to remove

chemisorbed volatile species. The main chamber of the Leybold–Heraeus

LHS10 spectrometer was equipped with an EA-200MCD hemispherical

electron analyzer with a dual X-ray source using AlKα (hv = 1486.6 eV) at 120

W, at 30 mA, with C(1s) as energy reference (284.6 eV).

4.2.8. Inductively coupled plasma mass spectrometry

Inductively coupled plasma mass spectrometry (ICP-MS) was used for

elemental analysis of the catalyst samples. The technique consists, as its name

indicates, in ionizing a sample with inductively coupled plasma followed by

separation and quantification of these ions in a mass spectrometer. ICP-MS

analyses of nickel-based catalysts were performed by the staff at the Central

Service for Research Support (SCAI) of the University of Córdoba (see Paper

II). ICP-MS analyses of sulfur-poisoned cobalt-based catalysts were

performed at ALS Scandinavia (see Paper III).


55

4.2.9. Scanning and transmission electron microscopy

The internal morphology of the cylindrical nickel-based catalyst pellet

was observed with a Hitachi TM3000 scanning electron microscope (SEM)

using an accelerating voltage of 15 kV. For that purpose, the pellet was cut in

half. An energy-dispersive X-ray spectroscopy (EDS) mapping was carried out

by means of a Bruker Quantax 70 EDS detector in order to evaluate the

homogeneity of nickel inside the pellet (see Paper IV).

Transmission electron microscopy (TEM) was used to analyze the

morphology of the catalysts and more precisely determine the size of the metal

particles. The TEM images were obtained using a JEOL JEM 1400

microscope. All samples were mounted on 3 mm holey carbon copper grids.

The average particle size was calculated by measuring 80–100 particles.

Figure 13 shows representative TEM images of a titania-supported nickel

catalyst after reduction and passivation in air. The images allow for a clear

distinction between the nickel crystallites and the support.

Figure 13. TEM images of a TiO2-supported nickel catalyst (from Paper II).

Chapter 5 – Catalytic testing

57

Chapter 5

Catalytic testing

5.1. Reactor setup

The catalysts were tested in a down-flow fixed bed reactor (i.d. 9 mm).

A detailed drawing of the reactor and its connections is presented in Figure

14. The reactor was heated by an electric oven and its temperature was

regulated by a cascade control system. For that purpose, a sliding

thermocouple was placed in a thermowell inside the catalyst bed, and another

thermocouple was placed in the oven. This system allowed for the

measurements of the axial temperature profile in the reactor. An aluminum

jacket was placed between the reactor tube and the oven in order to obtain a

more even temperature profile. This configuration, together with a high

catalyst dilution with SiC powder (average pellet size = 75 μm), allowed for

operation at nearly isothermal conditions. The temperature profiles along the

bed were usually within ± 2 °C of the set point. The reader is directed to

Papers I-IV for a more detailed information regarding the catalyst loading,

pellet size, reduction and reaction procedures since these varied depending on

the scope of each study.

The gases were purified before entering the system using cleaning traps.

Water was fed to the reactor by means of a high pressure HPLC Gilson 307

piston pump. The reaction products were separated by two consecutive traps.

The heavy hydrocarbons (FT waxes) and part of the water were condensed in

a first trap heated at 120 °C. The rest of the water and lighter hydrocarbons

were condensed in a second trap at room temperature. The remaining gases

were depressurized and analyzed on-line using an Agilent 6890 gas

chromatograph (GC). A detailed description of the reaction equipment and

instrumentation can be found in Figure 15.


58

Figure 14. Drawing of the reactor used for the catalytic tests and its connections.


59

Figure 15. P&I diagram of the catalytic testing equipment.


60

5.2. Gas analysis and data treatment

The GC used for gas analysis was equipped with a thermal conductivity

detector (TCD) and a flame ionization detector (FID). H2, N2, CO, CH4 and

CO2 were separated by a Carbosieve II packed column and analyzed in the

TCD. The hydrocarbons present in the gas phase (C1 to C8) were separated by

an Al2O3-plot column and analyzed on the FID. The TCD signal of each

compound was calibrated with a certified gas containing known amounts of

N2, CO, CH4 and CO2. In this manner, it was possible to determine the

response factors of CO, CH4 and CO2 relative to N2 (the internal standard).

Thereby, the flow rates of CO, CH4 and CO2 were calculated according to

equation 5.1.

𝐹𝑖 = 𝐹𝑁2∙

𝐴𝑖

𝐴𝑁2

∙ (𝑅𝑒𝑠𝑝𝑜𝑛𝑠𝑒 𝑓𝑎𝑐𝑡𝑜𝑟)𝑖 (5.1)

Here, “F [mol/s]” is the molar flow rate, “A [a.u.]” is the area resulting

from integration of the TCD signal and “i” represents one of the mentioned

compounds (CO, CH4 or CO2). The quantification of the other hydrocarbons

in the FID was carried out by relating their corresponding signals to the one

of CH4, assuming that the FID signal is proportional to the number of C atoms.

The flow rate of these hydrocarbons was calculated according to equation 5.2.

𝐹𝑗 = 𝐹𝐶𝐻4∙

𝐴𝑗∗

𝐴𝐶𝐻4

∗ ∙ 𝑛𝑗 (5.2)

where “j” represents any hydrocarbon different from methane, “A*

[a.u.]” is the area resulting from integration of the FID signal and “𝑛𝑗” is the

number of carbon atoms of compound “j”.

Throughout this work, the CO conversion (XCO) and selectivity (Sp) to

specific carbon-containing products “p” is defined according to equations 5.3

and 5.4, respectively.

𝑋𝐶𝑂 =(𝐹𝐶𝑂)𝑖𝑛 − (𝐹𝐶𝑂)𝑜𝑢𝑡

(𝐹𝐶𝑂)𝑖𝑛

(5.3)

𝑆𝑝 =(𝐹𝑝)

𝑜𝑢𝑡∙ 𝑛𝑝

(𝐹𝐶𝑂)𝑖𝑛 − (𝐹𝐶𝑂)𝑜𝑢𝑡

(5.4)


61

Another important parameter used throughout this work is the so-

called selectivity to C5+ products (SC5+) which, as its name indicates,

represents the selectivity to products containing 5 or more carbon atoms. The

SC5+ is widely used as a measurement of the selectivity to long-chain

hydrocarbons and is defined according to equation 5.5.

𝑆𝐶5+ = 1 − (𝑆𝐶𝐻4+ 𝑆𝐶2 + 𝑆𝐶3 + 𝑆𝐶4 + 𝑆𝐶𝑂2

) (5.5)

63

Part III

Results and discussion

Chapter 6 – Carbonyl-induced nickel particle sintering

65

Chapter 6

Carbonyl-induced nickel particle sintering

6.1. Support and nickel particle size effects on carbonyl-induced

sintering (Paper I)

The objective of this work was to study the stability of typical supported

nickel catalysts under low temperature and high CO partial pressure

methanation conditions. In addition, the influence of the initial nickel particle

size on carbonyl-induced sintering was assessed. For that purpose, a total of

six catalysts were prepared and tested in the reactor setup at 310 °C, 20 bar

and an inlet H2/CO ratio of 3. Four different supports were employed: γ-Al2O3,

α-Al2O3, SiO2 and TiO2. The catalysts were named according to the carrier

used. The number placed before the name indicates the mass fraction of

nickel. The physicochemical properties of these six catalysts are presented in

Table 12.

Table 12. Physicochemical properties of different supported nickel catalysts (adapted from Paper I).

Catalyst Surface area (m2/g)

Pore volume (cm3/g)

dpore (nm)a

d(Ni0)XRD

(nm)b

20Ni/α-Al2O3 30 0.53 71 25

20Ni/TiO2 batch 1 9 0.63 279 25 20Ni/TiO2 batch 2 13 0.82 248 13

10Ni/γ-Al2O3 175 0.43 10 *

20Ni/γ-Al2O3 146 0.34 7 10

20Ni/SiO2 264 0.41 5 13 aAverage pore diameter calculated as: 4·Pore Volume/Surface area. bAverage nickel particle size as estimated by XRD after reduction and passivation in air at room

temperature.

*The XRD peak corresponding to Ni0 was not visible due to the small size of the nickel particles.

As can be observed in Table 12, the catalysts differ significantly in nickel

particle size. It is actually very challenging to synthetize catalysts with the


66

same nickel crystal size, especially when the supports have very different

physical properties. Since both the support and the initial particle size can

have an effect on catalyst deactivation, it was found convenient to prepare two

batches of 20Ni/TiO2 with different nickel particle size in order to provide a

better comparison between the catalysts.

The performance of these catalysts in methanation is presented in

Figure 16. As can be seen, all the catalysts present a rapid and significant loss

of activity in 24 hours. The figure also reveals that 20Ni/TiO2 catalysts present

the lowest deactivation rates. Furthermore, the 20Ni/TiO2 batch 1, which

possesses a larger initial mean nickel particle size, is the most stable.

0 2 4 6 8 10 12 14 16 18 20 22 24

0

10

20

30

40

50

60

70

80

90

100

20Ni/-Al2O

3

20Ni/TiO2 batch1

20Ni/TiO2 batch2

20Ni/SiO2

20Ni/-Al2O

3

CO

co

nve

rsio

n (

%)

Time on stream (h)

Figure 16. CO conversion versus time on stream for different supported nickel catalysts at 310 °C, 20 bar and

an inlet H2/CO ratio = 3 (adapted from Paper I).

The spent catalysts were analyzed by XRD to estimate their mean nickel

particle size and thus, evaluate the resistance of these catalysts to carbonyl-

induced sintering. The nickel particle sizes obtained for these catalysts are

presented in Table 13. As can be observed, sintering of the nickel crystallites

and, consequently, the loss of metallic surface area, was severe for all the

samples. Sintering does not only lead to a loss of metallic surface area but also

to a reduction of the TOF [162]. The loss of activity for all the samples can


67

mainly be ascribed to sintering. Nevertheless, carbon formation is also likely

to contribute at these hostile methanation conditions (see Chapter 7).

Table 13. Ni particle size of the fresh and spent catalyst samples after methanation (adapted from Paper I).

Catalyst Fresh d(Ni0)XRD

(nm)

Spent d(Ni0)XRD

(nm) Loss of Ni area

(%)a

20Ni/α-Al2O3 25 39 36

20Ni/TiO2 batch 1 25 36 31 20Ni/TiO2 batch 2 13 34 62

20Ni/γ-Al2O3 10 44 72

20Ni/SiO2 13 46 77 Precision error of the XRD instrument: ±1 nm, at a confidence level of 95 %. aEstimated assuming the Ni particles to be spherical.

Table 13 also allows for a comparison of the sintering resistance of the

catalysts. It can be observed, by comparing the 20Ni/TiO2 batch 2 with the

20Ni/γ-Al2O3 and the 20Ni/SiO2 catalysts, that the sintering rate is lower for

titania-supported catalysts. This observation is in line with the work of

Vannice and Garten [173], who found that the rate of Ni(CO)4 formation was

one order of magnitude lower for Ni/TiO2 than for Ni/SiO2 catalysts.

Moreover, by comparing the Ni/α-Al2O3 and the Ni/TiO2 batch 2 catalysts it

can also be seen that the sintering rate was lower for the latter catalyst. The

difference is however very small in this case.

A possible explanation for this positive behavior of Ni/TiO2 catalysts is

their characteristic strong metal-support interaction (SMSI) [270]. The

reduction of Ni/TiO2 leads to the reduction of TiO2 into TiOx species (x<2)

which decorate the nickel crystals and modify their catalytic and chemical

properties. The extent of decoration is more pronounced when using catalysts

with small nickel crystals.

Finally, it can be observed that the loss of Ni surface area (i.e. crystal

growth) in 24 hours was lower for the 20Ni/TiO2 catalyst possessing larger

initial nickel particles. A possible explanation is that the Ni particles reached

the size at which these are thermodynamically stable (or almost stable) at

these severe conditions. Indeed, the difference in the final nickel crystallite

size between the two batches is very small and can be ascribed to the precision

of the instrument.

It must be mentioned that Munnik et al. [224] also found that the crystal

growth was reduced when using Ni/SiO2 catalysts with initially larger Ni

particles. Their scientific explanation was based on the fact that small Ni


68

particles lead to a supersaturation of Ni(CO)4 in the small pores. This

supersaturation results in the breakage and widening of the pores and so

allows further sintering. The reader may note that this explanation is unlikely

to justify the results of the present work since TiO2 is a macroporous material

and the size of its pores is not restricting the growth of the nickel particles.

6.2. Catalyst pellet size effects on carbonyl-induced sintering

(Paper IV)

The purpose of this work was to study the effect of the catalyst pellet

size on this peculiar type of sintering occurring at low temperatures. It is well

known that the influence of heat and mass transfer limitations on the observed

activity increases with increasing pellet size. Excellent mathematical models

can be found for estimating the observed catalyst activity in situations in

which the effect of heat and mass transfer is significant [304, 305].

Nevertheless, it is not so well understood how mass and heat transfer

limitations can affect the deactivation rate.

The motivation behind this work arises from a discrepancy between the

laboratory work performed by Shen et al. [163], and the catalyst performance

observed in the pilot reactor in the ADAM I plant [185]. The first adiabatic

methanation reactor in the ADAM I plant operated at an inlet temperature of

306 °C and a CO partial pressure of 154 kPa. According to the work of Shen

and coworkers, these conditions are undoubtedly in an unsafe operating

regime. Nevertheless, carbonyl-induced sintering was never encountered and

the reactor operated successfully for ca. 1500 h. The catalyst used in the ADAM

I plant was the MCR-2X nickel-alumina catalyst from Haldor Topsoe which

consists of 7-hole cylindrical pellets. The catalyst employed in the work of

Shen et al. [163] consisted of nickel-alumina but it was a fine powder.

In order to study this possible pellet size effect, a Ni/γ-Al2O3 catalyst

consisting of 30 wt% Ni was prepared. The catalyst consisted of cylindrical

pellets with a diameter of 3.6 mm and an average length of 5 mm. The average

Ni particle size was estimated to be 10 nm by means of X-ray diffraction.

Complete information about the catalyst’s physicochemical properties can be

found in Paper IV.

The catalyst was then crushed and cut to prepare catalyst pellets of

different sizes. A total of four pellets were prepared: a fine powder (pellet size:

53-90 μm), another powder (pellet size: 500-850 μm), pellets with a size equal

to a 1/8th cut of the original cylindrical pellet (named: 1/8th pellet) and the


69

original pellet (named: Full-size pellet). The pellets were then tested under

methanation conditions at 310 °C, 20 bar and an inlet H2/CO ratio = 3. The

performance of these samples is presented in Figure 17.

Figure 17. CO conversion versus time on stream (left) and relative activity versus time on stream (right) for

different catalyst pellet sizes (adapted from Paper IV).

As can be observed, the two smallest pellet sizes exhibited very high

deactivation rates, in line with the work presented for different catalysts in

Paper I. Nevertheless, the two largest pellets exhibited a surprisingly stable

performance. The figure also reveals that the deactivation rate decreases with

increasing pellet size.

The reaction was stopped after 24 hours in order to analyze the samples

by XRD and determine the size of the Ni crystallites. The XRD results are

presented in Table 14. As can be observed, sintering was severe for the two

smallest pellets. The two largest pellets exhibited almost no growth of the Ni

particles. The slightly higher Ni crystal size observed for the Full-size pellet,

compared to that of the 1/8th pellet, is ascribed to the precision of the

instrument and/or to possible small differences in the initial crystal size

between the samples tested.

The results show that there is an important effect of the pellet size on

the catalyst deactivation rate which is directly related to sintering of the Ni

particles. The most probable scientific explanation for this observation is that

the operation with large pellets leads to large temperature and CO

concentration gradients in the gas film and inside the pellet. Since the reaction

0 4 8 12 16 20 24

0.0

0.2

0.4

0.6

0.8

1.0

53-90 m

500-850 m

1/8th pellet

Full-size pellet

Re

lative a

ctivity

Time on stream (h)

0 4 8 12 16 20 24

0

10

20

30

40

50

60

53-90 m

500-850 m

1/8th pellet

Full-size pellet

CO

convers

ion (

%)

Time on stream (h)


70

is exothermic, the temperature at the catalyst external surface can greatly

exceed that measured in the bulk gas. In consequence, it is possible that the

catalyst surface is exposed to a sufficiently high temperature (and sufficiently

low CO partial pressure) for sintering not to occur.

Table 14. Ni particle size, loss of Ni area and loss of activity of the spent catalyst pellets (adapted from Paper IV).

Sample Spent d(Ni0)XRD

(nm)a

Loss of Ni area

(%)b

Loss of activity (%)

53-90 μm 34 71 >95

500-850 μm 35 71 >95

1/8th pellet 11 9 13 Full-size pellet 12 17 9

aPrecision error of the XRD instrument: ±1 nm, at a confidence level of 95 %. bThe loss of Ni surface area was calculated assuming the particles to be spherical and the initial Ni

particle size to be 10 nm for all the samples.

This explanation is actually in line with the differences in selectivity

observed for the four samples. The selectivity results are presented in Table

15. As can be seen the selectivity to long-chain hydrocarbons decreases with

increasing pellet size. This is a clear indication of important concentration and

temperature gradients in the gas film and/or inside the pellets.

Table 15. Pellet size effects on selectivity (adapted from Paper IV).

Sample SC1 (%)

SC2 (%)

SC3 (%)

SC4 (%)

SC5+

(%)

SCO2 (%)

53-90 μm 63.4 13.8 9.1 3.5 9.5 0.8

500-850 μm 71.3 14.6 9.9 3.7 0.0 0.5

1/8th pellet 92.8 3.8 1.0 0.2 0.0 2.1 Full-size pellet 96.3 1.5 0.2 0.0 0.0 2.0

In order to properly understand this pellet size effect it was necessary

to estimate the temperature and the CO partial pressure at the catalyst

external surface and inside the catalyst pellets. The concentration and

temperature at the external surface was estimated using classical chemical

engineering calculations (calculation of mass and heat transfer coefficients)

[304, 306] and other mathematical expressions for the estimation of the

Sherwood [307] and Nusselt numbers [308]. A summary of the most

important results from these calculations is presented in Table 16. The

detailed mathematical model is presented in Paper IV.


71

Table 16. Estimated mass and heat transfer parameters and conditions at the catalyst external surface (adapted from Paper IV).

Sample kc (m/s)a

h (W/(m2·°C))b

f mass transfer (%)c

Ts (°C)d

53-90 μm 6.06 5886 0.00 310

500-850 μm 0.65 1117 0.02 312

1/8th pellet 0.20 626 0.27 329 Full-size pellet 0.10 493 0.67 340

a Mass transfer coefficient. b Heat transfer coefficient. c Extent of mass transfer control defined as the relative difference between the concentration at the

catalyst surface and the concentration in the bulk gas. d Temperature at the catalyst external surface.

As can be seen in Table 16, the extent of mass transfer control is very

small for all the pellets (less than 1 %). However, the temperature variation

through the gas film is significant, especially for the Full-size pellet. The

presence of large temperature differences between the catalyst surface and the

bulk gas is a common phenomenon in gas phase reactions due to the low

thermal conductivity of the gas film [309, 310].

Unfortunately, this difference in temperature cannot justify the high

stability observed for the large pellets. The catalyst external surface is still

exposed to conditions at which carbonyl-induced sintering is problematic. For

clarification, these conditions are presented in Figure 18 which is the diagram

developed in the work of Shen et al. [163] for identifying safe and unsafe

operating conditions (also presented in Chapter 3).

It must be mentioned that the estimated temperature difference in the

gas film would have been higher if the catalyst pellets had not been diluted

with a fine SiC powder. In the present system, the gas film surrounding the

catalyst pellet comprises SiC particles [307]. This phenomenon enhances the

heat transfer rate and thus reduces the temperature difference between the

bulk gas and the catalyst surface [308]. In a pilot or industrial scenario in

which this dilution with SiC is not present (like in the ADAM I plant) the

temperature difference can be significantly higher. Therefore, it can be

speculated that the temperature at the catalyst surface in the ADAM I reactor

was high enough to expose the catalyst to a safe regime. In addition, the

catalyst pellets prepared by Haldor Topsoe are larger and may be more active

than the one employed in the present work, a fact that increases the

temperature difference in the gas film.


72

Figure 18. Map of safe operating regions for alumina-supported nickel catalysts (adapted from Shen et al. [163]).

The green point indicates the bulk gas inlet conditions in the first methanation reactor in the ADAM I plant. The red points indicate the bulk gas inlet conditions in the experiments conducted at KTH (lower point) and the conditions at the catalyst external surface for the Full-size pellet (higher point).

The temperature and CO partial pressure profiles inside the catalyst

pellets were also calculated in order to try to find an explanation of the results.

Internal mass and heat balances were solved simultaneously using the

software COMSOL Multiphysics v5.0. Detailed information can be found in

Paper IV regarding the assumptions and the equations employed. The

temperature and CO partial pressure profiles inside the different pellets are

presented in Figure 19.

As can be seen, the profiles indicate that the interior of the catalyst and

the external surface can be exposed to very different conditions. In particular,

the CO partial pressure can reach negligible levels in the interior of the

catalyst. This insignificant CO partial pressure can justify the stability of the

nickel crystals under methanation conditions. The figure also reveals that

most of the interior of the catalyst in the 1/8th pellet and the Full-size pellet is

exposed to safe operating conditions.


73

Figure 19. Internal slice temperature (T) and CO partial pressure (pCO) profiles for the four catalyst pellets.

Profiles estimated using a catalyst effective thermal conductivity equal to 8.37· 10−4 W/(cm·°C). Adapted from Paper IV.

It is believed that the explanation for the higher stability observed for

large pellets is this drastic difference between the conditions at the external

surface and the interior of the catalyst. The external surface is, nevertheless,

exposed to unsafe conditions. This could be an explanation for the observed

slow deactivation rates and the slight increase in Ni particle size.

Another possible explanation is that carbonyl species forming near the

external surface diffuse inwards where the concentration of these is lower and

where other nickel crystals are located (in order to form large and

thermodynamically stable nickel particles). However, these carbonyl species

may rapidly decompose as soon as they reach a region in the interior of the

pellet exposed to lower CO partial pressures and higher temperatures. As a

result, the sintering rate is dramatically reduced and practically limited to

particles located on the external surface.

Chapter 7 – Carbon formation in methanation

75

Chapter 7

Carbon formation in methanation

7.1. Support effects on carbon formation (Paper I)

The objective of Paper I was to study the stability of different

supported nickel catalysts under methanation conditions favorable for both

carbon formation and nickel carbonyl-induced sintering. In order to study the

amount and type of carbon formed after methanation, the spent samples were

analyzed by temperature-programmed hydrogenation (TPH). As explained

previously, this procedure consists in converting carbon into methane by

flowing H2. The methane signal is then monitored while increasing the

temperature allowing for the identification of different carbon species.

The TPH profiles obtained for the different supported nickel catalysts

are presented in Figure 20. The results show multiple peaks corresponding to

different carbon species. As explained in Chapter 3, atomic carbon (Cα) is

expected to hydrogenate at around 200 °C. Polymeric carbon (Cβ) should

present its peak temperature at around 400 °C. Other species such as whisker

carbon (Cv) react between 400 and 600 °C.

The first direct observation that can be deduced from the TPH profiles

is that the Ni/TiO2 catalysts presented low carbon formation rates compared

to the other catalysts. Moreover, the carbon species present on Ni/TiO2

catalysts are highly reactive. Carbon is removed from the Ni/TiO2 catalysts at

quite low temperatures. This observation, together with the results obtained

by XRD (see Chapter 6), help to explain the higher stability observed for

Ni/TiO2 catalysts.

The TPH profiles also offer information on the types of carbon present

on the surface. The spent Ni/γ-Al2O3 and Ni/SiO2 catalysts present carbon

species reacting at around 500-600 °C. These could be ascribed to whisker


76

carbon. Nevertheless, it cannot be discarded that these are polymeric carbon

species which are harder to hydrogenate. It must be noticed that the above

classification of species is based on the work of Bartholomew [214], which is a

review of his and his group’s previous work [265, 267] and the pioneering

work of McCarty and Wise [264]. Actually, McCarty and Wise [264], also

observed these hard species when decomposing ethylene on nickel catalysts.

However, these were ascribed to polymeric carbon (Cβ).

0 100 200 300 400 500 600 700 800 Isothermal

CH

4 s

ign

al (a

.u.)

Temperature (°C)

20Ni/-Al2O

3

20Ni/TiO2 batch 1

20Ni/TiO2 batch 2

20Ni/-Al2O

3

20Ni/SiO2

Figure 20. TPH profile for different supported nickel catalysts after 24 hours under methanation at 310 °C, 20 bar and an inlet H2/CO ratio = 3 (adapted from Paper I).

The presence of atomic carbon is a bit unclear. Only Ni/TiO2 catalysts

and Ni/SiO2 catalysts present species hydrogenating close to 200 °C.

Moreover, the formation of Cβ (reacting at around 400 °C) is apparent for all

the catalysts except for the Ni/TiO2 catalysts. Nevertheless, the TPH profile of

the Ni/TiO2 batch 2 insinuates the presence of a possible second peak

hydrogenating at around 300 °C. It is unfortunately not clear whether these

species correspond to lighter polymeric carbon species or to residual atomic

carbon present on the surface after reaction.


77

The nature of this intermediate species reacting at 300 °C was further

studied by repeating the methanation test with the 20Ni/TiO2 batch 2 catalyst.

However, in this second test the reaction was stopped after 2 hours on stream.

Since Cβ is a result from polymerization of Cα, it is expected that the relative

amounts of these two species change with time on stream. In other words, if

this carbon species hydrogenating at 300 °C is polymeric carbon, its amount

should increase with time on stream and its presence should be insignificant

at the beginning of the reaction. The TPH profiles of these two tests performed

with the 20Ni/TiO2 batch 2 catalyst are presented in Figure 21.

0 100 200 300 400 500 600

CH

4 s

ign

al (a

.u.)

Temperature (°C)

20Ni/TiO2 batch 2 after 2 hours

20Ni/TiO2 batch 2 after 24 hours

Figure 21. TPH profile of the 20Ni/TiO2 batch 2 catalyst after 2 h and 24 h under methanation at 310 °C, 20

bar and an inlet H2/CO ratio = 3 (adapted from Paper I).

As can be concluded from Figure 21, there are clearly two carbon

species. The species reacting at 200 °C can be ascribed to atomic carbon. The

species reacting at 300 °C, is ascribed to polymeric carbon since its amount

increases with time on stream. The formation of this light polymeric carbon

species is probably responsible for catalyst deactivation.


78

7.2. Nickel particle size effects on carbon formation (Paper I)

In the present work, the effect of the Ni particle size was evaluated with

Ni/TiO2 catalysts and Ni/γ-Al2O3 catalysts. The two 20Ni/TiO2 catalysts

batches, differing in Ni particle size, were used for this study. The 20Ni/γ-

Al2O3 catalyst was compared to the 10Ni/γ-Al2O3 catalyst which had smaller

Ni particles. For that purpose the catalysts were tested under methanation

conditions during 24 hours (310 °C, 20 bar, H2/CO ratio = 3) and analyzed by

TPH. The TPH profiles of these four tests are presented in Figure 22.

0 200 400 600 800

CH

4 s

ign

al / in

itia

l N

i su

rfa

ce

are

a (

a.u

./m

2)

Temperature (°C)

10Ni/-Al2O

3 d(Ni)

H= 10 nm

20Ni/-Al2O

3 d(Ni)

H= 14 nm

20Ni/TiO2 batch 1 d(Ni)

XRD= 25 nm

20Ni/TiO2 batch 2 d(Ni)

XRD= 13 nm

Figure 22. Ni particle size effect on carbon formation for Ni/TiO2 and Ni/γ-Al2O3 catalysts (adapted from Paper

I).

As can be seen, the rate of carbon formation is higher for larger Ni

particles in the case of Ni/γ-Al2O3 catalysts. Surprisingly, Ni/TiO2 shows the

opposite behavior. This result observed for Ni/γ-Al2O3 catalysts is in line with

the study of Yan et al. [311] in which they compared the carbon formation rate

for two Ni/SiO2 catalysts differing in their Ni crystal size. The results are also

analogous to small cobalt nanoparticles which exhibit higher H2 coverage and

lower polymerization rates in FTS [115].


79

This contradictory Ni particle size effect observed on Ni/TiO2 catalysts

is probably due to the SMSI characteristic of these catalysts. It is believed that

the presence of TiOx (x<2) species promote the dissociation of CO by accepting

oxygen atoms [312]. As a result, the activity and the polymerization activity of

the catalyst is increased [173, 174, 313]. The results observed in the present

study are actually in agreement with the work of van de Loosdrecht et al. [270]

in which they also compared the carbon formation rates on Ni/TiO2 catalysts

with different Ni crystal sizes. In their study they also found that Ni/TiO2

catalysts presented much larger carbon formation rates than Ni/SiO2

catalysts. However, the Ni/TiO2 catalysts employed in their work consisted of

smaller Ni crystallites (in the range of 1-5 nm). Apparently, it seems that an

excessive SMSI on Ni/TiO2 catalysts can have an adverse effect on carbon

formation.

7.3. Effect of the H2/CO ratio on carbon formation (Paper II)

A systematic study was performed on the 20Ni/TiO2 batch 1 catalyst in

order to evaluate the effect of different operating conditions on catalyst

deactivation and carbon formation. One of the operating parameters studied

was the H2/CO ratio. For that purpose, the catalyst was tested under

methanation conditions for 6 hours at 310 °C, 1 bar and using inlet H2/CO

ratios equal to 2, 3, 6 and 9. The spent samples were then analyzed by TPH in

order to evaluate the H2/CO effect on carbon formation.

The TPH results are presented in Figure 23. As expected, higher H2/CO

ratios lead to lower amounts of carbon. The experiment using a H2/CO ratio =

9 practically only shows formation of atomic carbon at temperatures even

below 200 °C. The profiles indicate the existence of two carbon species: one

hydrogenating between 400 °C and 500 °C (probably, Cβ) and another reacting

between 200 °C and 300 °C (ascribed to Cα and/or light polymeric carbon). It

is also interesting to see that the peaks shift towards higher hydrogenating

temperatures with decreasing H2/CO ratio. This suggests that low H2/CO

favors the formation of carbon species with less reactivity.


80

0 100 200 300 400 500 600

CH

4 s

ign

al (a

.u.)

Temperature (°C)

H2/CO = 2

H2/CO = 3

H2/CO = 6

H2/CO = 9

a

b

c

d

Figure 23. Effect of the H2/CO ratio on carbon formation for the 20Ni/TiO2 batch 1 catalyst: a) H2/CO=2, b)

H2/CO=3, c) H2/CO=6 and d) H2/CO=9. Adapted from Paper II.

This higher carbon formation rate observed when using lower H2/CO

ratios is in agreement with the results obtained from high-pressure

methanation tests (310 °C, 20 bar). These additional results are presented in

Figure 24 which shows the CO conversion as function of time on stream for

two consecutive operating periods. The numbers written inside the figure next

to the periods indicate the catalyst deactivation rate in terms of “loss of CO

conversion/hour”. As can be observed, the deactivation rate decreases when

increasing the H2/CO ratio from 3 to 6.


81

80 90 100 110 120

0

10

20

30

40

50

60

70

80

90

100

-0.18 %/h

inlet H2/CO=3

inlet H2/CO=6

CO

co

nve

rsio

n (

%)

Time on stream (h)

-0.23 %/h

Figure 24. Effect of the H2/CO ratio on the deactivation rate at 310 °C and 20 bar for the 20Ni/TiO2 batch 2

catalyst. Adapted from Paper II.

7.4. Effect of pressure on carbon formation (Paper II)

Another parameter evaluated in this study was the operating pressure.

For that purpose, methanation tests were carried out at 310 °C, 20 bar and

different H2/CO ratios. In this manner, the results obtained from TPH could

be compared with those presented in the previous section (carried out at

atmospheric pressure).

The TPH profiles obtained for different H2/CO ratios at 1 bar and 20 bar

are presented in Figure 25. As can be observed, the amount of carbon formed

is lower for the experiments carried out at 20 bar. This is quite interesting

because the formation of polymeric carbon should be enhanced by increasing

the partial pressure of CO. Nevertheless, increasing the pressure also involves

an increase in the partial pressure of H2. It may therefore be that the carbon

gasification rate is very sensitive to the partial pressure of H2, a fact that would


82

compensate the increase in the carbon formation rate when increasing the CO

partial pressure.

Another interesting observation is that the carbon species

hydrogenating at around 400-500 °C are practically not present (or not

visible) in the TPH profiles of the samples treated at high pressure. The

absence of this hardly-reacting carbon species may explain the slight decrease

in the deactivation rate when increasing the H2/CO ratio at high pressures

(presented in Figure 24).

0 100 200 300 400 500 600 700

CH

4 s

ign

al (a

.u.)

Temperature (°C)

H2/CO = 2 P = 1 bar

H2/CO = 2 P = 20 bar

H2/CO = 3 P = 1 bar

H2/CO = 3 P = 20 bar

H2/CO = 6 P = 1 bar

H2/CO = 6 P = 20 bar

Figure 25. Effect of pressure on carbon formation for the 20Ni/TiO2 batch 1 catalyst (adapted from Paper II).


83

7.5. Effect of temperature on carbon formation (Paper II)

The effect of temperature on carbon formation is actually less

understood than it may appear. According to Bartholomew [214], the

formation of polymeric carbon (Cβ) on Ni/γ-Al2O3 catalysts occurs in the

temperature range between 325 °C and 425 °C, approximately. However,

according to the pilot experience from Haldor Topsoe [196, 197], deactivation

due to Cβ formation occurs mainly below 330 °C. It may be worth mentioning

that the observations of Haldor Topsoe members are based on pilot

experiments using large catalyst pellets [196].

Three experiments were performed in order to study the effect of

temperature on carbon formation. The temperatures employed were 280 °C,

310 °C and 340 °C. The experiments were carried out at 1 bar and using a

H2/CO ratio = 3. The TPH profiles of the spent samples are presented in Figure

26. As can be seen, the carbon formation rate increases with increasing

temperature, especially for the species hydrogenating at around 500 °C.

0 100 200 300 400 500 600

CH

4 s

ign

al (a

.u.)

Temperature (°C)

T=280°C

T=310°C

T=340°C

Figure 26. Effect of temperature on carbon formation for the 20Ni/TiO2 batch 1 catalyst (adapted from Paper II).


84

The observation that the carbon formation rate is higher when using

higher temperatures is in line with the results obtained from high pressure

methanation tests (H2/CO=3 and 20 bar). These results are presented in

Figure 27. As can be seen, the deactivation rate increases significantly when

increasing the temperature from 310 °C to 340 °C.

120 125 130 135 140 145 150 155 160 165 170

0

10

20

30

40

50

60

70

80

90

100

-0.65 %/h

T = 310 °C

T = 340 °C

CO

co

nve

rsio

n (

%)

Time on stream (h)

-0.13 %/h

Figure 27. Effect of temperature on the deactivation rate for the 20Ni/TiO2 batch 2 catalyst at 20 bar and using

an inlet H2/CO=3 (adapted from Paper II).

The results show a decent agreement with those observed in the work

of Bartholomew [214] even though a different catalyst was used. Nevertheless,

the effect of temperature on carbon formation is a topic which deserves more

attention. Other parameters such as the presence of CO2 or steam in the feed

may shift the temperatures at which carbon formation is observed.


85

7.6. Effect of steam addition on carbon formation (Paper II)

The last studied operating parameter was the addition of steam to the

syngas feed. For that purpose, three experiments were carried out at 20 bar

and different H2/CO ratios but adding steam to the feed. For a proper

comparison with experiments in which steam is not added, three analogous

experiments were also performed, however, adding helium to the feed. The

TPH profiles are presented in Figure 28. The results indicate that steam

addition reduces the rate of carbon formation. These results were actually

expected since H2O should act as a carbon gasification agent.

0 100 200 300 400 500 600 700 800

CH

4 s

ign

al (a

.u.)

Temperature (°C)

H2/CO=2, 20% He in feed

H2/CO=2, 20% H

2O in feed


H2/CO=3, 5% H

2O in feed


H2/CO=6, 20% H

2O in feed

Figure 28. Effect of steam addition on carbon formation for the 20Ni/TiO2 batch 1 catalyst (adapted from Paper

II).

Just as in the previous sections, the effect of steam addition on the

catalyst deactivation rate was also evaluated by performing high-pressure

experiments at 20 bar, 310 °C and a H2/CO ratio = 3. This was done by

performing a sequence of varying operating conditions. The sequence

consisted of four periods. In the first period, the catalyst was exposed to 20

vol% He in the feed. In the second period, the feed syngas contained 10% He


86

and 10% H2O. In the third period, 20% H2O was added to the feed (no He

added). The last period was identical to the first one.

20 30 40 50 60 70 80

10

15

20

25

30

-0.15 %/h

-0.31 %/h

-0.24 %/h

20 % He

10 % He and 10 % H2O

20 % H2O

20 % He

CO

co

nve

rsio

n (

%)

Time on stream (h)

-0.17 %/h

Figure 29. Effect of steam addition on catalyst deactivation for the 20Ni/TiO2 batch 2 catalyst after reduction at

500 °C. Reaction conditions: 310 °C, 20 bar and inlet H2/CO=3. Adapted from Paper II.

The results of this sequence are presented in Figure 29. As can be

deduced, addition of steam enhances the catalyst deactivation rate. This is

somewhat surprising since steam clearly diminished the carbon formation

rate. In consequence, water is enhancing another deactivation mechanism.

Three deactivation causes were postulated:

- Nickel particle sintering (hydrothermal sintering).

- Metal-support compound formation.

- Destruction of the strong metal-support interaction (SMSI)

characteristic of Ni/TiO2 catalysts (oxidation of TiOx species).

The spent samples were analyzed by XRD in order to evaluate if water

enhanced the Ni crystal growth or if it enhanced re-oxidation and formation

of metal-support compounds (e.g. NiTiO3). Unfortunately, the samples

presented approximately the same mean Ni particle size after reaction as

before and no NiO nor NiTiO3 compounds were visible. However, this does


87

not discard the fact that metal-support compound formation occurs since the

crystal thickness of these compounds could be too thin to be observed by XRD.

In order to understand whether water can destroy the SMSI of Ni/TiO2

catalysts some additional experiments were performed. One experiment

consisted in evaluating the effect of steam addition on deactivation but

reducing the catalyst at 300 °C, a temperature at which this SMSI should be

weaker or nonexistent. Thus, water should not exhibit any increase in the

deactivation rate since no destruction of these species occurs. Another

experiment consisted in comparing the deactivation rate of the catalyst when

reducing it at 300 °C and 500 °C. If water destroys this SMSI, a faster

deactivation rate for the sample reduced at 500 °C is to be expected. These two

experiments are presented in Figure 30 and 31.

25 30 35 40 45 50 55

0

5

10

15

20

25

30

35

40

20 % He

10 % He and 10 % H2O

CO

co

nve

rsio

n (

%)

Time on stream (h)

-0.19 %/h

-0.16 %/h

Figure 30. Effect of water addition on the deactivation rate for the 20Ni/TiO2 batch 2 catalyst after reduction at 300 °C. Reaction conditions: 310 °C, 20 bar and inlet H2/CO ratio = 3. Adapted from Paper II.


88

0 4 8 12 16 20 24

0

20

40

60

80

100

20 Ni/TiO2 batch 2 after reduction at 500 °C

20 Ni/TiO2 batch 2 after reduction at 300 °C

CO

co

nve

rsio

n (

%)

Time on stream (h)

° ° °

Figure 31. Effect of the reduction temperature on the deactivation rate for the 20Ni/TiO2 batch 2 catalyst.

Reaction conditions: 310 °C, 20 bar and inlet H2/CO ratio = 3. Adapted from Paper II.

As can be observed, the addition of steam to the feed did not accelerate

the deactivation rate for the catalyst reduced at 300 °C. Moreover, the catalyst

reduced at 300 °C presents a lower deactivation rate than that reduced at 500

°C. These two observations suggest that this SMSI interaction, which enhances

the activity of the catalyst, is progressively destroyed during methanation.

Finally, a methanation experiment was carried out in situ in the H2-

chemisorption instrument (Micromeritics ASAP 2020) in order to observe the

variation of the metallic surface area with time on stream. As explained before,

this SMSI weakens the chemisorption of H2, a fact that results in an

underestimation of the metallic surface area. Nevertheless, if this SMSI is

destroyed, the observed metallic surface area observed by H2 chemisorption

should increase after methanation. Two consecutive methanation

experiments of a duration of 5 minutes with intermediate chemisorption

analyses were performed. The results are presented in Table 17.


89

Table 17. Variation of the metallic surface area after methanation for the 20Ni/TiO2 batch 1 catalyst. Reaction conditions: 310 °C, 1 bar and inlet H2/CO ratio = 3. Adapted from Paper II.

Time on stream (minutes) Metallic surface area (m2/g)

0 0.11 5 1.32 10 1.40

Metallic surface area estimated by H2 chemisorption after reduction at 500 °C.

The H2 chemisorption results also support the fact that this SMSI

disappears during methanation. This water effect on catalyst deactivation

requires nevertheless further study. As mentioned previously, other

deactivation causes such as the formation of metal-support compounds

should not be discarded.

Chapter 8 – Sulfur poisoning of cobalt Fischer-Tropsch catalysts

91

Chapter 8

Sulfur poisoning of cobalt Fischer-Tropsch

catalysts

8.1. Effect of sulfur on activity (Paper III)

The objective of this work was to study the effect of sulfur on the activity

and selectivity of cobalt-based Fischer-Tropsch catalysts. For that purpose, a

Pt-promoted Co/γ-Al2O3 catalyst was ex situ poisoned with different amounts

of sulfur. The procedure used to poison the catalyst was that described in the

work of Visconti et al. [272]. This procedure consists in reducing the catalysts

and, without exposing them to air, impregnating them with (NH4)2S solutions

in an inert atmosphere. The purpose of this method is to better simulate the

interaction between sulfur and Co0 (in reduced state).

A total of four catalysts were prepared containing 10, 100, 250 and 1000

ppmw of S. These amounts of sulfur correspond to 250, 2500, 6250 and

25000 hours on stream for a gas hourly space velocity (GHSV) of 2 NL/h-

gcatalyst and a sulfur concentration in syngas of 20 ppbv, assuming that all the

sulfur fed into the reactor chemisorbs irreversibly on the catalyst surface. The

catalysts were named according to their sulfur loading (i.e. S10, S100, S250

and S1000). The S-free catalyst was also included in the work and was named

S0. The samples were then characterized by means of TPR, XRD, H2

chemisorption and N2 physisorption and tested under relevant FT conditions.

An additional experiment was carried out in order to estimate the sulfur

capacity of the catalyst, i.e. the amount of sulfur in the catalyst when its

metallic surface area is completely occupied by S atoms. This was done by in

situ poisoning of the reduced catalyst with a H2S/H2 (H2S/H2=10-5) flow in a

quartz tube reactor at 350 °C for 2 days, inside a ventilated fume hood.

Afterwards, the S-passivated sample was analyzed by ICP-MS in order to

determine its sulfur content (i.e. the sulfur capacity).


92

A summary of the physicochemical properties of the fresh catalyst is

presented in Table 18. The catalyst is characterized by a relatively high degree

of reduction (due to its promotion with Pt) and large Co particle size.

Table 18. Physicochemical properties of the fresh Pt-promoted Co/γ-Al2O3 catalyst (adapted from Paper III).

Property Value

Degree of reduction a 77 % Metal dispersion b 4.7 % d(Co0)H c 15.7 nm d(Co3O4)XRD d 16.3 nm d(Co0)XRD e 12.3 nm Sulfur capacity f 5377 ppmw

aDegree of reduction as estimated by H2-TPR after catalyst reduction. bMetal dispersion estimated by H2 chemisorption after catalyst reduction. cAverage Co particle size estimated by H2 chemisorption after catalyst reduction. dAverage Co3O4 particle size estimated by X-ray diffraction after catalyst calcination. eAverage Co particle size estimated from the average Co3O4 particle size.

fSulfur capacity as estimated by ICP-MS after catalyst reduction and exposure to a H2S/H2 flow.

The Fischer-Tropsch experiments were carried out at 20 bar, 210 °C and

using an inlet H2/CO ratio 0f 2.1. In order to make an accurate comparison of

the different samples, the GHSV was adjusted in order to expose all the

catalysts to a CO conversion of 30 %. The tests lasted for 200-300 hours in

order to reach steady state conditions for both activity and selectivity.

The relative activity of the catalyst was studied both as function of the

sulfur loading in the catalyst and the sulfur coverage (i.e. sulfur loading/sulfur

capacity). A deactivation model was also proposed, based on the experimental

data, in order to estimate the loss of activity as function of the sulfur coverage.

The following deactivation expression (Equation 8.1) was used to fit the

experimental data:

𝑎 = (1 − 𝜃𝑠)𝛿 (8.1)

where “a” is the relative activity (activity/activity of the S-free sample)

and θs is the sulfur coverage. The activity as function of the sulfur loading and

the sulfur coverage is presented in Figure 32. As can be seen, sulfur has a very

detrimental effect on catalyst activity. The catalyst loses practically all its

activity long before its surface is completely blocked by sulfur atoms. The

deactivation order “δ” which best fitted the data was around 7, which is a very

high value compared to that obtained by sulfur poisoning of Ni-based catalysts

[281, 283].


93

0 1000 2000 3000 4000 5000

0.0

0.2

0.4

0.6

0.8

1.0

Rela

tive

activity (

a)

Sulfur loading [g S/g cat]

0.0 0.2 0.4 0.6 0.8 1.0

Experimental data points

Deactivation model

Sulfur coverage (s)

a=(1-s)

=7.06

Figure 32. Relative activity as a function of the sulfur coverage and the sulfur capacity (adapted from Paper III).

The relative activity was also studied and modelled as function of the

sulfur-to-cobalt active site ratio. This additional model was added to the study

because it has been found that the S/Co adsorption stoichiometry increases

with sulfur coverage [314]. This fact implies that a proper understanding of

this variation of the S/Co adsorption stoichiometry is required to determine

the available Co metallic area (or the real sulfur coverage) for each poisoned

sample. Therefore, the following expression (Equation 8.2), analogous to

Equation 8.1, was also used to fit the experimental data:

𝑎 = (1 −𝑆

𝐶𝑜∗)

𝛿

(8.2)

where “S/Co*” is the sulfur-to-cobalt active site ratio calculated using

the metal dispersion obtained with H2 chemisorption. The relative activity as

a function of the S/Co* is presented in Figure 33. The model assumes that the

activity of the catalyst is negligible at S/Co* = 1, as evidenced in other

experimental studies [277]. The deactivation order “δ” found using this

equation is 3.78. It may be worth mentioning that Figure 33 is in good


94

agreement with the in situ poisoning studies from Bartholomew et al. [276]

using Co/SiO2 catalysts.

0.0 0.2 0.4 0.6 0.8 1.0

0.0

0.2

0.4

0.6

0.8

1.0

a=(1-S/Co*)

=3.78

Experimental results

Deactivation model

Rela

tive a

ctivity (

a)

S/Co* (Sulfur atoms / Co active site)

Figure 33. Relative activity as a function of the sulfur-to-cobalt active site ratio (adapted from Paper III).

8.2. Effect of sulfur on selectivity (Paper III)

The effect of sulfur on selectivity is more difficult to evaluate than it may

appear. There are different issues which should be considered for a proper

assessment of this effect. Among others, it must be taken into consideration

that the FT selectivity is dependent on the CO conversion. This effect of

conversion on selectivity is particularly important because poisoning by sulfur

decreases the activity of the catalyst and so the conversion. Another important

aspect to take into consideration is the strong chemisorption of S on cobalt.

This issue is important in the case of studying poisoning in fixed bed reactors

by in situ procedures. In these situations, it is likely that sulfur chemisorbs

basically at the inlet of the catalyst bed leading to two reactor zones: one which

is totally deactivated and increases in size with time on stream and another

that behaves as a nearly S-free catalyst. As a result, the observed catalyst


95

selectivity is similar to that observed on a catalyst which has not been

poisoned.

The experiments in the present work were therefore carried out at

varying GHSV in order to adjust the CO conversion to 30 % for all the tests.

The selectivity to C1, C2, C3, C4, C5+ and CO2 is presented in Figure 34 for all

the samples. As can be seen, the selectivity to methane increases significantly

with increasing S content. Sulfur also gives a slight increase in the selectivity

to C2-C4 hydrocarbons. As a result the selectivity to long chain hydrocarbons

(C5+) decreases with increasing S concentration in the catalyst. The selectivity

to CO2 was found negligible for all the samples and presented no appreciable

trend with increasing S content.

0

10

20

30

40

50

60

70

80

90

100

SCO2

SC5+

SC4

SC3

SCH4

SC2

Sele

ctivity (

%)

S0

S10

S100

S250

S1000

Figure 34. Effect of sulfur on the selectivity to different hydrocarbons and CO2. Results obtained at 210 °C, 20 bar, inlet H2/CO=2.1 and at a CO conversion=30 %. Adapted from Paper III.

The effect of sulfur on the olefin-to-paraffin ratio was also studied for

the C2, C3 and C4 hydrocarbons. The results are presented in Figure 35. As

can be observed, the olefin-to-paraffin ratio decreases considerably with

increasing sulfur loading.


96

0 200 400 600 800 1000 1200

0.00

0.04

0.08

0.12

0.16

C2

C2

ole

fin

to

pa

raffin

ra

tio

Sulfur loading (ppmw)

0

1

2

3

4

C3

C4

C3

an

d C

4 o

lefin

to

pa

raffin

ra

tio

Figure 35. Effect of sulfur on the olefin-to-paraffin ratio. Results obtained at 210 °C, 20 bar, inlet H2/CO=2.1 and

at a CO conversion=30 %. Adapted from Paper III.

In conclusion, the results suggest that S increases the selectivity to short

chain hydrocarbons (methane, mainly) and favors olefin hydrogenation. A

possible explanation for this effect is the hydrogenation activity of cobalt

sulfide [315] and a possible electronic modification of neighboring S-free

cobalt atoms. However, it must be mentioned that this effect of S on selectivity

is still a debated topic and it should be clarified further to provide scientific

explanations.

Several researchers have performed similar studies. However, the

results are very contradictory. A small review of the main results of different

sulfur poisoning studies is presented in Table 19. This lack of scientific

agreement may be due to the use of different catalysts, poisoning agents,

poisoning methods, type of reactor and perhaps, an underestimation of CO

conversion effects on the FT selectivity, in some cases.


97

Table 19. Review of the effect of sulfur on selectivity according to different literature studies (author’s own

elaboration).

Reference Year Catalyst Poisoning Reactor SC5+ SCO2 o/pa

[276] 1985 Co/SiO2 In situ H2S

Fixed-bed

↑ ↓ n.a.

[272] 2007 Co/Al2O3 Ex situ (NH4)2S

Fixed-bed

↓ ↑ ↑

[316] 2008 Co/TiO2 Ex situ (NH4)2S

Fixed-bed

↓ n.a. n.a.

[314] 2010 Co/Al2O3 In situ (CH3)2S

CSTRb ↓ n.a. ↓

[277] 2011 Co/Al2O3 In situ H2S

Fixed-bed

0 0 ↓

[317] 2014 Co/SiO2 In situ C4H10S

Fixed-bed

↓ n.a. ↑

Paper III

2015 Pt-Co/Al2O3

Ex situ (NH4)2S

Fixed-bed

↓ 0 ↓

[318] 2016 Pt-Co/Al2O3

In situ H2S

CSTRb ↑ ↑ ↓

The symbol “↑” indicates that the selectivity increases with the addition of S.

The symbol “↓” indicates that the selectivity decreases with the addition of S.

The symbol “o” indicates that the selectivity is practically unaffected with the addition of S.

The abbreviation “n.a.” means that effect of S on that parameter was not assessed or not reported. a Olefin-to-paraffin ratio. b Continuously stirred tank reactor.

Chapter 9 – Conclusions and outlook

99

Chapter 9

Conclusions and outlook

The objective of this thesis was to investigate the deactivation of cobalt-

and nickel-based catalysts during Fischer-Tropsch synthesis and

methanation, respectively. Special focus was given to three deactivation

causes: carbonyl-induced nickel particle sintering, carbon formation on nickel

catalysts and sulfur poisoning of cobalt catalysts. The main results and

conclusions are summarized in the following sections.

Future work could be nonetheless oriented to other deactivation causes

which are also important. There are many aspects that could be studied as a

continuation of this doctoral thesis. Some of these are pointed out below:

- Investigation of the two possible sintering mechanisms in low-

temperature Fischer-Tropsch (carbonyl-induced cobalt sintering and

thermal sintering).

- Understanding the effect of steam on catalyst deactivation at low

temperatures.

- Studying the poisoning effect of biomass-derived impurities.

- Understanding how to prepare thermally resistant materials for high-

temperature methanation.

- Development of cost-effective catalyst regeneration methods.

9.1. Carbonyl-induced nickel particle sintering

Carbonyl-induced sintering was studied using different supported

nickel catalysts. Four supports were employed: TiO2, SiO2, γ-Al2O3 and α-

Al2O3. It was found that Ni/TiO2 catalysts exhibit lower sintering rates than

the other catalysts. The origin of this stability is not fully understood but may

be related to the strong metal-support interaction (SMSI) characteristic of

Ni/TiO2 catalysts.


100

The most astonishing result found on this topic was the effect of the

catalyst pellet size. The sintering rate is extremely reduced when using large

pellets even at the severe conditions employed in the present thesis (310 °C

and 500 kPa of CO). This observation was explained by the presence of

significant temperature and CO partial pressure differences between the bulk

gas and the interior of the catalyst. The interior of large catalyst pellets is

exposed to lower CO partial pressures and higher temperatures, a fact that

reduces the potential for Ni(CO)4 formation.

At present, the criterion used to identify safe and unsafe operating

regimes is only based on the CO partial pressure and the temperature. Future

work could address redefining this criterion and modelling the sintering rate

having in consideration other variables such as the initial Ni particle size and

mass and heat transfer parameters. A better understanding of the SMSI

present on Ni/TiO2 catalysts is vital to evaluate whether TiO2 supports (or

other supports coated with TiO2) can be used in the methanation of synthesis

gas.

9.2. Carbon formation in methanation

Carbon formation was also studied using Ni/TiO2, Ni/SiO2 and

Ni/Al2O3 catalysts. It was found that the carbon species present on the Ni

surface vary significantly depending on the carrier employed. Ni/TiO2

exhibited the lowest carbon formation rates. It was also found that the Ni

particle size has an important effect on carbon formation. The carbon

formation rate is lower for smaller Ni particles when using Ni/γ-Al2O3

catalysts. However, the opposite Ni particle size effect was observed for

Ni/TiO2 catalysts. This contradictory effect was ascribed to a stronger metal-

support interaction between Ni and TiO2 when using smaller particles.

Apparently, an excessive SMSI on Ni/TiO2 catalysts can have an adverse effect

on the carbon formation rate.

The effect of different operating conditions on carbon deposition was

also studied using Ni/TiO2 catalysts. It was found that increasing the H2/CO

ratio, the syngas pressure and the amount of steam in the feed decreased the

carbon formation rate. In contrast, increasing the temperature from 280 °C to

340 °C favored carbon deposition. The carbon formation rates were in line

with the observed deactivation in high-pressure experiments for all the

studied operating parameters except for steam addition, which exhibited a

clear enhancement of the deactivation rate. This adverse effect of steam


101

addition may be ascribed to a progressive destruction of the SMSI, i.e. the

oxidation of TiOx (x>2) species.

Future work could be focused on studying the effect of co-feeding CO2.

Carbon dioxide can act as a carbon gasification agent and therefore, the effect

of the inlet CO2 concentration on carbon formation and deactivation is of great

interest. Further efforts should also be aimed at understanding the

detrimental effect of steam on catalyst activity at low temperatures. The effect

of steam on other deactivation causes such as surface re-oxidation and metal-

support compound formation must be carefully studied. In situ

characterization techniques (e.g. X-ray photoelectron spectroscopy) may help

in understanding whether these deactivation phenomena are enhanced in the

presence of steam.

9.3. Sulfur poisoning of cobalt Fischer-Tropsch catalysts

A Pt-promoted Co/γ-Al2O3 catalyst was ex situ poisoned with known

amounts of sulfur in order to evaluate the effect of this poison on the activity

and selectivity of the catalyst. As expected, the activity decreases with sulfur

addition. A deactivation model was proposed in order to predict the decline in

activity as function of the sulfur coverage and the sulfur-to-cobalt active site

ratio.

It was found that sulfur decreases the selectivity to long-chain

hydrocarbons (mainly, by increasing the selectivity to methane) and the

selectivity to olefins. This effect of sulfur may be associated with the

hydrogenation activity of cobalt sulfide. Nevertheless, other researchers have

observed contradicting results. A better consensus between literature studies

on this topic is required before making any concluding remark. Future work

could address clarifying this controversial effect of S on selectivity. For that

purpose, it would be interesting to perform in situ poisoning tests with the

catalyst.

Nonetheless, research efforts focused on the development of cobalt

catalysts with higher sulfur resistance may be more valuable. Catalysts

tolerating higher sulfur concentrations in the feed gas may reduce costs

associated to syngas cleaning. In this sense, it could be interesting to study the

effect of sulfur scavengers (e.g. ZnO) as catalyst promoters or supports.

Acknowledgements

103

Acknowledgements

First and foremost I would like to thank my supervisors Associate

Professor Magali Boutonnet and Professor Emeritus Sven Järås for giving me

the opportunity to do a PhD at KTH, for their help and for their confidence in

me. I am very thankful for the high level of responsibility that you have given

to me in the FFW project. I really appreciate the possibilities you have given

me to travel to conferences and project meetings and to expand my contact

network.

I want to send an eternal thank you to my great friend Matteo Lualdi,

former PhD student at KTH working in the same field, for transferring his

knowledge and enthusiasm. You have helped me a lot not only in my work but

also in many other things. Thank you.

I want to thank Rodrigo Suárez París, Rodri, for his friendship and help

as a lab, flat and officemate during these years. Thank you for all our fruitful

discussions.

Thank you to Professor Lars J. Pettersson, Lasse, for your teachings, for

organizing such interesting PhD courses, for arranging such joyful events at

our division and for your constant optimism and sense of humor.

I would also like to thank Roberto Lanza for his teachings in catalysis,

chemistry and cooking! I learned lots of things with you during our lunch

times.

I want to thank all professors and researchers, former and present PhD

students, master thesis workers and others who have passed KT for your help

and for making my stay at KT a wonderful time.

Thank you to my two talented master thesis students (Niklas González

and Baldassarre Venezia) and our guest researchers (Vicente Montes and

Victoria Garcilaso de la Vega) who have worked with me and greatly

contributed to our presentations and publications.

I want to acknowledge Professor Jens. R. Rostrup-Nielsen for sharing

his knowledge on catalyst deactivation and scale up of catalytic processes at

KT. Our discussions have definitely enlightened the path to better research.

Acknowledgements

104

Many thanks to Christina Hörnell for her assistance in the English

language and revision of this thesis.

I would like to acknowledge the financial support of the European

Union Seventh Framework Programme (FP7/2013) which has made this work

possible within the FFW project. I would also like to thank all the partners

from the FFW project who I had the pleasure to work with. I have enjoyed and

learned a lot working with you.

Thank you to the all good friends I made in Stockholm and my

basketball team mates (Stockholm Rockets) for the great time I spent with you

these years. Thank you to all my friends from Barcelona for being always there

for me.

My deepest thank you to my dear Parisa. Thank you for filling my life

with so much joy and love. Thank you as well for your patience and

understanding during all this time. I would not have been able to do this

without your sweet and warm heart which provides me with the energy to

achieve anything I want.

Muchas gracias a mis padres por todo vuestro apoyo y amor en estos

difíciles y distantes años. Vuestro cariño es de enorme importancia. Un fuerte

abrazo.

List of abbreviations

105


A*: area resulting from the integration of the FID signal in the GC [a.u.].

A: area resulting from the integration of the TCD signal in the GC [a.u.].

a: relative activity.

ATCD: area resulting from integration of the TCD signal in the Micromeritics

AutoChem 2910 [a.u.].

a.u.: arbitrary units.

ASF: Anderson-Schulz-Flory.

ATR: autothermal reforming.

bbl/day: barrels per day (1 barrel = 42 US gallons ≈ 159 L).

BET: Brunauer, Emmett and Teller.

BEPC: Basin Electric Power Cooperative.

BFB: bubbling fluidized bed gasifier.

BTG: biomass-to-gas.

BTL: biomass-to-liquids.

Cc: graphitic carbon.

CFB: circulating fluidized bed gasifier.

CSTR: continuously stirred tank reactor.

CTG: coal-to-gas.

CTL: coal-to-liquids.

Cv: vermicular, whisker carbon.

Cα: atomic carbon.

Cβ: polymeric amorphous carbon.

Cϒ: metal carbide.

d(Co0)H: mean Co0 crystallite size as estimated by H2 chemisorption [nm].

d(Co0)XRD: mean Co0 crystallite size as estimated by XRD [nm].

d(Co3O4)XRD: mean Co3O4 crystallite size as estimated by XRD [nm].

d(Ni0)H: mean Ni0 crystallite size as estimated by H2 chemisorption [nm].

d(Ni0)XRD: mean Ni0 crystallite size as estimated by XRD [nm].

d(NiO)XRD: mean NiO crystallite size of as estimated by XRD [nm].

DFT: density functional theory.

DH: metal dispersion as estimated by H2 chemisorption.

DOR: degree of reduction.

dpore: pore diameter [nm].

DR: dry reforming.

EDS: energy-dispersive X-ray spectroscopy.

EF: entrained-flow gasifier.

f mass transfer: extent of mass transfer control.

f: calibration factor of the TCD signal in the Micromeritics AutoChem 2910

[moles H2/a.u.].


106

F: molar flow [moles/s].

FID: flame ionization detector.

FT: Fischer-Tropsch.

FTS: Fischer-Tropsch synthesis.

GC: gas chromatograph.

GTL: gas-to-liquids.

h: heat transfer coefficient [W/m2·°C].

HICOM: high combined shift methanation.

HTFT: high temperature Fischer-Tropsch.

ICP-MS: inductively coupled plasma mass spectrometry.

kc: mass transfer coefficient [m/s].

LFPT: light fraction process technology.

LTFT: low temperature Fischer-Tropsch.

Mmetal: molecular weight of the metal [g/mol].

MTFB: multitubular fixed bed reactor.

MTFT: middle temperature Fischer-Tropsch.

n: carbon number.

nC: number of carbon atoms.

NNPC: Nigerian National Petroleum Corporation.

o/p: olefin-to-paraffin ratio.

P&I: process and instrumentation.

P: pressure [bar].

pCO: partial pressure of CO [bar].

POX: partial oxidation.

QP: Qatar Petroleum.

S/Co*: sulfur-to-cobalt active site ratio.

SBCR: slurry bed reactor.

SC2: selectivity to hydrocarbons with 2 carbon atoms.



SC5+: selectivity to hydrocarbons with 5 or more C atoms.

SCH4: selectivity to methane.

SCO2: selectivity to CO2.

SEM: scanning electron microscopy.

SMSI: strong metal-support interaction.

SNG: synthetic natural gas.

SR: steam reforming.

T: temperature.

TCD: thermal conductivity detector.

TEM: transmission electron microscopy.

THüttig: Hüttig temperature [°C].


107

Tin: inlet temperature [°C].

Tmelting: melting point [°C].

TOF: turnover frequency [molecules converted / (active site · second)].

TPH: temperature-programmed hydrogenation.

TPR: temperature-programmed reduction.

TREMP: Topsoe Recycling Energy Efficient Methanation Process.

Ts: temperature at the catalyst surface [°C].

TTammann: Tammann temperature [°C].

WGS: water gas shift.

wn: mass fraction of hydrocarbons with ”n” carbon atoms.

XCO: CO conversion.

xmetal: mass fraction of the metal in the catalyst sample.

xn: molar fraction of hydrocarbons with “n” carbon atoms.

XPS: x-ray photoelectron spectroscopy.

XRD: x-ray diffraction.

α: chain growth probability.

δ: deactivation order.

θs: sulfur coverage.

λ: effective thermal conductivity [W/m·°C].

References

109

References

[1] IEA - International Energy Agency, Key World Energy Statistics, (2015). [2] BP - British Petroleum, BP Statistical Review of World Energy, (2015). [3] ENI - Ente Nazionale Idrocarburi, World Oil & Gas Review 2014, (2014). [4] V. Krishnamoorthy, S.V. Pisupati, A critical review of mineral matter related issues during gasification of coal in fixed, fluidized, and entrained flow gasifiers, Energies, 8 (2015) 10430-10463. [5] P. Wang, M. Massoudi, Slag behavior in gasifiers. Part I: Influence of coal properties and gasification conditions, Energies, 6 (2013) 784-806. [6] P. Basu, Biomass Gasification, Pyrolysis and Torrefaction: Practical Design and Theory, Second ed., Academic Press, Boston, 2013. [7] E4Tech, Review of Technologies for Gasification of Biomass and Wastes. NNFCC project 09/008, (2009). [8] R. Warnecke, Gasification of biomass: Comparison of fixed bed and fluidized bed gasifier, Biomass and Bioenergy, 18 (2000) 489-497. [9] N. Couto, A. Rouboa, V. Silva, E. Monteiro, K. Bouziane, Influence of the Biomass Gasification Processes on the Final Composition of Syngas, Energy Procedia, 36 (2013) 596-606. [10] A.A. Muleja, Y. Yao, D. Glasser, D. Hildebrandt, Effect of feeding nitrogen to a fixed bed Fischer-Tropsch reactor while keeping the partial pressures of reactants the same, Chemical Engineering Journal, 293 (2016) 151-160. [11] V.V. Ordomsky, A. Carvalho, B. Legras, S. Paul, M. Virginie, V.L. Sushkevich, A.Y. Khodakov, Effects of co-feeding with nitrogen-containing compounds on the performance of supported cobalt and iron catalysts in Fischer-Tropsch synthesis, Catalysis Today, (2015). [12] A.P. Savost’yanov, R.E. Yakovenko, G.B. Narochnyi, A.L. Lapidus, Effect of the dilution of synthesis gas with nitrogen on the Fischer–Tropsch process for the production of hydrocarbons, Solid Fuel Chemistry, 49 (2015) 356-359. [13] C.N. Hamelinck, A.P.C. Faaij, H. den Uil, H. Boerrigter, Production of FT transportation fuels from biomass; technical options, process analysis and optimisation, and development potential, Energy, 29 (2004) 1743-1771. [14] A. Jess, R. Popp, K. Hedden, Fischer–Tropsch-synthesis with nitrogen-rich syngas: Fundamentals and reactor design aspects, Applied Catalysis A: General, 186 (1999) 321-342. [15] A. Jess, R. Popp, K. Hedden, Production of Diesel Oil and Wax by Fischer-Tropsch-Synthesis Using a Nitrogen-Rich Synthesis Gas - Investigations on a Semi-Technical Scale, Oil Gas European Magazine, 24 (1998) 34-43. [16] E.G. Baker, L.K. Mudge, D.H. Mitchell, Oxygen/steam gasification of wood in a fixed-bed gasifier, Industrial & Engineering Chemistry Process Design and Development, 23 (1984) 725-728. [17] S. Albertazzi, F. Basile, J. Brandin, J. Einvall, C. Hulteberg, G. Fornasari, V. Rosetti, M. Sanati, F. Trifirò, A. Vaccari, The technical feasibility of biomass gasification for hydrogen production, Catalysis Today, 106 (2005) 297-300. [18] J.P. Ciferno, J.J. Marano, US DOE National Energy Technology Laboratory, Benchmarking Biomass Gasification Technologies for Fuels, Chemicals and Hydrogen Production, Pittsburgh, (2002). [19] C.H. Bartholomew, R.J. Farrauto, Fundamentals of Industrial Catalytic Processes: Second Edition, Wiley, 2010. [20] J.R. Rostrup-Nielsen, Syngas in perspective, Catalysis Today, 71 (2002) 243-247.

References

110

[21] K. Aasberg-Petersen, I. Dybkjær, C.V. Ovesen, N.C. Schjødt, J. Sehested, S.G. Thomsen, Natural gas to synthesis gas - Catalysts and catalytic processes, Journal of Natural Gas Science and Engineering, 3 (2011) 423-459. [22] J.R. Rostrup-Nielsen, New aspects of syngas production and use, Catalysis Today, 63 (2000) 159-164. [23] W.S. Ernst, S.C. Venables, P.S. Christensen, A.C. Berthelsen, Push syngas production limits, Hydrocarbon Processing, 79 (2000) 100C-100J. [24] Y.J. Lee, S.I. Hong, D.J. Moon, Studies on the steam and CO2 reforming of methane for GTL-FPSO applications, Catalysis Today, 174 (2011) 31-36. [25] J.J. Lerou, Enabling a small scale gas-to-liquid process, 11AIChE - 2011 AIChE Annual Meeting, Conference Proceedings, 2011. [26] C. Lowe, K. Gerdes, W. Stupin, B. Hook, P. Marriott, Impact of Syngas generation technology selection on a GTL FPSO, in: E. Iglesia, J.J. Spivey, T.H. Fleisch (Eds.) Studies in Surface Science and Catalysis, 136, 2001, pp. 57-62. [27] J.A. Velasco, L. Lopez, S. Cabrera, M. Boutonnet, S. Järås, Synthesis gas production for GTL applications: Thermodynamic equilibrium approach and potential for carbon formation in a catalytic partial oxidation pre-reformer, Journal of Natural Gas Science and Engineering, 20 (2014) 175-183. [28] J. Rostrup-Nielsen, L.J. Christiansen, Concepts in Syngas Manufacture, in: G.J. Hutchings (Ed.) Catalytic Science Series, 10, Imperial College Press 2011, [29] K. Aasberg-Petersen, T.S. Christensen, I. Dybkjær, J. Sehested, M. Østberg, R.M. Coertzen, M.J. Keyser, A.P. Steynberg, Synthesis gas production for FT synthesis, in: A. Steynberg, M. Dry (Eds.) Fischer-Tropsch Technology, 152, 2004, pp. 258-405. [30] T. Sperle, D. Chen, R. Lødeng, A. Holmen, Pre-reforming of natural gas on a Ni catalyst. Criteria for carbon free operation, Applied Catalysis A: General, 282 (2005) 195-204. [31] M.K. Gnanamani, W.D. Shafer, D.E. Sparks, B.H. Davis, Fischer–Tropsch synthesis: Effect of CO2 containing syngas over Pt promoted Co/γ-Al2O3 and K-promoted Fe catalysts, Catalysis Communications, 12 (2011) 936-939. [32] Haldor Topsoe A/S, SSK-10. Available at: http://www.topsoe.com/products/ssk-10. Accessed on 28th of April 2016. [33] V.M.H. van Wechem, M.M.G. Senden, Conversion of Natural Gas to Transportation Fuels Via the Shell Middle Distillate Synthesis Process (SMDS), in: H.E. Curry-Hyde, R.F. Howe (Eds.) Natural Gas Conversion II, Volume 81, Elsevier, 1994, pp. 43-71. [34] D.B. Bukur, X. Lang, Highly active and stable iron Fischer-Tropsch catalyst for synthesis gas conversion to liquid fuels, Industrial and Engineering Chemistry Research, 38 (1999) 3270-3275. [35] M.E. Dry, The Sasol Fischer-Tropsch Processes, in: B.E. Leach (Ed.) Applied Industrial Catalysis, 2, Academic Press, 1983, pp. 167-213. [36] F.G. Botes, J.W. Niemantsverdriet, J. van de Loosdrecht, A comparison of cobalt and iron based slurry phase Fischer–Tropsch synthesis, Catalysis Today, 215 (2013) 112-120. [37] M. Rudloff, The CHOREN BTL process status – progress – future prospects, (2008). [38] E. Rytter, E. Ochoa-Fernández, A. Fahmi, Biomass-to-Liquids by the Fischer–Tropsch Process, Catalytic Process Development for Renewable Materials, Wiley-VCH Verlag GmbH & Co. KGaA, 2013, pp. 265-308. [39] H. Weiss, Rectisol wash for purification of partial oxidation gases, Gas Separation and Purification, 2 (1988) 171-176. [40] E.I. Koytsoumpa, K. Atsonios, K.D. Panopoulos, S. Karellas, E. Kakaras, J. Karl, Modelling and assessment of acid gas removal processes in coal-derived SNG production, Applied Thermal Engineering, 74 (2015) 128-135. [41] A.A. Olajire, CO2 capture and separation technologies for end-of-pipe applications - A review, Energy, 35 (2010) 2610-2628.

http://www.topsoe.com/products/ssk-10

References

111

[42] P. Chiesa, S. Consonni, T. Kreutz, W. Robert, Co-production of hydrogen, electricity and CO2 from coal with commercially ready technology. Part A: Performance and emissions, International Journal of Hydrogen Energy, 30 (2005) 747-767. [43] A.M. Ribeiro, J.C. Santos, A.E. Rodrigues, Pressure swing adsorption for CO2 capture in Fischer-Tropsch fuels production from biomass, Adsorption, 17 (2011) 443-452. [44] T.C. Drage, J.M. Blackman, C. Pevida, C.E. Snape, Evaluation of activated carbon adsorbents for CO2 capture in gasification, Energy and Fuels, 23 (2009) 2790-2796. [45] N. Abdoulmoumine, S. Adhikari, A. Kulkarni, S. Chattanathan, A review on biomass gasification syngas cleanup, Applied Energy, 155 (2015) 294-307. [46] P.J. Woolcock, R.C. Brown, A review of cleaning technologies for biomass-derived syngas, Biomass and Bioenergy, 52 (2013) 54-84. [47] D.A. Bell, B.F. Towler, M. Fan, Chapter 6 - Sulfur Recovery, in: D.A. Bell, B.F. Towler, M. Fan (Eds.) Coal Gasification and Its Applications, William Andrew Publishing, Boston, 2011, pp. 113-136. [48] J. Skrzypski, I. Bezverkhyy, O. Heintz, J.P. Bellat, Low temperature H2S removal with metal-doped nanostructure ZnO sorbents: Study of the origin of enhanced reactivity in Cu-containing materials, Industrial and Engineering Chemistry Research, 50 (2011) 5714-5722. [49] Patent: US 6284807 B1, S.C. Leviness, C.J. Mart, W.C. Behrmann, S.J. Hsia, D.R. Neskora, Slurry hydrocarbon synthesis process with increased catalyst life, 2001. [50] S. Cheah, D.L. Carpenter, K.A. Magrini-Bair, Review of mid- to high-temperature sulfur sorbents for desulfurization of biomass- and coal-derived syngas, Energy and Fuels, 23 (2009) 5291-5307. [51] C. Li, K. Suzuki, Tar property, analysis, reforming mechanism and model for biomass gasification - An overview, Renewable and Sustainable Energy Reviews, 13 (2009) 594-604. [52] D. Sutton, B. Kelleher, J.R.H. Ross, Review of literature on catalysts for biomass gasification, Fuel Processing Technology, 73 (2001) 155-173. [53] M.M. Yung, W.S. Jablonski, K.A. Magrini-Bair, Review of catalytic conditioning of biomass-derived syngas, Energy and Fuels, 23 (2009) 1874-1887. [54] M.E. Dry, J.C. Hoogendoorn, Technology of the Fischer-Tropsch Process, Catalysis Reviews, 23 (1981) 265-278. [55] A.K. Dalai, B.H. Davis, Fischer-Tropsch synthesis: A review of water effects on the performances of unsupported and supported Co catalysts, Applied Catalysis A: General, 348 (2008) 1-15. [56] B.C. Enger, A. Holmen, Nickel and Fischer-Tropsch Synthesis, Catalysis Reviews, 54 (2012) 437-488. [57] E. Iglesia, S.L. Soled, R.A. Fiato, G.H. Via, Bimetallic Synergy in Cobalt Ruthenium Fischer-Tropsch Synthesis Catalysts, Journal of Catalysis, 143 (1993) 345-368. [58] Patent: EP0319625 B1, S.L. Soled, R.A. Fiato, E. Iglesia, Cobalt-ruthenium catalysts for fischer-tropsch synthesis, 1991. [59] E. Kikuchi, M. Matsumoto, T. Takahashi, A. Machino, Y. Morita, Fischer-Tropsch synthesis over titania-supported ruthenium catalysts, Applied Catalysis, 10 (1984) 251-260. [60] S.T. Sie, M.M.G. Senden, H.M.H. van Wechem, Conversion of natural gas to transportation fuels via the shell middle distillate synthesis process (SMDS), Catalysis Today, 8 (1991) 371-394. [61] R. Suárez París, L. Lopez, J. Barrientos, F. Pardo, M. Boutonnet, S. Järås, Chapter 3: Catalytic conversion of biomass-derived synthesis gas to fuels, in: J.J. Spivey, Y.-F. Han, K.M. Dooley (Eds.) Catalysis, 27, Royal Society of Chemistry, 2015, pp. 62-143. [62] M. Claeys, E. van Steen, Basic studies, in: A. Steynberg, M.E. Dry (Eds.) Studies in Surface Science and Catalysis, 152, Elsevier, 2004, pp. 601-680. [63] K.G. Anderson, J.G. Ekerdt, Study of Fischer-Tropsch synthesis over Fe SiO2: Effect of diethylamine on hydrocarbon and alcohol production, Journal of Catalysis, 95 (1985) 602-604. [64] M. Ojeda, R. Nabar, A.U. Nilekar, A. Ishikawa, M. Mavrikakis, E. Iglesia, CO activation pathways and the mechanism of Fischer-Tropsch synthesis, Journal of Catalysis, 272 (2010) 287-297.

References

112

[65] S. Shetty, R.A. van Santen, CO dissociation on Ru and Co surfaces: The initial step in the Fischer-Tropsch synthesis, Catalysis Today, 171 (2011) 168-173. [66] S. Shetty, R.A. van Santen, Hydrogen induced CO activation on open Ru and Co surfaces, Physical Chemistry Chemical Physics, 12 (2010) 6330-6332. [67] P. Biloen, W.M.H. Sachtler, Mechanism of Hydrocarbon Synthesis over Fischer-Tropsch Catalysts, Advances in Catalysis, 30 (1981) 165-216. [68] C.G. Visconti, E. Tronconi, L. Lietti, P. Forzatti, S. Rossini, R. Zennaro, Detailed kinetics of the Fischer-Tropsch synthesis on cobalt catalysts based on H-assisted CO activation, Topics in Catalysis, 54 (2011) 786-800. [69] R.B. Anderson, R.A. Friedel, H.H. Storch, Fischer-Tropsch reaction mechanism involving stepwise growth of carbon chain, The Journal of Chemical Physics, 19 (1951) 313-319. [70] G.P. van der Laan, A.A.C.M. Beenackers, Kinetics and Selectivity of the Fischer-Tropsch Synthesis: A Literature Review, Catalysis Reviews - Science and Engineering, 41 (1999) 255-318. [71] E. Iglesia, S.C. Reyes, R.J. Madon, S.L. Soled, Selectivity Control and Catalyst Design in the Fischer-Tropsch Synthesis: Sites, Pellets, and Reactors, Advances in Catalysis, 39 (1993) 221-302. [72] Z. Liu, S. Shi, Y. Li, Coal liquefaction technologies—Development in China and challenges in chemical reaction engineering, Chemical Engineering Science, 65 (2010) 12-17. [73] Y.-W. Li, Synfuels China, State Key Laboratory of Coal Conversion, Institute of Coal Chemistry, Chinese Academy of Science, Bejing, Clean Diesel Production from Coal Based Syngas via Fischer-Tropsch Synthesis: Technology Status and Demand in China, (2004). [74] J. Xu, Y. Yang, Y.W. Li, Fischer-Tropsch synthesis process development: Steps from fundamentals to industrial practices, Current Opinion in Chemical Engineering, 2 (2013) 354-362. [75] R.A. Dictor, A.T. Bell, Fischer-Tropsch synthesis over reduced and unreduced iron oxide catalysts, Journal of Catalysis, 97 (1986) 121-136. [76] S. Li, S. Krishnamoorthy, A. Li, G.D. Meitzner, E. Iglesia, Promoted iron-based catalysts for the Fischer-Tropsch synthesis: Design, synthesis, site densities, and catalytic properties, Journal of Catalysis, 206 (2002) 202-217. [77] M.D. Shroff, D.S. Kalakkad, K.E. Coulter, S.D. Kohler, M.S. Harrington, N.B. Jackson, A.G. Sault, A.K. Datye, Activation of Precipitated Iron Fischer-Tropsch Synthesis Catalysts, Journal of Catalysis, 156 (1995) 185-207. [78] A.P. Steynberg, M.E. Dry, B.H. Davis, B.B. Breman, Fischer-Tropsch reactors, in: A. Steynberg, M.E. Dry (Eds.) Studies in Surface Science and Catalysis, 152, Elsevier, 2004, pp. 64-195. [79] A. Steynberg, M. Dry, Fischer-Tropsch Technology, Elsevier Science, 2004. [80] A.P. Steynberg, R.L. Espinoza, B. Jager, A.C. Vosloo, High temperature Fischer-Tropsch synthesis in commercial practice, Applied Catalysis A: General, 186 (1999) 41-54. [81] D.A. Wood, C. Nwaoha, B.F. Towler, Gas-to-liquids (GTL): A review of an industry offering several routes for monetizing natural gas, Journal of Natural Gas Science and Engineering, 9 (2012) 196-208. [82] A.Y. Khodakov, W. Chu, P. Fongarland, Advances in the development of novel cobalt Fischer-Tropsch catalysts for synthesis of long-chain hydrocarbons and clean fuels, Chemical Reviews, 107 (2007) 1692-1744. [83] F. Rohr, O.A. Lindvåg, A. Holmen, E.A. Blekkan, Fischer-Tropsch synthesis over cobalt catalysts supported on zirconia-modified alumina, Catalysis Today, 58 (2000) 247-254. [84] G. Jacobs, J.A. Chaney, P.M. Patterson, T.K. Das, B.H. Davis, Fischer-Tropsch synthesis: Study of the promotion of Re on the reduction property of Co/Al2O3 catalysts by in situ EXAFS/XANES of Co K and Re LIII edges and XPS, Applied Catalysis A: General, 264 (2004) 203-212. [85] G. Jacobs, J.A. Chaney, P.M. Patterson, T.K. Das, J.C. Maillot, B.H. Davis, Fischer-Tropsch synthesis: Study of the promotion of Pt on the reduction property of Co/Al2O3 catalysts by in situ EXAFS of Co K and Pt LIII edges and XPS, Journal of Synchrotron Radiation, 11 (2004) 414-422.

References

113

[86] G. Jacobs, T.K. Das, Y. Zhang, J. Li, G. Racoillet, B.H. Davis, Fischer-Tropsch synthesis: Support, loading, and promoter effects on the reducibility of cobalt catalysts, Applied Catalysis A: General, 233 (2002) 263-281. [87] S.J. Park, S.M. Kim, M.H. Woo, J.W. Bae, K.W. Jun, K.S. Ha, Effects of titanium impurity on alumina surface for the activity of Co/Ti-Al2O3 Fischer-Tropsch catalyst, Applied Catalysis A: General, 419-420 (2012) 148-155. [88] B. Jongsomjit, J. Panpranot, J.G. Goodwin Jr, Effect of zirconia-modified alumina on the properties of Co/γ -Al2O3 catalysts, Journal of Catalysis, 215 (2003) 66-77. [89] J. Hong, W. Chu, P.A. Chernavskii, A.Y. Khodakov, Effects of zirconia promotion on the structure and performance of smaller and larger pore silica-supported cobalt catalysts for Fischer-Tropsch synthesis, Applied Catalysis A: General, 382 (2010) 28-35. [90] M. Lualdi, G. Di Carlo, S. Lögdberg, S. Järås, M. Boutonnet, V. La Parola, L.F. Liotta, G.M. Ingo, A.M. Venezia, Effect of Ti and Al addition via direct synthesis to SBA-15 as support for cobalt based Fischer-Tropsch catalysts, Applied Catalysis A: General, 443-444 (2012) 76-86. [91] S. Hinchiranan, Y. Zhang, S. Nagamori, T. Vitidsant, N. Tsubaki, TiO2 promoted Co/SiO2 catalysts for Fischer-Tropsch synthesis, Fuel Processing Technology, 89 (2008) 455-459. [92] J. Xiong, Y. Ding, T. Wang, Y. Lü, H. Zhu, H. Luo, L. Lin, Effect of La2O3 promoter on reaction performance of Fischer-Tropsch synthesis over Co/AC catalyst, Chinese Journal of Catalysis, 26 (2005) 874-878. [93] K. Shimura, T. Miyazawa, T. Hanaoka, S. Hirata, Fischer-Tropsch synthesis over alumina supported cobalt catalyst: Effect of promoter addition, Applied Catalysis A: General, 494 (2015) 1-11. [94] F. Morales, E. de Smit, F.M.F. de Groot, T. Visser, B.M. Weckhuysen, Effects of manganese oxide promoter on the CO and H2 adsorption properties of titania-supported cobalt Fischer-Tropsch catalysts, Journal of Catalysis, 246 (2007) 91-99. [95] F. Morales Cano, O.L.J. Gijzeman, F.M.F. de Groot, B.M. Weckhuysen, Manganese promotion in cobalt-based Fischer-Tropsch catalysis, in: X. Bao, Y. Xu (Eds.) Studies in Surface Science and Catalysis, 147, 2004, pp. 271-276. [96] F. Morales, F.M.F. de Groot, O.L.J. Gijzeman, A. Mens, O. Stephan, B.M. Weckhuysen, Mn promotion effects in Co/TiO2 Fischer-Tropsch catalysts as investigated by XPS and STEM-EELS, Journal of Catalysis, 230 (2005) 301-308. [97] Patent: US7365040 B2, P.J. van Berge, S. Barradas, Catalysts, 2008. [98] Patent: WO2007068731, R.J. Dogterom, C.M.A.M. Mesters, M.J. Reynhout, Process for preparing a hydrocarbon synthesis catalyst, 2007. [99] Patent: WO2005072866 A1, E. Rytter, T.H. Skagseth, H. Wigum, N. Sincadu, Fischer-Tropsch catalysts, 2005. [100] Patent: WO2007009680, G. Bellussi, L.C. Carluccio, R. Gennaro, P.G. Del, Process for the preparation of fischer-tropsch catalysts with a high mechanical, thermal and chemical stability, 2007. [101] Patent: WO2005099897, M. Ikeda, T. Waku, N. Aoki, Catalyst for fisher-tropsch synthesis and method for producing hydrocarbon, 2005. [102] Patent: WO2005058493, H.J. Robota, S. Goodman, Modified catalyst supports, 2005. [103] Patent: WO2005071044, J. Inga, P. Kennedy, S. Leviness, Fischer-tropsch process in the presence of nitrogen contaminants, 2005. [104] J.J.H.M. Font Freide, J.P. Collins, B. Nay, C. Sharp, A history of the BP Fischer-Tropsch catalyst from laboratory to full scale demonstration in Alaska and beyond, in: B.H. Davis, M.L. Occelli (Eds.) Studies in Surface Science and Catalysis, 163, 2007, pp. 37-44. [105] J.J. Font Freide, B. Nay, C. Sharp, Determination of the partial pressure of water during the activation of the BP GTL Fischer-Tropsch catalyst and its effects on catalyst performance, in: B.H. Davis, M.L. Occelli (Eds.) Studies in Surface Science and Catalysis, 163, 2007, pp. 273-287.

References

114

[106] Patent: US5268344, L.E. Pedrick, C.H. Mauldin, W.C. Behrmann, Draft tube for catalyst rejuvenation and distribution, 1993. [107] Patent: US6822008B2, N. Srinivasan, R.L. Espinoza, K.L. Coy, K. Jothimurugesan, Fischer-Tropsch catalysts using multiple precursors, 2004. [108] Patent: EP1567616 B1, M.J. Bowe, C.D. Lee-Tuffnell, Catalytic reactor and process, 2006. [109] Patent: CA2826734 A1, F. Daly, L. Richard, S. Rugmini, Catalysts, 2012. [110] J. van de Loosdrecht, F.G. Botes, I.M. Ciobica, A. Ferreira, P. Gibson, D.J. Moodley, A.M. Saib, J.L. Visagie, C.J. Weststrate, J.W. Niemantsverdriet, 7.20 - Fischer–Tropsch Synthesis: Catalysts and Chemistry, in: J. Reedijk, K. Poeppelmeier (Eds.) Comprehensive Inorganic Chemistry II (Second Edition), Elsevier, Amsterdam, 2013, pp. 525-557. [111] E. Rytter, A. Holmen, Deactivation and Regeneration of Commercial Type Fischer-Tropsch Co-Catalysts - A Mini-Review, Catalysts, 5 (2015) 478-499. [112] G. Prieto, A. Martínez, P. Concepción, R. Moreno-Tost, Cobalt particle size effects in Fischer-Tropsch synthesis: structural and in situ spectroscopic characterisation on reverse micelle-synthesised Co/ITQ-2 model catalysts, Journal of Catalysis, 266 (2009) 129-144. [113] G.L. Bezemer, J.H. Bitter, H.P.C.E. Kuipers, H. Oosterbeek, J.E. Holewijn, X. Xu, F. Kapteijn, A.J. Van Dillen, K.P. de Jong, Cobalt particle size effects in the Fischer-Tropsch reaction studied with carbon nanofiber supported catalysts, Journal of the American Chemical Society, 128 (2006) 3956-3964. [114] W. Ma, G. Jacobs, D.E. Sparks, M.K. Gnanamani, V.R.R. Pendyala, C.H. Yen, J.L.S. Klettlinger, T.M. Tomsik, B.H. Davis, Fischer–Tropsch synthesis: Support and cobalt cluster size effects on kinetics over Co/Al2O3 and Co/SiO2 catalysts, Fuel, 90 (2011) 756-765. [115] J.P. den Breejen, P.B. Radstake, G.L. Bezemer, J.H. Bitter, V. Frøseth, A. Holmen, K.P. de Jong, On the origin of the cobalt particle size effects in Fischer-Tropsch catalysis, Journal of the American Chemical Society, 131 (2009) 7197-7203. [116] J. Patzlaff, Y. Liu, C. Graffmann, J. Gaube, Studies on product distributions of iron and cobalt catalyzed Fischer-Tropsch synthesis, Applied Catalysis A: General, 186 (1999) 109-119. [117] V. Ponec, Some Aspects of the Mechanism of Methanation and Fischer-Tropsch Synthesis, Catal Rev, 18 (1978) 151-171. [118] S. Lögdberg, M. Lualdi, S. Järås, J.C. Walmsley, E.A. Blekkan, E. Rytter, A. Holmen, On the selectivity of cobalt-based Fischer-Tropsch catalysts: Evidence for a common precursor for methane and long-chain hydrocarbons, Journal of Catalysis, 274 (2010) 84-98. [119] B.W. Wojciechowski, Kinetics of the Fischer-Tropsch synthesis, Catalysis Reviews, 30 (1988) 629-702. [120] E. Iglesia, S.C. Reyes, R.J. Madon, Transport-enhanced α-olefin readsorption pathways in Ru-catalyzed hydrocarbon synthesis, Journal of Catalysis, 129 (1991) 238-256. [121] I. Puskas, R.S. Hurlbut, Comments about the causes of deviations from the Anderson-Schulz-Flory distribution of the Fischer-Tropsch reaction products, Catalysis Today, 84 (2003) 99-109. [122] E. Iglesia, S.L. Soled, R.A. Fiato, Fischer-Tropsch synthesis on cobalt and ruthenium. Metal dispersion and support effects on reaction rate and selectivity, Journal of Catalysis, 137 (1992) 212-224. [123] F.G. Botes, Proposal of a new product characterization model for the iron-based low-temperature Fischer-Tropsch synthesis, Energy and Fuels, 21 (2007) 1379-1389. [124] B. Todic, W. Ma, G. Jacobs, B.H. Davis, D.B. Bukur, CO-insertion mechanism based kinetic model of the Fischer-Tropsch synthesis reaction over Re-promoted Co catalyst, Catalysis Today, 228 (2014) 32-39. [125] Patent: WO2001096015 A2, P.J. van Berge, J. van de Loosdrecht, J.L. Visagie, Cobalt catalysts, 2001. [126] Patent: US7524787 B2, J.L. Visagie, H.M. Veltman, Producing supported cobalt catalysts for the Fisher-Tropsch synthesis, 2009.

References

115

[127] M. Lualdi, S. Lögdberg, G. Di Carlo, S. Järås, M. Boutonnet, A.M. Venezia, E.A. Blekkan, A. Holmen, Evidence for diffusion-controlled hydrocarbon selectivities in the Fischer-Tropsch synthesis over cobalt supported on ordered mesoporous silica, Topics in Catalysis, 54 (2011) 1175-1184. [128] Ø. Borg, S. Eri, E.A. Blekkan, S. Storsæter, H. Wigum, E. Rytter, A. Holmen, Fischer-Tropsch synthesis over γ-alumina-supported cobalt catalysts: Effect of support variables, Journal of Catalysis, 248 (2007) 89-100. [129] Ø. Borg, N. Hammer, S. Eri, O.A. Lindvåg, R. Myrstad, E.A. Blekkan, M. Rønning, E. Rytter, A. Holmen, Fischer-Tropsch synthesis over un-promoted and Re-promoted γ-Al2O3 supported cobalt catalysts with different pore sizes, Catalysis Today, 142 (2009) 70-77. [130] E. Rytter, D. Schanke, S. Eri, H. Wigum, T.H. Skagseth, N. Sinkadu, Optimization of statoil's Fischer-Tropsch Co/Re/alumina catalyst, ACS National Meeting Book of Abstracts, 2004. [131] E. Iglesia, Design, synthesis, and use of cobalt-based Fischer-Tropsch synthesis catalysts, Applied Catalysis A: General, 161 (1997) 59-78. [132] E. Iglesia, S.L. Soled, J.E. Baumgartner, S.C. Reyes, Synthesis and Catalytic Properties of Eggshell Cobalt Catalysts for the Fischer-Tropsch Synthesis, Journal of Catalysis, 153 (1995) 108-122. [133] J.W.A. de Swart, R. Krishna, S.T. Sie, Selection, design and scale up of the Fischer-Tropsch reactor, in: M.d. Pontes, R.L. Espinoza, C.P. Nicolaides, J.H. Scholtz, M.S. Scurrell (Eds.) Studies in Surface Science and Catalysis, 107, Elsevier, 1997, pp. 213-218. [134] R.A. Harris, Commercializing and deploying microchannel FT reactors for smaller scale GTL facilities, AIChE Process Development SymposiumHouston (USA), 2015. [135] Velocys, Projects. Available at: http://www.velocys.com/projects_overview.php. Accessed on 15th of May 2016. [136] J. McDaniel, Concept into Practical GTL Plant for FPSOs, Emerging FPSO forum, Galveston - Texas (USA), 2013. [137] J. McDaniel, Smaller scale FT enables biomass-to-liquds, Gasification Technologies Conference 2013, Colorado Springs, Colorado, 2013. [138] S. LeViness, Velocys Fischer-Tropsch Synthesis Technology – Comparison to Conventional FT Technologies, 13th Topical Conference on Gas Utilization AIChE 2013 Spring Meeting, San Antonio - Texas, 2013. [139] Shell Global, PEARL GTL - OVERVIEW. Available at: http://www.shell.com/about-us/major-projects/pearl-gtl/pearl-gtl-an-overview.html. Accessed on 5th of May 2016. [140] Qatar Petroleum, Oryx GTL. Available at: https://www.qp.com.qa/en/QPActivities/Pages/SubsidiariesAndJointVenturesDetails.aspx?aid=37. Accessed on 5th of May 2016. [141] Chevron, Transforming natural gas into superclean fuels. Available at: https://www.chevron.com/stories/gas-to-liquids. Accessed on 5th of May 2016. [142] D.J. Moodley, On the Deactivation of Cobalt-based Fischer-Tropsch Synthesis Catalysts, 2008, Technische Universiteit Eindhoven. Doctoral thesis. [143] A. de Klerk, Can Fischer-Tropsch Syncrude be Refined to On-Specification Diesel Fuel? Energy and Fuels, 23 (2009) 4593-4604. [144] A. de Klerk, E. Furimsky, Catalysis in the Refining of Fischer-Tropsch Syncrude, Royal Society of Chemistry, 2010. [145] A. de Klerk, Fischer-Tropsch Refining, Wiley, 2012. [146] A. de Klerk, P.M. Maitlis, What Can We Do with Fischer-Tropsch Products? Greener Fischer-Tropsch Processes for Fuels and Feedstocks, Wiley, 2013, pp. 81-105. [147] A. de Klerk, Fischer-Tropsch fuels refinery design, Energy and Environmental Science, 4 (2011) 1177-1205. [148] F. Regali, Hydroconversion of model Fischer-Tropsch wax over noble metal/silica-alumina catalysts, 2013, KTH - Royal Institute of Technology. Doctoral thesis.

http://www.velocys.com/projects_overview.php

http://www.shell.com/about-us/major-projects/pearl-gtl/pearl-gtl-an-overview.html

http://www.shell.com/about-us/major-projects/pearl-gtl/pearl-gtl-an-overview.html

http://www.qp.com.qa/en/QPActivities/Pages/SubsidiariesAndJointVenturesDetails.aspx?aid=37

http://www.chevron.com/stories/gas-to-liquids

References

116

[149] R. Suárez París, M.E. L’Abbate, L.F. Liotta, V. Montes, J. Barrientos, F. Regali, A. Aho, M. Boutonnet, S. Järås, Hydroconversion of paraffinic wax over platinum and palladium catalysts supported on silica–alumina, Catalysis Today, (2015), http://dx.doi.org/10.1016/j.cattod.2015.1011.1026. [150] C. Bouchy, G. Hastoy, E. Guillon, J.A. Martens, Fischer-Tropsch waxes upgrading via hydrocracking and selective hydroisomerization, Oil and Gas Science and Technology, 64 (2009) 91-112. [151] W. Böhringer, A. Kotsiopoulos, M. de Boer, C. Knottenbelt, J.C.Q. Fletcher, Selective Fischer-Tropsch wax hydrocracking - opportunity for improvement of overall gas-to-liquids processing, in: B.H. Davis, M.L. Occelli (Eds.) Studies in Surface Science and Catalysis, 163, 2007, pp. 345-365. [152] J. Scherzer, A.J. Gruia, Hydrocracking Science and Technology, Taylor & Francis, 1996. [153] J. Gao, Y. Wang, Y. Ping, D. Hu, G. Xu, F. Gu, F. Su, A thermodynamic analysis of methanation reactions of carbon oxides for the production of synthetic natural gas, RSC Advances, 2 (2012) 2358-2368. [154] J. Kopyscinski, T.J. Schildhauer, S.M.A. Biollaz, Production of synthetic natural gas (SNG) from coal and dry biomass – A technology review from 1950 to 2009, Fuel, 89 (2010) 1763-1783. [155] GL Noble Denton and Pöyry Management Consulting, Gas Quality Harmonisation Cost Benefit Analysis, (2012). [156] V.M. Lebarbier, R.A. Dagle, L. Kovarik, K.O. Albrecht, X. Li, L. Li, C.E. Taylor, X. Bao, Y. Wang, Sorption-enhanced synthetic natural gas (SNG) production from syngas: A novel process combining CO methanation, water-gas shift, and CO2 capture, Applied Catalysis B: Environmental, 144 (2014) 223-232. [157] J.H. Jensen, J.M. Poulsen, N.U. Andersen, Haldor Topsoe A/S, From coal to clean energy, Nitrogen+Syngas, 310, (2011). [158] G.H. Watson, I.C. Research, Methanation Catalysts, IEA Coal Research, 1980. [159] A. Williams, G.A. Butler, J. Hammonds, Sintering of nickel-alumina catalysts, Journal of Catalysis, 24 (1972) 352-355. [160] J. Liu, E. Wang, J. Lv, Z. Li, B. Wang, X. Ma, S. Qin, Q. Sun, Investigation of sulfur-resistant, highly active unsupported MoS2 catalysts for synthetic natural gas production from CO methanation, Fuel Processing Technology, 110 (2013) 249-257. [161] J. Gao, Q. Liu, F. Gu, B. Liu, Z. Zhong, F. Su, Recent advances in methanation catalysts for the production of synthetic natural gas, RSC Advances, 5 (2015) 22759-22776. [162] J.R. Rostrup-Nielsen, K. Pedersen, J. Sehested, High temperature methanation. Sintering and structure sensitivity, Applied Catalysis A: General, 330 (2007) 134-138. [163] W.M. Shen, J.A. Dumesic, C.G. Hill Jr, Criteria for stable Ni particle size under methanation reaction conditions: Nickel transport and particle size growth via nickel carbonyl, Journal of Catalysis, 68 (1981) 152-165. [164] J.R. Rostrup-Nielsen, Criteria for Carbon Formation (Steam Reforming and Methanation), NATO Advanced Study Institutes Series, Series E: Applied Sciences, 54, 1982, pp. 127-149. [165] J.H. Russell, L.L. Oden, J.L. Henry, Effects of Additives on Methanation Activity of Raney Nickel catalysts, Report of Investigations - United States, Bureau of Mines, (1980). [166] L.L. Oden, J.H. Russell, Methanation activity of Raney Nickel Catalysts: Effect of Proportion of NiAl3 and Ni2Al3 in Precursor Alloys, US Bur Mines Rep Invest, (1978). [167] S. Sane, J.M. Bonnier, J.P. Damon, J. Masson, Raney Metal Catalysts: I. Comparative Properties of Raney Nickel Proceeding from Ni-Al Intermetallic Phases, Applied Catalysis, 9 (1984) 69-83. [168] C.H. Bartholomew, R.J. Farrauto, Chemistry of nickel-alumina catalysts, Journal of Catalysis, 45 (1976) 41-53. [169] G. Li, L. Hu, J.M. Hill, Comparison of reducibility and stability of alumina-supported Ni catalysts prepared by impregnation and co-precipitation, Applied Catalysis A: General, 301 (2006) 16-24.


References

117

[170] S. Hwang, J. Lee, U.G. Hong, J.G. Seo, J.C. Jung, D.J. Koh, H. Lim, C. Byun, I.K. Song, Methane production from carbon monoxide and hydrogen over nickel-alumina xerogel catalyst: Effect of nickel content, Journal of Industrial and Engineering Chemistry, 17 (2011) 154-157. [171] C.H. Bartholomew, R.B. Pannell, R.W. Fowler, Sintering of alumina-supported nickel and nickel bimetallic methanation catalysts in H2/H2O atmospheres, Journal of Catalysis, 79 (1983) 34-46. [172] B.M. Vogelaar, A.D. van Langeveld, P.J. Kooyman, C.M. Lok, R.L.C. Bonné, J.A. Moulijn, Stability of metal nanoparticles formed during reduction of alumina supported nickel and cobalt catalysts, Catalysis Today, 163 (2011) 20-26. [173] M.A. Vannice, R.L. Garten, Metal-support effects on the activity and selectivity of Ni catalysts in CO/H2 synthesis reactions, Journal of Catalysis, 56 (1979) 236-248. [174] M.A. Vannice, Titania-supported metals as CO hydrogenation catalysts, Journal of Catalysis, 74 (1982) 199-202. [175] G. Zhang, T. Sun, J. Peng, S. Wang, S. Wang, A comparison of Ni/SiC and Ni/Al2O3 catalyzed total methanation for production of synthetic natural gas, Applied Catalysis A: General, 462–463 (2013) 75-81. [176] C. Guo, Y. Wu, H. Qin, J. Zhang, CO methanation over ZrO2/Al2O3 supported Ni catalysts: A comprehensive study, Fuel Processing Technology, 124 (2014) 61-69. [177] X. Yang, X. Wang, G. Gao, Wendurima, E. Liu, Q. Shi, J. Zhang, C. Han, J. Wang, H. Lu, J. Liu, M. Tong, Nickel on a macro-mesoporous Al2O3@ZrO2 core/shell nanocomposite as a novel catalyst for CO methanation, International Journal of Hydrogen Energy, 38 (2013) 13926-13937. [178] D. Hu, J. Gao, Y. Ping, L. Jia, P. Gunawan, Z. Zhong, G. Xu, F. Gu, F. Su, Enhanced investigation of CO methanation over Ni/Al2O3 catalysts for synthetic natural gas production, Industrial and Engineering Chemistry Research, 51 (2012) 4875-4886. [179] X. Yang, W.T. Zheng, G.G. Wang, S.L. Sun, Q. Sun, Effect of MgO on catalytic performance of Ni/Al2O3 catalyst for CO methanation, Xiandai Huagong/Modern Chemical Industry, 34 (2014) 90-94. [180] M.T. Fan, K.P. Miao, J.D. Lin, H.B. Zhang, D.W. Liao, Mg-Al oxide supported Ni catalysts with enhanced stability for efficient synthetic natural gas from syngas, Applied Surface Science, 307 (2014) 682-688. [181] D.W. Goodman, R.D. Kelley, T.E. Madey, J.T. Yates Jr, Kinetics of the hydrogenation of CO over a single crystal nickel catalyst, Journal of Catalysis, 63 (1980) 226-234. [182] M.P. Andersson, F. Abild-Pedersen, I.N. Remediakis, T. Bligaard, G. Jones, J. Engbæk, O. Lytken, S. Horch, J.H. Nielsen, J. Sehested, J.R. Rostrup-Nielsen, J.K. Nørskov, I. Chorkendorff, Structure sensitivity of the methanation reaction: H2-induced CO dissociation on nickel surfaces, Journal of Catalysis, 255 (2008) 6-19. [183] B. Höhlein, H. Niessen, J. Range, H.J.R. Schiebahn, M. Vorwerk, Methane from synthesis gas and operation of high-temperature methanation, Nuclear Engineering and Design, 78 (1984) 241-250. [184] S. Rönsch, J. Schneider, S. Matthischke, M. Schlüter, M. Götz, J. Lefebvre, P. Prabhakaran, S. Bajohr, Review on methanation – From fundamentals to current projects, Fuel, 166 (2016) 276-296. [185] B. Höhlein, R. Menzer, J. Range, High temperature methanation in the long-distance nuclear energy transport system, Applied Catalysis, 1 (1981) 125-139. [186] O. Görke, P. Pfeifer, K. Schubert, Highly selective methanation by the use of a microchannel reactor, Catalysis Today, 110 (2005) 132-139. [187] J.J. Lerou, A.L. Tonkovich, L. Silva, S. Perry, J. McDaniel, Microchannel reactor architecture enables greener processes, Chemical Engineering Science, 65 (2010) 380-385. [188] M.E. Frank, M.B. Sherwin, D.B. Blum, R.L. Mednick, Liquid Phase Methanation - Shift PDU Results and Pilot Plant Status, Proceedings of the 8th Synthetic Pipeline Gas Symposium, 1976, pp. 159-182.

References

118

[189] Chem Systems US DOE, Liquid phase methanation pilot plant operation and laboratory support work, (1977). [190] J. Kopyscinski, T.J. Schildhauer, S.M.A. Biollaz, Methanation in a fluidized bed reactor with high initial CO partial pressure: Part I—Experimental investigation of hydrodynamics, mass transfer effects, and carbon deposition, Chemical Engineering Science, 66 (2011) 924-934. [191] J. Kopyscinski, T.J. Schildhauer, S.M.A. Biollaz, Methanation in a fluidized bed reactor with high initial CO partial pressure: Part II — Modeling and sensitivity study, Chemical Engineering Science, 66 (2011) 1612-1621. [192] J. Kopyscinski, M.C. Seemann, R. Moergeli, S.M.A. Biollaz, T.J. Schildhauer, Synthetic natural gas from wood: Reactions of ethylene in fluidised bed methanation, Applied Catalysis A: General, 462-463 (2013) 150-156. [193] D. Cui, J. Liu, J. Yu, F. Su, G. Xu, Attrition-resistant Ni-Mg/Al2O3 catalyst for fluidized bed syngas methanation, Catalysis Science & Technology, (2015) 3119-3129. [194] J. Liu, W. Shen, D. Cui, J. Yu, F. Su, G. Xu, Syngas methanation for substitute natural gas over Ni–Mg/Al2O3 catalyst in fixed and fluidized bed reactors, Catalysis Communications, 38 (2013) 35-39. [195] Haldor Topsoe A/S, From solid fuels to substitute natural gas (SNG) using TREMP™. Available at: https://www.netl.doe.gov/File%20Library/research/coal/energy%20systems/gasification/gasifipedia/tremp-2009.pdf. Accessed on 20th of June 2016. [196] T.T.M. Nguyen, L. Wissing, M.S. Skjøth-Rasmussen, High temperature methanation: Catalyst considerations, Catalysis Today, 215 (2013) 233-238. [197] K. Pedersen, A. Skov, J.R. Rostrup-Nielsen, Catalytic aspects of high temperature methanation, ACS Division of Fuel Chemistry, Preprints, 25 (1980) 89-100. [198] H.G. Harms, B. Höhlein, E. Jorn, A. Skov, High-temp methanation tests run, Oil and Gas Journal, 78 (1980) 120-135. [199] F.W. Moeller, H. Roberts, B. Britz, Methanation of Coal Gas for SNG, Hydrocarbon Processing, 53 (1974) 69-74. [200] Patent: GB2077613A, J.R. Rostrup-Nielsen, K. Pedersen, E. Jorn, A. Skov, A method and catalyst for exothermal catalytic gas phase reactions, 1981. [201] Haldor Topsoe A/S, GCC-2. Available at: http://www.topsoe.com/products/gcc-2. Accessed on 19th of April 2016. [202] Haldor Topsoe A/S, World’s largest SNG plant goes on-stream in China with catalysts and process technology from Haldor Topsoe A/S. Available at: http://www.topsoe.com/content/world’s-largest-sng-plant-goes-stream-china-catalysts-and-process-technology-haldor-topsøe. Accessed on 18th of May 2016. [203] C.-J. Yang, R.B. Jackson, China’s synthetic natural gas revolution, Nature Climate Change, 3 (2013) 852-854. [204] Göteborg Energi AB, GoBiGas-projektets första etapp. Available at: http://gobigas.goteborgenergi.se/Svensk_version/Anlaggningen. Accessed on 18th of May 2016. [205] Dakota Gasification Company, At a Glance. Available at: https://www.dakotagas.com/About_Us/At_A_Glance/. Accessed on 18th of May 2016. [206] C.-J. Yang, China's precarious synthetic natural gas demonstration, Energy Policy, 76 (2015) 158-160. [207] Haldor Topsoe A/S, Huineng SNG plant goes on-stream transforming coal into clean energy in China. Available at: http://www.topsoe.com/news/2014/11/huineng-sng-plant-goes-stream-transforming-coal-clean-energy-china. Accessed on 18th of May 2016. [208] R. Razzaq, H. Zhu, L. Jiang, U. Muhammad, C. Li, S. Zhang, Catalytic methanation of CO and CO2 in coke oven gas over Ni-Co/ZrO2-CeO2, Industrial and Engineering Chemistry Research, 52 (2013) 2247-2256.

http://www.netl.doe.gov/File%20Library/research/coal/energy%20systems/gasification/gasifipedia/tremp-2009.pdf

http://www.netl.doe.gov/File%20Library/research/coal/energy%20systems/gasification/gasifipedia/tremp-2009.pdf

http://www.topsoe.com/products/gcc-2

http://www.topsoe.com/content/world’s-largest-sng-plant-goes-stream-china-catalysts-and-process-technology-haldor-topsøe

http://www.topsoe.com/content/world’s-largest-sng-plant-goes-stream-china-catalysts-and-process-technology-haldor-topsøe

http://gobigas.goteborgenergi.se/Svensk_version/Anlaggningen

http://www.dakotagas.com/About_Us/At_A_Glance/

http://www.topsoe.com/news/2014/11/huineng-sng-plant-goes-stream-transforming-coal-clean-energy-china

http://www.topsoe.com/news/2014/11/huineng-sng-plant-goes-stream-transforming-coal-clean-energy-china

References

119

[209] R. Razzaq, C. Li, S. Zhang, Coke oven gas: Availability, properties, purification, and utilization in China, Fuel, 113 (2013) 287-299. [210] Z. Qin, J. Ren, M. Miao, Z. Li, J. Lin, K. Xie, The catalytic methanation of coke oven gas over Ni-Ce/Al2O3 catalysts prepared by microwave heating: Effect of amorphous NiO formation, Applied Catalysis B: Environmental, 164 (2015) 18-30. [211] J.B. Hansen, Syngas Routes to Alternative Fuels from Renewable Sources, Conference of the European Biogas Association, Karlsruhe, Germany, 2014. [212] Haldor Topsoe A/S, From coke oven gas to SNG using TREMP™ process. Available at: http://www.haldortopsoe.ru/business_areas/gasification_based/Downloads.aspx. Accessed on 20th of June 2016. [213] R.W.R. Zwart, H. Boerrigter, High efficiency co-production of synthetic natural gas (SNG) and Fischer-Tropsch (FT) transportation fuels from biomass, Energy and Fuels, 19 (2005) 591-597. [214] C.H. Bartholomew, Mechanisms of catalyst deactivation, Applied Catalysis A: General, 212 (2001) 17-60. [215] J. Zarfl, D. Ferri, T.J. Schildhauer, J. Wambach, A. Wokaun, DRIFTS study of a commercial Ni/γ-Al2O3 CO methanation catalyst, Applied Catalysis A: General, 495 (2015) 104-114. [216] J. Wilson, C. de Groot, Atomic-Scale Restructuring in High-Pressure Catalysis, The Journal of Physical Chemistry, 99 (1995) 7860-7866. [217] M. Agnelli, M. Kolb, C. Mirodatos, CO Hydrogenation on a Nickel Catalyst: 1. Kinetics and Modeling of a Low-Temperature Sintering Process, Journal of Catalysis, 148 (1994) 9-21. [218] A.A. Blanchard, The Volatile Metal Carbonyls, Chemical Reviews, 21 (1937) 3-38. [219] C. Mirodatos, H. Praliaud, M. Primet, Deactivation of nickel-based catalysts during CO methanation and disproportionation, Journal of Catalysis, 107 (1987) 275-287. [220] J.E. Spice, L.A.K. Staveley, G.A. Harrow, The heat capacity of nickel carbonyl and the thermodynamics of its formation from nickel and carbon monoxide, Journal of the Chemical Society (Resumed), (1955) 100-104. [221] L.W. Ross, F.H. Haynie, R.F. Hochman, Thermodynamic functions of nickel carbonyl and iron pentacarbonyl, Journal of Chemical and Engineering Data, 9 (1964) 339-340. [222] I.M. Kalihovich, N.F. Mikhailova, D.N. Taganov, A.Y. Kipnis, Applied Kinetics of Heterogeneous Thermal Decomposition of Nickel Carbonyl, J Appl Chem USSR, 47 (1974) 42-47. [223] B.A. Tsylov, Heterogeneous and homogeneous thermal decomposition of nickel carbonyl, Soviet Powder Metallurgy and Metal Ceramics, 10 (1971) 362-366. [224] P. Munnik, M.E.Z. Velthoen, P.E. de Jongh, K.P. de Jong, C.J. Gommes, Nanoparticle growth in supported nickel catalysts during methanation reaction - Larger is better, Angewandte Chemie - International Edition, 53 (2014) 9493-9497. [225] W.M. Goldberger, D.F. Othmer, The kinetics of nickel carbonyl formation, I&EC Process Design and Development, 2 (1963) 202-209. [226] N.E. Tsakoumis, M. Rønning, Ø. Borg, E. Rytter, A. Holmen, Deactivation of cobalt based Fischer–Tropsch catalysts: A review, Catalysis Today, 154 (2010) 162-182. [227] A.M. Saib, D.J. Moodley, I.M. Ciobîcă, M.M. Hauman, B.H. Sigwebela, C.J. Weststrate, J.W. Niemantsverdriet, J. van de Loosdrecht, Fundamental understanding of deactivation and regeneration of cobalt Fischer-Tropsch synthesis catalysts, Catalysis Today, 154 (2010) 271-282. [228] D. Kistamurthy, A.M. Saib, D.J. Moodley, J.W. Niemantsverdriet, C.J. Weststrate, Ostwald ripening on a planar Co/SiO2 catalyst exposed to model Fischer-Tropsch synthesis conditions, Journal of Catalysis, 328 (2015) 123-129. [229] M. Claeys, M.E. Dry, E. van Steen, P.J. van Berge, S. Booyens, R. Crous, P. van Helden, J. Labuschagne, D.J. Moodley, A.M. Saib, Impact of process conditions on the sintering behavior of an alumina-supported cobalt Fischer-Tropsch catalyst studied with an in situ magnetometer, ACS Catalysis, 5 (2015) 841-852.

http://www.haldortopsoe.ru/business_areas/gasification_based/Downloads.aspx

References

120

[230] J.A. Moulijn, A.E. van Diepen, F. Kapteijn, Catalyst deactivation: is it predictable?: What to do? Applied Catalysis A: General, 212 (2001) 3-16. [231] J. Sehested, Four challenges for nickel steam-reforming catalysts, Catalysis Today, 111 (2006) 103-110. [232] G. Eriksson, A.D. Pelton, Critical evaluation and optimization of the thermodynamic properties and phase diagrams of the MnO-TiO2, MgO-TiO2, FeO-TiO2, Ti2O3-TiO2, Na2O-TiO2, and K2O-TiO2 systems, Metallurgical Transactions B, 24 (1993) 795-805. [233] J. Gao, C. Jia, J. Li, M. Zhang, F. Gu, G. Xu, Z. Zhong, F. Su, Ni/Al2O3 catalysts for CO methanation: Effect of Al2O3 supports calcined at different temperatures, Journal of Energy Chemistry, 22 (2013) 919-927. [234] A. Tavasoli, R.M. Malek Abbaslou, A.K. Dalai, Deactivation behavior of ruthenium promoted Co/γ-Al2O3 catalysts in Fischer-Tropsch synthesis, Applied Catalysis A: General, 346 (2008) 58-64. [235] G. Jacobs, P.M. Patterson, Y. Zhang, T. Das, J. Li, B.H. Davis, Fischer-Tropsch synthesis: Deactivation of noble metal-promoted Co/Al2O3 catalysts, Applied Catalysis A: General, 233 (2002) 215-226. [236] C.J. Bertole, C.A. Mims, G. Kiss, The effect of water on the cobalt-catalyzed Fischer-Tropsch synthesis, Journal of Catalysis, 210 (2002) 84-96. [237] G. Kiss, C.E. Kliewer, G.J. DeMartin, C.C. Culross, J.E. Baumgartner, Hydrothermal deactivation of silica-supported cobalt catalysts in Fischer-Tropsch synthesis, Journal of Catalysis, 217 (2003) 127-140. [238] G.Z. Bian, N. Fujishita, T. Mochizuki, W.S. Ning, M. Yamada, Investigations on the structural changes of two Co/SiO2 catalysts by performing Fischer-Tropsch synthesis, Applied Catalysis A: General, 252 (2003) 251-260. [239] W. Zhou, J.G. Chen, K.G. Fang, Y.H. Sun, The deactivation of Co/SiO2 catalyst for Fischer-Tropsch synthesis at different ratios of H2 to CO, Fuel Processing Technology, 87 (2006) 609-616. [240] L. Shi, J. Chen, K. Fang, Y. Sun, Effect of water on the deactivation of coprecipitated Co-ZrO2 catalyst for fischer-tropsch synthesis, in: F.B. Noronha, M. Schmal, E. Falabella Sousa-Aguiar (Eds.) Studies in Surface Science and Catalysis, 167, 2007, pp. 97-102. [241] G.W. Huber, C.G. Guymon, T.L. Conrad, B.C. Stephenson, C.H. Bartholomew, Hydrothermal stability of Co/SiO2 Fischer-Tropsch synthesis catalysts, in: J.J. Spivey, G.W. Roberts, B.H. Davis (Eds.) Studies in Surface Science and Catalysis, 139, 2001, pp. 423-430. [242] C.J. Weststrate, M.M. Hauman, D.J. Moodley, A.M. Saib, E. van Steen, J.W. Niemantsverdriet, Cobalt fischer-tropsch catalyst regeneration: The crucial role of the kirkendall effect for cobalt redispersion, Topics in Catalysis, 54 (2011) 811-816. [243] E. van Steen, M. Claeys, M.E. Dry, J. van de Loosdrecht, E.L. Viljoen, J.L. Visagie, Stability of nanocrystals: Thermodynamic analysis of oxidation and re-reduction of cobalt in water/hydrogen mixtures, Journal of Physical Chemistry B, 109 (2005) 3575-3577. [244] A.M. Hilmen, D. Schanke, K.F. Hanssen, A. Holmen, Study of the effect of water on alumina supported cobalt Fischer-Tropsch catalysts, Applied Catalysis A: General, 186 (1999) 169-188. [245] D. Schanke, A.M. Hilmen, E. Bergene, K. Kinnari, E. Rytter, E. Ådnanes, A. Holmen, Reoxidation and Deactivation of Supported Cobalt Fischer-Tropsch Catalysts, Energy and Fuels, 10 (1996) x-872. [246] A.M. Hilmen, D. Schanke, A. Holmen, Reoxidation of supported cobalt Fischer-Tropsch catalysts, in: M. Misono (Ed.) Studies in Surface Science and Catalysis, 107, 1997, pp. 237-242. [247] A.M. Saib, A. Borgna, J. van de Loosdrecht, P.J. van Berge, J.W. Niemantsverdriet, XANES study of the susceptibility of nano-sized cobalt crystallites to oxidation during realistic Fischer-Tropsch synthesis, Applied Catalysis A: General, 312 (2006) 12-19. [248] J.R. Rostrup-Nielsen, Catalytic Steam Reforming, Catalysis: Science and Technology, 5 (1984) 1-117.

References

121

[249] J. van de Loosdrecht, B. Balzhinimaev, J.A. Dalmon, J.W. Niemantsverdriet, S.V. Tsybulya, A.M. Saib, P.J. van Berge, J.L. Visagie, Cobalt Fischer-Tropsch synthesis: Deactivation by oxidation? Catalysis Today, 123 (2007) 293-302. [250] P.J. van Berge, J. van de Loosdrecht, S. Barradas, A.M. van der Kraan, Oxidation of cobalt based Fischer-Tropsch catalysts as a deactivation mechanism, Catalysis Today, 58 (2000) 321-334. [251] A. Kogelbauer, J.C. Weber, J.G. Goodwin Jr, The formation of cobalt silicates on Co/SiO2 under hydrothermal conditions, Catalysis Letters, 34 (1995) 259-267. [252] B. Jongsomjit, J. Panpranot, J.G. Goodwin Jr, Co-support compound formation in alumina-supported cobalt catalysts, Journal of Catalysis, 204 (2001) 98-109. [253] A. Sirijaruphan, A. Horváth, J.G. Goodwin Jr, R. Oukaci, Cobalt aluminate formation in alumina-supported cobalt catalysts: Effects of cobalt reduction state and water vapor, Catalysis Letters, 91 (2003) 89-94. [254] S. Helveg, J. Sehested, J.R. Rostrup-Nielsen, Whisker carbon in perspective, Catalysis Today, 178 (2011) 42-46. [255] J.R. Rostrup-Nielsen, J. Sehested, Whisker Carbon Revisited, in: J.J. Spivey, G.W. Roberts, B.H. Davis (Eds.) Studies in Surface Science and Catalysis, Volume 139, Elsevier, 2001, pp. 1-12. [256] O. Ducreux, J. Lynch, B. Rebours, M. Roy, P. Chaumette, In situ characterisation of cobalt based Fischer-Tropsch catalysts: A new approach to the active phase, in: A. Parmaliana, D. Sanfilippo, F. Frusteri, A. Vaccari, F. Arena (Eds.) Studies in Surface Science and Catalysis, 119, 1998, pp. 125-130. [257] D.J. Moodley, J. van de Loosdrecht, A.M. Saib, M.J. Overett, A.K. Datye, J.W. Niemantsverdriet, Carbon deposition as a deactivation mechanism of cobalt-based Fischer–Tropsch synthesis catalysts under realistic conditions, Applied Catalysis A: General, 354 (2009) 102-110. [258] G. Jacobs, Y. Zhang, T.K. Das, J. Li, P.M. Patterson, B.H. Davis, Deactivation of a Ru promoted Co/Al2O3 catalyst for FT synthesis, in: G.W. Roberts, B.H. Davis (Eds.) Studies in Surface Science and Catalysis, 139, 2001, pp. 415-422. [259] E. Iglesia, S.L. Soled, R.A. Fiato, G.H. Via, Dispersion, support, and bimetallic effects in Fischer-Tropsch synthesis on cobalt catalysts, in: H.E. Curry-Hyde, R.F. Howe (Eds.) Studies in Surface Science and Catalysis, 81, 1994, pp. 433-442. [260] J. Lahtinen, G.A. Somorjai, The effects of promoters in carbon monoxide hydrogenation on cobalt foil model catalysts, Journal of Molecular Catalysis A: Chemical, 130 (1998) 255-260. [261] K.F. Tan, J. Chang, A. Borgna, M. Saeys, Effect of boron promotion on the stability of cobalt Fischer–Tropsch catalysts, Journal of Catalysis, 280 (2011) 50-59. [262] J. He, Z. Liu, Y. Yoneyama, N. Nishiyama, N. Tsubaki, Multiple-functional capsule catalysts: A tailor-made confined reaction environment for the direct synthesis of middle isoparaffins from syngas, Chemistry - A European Journal, 12 (2006) 8296-8304. [263] X. Li, J. He, M. Meng, Y. Yoneyama, N. Tsubaki, One-step synthesis of H-β zeolite-enwrapped Co/Al2O3 Fischer-Tropsch catalyst with high spatial selectivity, Journal of Catalysis, 265 (2009) 26-34. [264] J.G. McCarty, H. Wise, Hydrogenation of surface carbon on alumina-supported nickel, Journal of Catalysis, 57 (1979) 406-416. [265] C.H. Bartholomew, Carbon deposition in steam reforming and methanation, Catalysis Reviews: Science and Egineering, 24 (1982) 67-112. [266] D.C. Gardner, C.H. Bartholomew, Kinetics of Carbon Deposition during Methanation of CO, Industrial & Engineering Chemistry, Product Research and Development, 20 (1981) 80-87. [267] C.H. Bartholomew, M.V. Strasburg, H.-Y. Hsieh, Effects of Support on Carbon Formation and Gasification on Nickel during Carbon Monoxide Hydrogenation, Applied Catalysis, 36 (1988) 147-162. [268] J.R. Rostrup-Nielsen, Equilibria of decomposition reactions of carbon monoxide and methane over nickel catalysts, Journal of Catalysis, 27 (1972) 343-356.

References

122

[269] C. Mirodatos, J.A. Dalmon, G.A. Martin, Nature and Role of Carbon Species during CO and CO-H2 Reactions over Ni/SiO2 Catalysts, in: S. Kaliaguine, A. Mahay (Eds.) Studies in Surface Science and Catalysis, 19, 1984, pp. 505-513. [270] J. van de Loosdrecht, A.M. van der Kraan, A.J. van Dillen, J.W. Geus, Metal-Support Interaction: Titania-Supported and Silica-Supported Nickel Catalysts, Journal of Catalysis, 170 (1997) 217-226. [271] D.R. Goodman, Handling and Using Catalysts in the Plant, in: M.V. Twigg (Ed.) Catalyst Handbook, Wolfe, London, 1989, pp. 140-188. [272] C.G. Visconti, L. Lietti, P. Forzatti, R. Zennaro, Fischer–Tropsch synthesis on sulphur poisoned Co/Al2O3 catalyst, Applied Catalysis A: General, 330 (2007) 49-56. [273] P.K. Agrawal, J.R. Katzer, W.H. Manogue, Methanation over transition metal catalysts. III. Co Al2O3 in sulfur-poisoning studies, Journal of Catalysis, 69 (1981) 327-344. [274] P.K. Agrawal, W.D. Fitzharris, J.R. Katzer, Sulfur Poisoning and Carbon Deactivation of Alumina-Supported Ni, Co, Fe and Ru Catalysts in CO Hydrogenation, in: B. Delmon, G.F. Froment (Eds.) Studies in Surface Science and Catalysis, 6, Elsevier, 1980, pp. 179-200. [275] V. Curtis, C.P. Nicolaides, N.J. Coville, D. Hildebrandt, D. Glasser, The effect of sulfur on supported cobalt Fischer–Tropsch catalysts, Catalysis Today, 49 (1999) 33-40. [276] C.H. Bartholomew, R.M. Bowman, Sulfur poisoning of cobalt and iron fischer-tropsch catalysts, Applied Catalysis, 15 (1985) 59-67. [277] Ø. Borg, N. Hammer, B.C. Enger, R. Myrstad, O.A. Lindvåg, S. Eri, T.H. Skagseth, E. Rytter, Effect of biomass-derived synthesis gas impurity elements on cobalt Fischer–Tropsch catalyst performance including in situ sulphur and nitrogen addition, Journal of Catalysis, 279 (2011) 163-173. [278] J.G. McCarty, H. Wise, Thermodynamics of sulfur chemisorption on metals. III. Iron and cobalt, The Journal of Chemical Physics, 76 (1982) 1162-1167. [279] H. Boerrigter, H. den Uil, H. Calis, Green Diesel from Biomass via Fischer-Tropsch synthesis: New Insights in Gas Cleaning and Process Design, Pyrolysis and Gasification of Biomass and Waste, Expert Meeting,Strasbourg, France, 2002. [280] E.N. Sickafus, Sulfur and carbon on the (110) surface of nickel, Surface Science, 19 (1970) 181-197. [281] W.D. Fitzharris, J.R. Katzer, W.H. Manogue, Sulfur deactivation of nickel methanation catalysts, Journal of Catalysis, 76 (1982) 369-384. [282] J.R. Rostrup-Nielsen, K. Pedersen, Sulfur poisoning of Boudouard and methanation reactions on nickel catalysts, Journal of Catalysis, 59 (1979) 395-404. [283] C.H. Bartholomew, J.R. Katzer, Sulfur Poisoning of Nickel in CO Hydrogenation, in: B. Delmon, G.F. Froment (Eds.) Studies in Surface Science and Catalysis, 6, Elsevier, 1980, pp. 375-396. [284] C.H. Bartholomew, G.D. Weatherbee, G.A. Jarvi, Sulfur poisoning of nickel methanation catalysts: I. in situ deactivation by H2S of nickel and nickel bimetallics, Journal of Catalysis, 60 (1979) 257-269. [285] E.J. Erekson, C.H. Bartholomew, Sulfur poisoning of nickel methanation catalysts: II. Effects of H2S concentration, CO and H2O partial pressures and temperature on reactivation rates, Applied Catalysis, 5 (1983) 323-336. [286] J.R. Rostrup-Nielsen, Sulfur-passivated nickel catalysts for carbon-free steam reforming of methane, Journal of Catalysis, 85 (1984) 31-43. [287] J.R. Rostrup-Nielsen, Sulfur Poisoning, in: J.L. Figueiredo (Ed.) Progress in Catalyst Deactivation: Proceedings of the NATO Advanced Study Institute on Catalyst Deactivation, Algarve, Portugal, May 18–29, 1981, Springer Netherlands, Dordrecht, 1982, pp. 209-227. [288] B. Legras, V.V. Ordomsky, C. Dujardin, M. Virginie, A.Y. Khodakov, Impact and detailed action of sulfur in syngas on methane synthesis on Ni/γ-Al2O3 catalyst, ACS Catalysis, 4 (2014) 2785-2791. [289] J.R. Rostrup-Nielsen, Some principles relating to the regeneration of sulfur-poisoned nickel catalyst, Journal of Catalysis, 21 (1971) 171-178.

References

123

[290] D. Wei, J.G. Goodwin Jr, R. Oukaci, A.H. Singleton, Attrition resistance of cobalt F-T catalysts for slurry bubble column reactor use, Applied Catalysis A: General, 210 (2001) 137-150. [291] D. Wei, Y. Zhang, J.G. Goodwin Jr, The effect of pretreatment on the attrition resistance of spray-dried alumina, Applied Catalysis A: General, 201 (2000) 129-138. [292] C. Perego, R. Bortolo, R. Zennaro, Gas to liquids technologies for natural gas reserves valorization: The Eni experience, Catalysis Today, 142 (2009) 9-16. [293] S. Barradas, E.A. Caricato, P.J. van Berge, J. van de Loosdrecht, Support modification of cobalt based slurry phase Fischer-Tropsch catalysts, in: E. Gaigneaux, D.E.D. Vos, P. Grange, P.A. Jacobs, J.A. Martens, P. Ruiz, G. Poncelet (Eds.) Studies in Surface Science and Catalysis, 143, 2002, pp. 55-65. [294] J. van De Loosdrecht, S. Barradas, E.A. Caricato, P.J. van Berge, J.L. Visagie, Support modification of cobalt based slurry phase Fischer-Tropsch catalysts, ACS Division of Fuel Chemistry, Preprints, 45 (2000) 587-591. [295] H. Schulz, Z. Nie, F. Ousmanov, Construction of the Fischer-Tropsch regime with cobalt catalysts, Catalysis Today, 71 (2002) 351-360. [296] I.M. Ciobica, R.A. Van Santen, Carbon monoxide dissociation on planar and stepped Ru(0001) surfaces, Journal of Physical Chemistry B, 107 (2003) 3808-3812. [297] H.R. Sadeghi, V.E. Henrich, SMSI in Rh/TiO2 model catalysts: Evidence for oxide migration, Journal of Catalysis, 87 (1984) 279-282. [298] G.B. Raupp, J.A. Dumesic, Effect of titania oxidation state on the chemisorptive properties of titania-supported nickel, Journal of Catalysis, 97 (1986) 85-99. [299] G. Prieto, M.I.S. De Mello, P. Concepción, R. Murciano, S.B.C. Pergher, A. Martınez, Cobalt-Catalyzed Fischer–Tropsch Synthesis: Chemical Nature of the Oxide Support as a Performance Descriptor, ACS Catalysis, 5 (2015) 3323-3335. [300] R.C. Reuel, C.H. Bartholomew, The stoichiometries of H2 and CO adsorptions on cobalt: Effects of support and preparation, Journal of Catalysis, 85 (1984) 63-77. [301] D. Tristantini, S. Lögdberg, B. Gevert, O. Borg, A. Holmen, Direct use of H2-poor bio-syngas model in Fischer-Tropsch synthesis over un-promoted and rhenium promoted alumina-supported cobalt catalysts, 2006. [302] M.P. Lemaitre, P.G. Menon, F.Delannay in: Delannay F. (ed), Characterization of Heteogeneous Catalysts, Marcel Dekker, New York, 1984. [303] D. Schanke, S. Vada, E.A. Blekkan, A.M. Hilmen, A. Hoff, A. Holmen, Study of Pt-Promoted Cobalt CO Hydrogenation Catalysts, Journal of Catalysis, 156 (1995) 85-95. [304] J.M. Thomas, W.J. Thomas, Principles and Practice of Heterogeneous Catalysis, Wiley, 1997. [305] P.B. Weisz, J.S. Hicks, The behaviour of porous catalyst particles in view of internal mass and heat diffusion effects, Chemical Engineering Science, 17 (1962) 265-275. [306] C.N. Satterfield, Mass transfer in heterogeneous catalysis, M.I.T. Press, 1970, pp. 78-128. [307] J.R.F. Guedes De Carvalho, M.A.M. Alves, Mass transfer and dispersion around active sphere buried in a packed bed, AIChE Journal, 45 (1999) 2495-2502. [308] R. Pešić, T. Kaluđerović Radoičić, N. Bošković-Vragolović, Z. Arsenijević, Ž. Grbavčić, Heat transfer between a packed bed and a larger immersed spherical particle, International Journal of Heat and Mass Transfer, 78 (2014) 130-136. [309] D.E. Mears, Tests for transport limitations in experimental catalytic reactors, Industrial and Engineering Chemistry: Process Design and Development, 10 (1971) 541-547. [310] D.E. Mears, Diagnostic criteria for heat transport limitations in fixed bed reactors, Journal of Catalysis, 20 (1971) 127-131. [311] X. Yan, Y. Liu, B. Zhao, Z. Wang, Y. Wang, C.-j. Liu, Methanation over Ni/SiO2: Effect of the catalyst preparation methodologies, International Journal of Hydrogen Energy, 38 (2013) 2283-2291.

References

124

[312] J. Zieliński, Interaction of carbon monoxide with supported nickel catalysts, Journal of Molecular Catalysis, 79 (1993) 187-198. [313] M.A. Vannice, The catalytic synthesis of hydrocarbons from H2CO mixtures over the group VIII metals: IV. The kinetic behavior of CO hydrogenation over Ni catalysts, Journal of Catalysis, 44 (1976) 152-162. [314] S.S. Pansare, J.D. Allison, An investigation of the effect of ultra-low concentrations of sulfur on a Co/γ-Al2O3 Fischer–Tropsch synthesis catalyst, Applied Catalysis A: General, 387 (2010) 224-230. [315] C.J.H. Jacobsen, E. Törnqvist, H. Topsøe, HDS, HDN and HYD activities and temperature-programmed reduction of unsupported transition metal sulfides, Catalysis Letters, 63 (1999) 179-183. [316] N.N. Madikizela-Mnqanqeni, N.J. Coville, The effect of sulfur addition during the preparation of Co/Zn/TiO2 Fischer-Tropsch catalysts, Applied Catalysis A: General, 340 (2008) 7-15. [317] A.S. Bambal, V.S. Guggilla, E.L. Kugler, T.H. Gardner, D.B. Dadyburjor, Poisoning of a silica-supported cobalt catalyst due to presence of sulfur impurities in syngas during Fischer-Tropsch synthesis: Effects of chelating agent, Industrial and Engineering Chemistry Research, 53 (2014) 5846-5857. [318] W. Ma, G. Jacobs, W.D. Shafer, V.R.R. Pendyala, Q. Xiao, Y. Hu, B.H. Davis, Effect of H2S in Syngas on the Fischer–Tropsch Synthesis Performance of a 0.5%Pt–25%Co–Al2O3 Catalyst, Catalysis Letters, (2016) 1-9.

deactivation of cobalt and nickel catalysts in fischer ...952256/fulltext01.pdf · gas and biomass)...

Documents