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Substitute Natural Gas Production with

direct Conversion of Higher Hydrocarbons

Erzeugung von Substitute Natural Gas mit direkter

Umsetzung von höheren Kohlenwasserstoffen

Der Technischen Fakultät der Universität Erlangen-Nürnberg

zur Erlangung des Grades

DOKTOR-INGENIEUR

vorgelegt von

Christoph Baumhakl

aus Graz

Als Dissertation genehmigt von der Technischen Fakultät der

Friedrich-Alexander-Universität Erlangen-Nürnberg

Tag der mündlichen Prüfung: 25.07.2014

Vorsitzende des Promotionsorgans: Prof. Dr.-Ing. habil. Marion Merklein

Gutachter: Prof. Dr.-Ing. Jürgen Karl

Prof. Dr. Wilhelm Schwieger

III

Vorwort/Acknowledgement

Die vorliegende Arbeit entstand im Zuge meiner Tätigkeit als wissenschaftlicher Mitarbeiter am

Institut für Wärmetechnik der Technischen Universität Graz und am Lehrstuhl für

Energieverfahrenstechnik der Friedrich-Alexander-Universität Erlangen-Nürnberg.

Ich möchte an dieser Stelle ganz herzlich Herrn Prof. Jürgen Karl für seine Betreuung und

Unterstützung bei dieser Arbeit bedanken. Im Besonderen aber auch für das optimale Arbeitsumfeld

und den Freiraum, der mir im Rahmen meiner Arbeit geboten wurde und maßgeblich zum Gelingen

beitrug. Besonderer Dank gilt auch Herrn Prof. Wilhelm Schwieger für die Übernahme des

Zweitgutachtens.

Großen Dank schulde ich meinem Mitstreiter im Kampf mit Vergasungsanlagen, Dr. Thomas

Kienberger, für die Einführung in die Vergasung und Methansynthese und die langen Diskussionen.

Meinen ehemaligen Kollegen, Dr. Lorenz Griendl, Dr. Martin Hauth, Bernhard Gatternig und

Dr. Andreas Schweiger danke ich ganz herzlich für ihre Hilfestellung beim Bau und Betrieb von

Versuchsanlagen.

Natürlich danke ich auch allen meinen anderen Kollegen für ihre Hilfe und offenen Ohren und das

freundschaftliche Verhältnis.

So eine Arbeit wäre ohne die Mitarbeit von Studenten im Rahmen von verschiedensten

Abschlussarbeiten nicht durchführbar. Dafür möchte ich ihnen auch ganz herzlich Danke sagen.

Großer Dank gilt natürlich meinen Eltern und meiner Familie, durch deren Motivation und

finanziellen Unterstützung meiner Ausbildung, ich erst so weit kommen konnte um diese

Doktorarbeit zu erstellen.

Dir Katrin gilt mein größter Dank; dass Du mich immer aufgeheitert und auf andere Gedanken

gebracht hast und ich durch Dich immer wieder aufs Neue zu meiner Arbeit motiviert worden bin.

Nürnberg, im April 2014 Christoph Baumhakl

A part of the research leading to these results has received funding from the European Community’s

Research Fund for Coal and Steel (RFCS) under grant agreement n° RFCR-CT-2009-00003.

Abstract

V

Abstract

This thesis gives a contribution to develop a methanation process for production of Substitute

Natural Gas (SNG) in small-scale, decentralized facilities. Smaller plant sizes require a reduction of

the plant complexity. Therefore, a reduced gas cleaning and simplified methanation process is

proposed. A reduced gas cleaning effort results in remaining of certain contaminations in the

synthesis gas. Consequently, the methanation catalyst must be able to deal with these species.

To investigate the influence of contaminations on the methanation, suitable test setups were

constructed to validate these influences experimentally. The tests were performed with artificial,

bottle-mixed synthesis gas as well as with real synthesis gas from allothermal gasification of biomass

and lignite. The gas composition for the tests with bottle-mixed syngas bases mainly on results from

gasification tests.

In a first step, bench-scale methanation tests with clean, bottle-mixed synthesis gas prove the

proposed polytropic fixed bed reactor concept for methanation. Results from long-term tests show a

full-conversion respectively yield down to 230°C without deactivation of the catalyst. Due to the

polytropic operation of the reactor, a temperature peak originates at its inlet. It is assumed that this

temperature peak provides the required heat for conversion of higher hydrocarbons.

The lab-scale tests with contaminated synthesis gas investigate the influence of typical synthesis gas

contaminations such as ethylene, tars and hydrogen sulfide. The tests confirm that higher

hydrocarbons are directly converted within methanation. Conversion tests with ethylene and tars

showed that they fully convert within the first centimeters of the reactor. Main problem thereby is

the coking of the catalyst.

Addition of higher 0.5 vol. % ethylene results in severe coking, whereas only minor coking occurred

by addition of a representative tar mixture with a concentration of 6-12 g/Nm³. The amount of

deposited carbon depends on the reactor temperatures and the water content of the syngas.

A combined conversion of ethylene and tars showed lower coking compared to conversion of

ethylene only. A further lowering as well as prevention of carbon deposition is possible by addition of

traces (< 1 ppm) of hydrogen sulfide.

In the last step, the whole SNG production process, containing gasification, gas cleaning and

methanation is demonstrated in bench-scale. The simplified gas cleaning removes sufficiently dust,

alkalis and sulfur species such as H2S and COS, but has probably weaknesses with organic sulfur.

Therefore, the measured catalyst deactivations are high, which requires further improvements.

Promising is the almost full conversion of tars during methanation with real synthesis gas.

Kurzfassung

VI

Kurzfassung

Diese Arbeit beschäftigt sich mit verschiedenen Aspekten zur Entwicklung eines Prozesses zur

Erzeugung von Substitute Natural Gas (SNG) in kleinen, dezentralen Anlagen. Um auch im kleineren

Leistungsbereich wirtschaftlich sein zu können, muss die Anlagenkomplexität reduziert werden.

Daher wird in dieser Arbeit ein Prozess mit einer reduzierten Gasreinigung und vereinfachter

Methanisierung vorgeschlagen. Durch den reduzierten Gasreinigungsaufwand verbleiben bestimmte

Verunreinigungen im Synthesegas und beeinflussen die Methanisierung.

Zur Untersuchung des Einflusses von Verunreinigungen auf die Methansynthese wurden geeignete

Versuchsanordnungen aufgebaut und die Einflüsse experimentell ermittelt. Dabei sind sowohl

künstliche, flaschengemischte Synthesegase aber auch reale Synthesegase aus der thermischen

Vergasung verwendet worden. Die Gaszusammensetzung bei den Tests mit flaschengemischten

Gasen basiert hauptsächlich auf Ergebnissen aus Vergasungstests.

In einem ersten Schritt bestätigen Tests mit sauberen, flaschengemischten Synthesegasen die

Eignung des vorgeschlagenen polytropen Festbettreaktor Konzepts für die Methanisierung.

Langzeittests zeigen eine Aktivität des gewählten Katalysators bis runter zu 230°C, wobei ein

vollständiger Umsatz ins thermodynamische Gleichgewicht möglich ist. Zeichen für eine

Deaktivierung des Katalysators waren dabei nicht erkennbar. Durch die polytrope Betriebsweise des

Reaktors bildet sich ein Temperaturpeak in der Eintrittszone des Reaktors aus. Es wird vermutet, dass

dieser Temperaturpeak genügend Wärme für den Umsatz von höheren Kohlenwasserstoffe im Zuge

der Methanisierung liefert.

Labortests mit flaschengemischten Synthesegas unter Zugabe verschiedener Verunreinigungen wie

Ethylen, Teere und Schwefelwasserstoff zeigen den Einfluss dieser Komponenten auf die

Methanisierung. Die Ergebnisse bestätigen, dass höhere Kohlenwasserstoff direkt im Zuge der

Methansynthese umgesetzt werden. Untersuchungen des Umsatzes zeigen, dass dies innerhalb der

ersten Zentimeter des Reaktors geschieht. Hauptproblem dabei ist aber die Verkokung des

Katalysators.

Die Zugabe von mehr als 0.5 vol. % Ethylen führt zu starker Verkokung, wohingegen Teere in

Konzentrationen von 6-12 g/Nm³ nur zu geringen Kohlenstoffablagerungen am Katalysator führten.

Die Menge des abgelagerten Kohlestoffs hängt von den Reaktortemperaturen aber auch dem

Wasseranteil des Synthesegases ab.

Ein kombinierter Umsatz von Ethylen und Teeren zeigte, im Vergleich zum Umsatz von Ethylen

alleine, geringere Verkokung. Diese Verkokung lässt sich weiter reduzieren, beziehungsweise

vermeiden, durch die Zugabe von geringen Mengen (< 1 ppm) Schwefelwasserstoff.

Im letzten Teil der Arbeit wurde die gesamte Prozesskette der SNG-Produktion, von der Vergasung,

über die Gasreinigung, bis zur Methanisierung, im Labormaßstab erprobt. Die vereinfachte

Gasreinigung entfernt effektiv Staub, Alkalien und Schwefelverbindungen wie H2S und COS, hat aber

wahrscheinlich Schwächen bei der Abscheidung von organischen Schwefelverbindungen. Das zeigt

sich auch in den noch recht hohen Deaktivierungsraten des Katalysators. Eine weitere Optimierung

der Entschwefelung ist deshalb erforderlich. Vielversprechend ist der fast vollständige Umsatz von

Teere auch in den Tests mit realem Synthesegas.

Content

VII

Content

1. Introduction ......................................................................................................................... 1

1.1. Motivation ............................................................................................................................... 1

1.2. Objectives ................................................................................................................................ 3

2. State-of-the-Art .................................................................................................................... 5

2.1. Reactor concepts for methanation ......................................................................................... 6

2.2. Large SNG plants and projects in operation ............................................................................ 9

2.2.1. Large-scale coal-to-SNG plants ........................................................................................ 9

2.2.2. Biomass-to-SNG projects ............................................................................................... 10

2.2.3. Future large SNG plants and projects ............................................................................ 13

3. Theoretical Background ...................................................................................................... 15

3.1. Gasification ............................................................................................................................ 15

3.1.1. Allothermal gasification with water steam ................................................................... 16

3.1.2. Tar problematic of thermal gasification ........................................................................ 18

3.1.3. Contaminations in the product gas from allothermal gasification ............................... 19

3.2. Hot gas cleaning for sulfur and chlorine removal ................................................................. 21

3.2.1. Adsorptive desulfurization with metal oxides ............................................................... 21

3.2.2. Desulfurization with activated carbons ......................................................................... 23

3.3. Methanation .......................................................................................................................... 24

3.3.1. Thermodynamics ........................................................................................................... 25

3.3.2. Reaction kinetics and mechanisms ............................................................................... 27

3.3.3. Reforming of higher hydrocarbons ............................................................................... 29

3.3.4. Theoretical and practical aspects for the reactor design .............................................. 33

4. Catalyst Deactivation and Carbon Deposition ...................................................................... 35

4.1. Deactivation mechanisms ..................................................................................................... 35

4.2. Carbon deposition ................................................................................................................. 36

4.2.1. Types of carbon deposits and reactions ........................................................................ 36

4.2.2. Thermodynamics of carbon formation ......................................................................... 39

4.2.3. Possibilities for regeneration of carbon deposits .......................................................... 43

4.2.4. Measurement methods for carbon deposition ............................................................. 44

4.3. Poisoning ............................................................................................................................... 49

Content

VIII

4.3.1. Poisoning by sulfur ........................................................................................................ 49

4.3.2. Regeneration of sulfur-poisoned catalysts .................................................................... 50

4.4. Thermal degradation ............................................................................................................. 50

4.5. Evaporation – nickel tetracarbonyl ....................................................................................... 51

5. Bench-Scale Methanation Tests with Clean Syngas – Polytropic Reactor Concept.................. 53

5.1. Experimental setup ............................................................................................................... 53

5.2. Catalysts for methanation ..................................................................................................... 55

5.3. Test procedure ...................................................................................................................... 56

5.4. Methanation tests with different catalysts ........................................................................... 57

5.4.1. Basic performance screening ........................................................................................ 57

5.4.2. Detailed catalyst screening ........................................................................................... 59

5.4.3. Long-term performance of catalysts ............................................................................. 62

5.5. Conclusion bench-scale methanation tests........................................................................... 63

6. Experimental Investigations with Bottle-Mixed Contaminated Syngas – Experimental Setup 65

6.1. Investigation focus and program ........................................................................................... 65

6.1.1. Definition of investigation parameters ......................................................................... 65

6.1.2. Test program and procedure......................................................................................... 67

6.2. Test rig assembly ................................................................................................................... 68

6.2.1. Gas mixing station with tar conditioning unit ............................................................... 68

6.2.2. Methanation reactor test rig ......................................................................................... 71

6.2.3. Gas and tar analysis and measurement techniques ..................................................... 73

7. Experimental Investigations with Bottle-Mixed Contaminated Syngas – Results ................... 81

7.1. Parameter variations with non-contaminated synthesis gas ................................................ 81

7.2. Parameter variations with aliphatic hydrocarbons – Ethylene ............................................. 83

7.2.1. Behavior of carbon on the catalyst ............................................................................... 83

7.2.2. Definition of a critical/acceptable carbon content ....................................................... 86

7.2.3. Influence on ethylene-promoted carbon deposition .................................................... 89

7.3. Parameter variations with representative tar mixtures........................................................ 91

7.4. Reduction of carbon deposition by addition of sulfur species .............................................. 97

7.5. Visual evaluation of carbon deposits .................................................................................. 100

7.6. Summary and conclusion bottle-mixed syngas tests .......................................................... 102

Content

IX

8. Bench-Scale Tests with Real Synthesis Gas Produced in Allothermal Gasification ................. 103

8.1. Investigation focus and program ......................................................................................... 103

8.2. Test rig assembly and setup ................................................................................................ 103

8.2.1. Test rig assembly ......................................................................................................... 103

8.2.2. Test setup and operating conditions ........................................................................... 107

8.3. Results ................................................................................................................................. 110

8.3.1. Gasification .................................................................................................................. 110

8.3.2. Adsorptive hot gas cleaning ........................................................................................ 113

8.3.3. Methanation ................................................................................................................ 114

9. Conclusion ......................................................................................................................... 121

10. References ......................................................................................................................... 124

Content

X

Figures

XI

Figures

Figure 1.1: General process steps for the production of SNG .............................................................. 1

Figure 1.2: Simplified flow sheet for the proposed SNG production process with hot gas cleaning ... 2

Figure 2.1: First patent for a catalytic methanation apparatus from Elworthy, 1905 .......................... 5

Figure 2.2: Different reactor concepts and processes for methanation of synthesis gas .................... 6

Figure 2.3: Simplified flow sheet of the DGC Great Plains synfuels plant ............................................ 9

Figure 2.4: Simplified flow sheet of SNG production in the FICFB gasification plant ......................... 11

Figure 2.5: Conceptual design of the HPR and idea for decentralized SNG production..................... 12

Figure 3.1: Simulation of the influence of σ on the permanent gas composition (dry basis) for allothermal gasification .................................................................................................... 17

Figure 3.2: Typical concentrations of gaseous contaminates from gasification of woody biomass and lignite with the measured contaminates from the lab-scale allothermal gasifier .... 20

Figure 3.3: H2S equilibrium desulfurization concentrations for different sorbents with standard synthesis gas composition with 100 ppm H2S .................................................................. 21

Figure 3.4: Influence of temperature on the equilibrium composition of H2/CO=3 at 1 bar. ............ 25

Figure 3.5: Influence of pressure on the equilibrium composition of an H2/CO=3 at 300°C.............. 26

Figure 3.6: Influence of temperature on the equilibrium composition of the standard synthesis gas composition used and on the chemical efficiency ..................................................... 26

Figure 3.7: Influence of the H2O content on the equilibrium composition of the standard synthesis gas composition used at 250°C and atmospheric pressure .............................. 27

Figure 3.8: Model of the Langmuir-Hinshelwood approach for the methanation reaction ............... 28

Figure 3.9: Model of a combined L-H and E-R approach for the water-gas-shift reaction................. 29

Figure 3.10: Model of the reaction mechanism for the reforming of ethane ...................................... 30

Figure 3.11: Model for hydrocracking of phenanthrene and naphthalene .......................................... 31

Figure 3.12: Reforming of benzene, toluene and naphthalene in model gas over a Ni catalyst .......... 32

Figure 4.1: Forms of carbon deposits on Ni surfaces.......................................................................... 36

Figure 4.2: Reaction paths for formation, gasification and transformation of coke and carbons ..... 37

Figure 4.3: Steps of growth of carbon filaments ................................................................................ 38

Figure 4.4: C-H-O ternary plot with phase equilibrium lines for solid carbon. ................................... 39

Figure 4.5: Rates of formation and hydrogenation of C and C species ........................................... 41

Figure 4.6: Temperature dependency of carbon deposition on Ni, 1-butene propene in hydrogen . 42

Figure 4.7: Typically observed reactor differential pressure trends resulting from coking ............... 44

Figure 4.8: SEM-photos of different carbon deposit forms ................................................................ 45

Figure 4.9: Flow sheet of the TPO setup to determine carbon deposits ............................................ 45

Figure 4.10: Results of TPO analysis of a methanation catalyst without carbon deposits ................... 47

Figure 4.11: Results of TPO analysis of a methanation catalyst with severe carbon deposits ............. 47

Figure 4.12: Displacement of the reactor temperature profile due to selective deactivation ............ 49

Figure 4.13: Equilibrium concentration for Ni(CO)4 for different CO concentrations .......................... 51

Figure 5.1: Simplified flow sheet of the bench-scale methanation test rig ........................................ 53

Figure 5.2: 3D drawing of the tube reactor and sketch with positions of thermocouples................. 54

Figure 5.3: Temperature profiles of the tested catalysts ................................................................... 57

Figures

XII

Figure 5.4: Gas composition measured at various points of the reactor compared to temperature-related equilibrium gas compositions for EVT05 ........................................ 58

Figure 5.5: H2 content in the product gas for different catalysts at varying synthesis gas H2O contents .................................................................................................................... 59

Figure 5.6: H2 content in the product gas in dependency of the reactor outlet temperature and the water content with EVT01 .......................................................................................... 59

Figure 5.7: H2 content in the product gas in dependency of the reactor outlet temperature and the water content with EVT05 .......................................................................................... 60

Figure 5.8: H2 content in the product gas in dependency of the GHSV and the water content with EVT01 ............................................................................................................................. 61

Figure 5.9: H2 content in the product gas in dependency of the GHSV and the water content with EVT05 ............................................................................................................................. 61

Figure 5.10: Reactor temperatures for a long-term test with EVT01 ................................................... 62

Figure 5.11: Gas composition for a long-term test with EVT01 ............................................................ 63

Figure 6.1: Parameters influencing methanation ............................................................................... 66

Figure 6.2: Photo of the test rig for tests with bottle-mixed, contaminated synthesis gases ............ 68

Figure 6.3: Flow sheet of the gas mixing station with tar conditioning unit ...................................... 69

Figure 6.4: User interface of the control system ................................................................................ 71

Figure 6.5: Flow sheet of the methanation reactor test rig ............................................................... 72

Figure 6.6: 3D-drawing of the reactor oven with the reactors ........................................................... 72

Figure 6.7: Flow sheet of gas analyzing unit ....................................................................................... 73

Figure 6.8: UV absorption tar measuring cell ..................................................................................... 77

Figure 6.9: Configuration for SPA sampling ........................................................................................ 79

Figure 7.1: Measured axial temperature profiles over the reactor at different reactor oven temperatures .................................................................................................................... 82

Figure 7.2: Resulting reactor temperatures in dependency of the reactor oven temperatures ........ 82

Figure 7.3: Carbon deposition on the catalyst at 300°C using different C2H4 contents ..................... 84



Figure 7.6: Temperature profiles of a test with high carbon deposition ............................................ 85

Figure 7.7: Differential pressure across the reactor during a test with high carbon deposition ....... 86

Figure 7.8: Relation of differential pressure and the amount of carbon deposited in the reactor .... 87

Figure 7.9: Photographs of catalyst samples with different amounts of deposited carbon .............. 88

Figure 7.10: Specific amounts of catalyst consumption and cost ........................................................ 89

Figure 7.11: Influence of temperature on the amount of deposited carbon ....................................... 90

Figure 7.12: Deposited carbon in dependency of the temperature and the C2H4 content .................. 91

Figure 7.13: Amount of deposited carbon in dependency of the temperature and the H2O content, syngas with standard tar concentration ........................................................................... 92

Figure 7.14: Amount of deposited carbon in dependency of the temperature and the tar concentration .................................................................................................................... 93

Figure 7.15: Tar conversion during a methanation test in dependency of the reactor temperature .. 93

Figure 7.16: Tar conversion during a methanation test with reduced catalyst filling .......................... 94

Figure 7.17: Influence of methanation conditions on the conversion of toluene ................................ 95

Figures

XIII

Figure 7.18: Amount of deposited carbon resulting from methanation with simultaneous addition of C2H4 and tars compared to separate addition .............................................................. 95

Figure 7.19: Measured catalyst degradation from poisoning with H2S for EVT05 .............................. 98

Figure 7.20: Specific catalyst consumption and related catalyst cost due to poisoning with H2S ....... 98

Figure 7.21: Influence of C2H4 and H2S on the amount of deposited carbon ....................................... 99

Figure 7.22: Influence of a C2H4, tars and H2S on the amount of deposited carbon ............................ 99

Figure 7.23: States of polymeric carbon coverage on a catalyst pellet .............................................. 100

Figure 7.24: SEM photos of polymeric carbon deposits on the catalyst resulting from C2H4 ............ 100

Figure 7.25: SEM photos of polymeric carbon filaments resulting from C2H4 .................................... 101

Figure 7.26: SEM photos of polymeric carbon layers resulting from C2H4 ......................................... 101

Figure 8.1: Photo of the bench-scale test rig for SNG production with real synthesis ..................... 104

Figure 8.2: Flow sheet of the indirectly heated, fluidized bed gasifier ............................................ 105

Figure 8.3: Flow sheet of the bench-scale hot gas cleaning and methanation unit ......................... 106

Figure 8.4: Flow sheet of the gas analysis unit for methanation and gasification tests ................... 107

Figure 8.5: Mean dry permanent gas composition from gasification of biomass and lignite .......... 110

Figure 8.6: Mean wet permanent gas composition from gasification of biomass and lignite ......... 111

Figure 8.7: Mean C2-C4 content from gasification of woody biomass and lignite ............................ 111

Figure 8.8: Mean tar concentrations from gasification of woody biomass and lignite .................... 112

Figure 8.9: Mean contaminations from gasification of woody biomass and lignite ........................ 113

Figure 8.10: Comparison of the mean contaminations resulting from gasification of lignite before and after hot gas desulfurization with ZnO .................................................................... 113

Figure 8.11: Trend of the permanent gas composition after methanation ....................................... 114

Figure 8.12: Measured tar concentration after methanation and the related tar conversion .......... 115

Figure 8.13: Trend of the differential pressure across the methanation reactor .............................. 115

Figure 8.14: Measured catalyst carbon contents at different points of the methanation reactor .... 116

Figure 8.15: SEM-photos of polymeric carbon filaments on a catalyst sample taken after longtime real gas tests ................................................................................................................... 117

Figure 8.16: SEM-photos of laminar (graphitic) carbon deposits on a catalyst sample taken after longtime real gas tests .................................................................................................... 117

Figure 8.17: SEM-photo of cracks on a catalyst tab after 200 h runtime with real synthesis gas ...... 117

Figure 8.18: Axial temperature trends in the methanation reactor for different runtimes ............... 118

Figure 8.19: Measured catalyst degradation for tests 1-5.................................................................. 119

Figure 8.20: Measured specific catalyst consumptions for tests 1-5 ................................................. 120

Figure 9.1: Influence of contaminations on the specific amount of catalyst consumption ............. 121

Figure 9.2: Influence of sulfur concentration and ethylene content on catalyst consumption ....... 122

Tables

XIV

Tables

Table 2.1: Typical gas compositions of the FICFB gasifier, operated with wood chips ...................... 10

Table 2.2: Overview of gasification systems proposed for specific SNG projects .............................. 13

Table 3.1: Standard synthesis gas composition for the methanation tests ....................................... 17

Table 3.2: Tar classes according to ECN ............................................................................................. 18

Table 3.3: Overview of commercially available impregnated activated carbons............................... 23

Table 4.1: Overview of mechanisms of catalyst deactivation ............................................................ 35

Table 4.2: Carbon species formed on Ni catalyst ............................................................................... 37

Table 4.3: TPO method for a quantitative and qualitative analysis of carbon deposits .................... 46

Table 5.1: Overview of the catalysts used for the methanation tests ............................................... 55

Table 5.2: Standard reducing procedure ............................................................................................ 56

Table 6.1: Overview of parameters for the methanation tests .......................................................... 67

Table 6.2: Constants for Antoine equations of different tar species ................................................. 70

Table 6.3: Overview of used µ-GC-modules ....................................................................................... 75

Table 6.4: Parameters for the standard µ-GC method ....................................................................... 75

Table 6.5: Overview of the most commonly used detector tubes ..................................................... 78

Table 8.1: Fuel parameters for the used lignite and biomass .......................................................... 108

Table 8.2: Operating parameters for the real gas methanation tests .............................................. 108

Nomenclature

XV

Nomenclature

Abbreviations

BTX Benzene toluene xylene

CHP Combined heat and power

CNG Compressed natural gas

DGC Dakota Gasification Company

DVGW Deutscher Verein des Gas-und Wasserfaches

ECN Energy Research Center of the Netherlands

EDX Energy -dispersive X-ray analysis

E-R Eley-Rideal

EVT Institute for Energy Process Engineering, University Erlangen-Nurenberg

FICFB Fast internal circulating fluidized bed

FID Flame ionization detector

GA Gas analyzer

GC Gas chromatograph

GHSV Gas hourly space velocity

GoBiGas Gothenburg Biomass Gasification

HGR Hot gas recycle

HPR Heatpipe-Reformer

IGCC Integrated gasification combined cycle

IGT Institute of Gas Technology (GTI)

IPA Isopropyl alcohol

LED Light emitting diode

L-H Langmuir-Hinshelwood

LHV Lower heating value

MFC Mass flow controller

NDIR Non-dispersive infrared

PAH Polycyclic aromatic hydrocarbons

PSI Paul-Scherrer-Institute

RFCS Research Fund for Coal and Steel

RME Rapeseed methyl ester

S/C Steam to carbon

SEM Scanning electron microscopy

SNG Substitute natural gas

SPA Solid phase adsorption

SPE Solid phase extraction

TGA Thermo gravimetric analysis

TOF Turnover frequency

TPO Thermo programmed oxidation

TREMP Topsøes recycle methanation process

TU Technical University

TWR Tube wall reactor

UV Ultra violet

WGS Water-gas shift

Nomenclature

XVI

Latin symbols

aCatalyst Active catalyst area

an,Catalyst Normalized active catalyst area

Aλ Absorbance [-]

b Optical path length [m]

c Molar concentration [mol/l]

CC Carbon content [mgCarbon/gCatalyst]

cCat. Specific catalyst costs [€ct/kWhSyngas]

CCat. Catalyst costs [€/kg]

dP Catalyst particle diameter

dR Inner reactor diameter

GHSV Gas hourly space velocity [h-1]

I0 Intensity of the incident light [W/m²]

I1 Intensity of the transmitted light [W/m²]

L Reactor length

LHV Lower heating value [kJ/kg] or [kJ/mol]

mCat. Catalyst mass [g]

ni Mole content

ni Molar flow rate

p Pressure [Pa]

pi Partial pressure [Pa]

PSG Synthesis gas power [kW]

T Temperature [°C]

t Time [h] tOp. Catalyst operation time [h]

VReactor Reactor volume [m³]

VStd Standard volume flow [m³/h]

xH2O,min Minimal required mass of water [kgH2O/kgFuel]

X Conversion [-]

Greek symbols

ΔHR Reaction enthalpy [kJ/mol]

ελ Molar absorption coefficient [l mol-1 m-1]

ηMeth Chemical efficiency for methanation [-]

λ Air ratio [-]

σ Excess steam ratio [-]

Introduction

1

Chapter 1

1. Introduction

1.1. Motivation

The worldwide increasing consumption of energy requires new or improved technologies for

supplying the demand. Besides security of supply and public acceptance, the influence on the

environment and world climate in particular, has become a central issue. Production of Substitute

Natural Gas (SNG) from biomass could be one way of addressing these issues. SNG has several

advantages, such as the high conversion efficiency and the possibility to use the existing gas grid for

distribution and storage. Typically, larger amounts of biomass are available in rural areas with a

relatively low population density and therefore low energy demand. Since biomass has a low energy

density, transportation over long distances is hardly economical and ecologically sustainable. Thus,

the maximum plant size for biomass applications is limited [1]. In addition, direct transportation of

biomass into areas with higher energy needs, such as urban areas, is not convenient. Small-scale,

decentralized production of SNG and feed-in into the gas grid enables indirect transportation of solid

fuels via the gas grid. Additionally, smaller plants facilitate the utilization of waste heat.

Energy production processes are highly influenced by the economies-of-scale principle [2], which

implies that increased plant power significantly reduces production costs. Since plant power for

biomass applications is limited, alternative ways of making them economically viable need to be

found. A simplified downscaling of state-of-the-art, large-scale processes is generally not possible. It

is also necessary to reduce the complexity of plants. Therefore, this thesis proposes a simplified

process for the production of SNG.

Figure 1.1: General process steps for the production of SNG

The production process of SNG typically consists of four essential steps (figure 1.1): synthesis gas

production, gas cleaning, methanation and gas conditioning. Synthesis gas is produced by gasifying a

solid fuel. In the gas cleaning step contaminates such as particles, alkalis or sulfur, are removed prior

to methanation. In the methanation step, the synthesis gas is then catalytically converted to

raw-SNG. Before feed-in into the gas grid or alternative usage, it is necessary to condition the gas,

e.g. by removing CO2 and H2O or by odorization.

Thermalgasification

Gas cleaning MethanationGas

conditioning

Fuel H2O

Heat/O2

H2O/CO2

SNG

Introduction

2

State-of-the-art large-scale plants for SNG production, such as the Dakota Gas synfuels plant [3], put

a lot of effort into removing all impurities from the synthesis gas. This is only feasible by using

cold/wet gas cleaning methods such as Rectisol scrubbing. Operators of the Dakota Gas synfuels

plant described this as the ‘bottleneck’ in SNG production as it is the largest utility consumer in the

plant [3]. The most demonstrated or planned gas cleaning technology for biomass-scale SNG plants is

a combination of tar scrubbing (cold with bio oil) and adsorptive dechlorination and desulfurization

[4]. This is quite complex too and has not reached a commercial state yet.

To achieve a reduction of plant complexity, this thesis proposes the usage of hot gas cleaning

techniques. A reduction of exergetic losses, which are due to the high temperature level of waste

heat, is also advantageous. The main steps in hot synthesis gas cleaning are particle filtration and

different catalytic and adsorptive processes. The bioliq-plant [5] demonstrates such a process. Due to

the tar-free gasification by means of an entrained flow gasifier, tars do not need to be considered in

the gas cleaning process. Other gasification systems do not have this advantage. The SNG production

concept proposed in this thesis is based on indirectly heated fluidized bed gasification. Therefore, the

presence of tars and other higher hydrocarbons is an important issue to consider. The approach of

this thesis is to not remove tars and other hydrocarbons from the synthesis gas. Thus, the

methanation catalyst must be able to deal with these components. For methanation, this thesis

proposes a partially cooled, tubular fixed bed reactor. Due to the polytropic temperature profile

inside the reactor, the temperature peak that occurs at the inlet provides the heat needed for the

conversion of the hydrocarbons.

Figure 1.2 shows a possible option for SNG production by means of hot gas cleaning. After

gasification, a hot gas filter (ceramic or sinter metal) removes particles and ash. If the filter

temperature is sufficiently low (550-350°C), alkalis will condensate on the filter cake. Sulfur and

chlorine components can be adsorbed by means of different adsorptive materials such as zinc oxide

or activated carbons. Afterwards, the still tar-loaded synthesis gas is fed into the methanation

reactor. The central issue of the proposed SNG production concept is how the methanation catalyst

performs with higher hydrocarbons that are present in the synthesis gas.

Figure 1.2: Simplified flow sheet for the proposed SNG production process with hot gas cleaning

The proposed concept was developed on the basis of allothermal biomass gasification, as used in the

Heatpipe-Reformer (HPR) [6], [7] or in the FICFB-gasifier (Güssing) [8], [9]. However, the results of

this thesis do not only apply to allothermal biomass gasification systems. By considering the

boundary conditions, they can be also transferred, or partially transferred, to other concepts of

synthesis gas production.

Gasifier

Steam

Fuel

Heat

Ash

Hot gas filter

Adsorptive gas cleaning

Methanation Gas conditioning

SNG

CO2H2O

Introduction

3

1.2. Objectives

The main objective of this thesis is to gain a better understanding of the methanation process and

the influences of higher hydrocarbons and tars on the methanation catalysts for simplified systems

with hot gas cleaning in particular. The methanation process has been investigated and used for

more than 100 years. However, previous applications, like CO removal from town gas or methanation

of synthesis gas from oxygen-blown gasifiers, had other aspects to consider than the methanation

step of the concept proposed here. Therefore, additional investigations are necessary.

Bench-scale methanation tests with clean, bottle-mixed synthesis gas

In a first step methanation tests with clean, bottle-mixed synthesis gas prove the polytropic reactor

concept for methanation. By screening of different commercial as well as experimental nickel-based

catalysts, an appropriate catalyst was chosen for detailed investigations (chapter 5).

Influence of higher hydrocarbons and syngas contamination

The major part of this work is dedicated to a detailed analysis of the influence of higher

hydrocarbons on methanation. Previous investigations by Kienberger [10] showed that a direct

conversion of hydrocarbons is possible during methanation, but is accompanied by deactivation of

the catalyst and coking. To reduce or prevent such negative effects, a more detailed understanding

of the processes is helpful. Therefore, methanation tests with bottle-mixed synthesis gas with

addition of different hydrocarbons were performed (chapter 7). The evaluation of these tests is

based on measured deactivation rates, amounts of carbon deposited as well as conversion rates. A

first test series investigates the manner of carbon deposition, and, in particular, how runtime and the

amount of deposited carbon correlate. If this correlation is known, the number of long-term tests can

be significantly reduced by substituting them with short-term (e.g. one-day) tests. From the test

results, conclusions can also be drawn about the amount of carbon that is acceptable on the catalyst.

With the information from this first test series, further tests analyze the influence of different

parameters on the amount of deposited carbon. Parameters to vary are the reaction temperature,

the concentration of hydrocarbons in the feed and the amount of water. The permanent gas

composition, derived from allothermal gasification of biomass, was kept constant for all tests. The

results enable the definition of operating limits for methanation with sufficiently low coking and

deactivation of the catalyst, and they also provide options for reducing carbon deposits.

Bench-scale demonstration of the SNG process

The whole SNG production process, consisting of gasification, gas cleaning and methanation, is

demonstrated in a bench-scale (chapter 8). These experimental validations focus on the performance

of the catalyst using real synthesis gas from allothermal gasification. The operating conditions for

these tests are set in accordance with the results of the detailed investigations made with bottle-

mixed synthesis gas. The results provide information about catalyst degradation rates, amounts of

deposited carbon as well as gas composition, including contaminates, at all points of the process.

For all tests it was necessary to develop, design and construct suitable test rigs and setups.

Additionally, methods of analyzing different deposits and contaminations, e.g. coke, sulfur,

hydrocarbons or tars had to be developed or applied.

Introduction

4

State-of-the-Art

5

Chapter 2

2. State-of-the-Art

The production of combustible gases from a solid fuel, mainly coal, has a long tradition. At the

beginning 19th century, the first gas grids were established in several European and North American

cities. The first commercial gas works started its operation in London in 1812 [11]. In the early years

town gas was mainly the product of gasified pyrolysis gases. The invention of the water-gas

generator by Carl Wilhelm Siemens in the middle of the 19th century, allowed the utilization of coke

too. At end of the 19th century, a town gas composition of around 50 vol. % hydrogen, 25 vol. %

methane, 10 vol. % carbon monoxide and various amounts of carbon dioxide, nitrogen, oxygen and

hydrocarbons had become established [12].

In 1897, Bone et al. [13] published first experiments of the formation of methane from carbon and

hydrogen. They reported that, at a temperature of around 1200°C, carbon unites directly with

hydrogen to form methane without formation of other hydrocarbons. In 1902, Sabatier and

Senderens [14] reported the first catalytic methanation. They discovered that a mixture of one part

carbon monoxide and three parts hydrogen undergoes complete conversion into methane and water

when passing reduced nickel at 250°C. The same happened with carbon dioxide and methane at

higher temperatures. Elworthy [15] identified the commercial potential of this discovery and applied

for several patents (figure 2.1) for the technical implementation of methanation.

Figure 2.1: First patent for a catalytic methanation apparatus from Elworthy, 1905 [15]

The commercial exploitation of Elworthy’s inventions failed due to a lack in demand for methane.

Feed-in into existing gas grids would have required replacing or modifying the majority of utilization

devices. Fischer and Tropsch were investigating the methanation of coal-derived synthesis gas when

State-of-the-Art

6

they discovered the formation of long-chain hydrocarbons, the basis for the Fischer-Tropsch process

[16]. Until the early 1960s several investigations into methanation were carried out; however, the

focus of both research and commercial activities remained on the liquefaction of solid fuels.

From the 1960s, many states began to switch from town gas to natural gas. Even in those days there

was already an awareness of the finiteness of oil and natural gas. Furthermore, many countries did

not want to become too dependent on other countries in terms of energy supply and began to

facilitate the utilization of domestic coal and lignite. At that time methanation moved into the focus

of research. The first energy crisis in 1973 also led to the emergence of commercial interests. The

majority of the different methanation concepts have their origins in the 1970s and early 80s.

2.1. Reactor concepts for methanation

Basically, four different concepts have been demonstrated for methanation (figure 2.2): adiabatic

fixed bed reactors, cooled reactors (isothermal reactors), three-phase methanation and fluidized bed

methanation. Kopyscinski et al. [4] and Karl et al. [12] give a detailed review on methanation

concepts and concepts for SNG production.

Figure 2.2: Different reactor concepts and processes for methanation of synthesis gas, according to [12]

3-Phasen-Wirbelschicht

Synthesegas

(CO, H2)

methanreiches Produktgas

Synthesegas/ Produktgas

(gasförmig)

Katalysator-Partikel

(fest)

ca. 70 bar

Zwischenkühlung

inertes Trägeröl

(flüssig)

Trägeröl-

Rezirkulation

Separatordruckaufgeladene Wirbelschicht

Synthesegas

(CO, H2)



(gasförmig)


(fest)

bis ca. 70 bar

Kühlung

adiabate Festbett-

katalysatoren

Synthesegas

(CO, H2)


Produktgas-rezirkulation

gestufte Edukt-zugabe

Zwischenkühlung

Zwischenkühlungca. 300-400°C

bis 600°C


Beispiel:

Synthane Hot Tube ReactorThermoöl-

Kühlung

Synthesegas

(CO, H2)


ca. 400°C

adiabate Festbettreaktoren gekühlte Reaktoren Beispiel:

Synthane Hot Tube Reactor

Beispiel:

Lurgi Methanation

Beispiel:

Thyssen Comflux-Verfahren

Beispiel:

Chem Systems Liquid Phase Methanation

Flüssigphasen-Methanierung Wirbelschicht-Methanierung

adiabatic fixed beds

syngas

recycle compressor

intercooling staged feed injection

staged feed injection

intercooling

raw-SNG

ca. 300-400°C

up to 600-700°C

Adiabatic fixed bed reactors

e.g. Lurgi methanation process


Synthesegas

(CO, H2)



(gasförmig)


(fest)

ca. 70 bar

Zwischenkühlung

inertes Trägeröl

(flüssig)

Trägeröl-

Rezirkulation


Synthesegas

(CO, H2)



(gasförmig)


(fest)

bis ca. 70 bar

Kühlung

adiabate Festbett-

katalysatoren

Synthesegas

(CO, H2)




Zwischenkühlung


bis 600°C


Beispiel:


Kühlung

Synthesegas

(CO, H2)


ca. 400°C



Beispiel:

Lurgi Methanation

Beispiel:


Beispiel:



syngas

raw-SNG

heat-transfer-fluid cooling

ca. 400°C

Cooled reactors

e.g. Synthane hot tube reactor


Synthesegas

(CO, H2)



(gasförmig)


(fest)

ca. 70 bar

Zwischenkühlung

inertes Trägeröl

(flüssig)

Trägeröl-

Rezirkulation


Synthesegas

(CO, H2)



(gasförmig)


(fest)

bis ca. 70 bar

Kühlung

adiabate Festbett-

katalysatoren

Synthesegas

(CO, H2)




Zwischenkühlung


bis 600°C


Beispiel:


Kühlung

Synthesegas

(CO, H2)


ca. 400°C



Beispiel:

Lurgi Methanation

Beispiel:


Beispiel:



syngas

raw-SNGheat exchanger

separatorreactants(gaseous)

inert fluid(liquid)

catalyst particles (solid)

three-phase fluidized bed

Three-phase methanation

e.g. Chem systems liquid phase methanation


Synthesegas

(CO, H2)



(gasförmig)


(fest)

ca. 70 bar

Zwischenkühlung

inertes Trägeröl

(flüssig)

Trägeröl-

Rezirkulation


Synthesegas

(CO, H2)



(gasförmig)


(fest)

bis ca. 70 bar

Kühlung

adiabate Festbett-

katalysatoren

Synthesegas

(CO, H2)




Zwischenkühlung


bis 600°C


Beispiel:


Kühlung

Synthesegas

(CO, H2)


ca. 400°C



Beispiel:

Lurgi Methanation

Beispiel:


Beispiel:



pressurized fluidized bedup to 70 bar

reactants (gaseous)

catalyst particles (solid)

syngas

raw-SNG

cooling

Fluidized bed methanation

e.g. Thyssengas Comflux process

State-of-the-Art

7

Since methanation is highly exothermic, the thermodynamic equilibrium demands low temperatures

and high pressures for a maximum methane yield. In the majority of processes demonstrated, nickel-

based catalysts were used. The main issue with these is the removal of the heat of reaction that is

released, with the aim of achieving the appropriate gas properties and preventing the destruction of

the catalyst.

Methanation with adiabatic fixed bed reactors

Methanation with adiabatic fixed bed reactors is the state-of-the-art concept for the production of

SNG. The common feature of all concepts with adiabatic fixed bed reactors is that they exist of two

to six reactors with intermediate cooling and recycle of the product gas or staged injection of feed.

The methanation concept with probably the highest amount of SNG output to date is the Lurgi

methanation process [17]. It consists of three reactors with intermediate cooling and a recycle of

product gas of around 70-85 % from the outlet of the second stage to the first one. A bypass of the

first reactor enables a staged feed injection of 10-60 % into the second reactor. The outlet gas from

the first reactor has typically a temperature of around 650°C and a CH4 content of 60-70 vol. %. In the

final stage the outlet gas temperature is around 290-400°C and the CH4 content between

85-95 vol. %. BASF is the exclusive supplier of the catalysts (e.g. BASF G1 85) used in the Lurgi

methanation process, which was first demonstrated in two semi-commercial pilot plants in South

Africa (for SASOL) and in the petroleum refinery Schwechat in Austria [18]. The DGC Great Plains

synfuels plant, which was the first – and for a long time only - large-scale commercial SNG plant, is

also based on the Lurgi methanation process [3].

Another important fixed bed methanation process is the Haldor Topsøe TREMP process [19], which is

quite similar to the Lurgi methanation, but tries to minimize the recycle ratio. This is possible due to

the usage of high temperature methanation catalysts (Topsøe MCR), which resist temperatures up to

700°C. The high process temperatures allow the efficient usage of waste heat by production of

superheated steam with typically 540°C at 100 bars. The TREMP process originates from the first

ADAM and EVA project [20]. The idea was to use chemically bound energy for long-distance

transportation of nuclear energy. A high temperature nuclear reactor provides the heat for the

reformation of methane (EVA). The produced syngas can be transported via pipelines to the

methanation plant (ADAM). The methanation plant re-converts the syngas to methane by release of

high temperature process heat. For a high efficiency, high process heat temperatures and therefore a

high temperature methanation process, TREMP, is required. The second-largest SNG plant in the

world, located in Yining in the province of Xinjiang and operated by the Chinese Qinghua Group, uses

the TREMP process for the production of SNG from coal [21]. In addition, several other large-scale

coal-to-SNG projects in China and Korea as well as the largest biomass-to-SNG project (GoBiGas)

intend to use the TREMP technology for methanation [21].

Besides the two most common processes, pilots of several other methanation concepts with fixed

bed reactors have been developed and demonstrated, such as the Conoco/Westfield process [22],

IGT Hygas process [23], RMP process [24], and the HICOM process [4].

State-of-the-Art

8

Methanation with cooled reactors

The main idea of cooled reactor concepts is a reduction of process complexity by reducing the

amount of reactors to, ideally, a single reactor. The challenge for cooled reactor concepts is the

removal of exothermic heat of the methanation process. In fixed bed reactors this heat removal is

limited by the high thermal resistance of the bed. Therefore, classical fixed bed concepts are not

suitable.

One approach is the usage of catalytically coated heat transfer elements, as demonstrated within the

Synthane hot-gas recycle (HGR) and tube-wall reactor (TWR) [25]. The TWR reactor consists of tubes

which are coated with Raney nickel on their inside. A heat transfer fluid removes the heat of reaction

and keeps the reaction temperature below 390°C.

Another idea was to also use the Linde isothermal reactor for methanation [4]. The Linde isothermal

reactor is a fixed bed reactor with a large number of cooling tube bundles in the catalyst bed.

However, there are no reports of its actual use for large-scale methanation.

A newer concept, which might allow nearly isothermal operation, bases on honeycomb catalysts.

Catalytically coated honeycomb carriers are easier to equip with heat transfer elements. Within the

RFCS research project ‘CO2freeSNG’, the usage of honeycomb methanation catalysts has been

investigated [26].

Three-phase methanation

Another option for isothermal methanation is the three-phase, or liquid-phase, reactor, in which a

bubble column-like reactor contains an inert heat transfer fluid and a solid catalyst. The gas passing

through the reactor fluidizes the catalyst, thus creating a three-phase fluidized bed.

Such a concept was demonstrated by Chem Systems [27] within their liquid-phase methanation/shift

process. The operating pressure was up to 70 bars. Mineral oil was used as a heat carrier and nickel

containing balls as a catalyst.

Nowadays, the usage of liquid-phase methanation is investigated for methanation of hydrogen and

carbon dioxide for power-to-gas production [28]. The advantage of the liquid-phase methanation is

the higher degree of flexibility offered during dynamical operation, as the fluid is easier to keep on

temperature.

Fluidized bed methanation

Fluidized bed methanation allows a simpler removal of exothermic heat from the reactor and

therefore a nearly isothermal operation. Furthermore, the catalyst can be replaced or partially

replaced more easily also during operation.

The Bi-Gas project [29] of Bituminous Coal Research Inc. demonstrated the methanation of coal-

derived syngas in a fluidized bed with a diameter of 150 mm and a reaction zone length of 2.4 m.

Around 3 m² of finned cooling tubes, cooled with a heat transfer fluid, remove the heat from the

reactor. The main challenge with fluidized bed methanation is the attrition of the catalyst. Since

according to reports, the number of fine particles increases during operation whereas the number of

coarse particles decreases. Attrition-stable catalysts are required for fluidized-bed methanation.

The largest demonstration plant for fluidized bed methanation was the Comflux process built by

Thyssengas [30]. The pre-commercial plant had a power of up to 20 MWSNG. One special feature of

this plant was the combination of methanation and water-gas shift within the same reactor.

The Paul-Scherrer-Institute (PSI) and partners developed a 1 MWSNG demonstration project for

methanation of biomass-derived synthesis gas at the FICFB gasification plant in Güssing

(chapter 2.2.2) [4].

State-of-the-Art

9

2.2. Large SNG plants and projects in operation

After the energy crises of the 1970s had been overcome, gas prices remained at a level with which

SNG from coal could not compete. Therefore, the majority of SNG research and demonstration

projects were stopped in the 1980s. Bucking this general trend, the DGC Great Plains synfuels plant,

which until 2013 remained the only large-scale SNG plant, started its operation in 1984.

With ever-increasing energy demand of countries like China or India, several new coal-to-SNG plants

are now being planned and constructed; the first large-scale plant in China, in the Inner Mongolia

region started to operate in 2012. In the last few years, with the favoring of renewable sources of

energy in Europe, SNG production from biomass has become an interesting option, as demonstrated

in the 1 MWSNG biomass-to-SNG/CNG project developed by PSI for the Güssing gasification plant. The

largest project currently under commissioning (Jan. 2014) is the GoBiGas project in Goteborg with a

total SNG capacity of 20 MW.

2.2.1. Large-scale coal-to-SNG plants

DGC Great Plains synfuels plant

The Dakota Gasification Company´s (DGC) Great Plains synfuels plant in North Dakota [3] was the

first large-scale commercial plant for the production of SNG from coal. Its 14 Lurgi-Mark-IV gasifiers

convert around 18,000 tons of coal per day. After gasification, around 2/3 of the gas is cooled and

removed from condensed water (figure 2.3). As the condensed process water contains valuable

by-products, separate by-product processing extracts phenol, dephenolized cresylic acid and

ammonium sulfate. The remaining 1/3 of the gas passes a shift conversion unit until it reunites with

the cooled gas stream. A Rectisol scrubber then removes contaminations and CO2 from the syngas.

Since 1999 the compressed CO2 has been fed into a pipeline and transported to oilfields for enhanced

oil recovery. Having passed the Rectisol unit, the cleaned synthesis gas is converted to SNG by a Lurgi

fixed bed methanation process. Before feed-in into the gas grid, water is removed and the dry SNG is

compressed to pipeline pressure. The average amount of SNG produced every day is around

4.33 mil. m³ with a heating value of 36.3 MJ/m³ and a corresponding SNG power of 1.82 GW [31].

To increase the income an ammonia plant was added which uses a slipstream of the synthesis gas. In

2012 the revenue from sales of SNG was $ 252.4 million and of by-and co-products at $ 295.3 million

[31]. The fact, that more than the half of the revenue comes from the selling of by-products, shows

how important the efficient usage of by-products is for large-scale plants.

Figure 2.3: Simplified flow sheet of the DGC Great Plains synfuels plant, adapted from [3]

ASUAir

Ash lock

Lurgi gasifiers

Coal lock

Coal Gas cooling

Shift conversion

Rectisol unit

Fixed bed methanation

Gravity separation

Tar oil By-product processing

Phenol, Cresol, Ammonium sulfate

SNG

CO2 for enhanced oil recoveryN2, Xe,

KrAsh

Steam

Naphta

Condenser

Water

Syngas to ammonia plant

O2

State-of-the-Art

10

Datang Keshiketeng (Hexigten) SNG plant

China`s first SNG plant, the Datang Keshiketeng SNG plant [32], started its operation in September

2012. It is located in Keshiketeng (Hexigten) in the Inner Mongolia region. The first phase has a plant

capacity of 1.33 bn. Nm³ SNG per year (around 1.4-1.6 GWSNG). It is planned to add two plants of the

same size in phases two and three after successful operation of phase one. The gasification

technology is provided by SEDIN (Second Design Institute of Chemical Industry) [33]. The plant uses

an SNG process (fixed bed) from Davy Process Technology with purification and methanation

catalysts from Johnson Matthey.

Yining SNG plant

The second large-scale SNG plant in China, operated by the Chinese Qinghua (Kingho) Group and

located in Yining/Yili in the province of Xinjiang, started its operation in 2013 [21]. Its planned SNG

output of the first phase is around 1.4 bn. Nm³ per year, which is equivalent to an SNG power of

1.5-1.7 GW, depending on the heating value of the gas. The final phase will have an SNG output of

around 5.5 bn. Nm³ per year. The majority of the produced SNG will be fed into a pipeline and

transported to the more densely populated eastern part of China. The plant uses the Haldor Topsøe

TREMP technology for methanation of the coal-derived synthesis gas. SEDIN provides the sixteen

gasifiers [33].

2.2.2. Biomass-to-SNG projects

Güssing/PSI

The fast internally circulating fluidized bed (FICFB) gasifier [34] installed in Güssing, Austria, which

was developed at Vienna University of Technology and constructed by Repotec, has operated

commercially since 2002. The main purpose of the Güssing gasifier is combined heat and power

generation (CHP). The total fuel capacity is 8 MW (wood chips) and the electrical output around

2 MW. The FICFB gasifier consists of two zones, a gasification zone and a combustion zone. In the

combustion zone, bed material (olivine) is heated up by combustion of biomass with air. The hot bed

material circulates to the gasification zone, where it provides the heat for the endothermic

gasification of biomass by means of water steam. This allows the production of synthesis/producer

gas, which contains next-to no nitrogen. Table 2.1 shows the typical gas compositions for the Güssing

gasification plant.

Table 2.1: Typical gas compositions of the FICFB gasifier, operated with wood chips, according to [35]

Permanent gases

H2 CO CO2 CH4 N2 35-45 vol. % 19-23 vol. % 20-25 vol. % 9-11 vol. % ≈1 vol. %

Higher hydrocarbons

C2H4 C2H6 C3H8 BTX Tars 2-3 vol. % ≈0.5 vol. % ≈0.5 vol. % 10 g/Nm³ 1-5 g/Nm³

Contaminations

H2S COS org. S HCl NH3 ≈150 ppm ≈5 ppm ≈37 ppm ≈3 ppm 500-1500 ppm

State-of-the-Art

11

The favorable gas compositions and the high availability were the reasons for demonstrating SNG

production from biomass on a slipstream of the FICFB gasifier. After bench-scale tests, the Paul-

Scherrer-Institute (PSI), in cooperation with Conzepte Technik Umwelt AG (CTU) and Repotec,

developed and constructed the whole process chain for a 1 MWSNG demonstration [36].

After gasification, particles are removed and the gas passes an RME scrubber (figure 2.4). The

majority of the gas goes to a gas engine. A slipstream is further processed for the SNG production.

Before methanation, the gas is compressed and cleaned of H2S. The methanation takes place in a

fluidized bed reactor, adapted from the Comflux process [4]. To meet the specifications for further

applications, NH3, H2O, CO2 and H2 are removed step by step from the raw-SNG.

Figure 2.4: Simplified flow sheet of SNG production in the FICFB gasification plant, adapted from [4]

The main challenge for methanation is the high ethylene concentration in the synthesis gas and the

resulting coking of the catalyst. Kopyscinski [37], [38] reported that the advantage of fluidized bed

methanation is the internal regeneration of the catalyst. Measurements showed strong carbon

exchange processes between gas phase and carbon species on the catalyst surface. The gas

compositions change over the height of the fluidized bed. Therefore the atmosphere of the upper

part of the fluidized bed allows a removal of deposited carbon from the catalyst.

GoBiGas

The largest European biomass-to-SNG project under commissioning (Jan. 2014) is the Gothenburg

Biomass Gasification (GoBiGas) project [39]. It will demonstrate the commercial production of SNG

from biomass with an SNG power of 20 MW. The gasifier used in Gothenburg is based on the FICFB

technology and is being constructed by Metso under a Repotec license. As in the plant in Güssing, the

first gas cleaning steps are the removal of dust trough textile filters and the removal of tars and other

solvable pollutants in an oil scrubber. Before methanation, the gas is compressed and passes a sulfur

removal unit, a water-gas-shift unit and a CO2 removal unit. For methanation, the Haldor Topsøe

TREMP process was chosen. After cooling and drying, the produced SNG is fed into the Swedish gas

grid.

The results of the first GoBiGas project should constitute the basis for a 100 MW follow-up project.

Ash

Filter Fluidized bed methanation

FICFBgasifier

RME

RME scrubber

Bulk H2S removal

Pre-heating

Biomass

Air

Steam Tars/tiophene

Fine H2S removalFlue

gas

NH3/H2O removal

CO2/H2 removal

SNG

CO2

H2Water

To gas engine

State-of-the-Art

12

Heatpipe-Reformer

The Heatpipe-Reformer (HPR) [6], which was developed at the University of Technology Munich

consists of two separated fluidized beds, one for combustion and one for gasification (figure 2.5 a).

The heat required for endothermic gasification is transported from the combustion chamber to the

reformer by means of heat pipes. These are closed tubes, filled with a small amount of sodium or

potassium. Due to the evaporation and condensation of the fluid, high heat fluxes can be achieved.

The reformer is pressurized during operation and uses water steam for fluidization and gasification.

Two prototypes were erected and tested under the European ‘Biomass Heat Pipe Reformer’ project

[40]. Pilot plants with a thermal input of 500 kW were developed by Agnion Energy Inc. in

Pfaffenhofen [7] and by HS Energieanlagen GmbH in Freising (both in Germany). Agnion erected the

first commercial plant with a thermal input of 1.3 MW in Grassau, Germany in 2012. In it, the syngas

produced fuels a gas engine for CHP.

As the gas quality achieved there is similar to that of the Güssing gasifier, the HPR is ideal for

synthesis applications. One of the possible applications of the HPR is the decentralized production of

SNG from biomass [41] (figure 2.5 b), the idea being that of using small-scale units to generate SNG

in rural areas, close to where the biomass resources are. The existing gas infrastructure allows the

transporting of the bio-SNG to areas with a higher demand. The waste heat from the process can be

used in local heat grids. However, small-scale units require lower complexities, as an economical

operation using the demonstrated state-of-the-art process chains for SNG production is not possible.

The European research project ‘CO2freeSNG’ investigated an upscale of the HPR technology for the

production of SNG from coal and lignite [26]. Besides an experimental validation, concept studies for

SNG production in a 50 MW range were developed.

Figure 2.5: a. Conceptual design of the HPR, b. idea for decentralized SNG production, according to [41]

FuelSteamSyngas

Flue gas

Combustion chamber

Reformer

Heat pipes

State-of-the-Art

13

2.2.3. Future large SNG plants and projects

In the last few years, plans for constructing a number of large-scale coal-to-SNG plants have been

announced. SNG from coal is attractive for countries with substantial domestic coal resources but

little natural gas. New projects are therefore mainly being considered in countries that are highly

dependent on imports of natural gas, such as China and South Korea. With the increasing gas prices

from 2000 to 2008, interest for new coal-to-SNG projects also began to grow in the United States.

Some of the projects were already quite specific, until sharply dropping gas prices stopped all

activities. Different press releases also mentioned plans for SNG plants in the Ukraine [42] and

Indonesia [43].

Currently (December 2013), concrete activities for erection of further large-scale coal-to-SNG plants

can be found only in Korea and China [33].

Numerous different technologies are planned or proposed for the different SNG projects. Table 2.2

gives an overview of the different gasification systems proposed for specific SNG projects.

Table 2.2: Overview of gasification systems proposed for specific SNG projects, adapted from [33]

Name Type Project

Siemens [44] entrained flow CPI Yinin (CN), Decatur SNG plant (US) SES U-Gas fluidized bed Jiangxi SNG (CN) [45] CB&I E-Gas entrained flow Posco Gwangyang (KR) GreatPointEnergy Bluegas hydromethanation Wanxiang Turpan (CN) [46] TPRI gasification entrained flow CHNG Xinjiang (CN) SEDIN fixed bed ? Datang Keshiketeng (CN), Yining SNG (CN)

For methanation, mainly systems from Haldor Topsøe, Davy - Johnson Matthey and Foster Wheeler -

Clariant are proposed. All these systems base on fixed bed methanation.

Posco Gwangyang SNG, Korea

The first South Korean large-scale coal-to-SNG plant is currently under construction in Gwangyang.

The start-up is estimated for 2014 and SNG production output is expected to be 500,000 t per year.

This equals an SNG power of around 550-700 MW, depending on the heating value of the SNG. Three

ConocoPhillips (CB&I) E-gas gasifiers (one back up) will produce the synthesis gas. The gas will be

cleaned by means of a Rectisol unit delivered by Linde. The plant will use the Haldor Topsøe TREMP

technology for the methanation of the synthesis gas. [47]

CPI Yinin, China

The China Power Investment Corporation (CPI) is planning to erect a 2 billion Nm³/year

(1.9-2.2 GWSNG) coal-to-SNG plant in Yinin/Yili in Xinjiang province. Siemens will deliver eight Siemens

SGF-500 entrained flow gasifiers with a thermal power of 500 MW each, while Haldor Topsøe will

supply the methanation. [48]

Other Chinese SNG projects

Besides the already constructed Datang and Yining SNG plant and the CPI Yinin plant, six other coal-

to-SNG plants have been approved by the Chinese government [49]. The total SNG capacity of the

approved SNG plants is 37.1 bn. Nm³ per year [49], which is equivalent to an SNG power of around

40-45 GW. However, only little information on the project and construction progress is available.

State-of-the-Art

14

Theoretical Background

15

Chapter 3

3. Theoretical Background

3.1. Gasification

The thermal gasification of a solid feed-stock is the essential step for the production of synthesis gas

and therefore for the production of SNG. After drying and pyrolysis, the products from the pyrolysis

step are gasified at temperatures above 700°C. The solid pyrolysis coke reacts in heterogeneous

gasification reactions (equations 3.1 to 3.5) with the gasification medium to form gaseous

components, whereas the gaseous pyrolysis products react in homogenous reactions (equations 3.6

to 3.10). Although gasification is the term used to describe the third step in the production of gas

from a solid feedstock, also the whole process, including drying and pyrolysis, is referred to as

gasification.

Heterogeneous gasification reactions

3.1

3.2

3.3

3.4

3.5

The heterogeneous combustion reactions (equations 3.1 and 3.2) produce the heat for the

endothermic gasification reactions. The combustion reactions only occur in autothermal gasification

with air or oxygen. The heterogeneous water-gas reaction (equation 3.3) and the Boudouard reaction

(equation 3.4), produce the main synthesis gas components H2 and CO.

Homogeneous gasification reactions

3.6

3.7

3.8

3.9

3.10


16

Similar to the heterogeneous reactions, the homogeneous combustion reactions (equations 3.6 to

3.8) only occur in the autothermal gasification process. The CH4 produced in the pyrolysis step, or in

smaller amounts via the hydrothermal gasification reaction (equation 3.5), is converted to H2 and CO

via the methane-reforming reaction (equation 3.10). The CO generated can further react via the

water-gas-shift reaction (equation 3.9) by increasing the H2 content of the synthesis gas.

Only four of the ten gasification equations are independent [1]. Therefore the entire thermodynamic

process can be described by these four equations: the Boudouard reaction, the water-gas-shift

reaction, the methane-reforming reaction, and, in case of gasification with air or oxygen, the

oxidation reaction of CO (equation 3.6).

The thermodynamic equilibrium of the reactions determines the reachable product gas composition

in an ideal case. Usually the residence time in the gasifier is not sufficiently long to reach this

thermodynamic equilibrium. As a result, certain quantities of the pyrolysis products - mainly tars and

CH4 - are present in the product gas.

The theoretically reachable composition of gas depends primarily on reaction temperature and

pressure as well as fuel composition and the gasification medium. According Le Chatelier’s principle

high temperatures favor endothermic reactions, whereas higher pressures favor volume-reducing

reactions. Consequently the amount of H2 and CO increases with an increase in temperature and the

amount of CO2 and CH4 increases with an increase in pressure.

3.1.1. Allothermal gasification with water steam

The allothermal gasification with water steam is an efficient method of producing synthesis gas with

high amounts of H2, which is ideal for methanation. Contrary to gasification with oxygen or air, an

external heat source provides the heat of reaction for the endothermic gasification reactions.

The equation for the general reforming reaction of a hydrocarbon (equation 3.11) enables the

calculation of the amount of stoichiometric molar water needed for the reformation of a fuel.

( ) (

) 3.11

The minimum mass of water xH2O,min required for complete reformation can be calculated according

to equation 3.12. The minimum mass of water for wood pellets, used within this work, with the

molar formula (wet basis) CH1.646O0.722 is xH2O,min=0.199 kgH2O/kgFuel; for the used lignite (RWE Power

split), CH1.341O0.506, it is xH2O,min=0.415 kgH2O/kgFuel.

( )

( ) [

]

3.12

Typically, the gasification is performed at higher amounts of water, as required by the stoichiometry.

The excess steam ratio σ (equation 3.13) can be calculated analogously to the excess air ratio λ.

3.13


17

The excess steam ratio allows an adjustment of the H2/CO ratio of the synthesis gas. A high σ leads to

higher amounts of hydrogen in the product gas. Figure 3.1 shows a simulated permanent gas

composition on a dry basis as well as the water content for allothermal gasification of wood pellets

(ENplus-A1) with varying σ. The thermodynamic simulation was performed with Aspen Plus by using

an equilibrium approach with restriction of the CH4 content. According to the thermodynamic

equilibrium only a minor amount of CH4 would be present at typical gasification conditions. A

restriction of the CH4 content to realistic values allows therefore simulations that are more precise.

The main parameters and assumptions for the simulation were a gasification temperature of 800°C, a

pressure of 2 bars and total carbon conversion.

It can be seen that an increase in the amount of water leads to an increase in the H2 and CO2 content

and to a decrease in the CO content, as a result of pushing the shift-reaction (equation 3.9) to the

product side.

Figure 3.1: Simulation of the influence of σ on the permanent gas composition (dry basis) for allothermal

gasification: wood pellets, 800°C, 2 bars, equilibrium with restriction on CH4, Aspen Plus simulation

Such thermodynamic simulations, as well as validations on real gasifiers, provided the basis for the

definition of the standard synthesis gas composition (table 3.1) for the methanation tests within this

thesis.

Table 3.1: Standard synthesis gas composition for the methanation tests

Dry [vol. %] Wet [vol. %]

H2 51.6 31

CO 18.2 10.9

CO2 23.3 14

CH4 6.9 4.1

H2O 40

0.0

0.1

0.2

0.3

0.4

0.5

0.6

2 4 6 8 10

Gas

co

mp

osi

tio

n [

mo

le f

ract

ion

]

Excess steam ratio σ [-]

H2Owet

CO2,dry

COdry

H2,dry

CH4,dry


18

3.1.2. Tar problematic of thermal gasification

One product formed during the pyrolysis step is tar, a mixture of numerous organic components. Its

composition strongly dependents on the feedstock and the type of origination. From an operational

point of view tar is defined as a condensable product of organics in the producer gas stream [50]. The

literature proposes numerous classifications for tar; these are usually adapted for a particular

purpose. One of the most common methods is the classification according to ECN [51] (table 3.2),

which focuses on the properties, in particular the detection properties and the condensation

behavior of the tar species. Benzene is no tar species according to the ECN-classification.

Nevertheless, this thesis uses the term ‘tar’ for all hydrocarbons greater than or equal to benzene,

including benzene.

Table 3.2: Tar classes according to ECN [51]

Class Class name Properties Species e.g.

1 GC undetectable

Very heavy tars, gravimetric tar 7-rings and higher

2 Heterocyclic Cyclic hydrocarbons with heteroatoms, water soluble

phenol, cresol, pyridine

3 Light aromatic Compounds that usually do not pose problems regarding condensation or water solubility

xylene, styrene, toluene

4 Light polyaromatic

2/3-ring compounds that condense at intermediate temperatures at higher concentrations

naphthalene, acenaphthene, anthracene

5 Heavy polyaromatic

4-to-6-ring compounds that condense at high temperatures at low concentrations

fluoranthene, pyrene, chrysene

Ideally tar is just an intermediate product, which, in the gasification step, further reacts to

permanent gases. However, since this conversion is not complete, a certain amount of tar remains in

the synthesis gas, the exacts amount primarily depending on the gasification system, the type of fuel

and the reaction conditions.

Milne und Evans [50] classified tars according to their origination into primary, secondary and

tertiary products. The oxygen-rich primary tars, e.g. substituted phenols, are produced during

pyrolysis at 200-500°C. The majority of primary tars with increasing temperature react to permanent

gases, olefins and secondary tars, e.g. xylene, cresol, phenol or toluene. High temperatures

respectively increased reaction severities facilitate the formation of tertiary tars. Typically, tertiary

tars are polycyclic aromatic hydrocarbons (PAHs) without substitutes, like benzene, naphthalene,

anthracene or pyrene. Tertiary tars are formed by recombination of smaller molecular fragments

[52]. Primary and tertiary tars are generally not present at the same time, due to their nature of

formation and destruction. Primary products are destroyed before tertiary products appear. In

general, the total amount of tar significantly decreases with increasing gasification temperature and

reaction time and also the type of tar species changes. Higher temperatures lead to increased

formation of tertiary tars, which are more stable and might be more difficult to crack and remove

than primary or secondary products [50]. The assumption that, at higher temperatures, tars

thermally crack to permanent and other lighter gases is true for primary products. However, this is


19

not valid for tertiary tars, which grow in molecular weight with temperature and gas phase resistance

time [50]. Thus, the end-use of the synthesis gas should be considered when choosing the

gasification concept and operating conditions.

Downdraft gasifiers usually have lower concentrations of, what are mainly tertiary tars as the

products have to pass the hot oxidation zone. Contrary to that, producer gas from updraft gasifiers

contains large amounts of primary products. The tar yield of fluidized beds lies between that of

downdraft und updraft gasifiers. Entrained-flow gasifiers can be assumed as being almost tar-free.

Typical tar concentrations for updraft gasifiers are in a range of 20-100 g/Nm³, for downdraft

gasifiers of 0.1-1 g/Nm³ and for fluidized beds in a range of 2-20 g/Nm³ [50].

Methods for qualitative and quantitative tar analysis are presented in the experimental part of this

thesis (chapter 6.2.3).

3.1.3. Contaminations in the product gas from allothermal gasification

Besides tars, several other contaminations are present in the producer gas generated in allothermal

gasification such as particles, alkalis, sulfur and chlorine species and nitrogen containing

contaminations.

Particles and alkalis

The amount of particles in the producer gas is strongly dependent on the type of gasifier that is used.

Typical quantities for the FICFB gasifier in Güssing, as the state-of-the-art representative for

allothermal, fluidized bed gasification, are 30-100 g/Nm³ [35]. In case of fluidized bed gasification,

the particle fraction is a mixture of fuel ash and attrited bed material. Particles can cause plugging of

downstream applications if they remain in the synthesis gas.

Main alkali metals from biomass gasification are sodium (Na) and potassium (K) with a maximum

concentration of a few ppm [53]. Alkalis typically condense at temperatures below 600°C. They are

deposited as a sticky film on metal surfaces and adhere particular matter by forming ash deposits.

These alkali deposits are assumed to be corrosive to metal surfaces [54].

Sulfur

The main sulfur components of the synthesis gas are hydrogen sulfide (H2S) and in general, according

to the thermodynamic equilibrium, with a one to two powers lower amount, the organic species

carbonyl sulfide (COS) and carbon disulfide (CS2). The amount of sulfur in the feedstock is the main

influence for the rate of H2S released to the synthesis gas [55]. Allothermal, fluidized bed gasification

of biomass pellets, which contain relatively small amounts of sulfur, lead to H2S concentrations of

around 15-25 ppm ( [7] and own results), whereas the use of wood chips results in concentrations of

30-150 ppm ( [35] and own results).

Additionally, amounts of other organic sulfur species, like thiols (e.g. ethyl mercaptan), tiophene and

aromatic sulfur species, can be found in producer gas. There is, however, hardly any documented

evidence in the literature of the amount of organic sulfur species present after biomass gasification,

probably due to the complex method of measurement required.

Measurements by Kienberger and Zuber [56] (dotted area in figure 3.2) of producer gas generated

through allothermal water-steam gasification of wood pellets have shown that organic sulfur species


20

are in a range of about 10 % (≈1-3 ppm) of H2S. The operating conditions – and reaction temperature,

in particular – are the main factors influencing the concentration of organic sulfur. Higher gasification

temperatures, as a rule, decrease the amount of organic sulfur in the gas [57].

Figure 3.2 shows typical concentrations for sulfur contaminations and other gaseous contaminates in

producer gas from thermal gasification of woody biomass and lignite. Additionally, results of

measurements taken at the lab-scale gasifier that was used for the real gas methanation tests within

this work are depicted.

Sulfur is the main poison for nickel-based methanation catalysts. To guarantee a long operating time

catalyst specifications require a sulfur concentration in a low ppb range (e.g. < 100 ppb).

Chlorine and nitrogen species

Chlorine compounds are present in most biomass feedstocks, but only in a low concentration in

woody biomass. Chlorines usually appear in producer gas in the form of hydrochloric acid (HCl) [35].

In own measurements HCl was not detectable (below the detection limit of 1 ppm) in the raw

synthesis gas. HCl is listed as a catalyst poison for nickel catalysts.

Ammonia (NH3) is the main nitrogen-containing contamination. The amount of NH3 released

primarily depends on the nitrogen content of the fuel as well as the process conditions [35].

According to the thermodynamic equilibrium, higher temperatures and longer resistance times lead

to an increased conversion of NH3 to N2 [58]. Most suppliers of Ni-based methanation catalysts

specify NH3 as catalyst poison, but probably only as a precaution. Ni catalysts are also used to

catalyze NH3 decomposition, although at higher temperatures than in methanation. Own results, as

well as thermodynamic calculations, showed no discernible interaction between Ni and NH3 at

methanation conditions, but ammonia can become a problem for the later usage of the produced

SNG as it is corrosive for downstream applications and, if combusted, increases NOX emissions.

Figure 3.2: Typical concentrations of gaseous contaminates from gasification of woody biomass and lignite

with the measured contaminates from the EVT lab-scale allothermal gasifier

Measured contaminates lab scale gasifier

Woody Biomass Lignite (higher quality)

Co

nce

ntr

atio

n [

pp

m]

or

[g/N

m³]

fo

r B

TX/T

ar

0.1

1

10

100

1000

10000

H2S COS CS2 org. S* HCl NH3 BTX Tar**

*excluding COS and CS2; **Total tar amount including BTX;


21

3.2. Hot gas cleaning for sulfur and chlorine removal

The concept proposed in this thesis requires a removal of sulfur contaminations and perhaps a

removal of chlorines too, from the hot synthesis gas. The temperature for this cleaning process has

to be above the condensation temperature of the tars (> 300°C). Therefore, adsorptive hot gas

cleaning is a suitable option.

3.2.1. Adsorptive desulfurization with metal oxides

There are several metal oxides that are capable of adsorbing sulfur compounds from the hot

synthesis gas. Equations 3.14 and 3.15 show the general reactions of the two major sulfur

contaminates in synthesis gas, H2S and COS, with metal oxides. Oxides from the group of transition

metals such as Co, Cu, Fe, Mn, Mo, V, W and Zn have particularly good properties for adsorptive

desulfurization [59]. Furthermore, oxides of Ba, Ca, Ce and Sn also showed to be sufficient for

adsorption of sulfur [59]. The most common sorbents for hot gas applications are ZnO, CuO and CaO.

3.14

3.15

Figure 3.3: H2S equilibrium desulfurization concentrations for different sorbents with standard synthesis gas

composition (table 3.1) with 100 ppm H2S, upper limit with 40 vol. % H2O, lower limit dry, Aspen simulation

Zinc oxide (ZnO)

Zinc oxide is probably the most widespread adsorbent used for removing H2S in hot conditions. ZnO

allows desulfurization to a low ppb-range, but is strongly dependent on the water content of the

synthesis gas. Due to the formation of zinc vapor the operating temperature must not exceed around

700°C [60]. Strong reducing atmospheres additionally reduce this maximum temperature.

0.001

0.01

0.1

1

10

100

200 300 400 500 600 700 800 900

Temperature [°C]

Fe3O4

Ni

MnO

CaO

ZnO

H2S

con

cen

trat

ion

[pp

m]


22

Figure 3.3 depicts equilibrium H2S concentrations after desulfurization with different metal oxides

and Ni of the standard synthesis gas (table 3.1) with 100 ppm H2S. The upper limit is calculated for a

synthesis gas containing 40 vol. % H2O, whereas the lower limit is for the dry case. It shows that ZnO

enables a removal to 0.1 ppm H2S at 300°C of the wet standard synthesis gas composition.

The literature reports sulfur capacities of up to 34 wt. % for ZnO based sorbents [61], [59]. Higher

temperatures increase the adsorption capacity for sulfur [61]. Sorbents on ZnO basis are

commercially available, e.g. Clariant/Süd-Chemie ActiSorb S2 (capacity according to supplier of

32 wt. % S [62]) or BASF R5-12 (29 wt. % S [63]).

Own results for desulfurization tests with the ZnO sorbent Clariant-Südchemie ActiSorb S2 show a

sulfur adsorption capacity of around 28 wt. %. These tests were performed with the standard

synthesis gas composition with 40 vol. % H2O with addition of 500 ppm H2S, 25 ppm COS and 16 ppm

CS2 at a adsorption temperature of 250°C. Besides H2S, ActiSorb S2 also showed activity for removal

of COS and CS2. It is not clear if COS directly reacts with ZnO or if it first converts to H2S. Only few

literature sources report a direct reaction according to equation 3.19. Zinc sulfide (ZnS), produced

from adsorption of H2S on ZnO, catalyzes the COS conversion via the hydrogenation reaction. Thus, if

the ZnO bed already contains ZnS, COS can convert to H2S and be subsequently adsorbed [57].

Tests with 99.9 % ZnO powder carried out in the same conditions as the tests with ActiSorb S2, did

not show the removal of COS or CS2, whereas H2S was adsorbed. The ZnO powder reached a sulfur

load of around 9 wt. %.

Copper oxide (CuO)

Copper oxide has excellent thermodynamic properties for the adsorption of H2S. The capacity for

sulfur is lower than that of ZnO, which is way it is mainly used for deep desulfurization. Sulfidation

proceeds according to equation 3.16 and 3.17. [64]

The calculation of the reachable equilibrium concentration of H2S for desulfurization with CuO

showed that concentrations are in a low ppb range. Since this calculation was only possible in inert

gas conditions (H2S in N2), it is not shown in figure 3.3. In reducing synthesis gas conditions, CuO

always reduced to Cu, which has a much lower affinity to sulfur.

The maximum operating temperature proposed for CuO is 750°C [65]. However, since CuO strongly

tends to reduce to metallic copper in reducing atmospheres, the operation temperature is

limited [66]. Pure CuO adsorbents are not normally used for synthesis gas purification. The addition

of other metal oxides can stabilize the CuO and prevent it from reduction. Examples therefore are

copper aluminates (CuAl2O4, CuAlO2) or combinations with iron oxide [64]. Mixed oxides from copper

and manganese, like CuMnO2 and CuMn2O4, are very promising too. Cu-Mn sorbents are already

commercially available, e.g. Clariant/Süd-ChemieActiSorb 310 or FCDS-GS6. FCDS-GS6 allows the

removal of H2S and also of COS and other organic sulfur species up to a temperature of 400°C. Own

results proved the desulfurization ability of synthesis gas containing H2S, COS and CS2. These tests

also showed hydrogenation of COS and CS2 to H2S, despite full sulfidation of the sorbent.

3.16

3.17


23

Other metal oxides

Two sorbents for coarse desulfurization are calcium oxide (CaO) and calcium carbonate (CaCO3) or

their naturally occurring forms dolomite and calcite. Due to their thermodynamic properties

calcium-based sorbents are not suitable for desulfurization in a low ppm range (figure 3.3). Common

applications are in-situ desulfurization during gasification [67] or desulfurization in IGCC plants [68].

Manganese oxide (MnO) allows desulfurization up to a temperature of 1000°C even in strongly

reducing atmospheres. Apart from H2S, MnO also adsorbs COS according to equation 3.15 [69].

Iron oxide (Fe3O4) was one of the first materials used for removing sulfur. It is readily available and

therefore cheap and has a high sulfur adsorption capacity. Typical operating temperatures are

between 330-660°C. Fe3O4 is, however, only suitable for a coarse desulfurization (figure 3.3) and can

catalyze unwanted reactions, e.g. Boudouard-reaction. [64]

The literature reports numerous other types of sorbents for desulfurization, such as ceria (CeO2) and

mixtures of cerium with zircon, copper or lanthanum [70] or zinc-ferrite, zinc-titanate [64].

3.2.2. Desulfurization with activated carbons

Adsorption on activated carbons is one of the state-of-the-art methods for fine desulfurization of

biogas. Commercially available activated carbons are made from wood, lignite and hard coal as well

as coconut shells. By applying different impregnations the activity and the sulfur adsorption capacity

can be increased. In addition to the effect of physisorption from pure activated carbon, impregnation

enables adsorption by chemisorption. Typical applications for gas cleaning with activated carbons

operate at ambient or near ambient temperatures. This is due to the decreased adsorption

performance of the activated carbon itself. However, impregnations can overcome this problem.

Table 3.3 gives an overview of commercially available impregnated activated carbons for sulfur

removal. [71]

It can be seen that, contrary to most metal oxide sorbents, activated carbons also allow the removal

of organic sulfur, which makes them an interesting option for the removal of organic sulfur from hot

synthesis gas.

Table 3.3: Overview of commercially available impregnated activated carbons, adapted from [71]

Impregnation Amount Applications Products

Potassium carbonate K2CO3

10-20 wt. % Sour gases (H2S, HCl, HF, SO2, NO2), CS2

Carbotech D47/3-KC10

Iron oxide 10 wt. % H2S, mercaptans, COS -

Potassium iodide KI 1-5 wt. % H2S, PH3, Hg, AsH3, radioactive gases, mercaptans

Norit ROZ 3

Potassium permanganate KMnO4

5 wt. % H2S (without O2), aldehydes -

Potassium hydroxide KOH

10 wt. % Sour gases (H2S, HBr, HCl, HF, SO2, NO2), mercaptans

Donau Carbon Desorex K43Na

Sodium hydroxide NaOH

10 wt. % H2S, mercaptans -


24

In the recent years several investigations have addressed the use of activated carbon for hot gas

cleaning applications.

Desulfurization tests [72] with activated carbons with KOH, NaOH, Na2CO3 and KI impregnation

showed no large influence on the adsorption capacity for H2S at room temperature; at higher

temperatures (up to 550°C), however, the adsorption capacity increased significantly. It can be

assumed that carbon binds sulfur physically at lower temperatures (< 130°C), whereas at higher

temperatures, chemisorption is the predominating effect. The measured H2S adsorption capacity of

these activated carbons was between 0.4 and 4 wt. %.

Tests by Sakanishi [73] examined the simultaneous removal of H2S and COS over iron-impregnated

activated carbons. The results showed a higher capacity for COS removal than for H2S, due to the

partial decomposition of COS to CO. These tests also indicate that H2S may be removed mainly

through reaction with metal to produce metal oxide, whereas COS may be preferably adsorbed as

COS itself in the pores and decompose further.

Only few authors reported investigations under realistic synthesis gas conditions. Cal et al. [74]

studied the influence of the different synthesis gas components on hot desulfurization with activated

carbons. CO2 and H2O were favorable for H2S adsorption, whereas CO and H2 showed a contrary

effect.

In summary it can be said that the use of impregnated activated carbons for desulfurization of hot

synthesis gas has so far been investigated insufficiently. However, the results reported show a high

potential, especially for applications for which standard metal oxide sorbents are not suitable, e.g.

the removal of organic sulfur.

3.3. Methanation

The French chemists Sabatier and Senderens discovered the catalytic reaction of hydrogen and

carbon monoxide to methane in 1902 [14]. They reported a complete conversion of three parts H2

and one part CO to CH4 and H2O by reaction over reduced nickel at 250°C.

The methanation reaction (equation 3.18) is highly exothermic and reduces the gas volume by half.

Due to the stoichiometry, the reaction requires an H2/CO ratio of three. When using real synthesis

gas from thermal gasification, the ideal stoichiometric ratio of H2/CO = 3 cannot be expected.

Advantageously, the water-gas-shift reaction (equation 3.19) can adjust the H2/CO ratio. Nickel also

catalyzes the water-gas-shift reaction, which implies that if sufficient H2O or CO2 is available,

synthesis gases with a wide range of H2/CO ratios can be fully converted in the methanation reactor.

O 3.18

3.19

3.20

3.21

Instead of CO and H2, also methanation with CO2 and H2, according to the Sabatier reaction (equation

3.20) is possible. This reaction is a combination of methanation and the water-gas-shift reaction. If

CO and CO2 are present in the synthesis gas, CO2 conversion does not start until almost all CO has

been converted [25].

Another reaction to consider for methanation processes is the Boudouard reaction (equation 3.21).

Depending on the reaction conditions, it can lead to carbon deposition on the catalyst.


25

3.3.1. Thermodynamics

The achievable equilibrium product gas composition is, of course, determined by the stoichiometry,

but also by the composition of the inlet gas, the temperature and the pressure. The equilibrium

composition of a gas mixture, resulting from the reactions equilibrium constant K, can be calculated

minimizing the free enthalpy.

The simultaneous equations, necessary for the calculation of the equilibrium gas composition, can be

devised manually (e.g. according to [75]), or by means of thermodynamic calculation software with

already implemented equations (e.g. FactSage, AspenPlus). For the following thermodynamic

calculations, the software AspenPlus was used.

Figure 3.4 shows the influence of the reaction temperature on an ideal stoichiometric H2/CO mixture

of three at atmospheric pressure (1 bar). Due to the exothermic nature of the methanation reaction

higher methane conversions are favored at lower temperatures. This also implies that a cooling of

the reactor is necessary to achieve suitable conversion ratios. For that purpose, state-of-the art

concepts use cooled reactors or more common multiple reactors with intermediate cooling and high

product gas recirculation.

Figure 3.4: Influence of temperature on the equilibrium composition of an H2/CO=3 mixture at 1 bar,

FactSage simulation

Elevated pressure also promotes the conversion of methane, due to the reduction of molar volume

by the methanation reaction (figure 3.5). As the shift reaction is equimolar, reaction pressure does

not influence it. The major influence of pressure is in the range of 1 bar (atmospheric pressure) and

10 bars. An increase in pressure from 1 bar to 5 bars reduces the H2 content in the product gas by

half (from 6 to 3 vol. %) and increases the CH4 content by more than 1.5 vol. %, whereas an increase

from 5 to 20 bars only reduces the H2 content by about 1 vol. % and increases the CH4 content by

0.7 vol. % (figure 3.5). Considering this as well as the greater technical effort for high-pressure

applications, a medium pressure range of up to 5 bars during operation is suggested. A doubling of

the pressure can compensate for a reaction temperature increase of 20°C [10] and is a suitable

method for reaching higher conversions, especially if the temperature of the catalyst is already at the

lower limit for the catalytic activity.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0 200 400 600 800 1000

Gas

co

mp

osi

tio

n [

mo

le f

ract

ion

]

Reaction temperature [°C]

H2

CO

CH4

H2O

CO2


26

Figure 3.5: Influence of pressure on the equilibrium composition of an H2/CO=3 mixture at 300°C, FactSage

Since real synthesis gas mixtures always contain certain amounts of other gaseous components, like

CO2, CH4 and H2O, they also influence the achievable equilibrium composition. Figure 3.6 shows the

equilibrium composition of the standard synthesis gas used and its dependency on the reaction

temperature. An important parameter, especially for feed-in into the gas grid, is the amount of H2

remaining in the raw-SNG as different national regulations and technical requirements restrict the

H2 content permissible in the natural gas grid. To avoid the high technical effort of downstream H2

separation, the operating conditions for methanation should be carefully chosen. Figure 3.6 and

figure 3.7 show the main parameters influencing the reachable equilibrium gas composition. At an

outlet temperature of 250°C and an operating pressure of 5 bars the standard gas composition used

results in an H2 content of 3.3 vol. % in the SNG (after removal of CO2 and H2O).

Figure 3.6: Influence of temperature on the equilibrium composition of the standard synthesis gas

composition used (H2=0.3096, CO=0.1092, CO2=0.1398, CH4=0.0414, H2O=0.4; in mole fraction) and on the

chemical efficiency for methane conversion at atmospheric pressure (1 bar), FactSage simulation

0

0.1

0.2

0.3

0.4

0.5

0.1 1 10 100

Gas

co

mp

osi

tio

n [

mo

le f

ract

ion

]

Reaction pressure [bar]

H2

CO

CH4

H2O

CO2

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

200 300 400 500 600

Ch

em

ical

eff

icie

ncy

ηM

eth

Gas

co

mp

osi

tio

n [

mo

le f

ract

ion

]

Reaction temperature [°C]

H2 CO

CH4

H2O

CO2

ηMeth


27

The chemical efficiency (equation 3.22) is the ratio between the chemically bounded energy of the

produced methane stream and the chemically bounded energy of the synthesis gas stream. Since

lower temperatures increase the methane yield, also the chemical efficiency increases with lower

temperatures up to its maximum of 84.5 % for the standard synthesis gas composition.

[ ]

3.22

Figure 3.7: Influence of the H2O content on the equilibrium composition (without carbon forming reactions) of

the standard synthesis gas composition used (H2=0.516, CO=0.182, CO2=0.233, CH4=0.069; dry basis in mole

fraction) at 250°C and atmospheric pressure (1 bar), FactSage simulation

3.3.2. Reaction kinetics and mechanisms

Different group VIII metals are known to catalyze both methanation and the water-gas shift reaction.

An experimental study [76] which sorted the metals according to their activity and selectivity for

methanation found that due to its high activity, relatively high selectivity and reasonable price, nickel

is the favored catalyst for methanation. Typical methanation catalysts have an active compound that

is finely dispersed on a catalyst support with a large surface. The support material also has a

significant influence on the kinetics. Bartholomew, who investigated the influence of different

support materials, reported the highest reaction rate for TiO2, followed by Al2O3 and SiO2 [77].

The structure of the nickel crystals also affects reactivity. Stepped crystal faces, e. g. Ni(211), are

generally more reactive than close-packed faces, e. g. Ni(111) [78]. Additionally, different promoters

like Pt or Ru improve catalytic activity [79].

There is no agreement in the literature about the reaction steps for the methanation reaction of CO.

The most common theory (figure 3.8), which is based on a Langmuir-Hinshelwood (L-H) approach,

uses a stepwise hydrogenation of adsorbed surface carbon [80], [81].

0

0.1

0.2

0.3

0.4

0.5

0.6

0 0.1 0.2 0.3 0.4 0.5 0.6

Dry

gas

co

mp

osi

tio

n [

mo

le f

ract

ion

]

H2O content of synthesis gas [mole fraction]

H2

CO

CH4

CO2


28

The first step is the dissociative adsorption of the reactants on the active sites of the catalyst (figure

3.8: steps 1, 2 and 3). In the methanation reaction the dissociation of CO can proceed by two

different routes (steps 2+4 and steps 3+5) [80]. The hydrogenation of the CH intermediate to CH2

(carbene) is assumed to be the rate-determining step (step 6), whereas the subsequent

hydrogenation to methane (steps 7+8) assumed to be very fast [80], [81], [82]. The removal of

adsorbed oxygen with hydrogen also proceeds rapidly (steps 9+10) [80].

This reaction mechanism also implies that with an increasing CO/H2 ratio, more and more adsorbed C

accumulates on the surface due to the decreasing amount of adsorbed H and thus slows the

hydrogenation of C (step 5) [80]. This accumulation of C can result in deactivation of the catalyst due

to carbon formation, as will be described in chapter 4.2.

Figure 3.8: Model of the Langmuir-Hinshelwood approach for the methanation reaction,

according to [80], [81]

For the water-gas shift reaction, too, different pathways are discussed in the literature [83], [84],

[85], [86]. One mechanism often suggested is based on a redox mechanism with use of surface

oxygen as an intermediate (figure 3.9) [83], [85]. H2O adsorbs on the catalyst surface (figure 3.9:

step 2) and dissociates to atomic, adsorbed H and O (steps 3+4). It is not clear yet, if CO2 forms via an

Eley-Rideal (E-R) mechanism or an L-H mechanism (step 5).

OCHH

H H

OC C O

H

H C O

HH

H HC

HHO

H H

OC

H H C O

H

H CH H H H

H HC H

H

H HC

H O H HH

O

1 2 3

4

5

6 7 8

9 10


29

Figure 3.9: Model of a combined L-H and E-R approach for the WGS reaction, according to [83], [85]

The ability of nickel to catalyze the water-gas shift reaction allows complete methanation of a wide

range of different gas compositions. Synthesis gases from thermal gasification generally contain a

certain amount of CO2. If there is an insufficient amount CO in the synthesis gas, the reverse water-

gas shift reaction can use CO2 to produce CO. Due to the stronger adsorption of CO molecules than of

CO2 molecules the reaction is kinetically limited and takes places only until almost all CO has been

converted [25], [87].

3.3.3. Reforming of higher hydrocarbons

Reforming of higher hydrocarbons is one of the most widespread chemical processes. The steam

reforming of methane for the production of hydrogen-rich synthesis gas is a particularly important

step in many chemical processes. Usually, steam reforming of natural gas, which is the major steam

reforming application, takes place in tubular reformers filled with catalyst material and at

temperatures of between 400-550°C at the inlet and up to 900°C at the outlet [88].

This work proposes a direct conversion of higher hydrocarbons (tars) on the methanation catalyst.

The strongly exothermic methanation reaction and adequate management of reactor heat create a

temperature peak at the inlet of the reactor, which allows the catalytically supported conversion of

hydrocarbons, as shown by Kienberger [10].

In the literature there is, due to the complexity and variety of hydrocarbon and tar species, no

generally valid model for the reforming reaction to be found.

For the reforming of methane, which is the simplest reforming reaction, Xu and Froment developed

and validated an often-cited model [89]:

1. H2O reacts with the Ni atoms of the catalyst, yielding adsorbed oxygen and gaseous

hydrogen.

2. CH4 adsorbs on Ni atoms of the catalyst and either reacts with the adsorbed oxygen or

dissociate to form chemisorb radicals (CH3, CH2, CH and C).

3. The adsorbed oxygen and the carbon-containing radicals react to form adsorbed CH2O, CHO,

CO and CO2.

4. The hydrogen, carbon monoxide and carbon dioxide formed are directly released into the

gas phase, regarding the adsorption equilibrium.

The overall reaction of H2O with CH4 and H2O with CO is assumed to be the rate-determining

step/reaction within this model.

OC HHO

HHO

OC H

O

OC H O C HO H HOO C H

HHOO COC

1 2

3 4 5

6 7E-R

L-H


30

Rostrup-Nielsen proposed a model for the reforming of higher hydrocarbons, based on ethane

reforming [90]. Figure 3.10 shows a simplified version (in which not all reaction steps are shown) of

the model for ethane reforming. In contrast to the model of Xu, water adsorbs on the catalyst

support (figure 3.10: step 1). H2O dissociates on Ni atoms to adsorbed O and gaseous H2 (steps 2 and

3). C2H6 initially adsorbs on a dual site on the surface of the catalyst, involving a dehydrogenation

(step 4) followed by a rupture of the C-C bond to form adsorbed radicals (step 5). Next, the carbon-

containing radicals react with the adsorbed oxygen to form CO and CO2 by releasing gaseous H2

(steps 6, 7 and 8).

It is not clear yet, whether the rupture of the carbon-carbon bond or desorption of the products is

the rate-determining step [90].

Figure 3.10: Model of the reaction mechanism for the reforming of ethane, according to [90]

Korre et al. [91] proposed a model containing four essential steps for the reforming of polycyclic

aromatic hydrocarbons. Figure 3.11 shows one possible path for hydrocracking according to this

model for phenanthrene and naphthalene. In it, the first step is the hydrogenation of the aromatic

compound. The first hydrogenation product of naphthalene is tetrahydronaphthalene, which further

hydrogenates to decahydronaphthalene [92]. The next step is the isomerization of the compound.

One isomerization product of naphthalene is methylindane. The third step is the ring opening, which,

for example for naphthalene, forms butylbenzene. The last dealkylation step gradually reduces the

alkyl groups until, in the case of naphthalene, only benzene remains. [91], [93]

With increasing numbers of aromatic rings, also the number of possible pathways for reforming

increases [91].

HHO

HHO O H H O H H

HH

H

H

H

C

H

H

H

C

H

H

C H

H

C H H

HH

H

H

C H

H

C H H

H

H

C O

OC

OC

C OH H

HH

1

2 3

4

5

6 7

8


31

Figure 3.11: Model for hydrocracking of phenanthrene and naphthalene, according to [91] and [93]

Ising detected a relation between the activation energies of typical tar components (benzene,

naphthalene and phenanthrene) and the resonance-energy differences between the aromatic basic

state and the 1,2-hydrogenation. This relation implies that the coordination on the nickel surface is

the rate-determining step. [94]

The conversion rate of higher hydrocarbons on nickel catalysts depends on many different

influencing factors such as the reaction temperature in particular, but also the water content of the

gas, the retention time respectively the flow rate, the gas composition and interactions between

different tar species [52]. It has been reported, that the reforming activity increases with increasing

temperature and with increasing water content of the synthesis gas [95], [96].

Coll et al. investigated the reactivity for the reforming of different tar components and ranked the

compounds accordingly [95]. Benzene showed a higher reactivity than toluene, whereas that of

anthracene was much lower than that of toluene. Pyrene and naphthalene had the lowest reactivity

of the tar species investigated. Tendentially, the reactivity for reforming increases the lower the ring

number of the aromatics is, except for naphthalene.

A study [97] with real tar compositions from thermal biomass gasification also confirms the low

reactivity of naphthalene, but contrary to the investigations of Coll et al., benzene showed a much

lower conversion rate. At 800°C, 97 % of naphthalene and only 86 % of benzene are actually

converted. One reason for this observation could be that benzene was produced as intermediate

from the reforming of other tar species [95], as already presented in the model in figure 3.11. A

second reason could be cross-influences and interaction of different tar components.

Jess [96] reported that in case of a naphthalene-benzene-methane mixture, only naphthalene is

converted up to a temperature of about 750°C. This can be explained by the fact that naphthalene

strongly adsorbs on the catalytic surface and thereby decreases the conversion of benzene and

methane. Methane and benzene adsorb only weakly and therefore do not influence the catalytic

conversion of each other. More detailed information can be found in a summarizing review on

catalytic biomass tar removal by Dayton [98].

Hydrogenation Isomerization Ring Opening Dealkylation


32

Low temperature tar reforming and in-situ tar reforming during methanation

Vosecký et al. [99] investigated tar removal by means of steam reforming on a nickel catalyst at a

temperature of 500°C. In tests with a model gas composition containing the permanent gases H2, CO,

CO2, CH4, N2 and the tar components toluene, benzene and naphthalene and H2O, a tar conversion of

80 % at 350°C and 95 % at 500°C was achieved (figure 3.12). Due to the high GHSV (25000-30000 h-1)

used for these tests, a higher conversion would have been possible by operating at a lower GHSV.

Tests carried out with real synthesis gas from thermal biomass gasification confirm this assumption,

as the tar conversion reached > 99 %, also at lower S/C-ratios of 5-6. [99]

Figure 3.12: Reforming of benzene, toluene and naphthalene in model gas containing N2, H2, CO, CO2, CH4

and H2O over a Ni catalyst with an S/C-ratio of 18.1 [99]

Kienberger [10], [100] investigated tar conversion in-situ methanation on a commercial Ni-based

methanation catalyst. In the used polytropic fixed bed reactor concept, which is very similar to the

reactor concept proposed in this work, a temperature peak originates at the inlet zone of the reactor,

which provides enough exothermic heat for the conversion of higher hydrocarbons. In tests with real

synthesis gas from an allothermal biomass gasifier, tar conversion rates of 97.9 % were achieved. The

synthesis gas used, which was produced by the gasifier, mainly contained the permanent gases H2,

CO, CO2, CH4 and N2, had a water content of 35 vol. % and a total tar load (without BTX) of 6 g/Nm³.

In the tests the methanation reactor operated at a GHSV of around 3200 h-1 and an inlet temperature

of 350°C, a peak temperature of around 480°C and an outlet temperature of 275°C. [10]

200 250 300 350 400 450 500

Temperature [°C]

100%

80%

60%

40%

20%

0%

CnH

mco

nve

rsio

n[%

]

Benzene Naphthalene Toluene


33

3.3.4. Theoretical and practical aspects for the reactor design

The reaction mechanisms provide a basis for developing models that allow calculating the reaction

rate respectively the rate constant for the reactions taking place in the methanation reactor. By

knowing the reaction properties, the chemical reaction rate and diffusion limitations, a layout model

for the reactor can be developed. Numerous different calculation models can be found in the

literature. Kopyscinski [4] provides a good overview of different kinetic models for methanation and

for the water-gas shift reaction. Since all these models have been developed for a particular catalyst

and for particular operating conditions, it is, in general, not possible to directly transfer them to

other applications.

A practical approach for reactor design is to follow the recommendations for operating conditions

made by the catalyst manufacturer. A commonly used parameter is the gas hourly space velocity

(GHSV) (equation 3.23). It can be used to calculate the reactor volume (VReactor) from a given synthesis

gas volume flow (Vstd). The GHSV for the methanation catalysts used in this work should be below

4000-6000 h-1, depending on the reaction temperatures. Unfavorable operating conditions, e.g. high

water contents or low operating temperatures, can require much lower space velocities respectively

higher retention times.

[ ] 3.23

Besides the GHSV, other parameters also have to be considered in reactor design. Important

parameters to achieve plug-flow conditions are the L/dR-ration (catalyst bed length / reactor

diameter), the dR/dP-ration (reactor diameter / catalyst particle diameter) and the related L/dP-ratio

(catalyst bed length / catalyst particle diameter).

Due to their strong dependency on various operating conditions, such as volume flow or density of

the catalyst bed, there are no generally valid limits for L/dR, dR/dP and L/dP ratios. Different

mathematical approaches, summarized e.g. in [101], [102], [103], allow a validation if plug-flow

conditions can be reached in the reactor.

Typical values for the minimum dR/dP ratio required for minimizing wall effects are in the range of

8-15 [103]. If dR/dP ratios are too high, this has a negative impact on heat removal from the reactor. If

isothermal conditions are required, the dR/dP ratio should be below 5-6.

Typical values for minimum L/dP ratios are in the range of 25-350 [103]. For experimental fixed bed

reactors Mears proposed treating them as plug-flow reactors if the dR/dP ratio is > 10 and the

L/dP-ratio is > 30 [104].


34

Catalyst Deactivation and Carbon Deposition

35

Chapter 4

4. Catalyst Deactivation and Carbon Deposition

4.1. Deactivation mechanisms

Catalyst deactivation is one of the major concerns in industrial catalytic processes. While a slow and

controlled loss of activity is common, a rapid and unpredictable deactivation has to be avoided. Fast

deactivation processes are typical symptoms of wrong operating conditions, the presence of

impurities in the gas or improperly designed processes.

The deactivation rate for processes is generally a matter of economy. Whereas large industrial

application can require a catalyst lifetime of several years, small- or medium-scale applications can

also allow a much shorter lifetime, if this helps to reduce process complexity.

Table 4.1 provides an overview of the different catalyst deactivation mechanisms. In methanation

with nickel-based catalysts, poisoning, and fouling due to carbon deposition are the two most

common causes of catalyst deactivation.

Several papers contain a detailed review of catalyst deactivation mechanisms: [105], [88], [106],

[107].

Table 4.1: Overview of mechanisms of catalyst deactivation, according to [105]

Mechanism Type Description

Poisoning Chemical Strong chemisorption of species on catalytic sites, thereby blocking sites for catalytic reaction

Fouling Mechanical Physical deposition of species from fluid phase onto the catalytic surface and in catalyst pores

Thermal degradation Thermal Thermally induced loss of catalytic surface area (sintering)

Vapor formation Chemical Reaction of gas with catalyst phase to produce volatile compounds, e.g. Ni(CO)4

Vapor-solid and solid-solid reactions

Chemical Reaction of fluid, support or promoter with catalytic phase to produce inactive phase

Attrition/Crushing Mechanical Loss of catalytic material due to abrasion or crushing of catalyst particles


36

4.2. Carbon deposition

Carbon deposition is one of the major challenges for catalytic methanation of synthesis gas. It occurs

when carbon from the gaseous phase deposits on catalytic surfaces. Carbon deposition can result in

the destruction of the catalyst (figure 4.1), by an increase of the pressure drop across the reactor by

plugging of the reactor voids (a, c) and/or a loss of activity due to blockage of the active sites (a, b).

Due to the high cost of catalyst replacement, it is important to avoid carbon deposition in large-scale

industrial applications.

Figure 4.1: Forms of carbon deposits on Ni surfaces: (a) encapsulating film, (b) plugging of pores, (c) whisker

carbon, adapted from [108], [10] and [109]

According to Bartholomew [105] carbon deposits are distinguished by their origin as either carbon or

coke. Carbon is a product of CO dissociation, while coke is produced by decomposition or

condensation of hydrocarbons on catalyst surfaces. Coke forms vary and range from high-molecular-

weight hydrocarbons, such as condensed polyaromatics, to primary carbons, such as graphite,

depending on their formation and aging conditions. Within this work, the terms coking and carbon

deposition will be used synonymously.

4.2.1. Types of carbon deposits and reactions

As already mentioned in chapter 3.3.2, different studies have shown that the methanation reaction

consists of two main steps, the dissociative adsorption of CO and the hydrogenation of the adsorbed

species [81], [110], [111]. The first step of this reaction is the formation of intermediate carbon,

defined as C, from the dissociation of CO (equation 4.1). In the second step, the intermediate

carbon and the adsorbed oxygen are hydrogenated to CH4 and H2O (equations 4.2 and 4.3).

( ) ( ) 4.1

( ) 4.2

( ) 4.3

A distinction of carbon species found on Ni catalysts can be made according to their reactivity [112],

[113]. Table 4.2 shows the five major carbon species during methanation on Ni catalysts with their

temperature of formation from CO and C2H4 decomposition and their temperature (peak

temperature) with the highest rate for hydrogenation.

Nickel

Carbon(a)

(b)(c)


37

Table 4.2: Carbon species formed on Ni catalyst, adapted from [113] and [112]

Carbon species Formation

temperature from CO decomposition

Formation temperature from

C2H4 decomposition

Peak temperature for reaction with H2

C Adsorbed, atomic (surface carbide)

200-400°C 100-330°C

(Peak 220°C) 200°C

C Polymeric, amorphous films or filaments

250-500°C 330-620°C

(Peak 430°C) 400°C

CV Vermicular (polymeric, amorphous) filaments, fibres or whiskers

300-1000°C - 400-600°C

C Nickel carbide 150-250°C 230-330°C

(Peak 270°C) 275°C

CC Graphitic (crystalline) platelets or films

500-550°C - 550-850°C

Adapted from [113], [108], the reaction paths of carbon species listed in table 4.2 are shown in figure

4.2.

Adsorbed carbon C can be formed either by dissociation of CO (figure 4.2 route 1) or by

decomposition/ dehydrogenation of hydrocarbons/CH4 (2), (4). C is the essential intermediate for

the methanation reaction, but also for the formation of solid carbon. The methanation reaction (1-2-

3 and 4-2-3) proceeds quite fast when no diffusion limitations exist [108]. These reaction routes are

the desired ones for methanation of hydrocarbon-loaded synthesis gas. If more C remains on the

surface than is reacted, it can polymerize to less reactive solid C (10).

Figure 4.2: Reaction paths for formation, gasification and transformation of coke and carbons, adapted from

[113] and [108]

CxHy

CO

CxHy

CH4

CH4

C

C Cv

C

Cc

+H2Ovia H2/CO2

+H2

+/- H2

via CHn

-H2

1

2

3

4

5

6 7 8

9

10

1112 13

SolidAdsorbedGaseous

fast reaction

kinetically limited reaction

+H2Ovia H2/CO2

via CHn


38

C can be gasified with H2O or hydrogenated with H2 (9), but the reaction speeds are around two

orders of magnitude lower compared to reactions 1-3 [108]. C can further transform to graphitic

carbon Cc (11), especially at higher temperatures which are not typically reached in methanation

processes. Remaining on the catalyst, C can encapsulate the Ni crystallites or lead to blockage of

the pores.

When C dissolves in nickel, it can form vermicular carbon CV. The steps of filament growth can be

seen in figure 4.3. Amorphous, flocculent carbon is assumed to be a precursor for filament formation

(figure 4.3, step 2). The dissolved C diffuses along a thermal gradient (4) resulting from heat

released by the decomposition of CO/hydrocarbons (3), before being deposited. If more carbon

deposits on the particle surface than is removed for forming the filaments, the free particle surface

decreases. As a result, the decomposition of the reactants decreases, the particle temperature drops

and the rate of growth of filaments is slowed down (5) [109]. Filament growth does not necessarily

lead to a loss of catalyst activity, unless they are formed in such great quantities as to cause plugging

of the reactor voids or loss of nickel by removing the carbon fibers during regeneration. In fact,

filaments can even increase the activity of the catalyst by re-dispersion of Ni on the carbon support

[105]. Bartholomew [113] reported that during methanation at 400-500°C, amorphous films as well

as amorphous vermicular carbon were observed. This indicates that there is some overlapping in the

Cβ and CV forms.

Figure 4.3: Steps of growth of carbon filaments, adapted from [109]

At lower temperatures, C can also form nickel carbide C. There is some uncertainty if nickel carbide

is a short-living intermediate in the process of carbon dissolving in nickel as a precursive process for

the formation of CV [114], [115]. Under typical methanation conditions the formation of fixed C is

unlikely as it is not stable at such temperatures [108], [116].

Another route to solid carbon deposits on the surface of the catalyst is the formation of coke from

hydrocarbons (figure 4.2, route 6). Coke originates from condensed/adsorbed higher hydrocarbons

or from polymerization of hydrocarbons and methane intermediates (7 upwards). Coke can

transform into C by dehydrogenation (8) and can be gasified with water to H2 and CO2 before finally

being transformed to CH4 (7 downwards). The reaction rates for the transformation of coke are

assumed to be rather slow compared to the methanation reaction [108], [117].

Coke in the form of soot can also be produced by homogenous reactions from C2-C4 hydrocarbons at

temperatures above 650°C, whereas BTX and higher hydrocarbons need even higher

temperatures [118]. Due to the high temperatures, homogeneous reactions play no role for

carbon/coke formation under methanation conditions.

Support

Ni

Support

NiAmorphouscarbon

Support

Ni

CO / CxHy

C C

1 2

34 5


39

4.2.2. Thermodynamics of carbon formation

Equilibrium calculations are useful for estimating the amount of solid carbon deposited in

equilibrium in dependency of pressure and temperature. The calculations are based on the three

independent reactions, methanation (equation 3.18), water-gas shift (equation 3.19) and Boudouard

(equation 3.21). The results of these calculations can be illustrated in a C-H-O ternary diagram (figure

4.4). For each temperature, a phase equilibrium line for solid carbon can be plotted; in equilibrium,

carbon deposition can only occur in the area above the lines. As plotted for lignite and biomass,

carbonaceous fuels are typically deep inside carbon deposition area. The synthesis gas, gained by

gasifying fuel by means of steam (allothermal gasification) or oxygen and steam (autothermal

oxygen-blown gasification), is located much closer to the equilibrium lines, but usually still within in

the carbon deposition zone. Therefore, additional water – the more the lower the temperature – is

needed to prevent carbon deposition in equilibrium. With the lignite used for the real gas

methanation tests the amount of water needed in the synthesis gas to prevent carbon deposition in

equilibrium was about 35 vol. % at 300°C and about 40 vol. % at 250°C.

Figure 4.4: C-H-O ternary plot with phase equilibrium lines for solid carbon at different temperatures at 1 bar

Although ternary plots are useful as a rough guideline for checking whether the process conditions

carry a higher or lower risk of carbon being deposited, they do not allow making precise predictions.

Due to the nature of reaction processes, equilibrium conditions are not always reached. Therefore,

0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0

0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.00.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

Hydrogen H Oxygen O

Carbon C

600°C500°C

400°C300°C

200°C

Phase equilibrium lines for solid carbon

Lignite CH0.87O0.27

BiomassCH1.36O0.61

Syngas dry

Syngas with 40 vol. % H2O

H2O


40

carbon deposition may not occur although one point, e.g. ‘syngas dry’ in the ternary plot, is clearly

above the boundary line.

In addition, carbon deposition can occur in regions were no carbon was predicted in equilibrium, e.g.

syngas with 40 vol. % water. These kinetic limitations and deviations from the equilibrium are even

higher if the synthesis gas contains other gaseous components, in particular C2-C5 hydrocarbons, BTX

and other higher hydrocarbons [119]. The carbon formation tendency decreases with increasing

saturation of the hydrocarbons: alkynes > alkenes > alkanes [118]. Other kinetic investigations

reported that the amount of carbon formed when using benzene and toluene is several magnitudes

higher than when using CO [118], [120].

The mechanisms how these hydrocarbons promote carbon deposition are not fully clear yet. Gates et

al. [121] suggested that coking from olefins might proceed via olefin polymerization, olefin cyclization

to substituted benzenes, and subsequent formation of polycyclic aromatics from benzene. If

aromatics are present, one step could be the dehydrogenation to olefins. All these steps use

carbonium ions as intermediates and are catalyzed by Brønsted acid sites. The hypothesis of

carbonium ion chemistry would also explain why coke forms faster in the presence of hydrogen

acceptors such as olefins [106].

The main influences on the kinetics of carbon formation, apart from the presence of hydrocarbons in

the feed, are the steam-to-carbon ratio (S/C), the H2/CO ratio, the temperature, the presence of

pollutants in the synthesis gas, and, of course, the catalyst type, its surface structure and the support

material used.

Effects of H2 and H2O on kinetics

Detailed investigations under typical methanation conditions have shown that the rates of carbon

formation decreases with increasing S/C ratios and H2/CO ratios [122]. This is due to the fact that

adsorbed H2O or H2 reacts with adsorbed carbon and coke precursors formed by the dissociation of

CO or by the decomposition of hydrocarbons, and, in doing so, removes it. If sufficient H2 or H2O is

present, the residence time of carbon and coke precursors is too short to allow transformation to

more inactive carbon forms as shown in figure 4.2 [113]. A complete prevention of carbon deposition

only by adjusting the H2 and H2O content is not possible, as other influencing factors also have to be

in appropriate conditions.

Effects of temperature on the kinetics

It is a well-known fact that the temperature has a significant influence on kinetics. Figure 4.5

provides a summary of data on the formation and hydrogenation of atomic carbon Cα and

amorphous carbon Cβ under methanation conditions [113]. It clearly shows that the rate of

hydrogenation of Cα exceeds the rate of formation below 325°C. Therefore, atomic carbon is

removed faster than it is produced and no carbon should be deposited in methanation below 325°C.

However, depending on the methanation concept, temperatures above this point will be reached. If

methanation takes place far below 325°C and hydrocarbons are present in the feed, coke formation

due to condensation of higher hydrocarbons can occur.

Above 325°C, the rate of Cα formation is higher than the rate of hydrogenation, which results in an

accumulation of Cα. If sufficient Cα accumulates, the rate of conversion of Cα to Cβ becomes

significant [113]. Above 425°C, hydrogenation - and therefore the minimization/removal of Cβ - is


41

faster than the transformation of Cα to Cβ; however, the formation of filamentous carbon CV also

increases significantly at temperatures above 425°C [123].

Kinetic data such as those shown in figure 4.5 are only valid for a specific catalyst under specific

operating conditions and therefore not directly transferable to other systems. However, the

tendency and general behaviors should be represented anyway.

Figure 4.5: Rates of formation and hydrogenation of C and C species [113]

The temperature dependencies become even more complex if hydrocarbons are present in the

synthesis gas. Different studies have investigated carbon formation from hydrocarbons [118], [119],

[124], [125], [126]. Figure 4.6 shows carbon formation rates in relation to the temperature, for

example, for 1-butene and propene. With increasing temperature, the formation of carbon increases

until it reaches a maximum at 500-550°C. At temperatures above that level, the rate decreases until a

minimum has been reached. This decrease in the reaction rate and resulting apparent negative

activation energy is probably due to the effects caused by the relative magnitude of the activation

energy and the heat of adsorption of reactants. In addition, gasification of carbon and encapsulation

of nickel by carbon may also contribute to the decrease [119].

1 1.5 2 2.5

-2

0

2

4

6

8

10

1 1.5 2 2.5

Rat

e o

f fo

rmat

ion

ln(N

x 1

0³)

Reciprocal Temperature 1/T [10-3 K]

700 600 500 400 300 200 150

Temperature [°C]

C+HCH

C C

C+H CH

CO C+O


42

Figure 4.6: Temperature dependency of carbon deposition on Ni; 1-butene=133 kPa in hydrogen=33 kPa

[118]; propene=42.5 kPa in hydrogen=42.5 kPa [119]

The increase of the formation rate at higher temperatures (after the minimum) is caused by

homogenous reactions which lead to the formation of coke (soot) [125]. Due to the high

temperatures, this is not relevant for methanation.

Behaviors similar to those shown in figure 4.6 for 1-butene and propene, have also been obtained for

other hydrocarbons, e.g. C2-C5 [119], C5-C6 and benzene [126].

Effects of catalysts and poisons

Catalyst manufactures put a lot of effort in the development and improvement of catalysts with

enhanced activity, selectivity and durability by adding different promoters or modifying the surface

structure and the catalyst support. Such modifications can also be made to optimize the catalyst’s

resistance to carbon and coke formation. Different studies under methanation conditions found out

that addition of molybdenum weakens, and addition of platinum, iridium, bismuth or copper

improves resistance to carbon formation on a Ni/Al2O3 (SiO2 for copper) [122], [127], [114], [110].

Besides increasing the hydrogenation/gasification rate of carbon and its precursors, promoters can

also reduce the mobility and/or solubility for carbon in Ni [105].

Contaminations/poisons in the synthesis gas can also affect the rate of carbon deposition. For sulfur

(H2S) - one of the major contamination in synthesis gas - negative as well as positive effects on

carbon formation have been reported.

Tests in which H2S was added in concentrations of < 10 ppm to the feed stream have shown that

sulfur enhances the transformation of Cα to less active, polymerized Cβ, either by catalyzing the

transition or by preventing the dissociative adsorption of H2 [128].

On the other hand, it has been demonstrated that coking during steam reforming can be minimized if

traces of sulfur are added to the feed. Sulfur, as one of the strongest poisons for nickel catalysts,

chemisorbs on the nickel surface and deactivates it. But in low concentrations (H2S/H2=0.75 ppm)

700 600 500 400

Temperature T [°C]

1

10

100

0.9 1 1.1 1.2 1.3 1.4 1.5

Rat

e o

f d

ep

osi

tio

n [

µg/

min

cm

²]

Reciprocal Temperature 1/T [10-3 K]

1-Butene

Propene


43

some delineated zones, where no sulfur is adsorbed, will remain. If the size of the remaining zones is

in an order of 5 atoms, it will inhibit carbon formation, while enabling the steam reforming

reaction [129]. These investigations have been industrially implemented in the Haldor Topsøe SPARG

process [130], in which a pre-desulfurized catalyst is used for steam reforming.

Effects of the surface structure and the catalyst support have been observed in several studies. It has

been reported that carbon formation occurs at different rates on different Ni-crystal faces [131]. The

formation of carbon is favored on small particles having a high frequency of rough planes [113]. The

catalyst support mainly influences the carbon formation by effecting the hydrogenation of carbon. It

was found that hydrogenation of adsorbed carbon occurred 21 % faster on Ni/TiO2 than on Ni/Al2O3

and 43 % faster than on Ni/SiO2 [132]. Furthermore, also the dissociation of CO to C was found to

occur more rapidly in Ni/TiO2 than in Ni/Al2O3 and Ni/SiO2 [113].

All these complex dependencies of different influences on carbon and coke formation make a

detailed, application-related investigation necessary.

4.2.3. Possibilities for regeneration of carbon deposits

Generally, the two options for removing carbon deposits from Ni catalysts are either

gasification/hydrogenation with H2O/CO2/H2 or oxidation with oxygen or oxygen-containing

compounds [113].

The gasification rate between 500-700°C is slow until encapsulations from the Ni catalyst have been

removed. The rate significantly increases when surface reactions with Ni can occur [118]. Therefore,

partial blockages are easier to regenerate, while the regeneration of encapsulations or plugged pores

is much more difficult and requires harsher regeneration conditions [108]. Hydrogenation with H2 is

also slower than gasification with H2O [118].

As regeneration by gasification/hydrogenation requires high temperatures, it is not practical for most

industrial applications. Regeneration by oxidation is possible at lower temperatures. Studies [133]

have shown that using a mixture of 1-3 % air in N2 or 1-4 % O2 in N2 allows the removal of carbon

deposits from Ni catalysts at 300°C, but also leads to loss of active surface by sintering and loss of

catalyst crystallites due to removal of filamentous carbon. According to the results, it is impractical to

regenerate carbon-fouled catalysts more than 2-3 times [133].

An innovative option for regeneration could be microwave-enhanced regeneration with H2O [134], a

process in which carbon is selectively heated by means of microwaves and gasified with water steam.

Due to the properties of microwaves, carbon heats up much faster than the Ni catalyst. The

challenges with this approach lie in the limitation of the surface temperature to prevent sintering

and in process integration. However, the use of this regeneration technique cannot prevent the loss

of catalyst crystallites as a result of removal of filamentous carbon, either.

Due to all these limitations and challenges, the regeneration of industrial methanation catalysts is

not common. In practice, deactivated methanation catalysts are replaced. Typically, these

replacement periods are in an order of several years for large-scale applications, but can also be

shorter if this is economically advantageous.


44

4.2.4. Measurement methods for carbon deposition

Measuring carbon deposits on catalysts is crucial for minimizing carbon deposition during

methanation. An online and in-situ detection method would be the ideal tool for that purpose.

Mueller [135] presented a newly developed in-situ method for detecting coking on single Al2O3

catalysts particles, in which the catalyst particles are electrically contacted and characterized by

impedance spectroscopy. This approach relies on the fact that measured impedance changes with

the amount of coke on the catalyst. Unfortunately, such an application is neither available nor state-

of-the-art.

Within this work, carbon deposits were detected and analyzed using four different methods.

Increase of differential pressure

Strong carbon deposition results in a blockage of the fixed-bed methanation reactor, typically in the

inlet zone. The resulting increase in the differential pressure across the reactor provides an indication

during operation if carbon is being deposited in large amounts. The graph of the differential pressure

also indicates the amount of deposited carbon and the area of the reactor affected by it. Figure 4.7

shows typically observed trends during methanation tests. Strong coking with small axial distribution

within the reactor results in a fast exponential-like increase over the runtime. A slower, steadier

increase indicates more widespread coking with lower deposition rates. This method, however,

detects only severe coking and is therefore only helpful as a shut-off criterion for the reactor.

Figure 4.7: Typically observed reactor differential pressure trends resulting from coking

Visual evaluation of carbon deposits

Visual inspection of a representative sample can provide the first clues as to the amount of carbon

deposited on the catalyst after methanation. By classifying the catalyst particles into different groups

(e.g. no carbon, partially/fully covered with carbon) it is possible to calculate a dimensionless number

for comparison of different samples. This method allows only the detection of visible surface carbon

and is only suitable for a rough estimation.

0

20

40

60

80

100

0 20 40 60 80 100

Re

acto

r d

iffe

rnti

al p

ress

ure

[m

bar

]

Runtime [h]

No or low coking


45

Qualitative analysis of carbon deposits by means of scanning electron microscopy (SEM)

Scanning electron microscopy (SEM) is a standard method for analyzing surface characteristics. Thus,

it allows the definition of surface carbon deposits. Figure 4.8 shows the three major forms of carbon

deposits identifiable using SEM: filamentous carbon, carbon films/layers and graphitic platelets.

Figure 4.8: SEM-photos of different carbon deposit forms: a) filamentous, b) film, c) graphitic platelets

A combination with energy dispersive X-ray spectroscopy (EDX) also enables quantitative analysis.

Within this work, a Zeiss Gemini Ultra 55 SEM with EDX detector from the Institute of Particle

Technology of the Friedrich-Alexander-Universität Erlangen-Nürnberg was used for qualitative

analysis of several catalyst samples. It used an accelerating voltage of 20 kV and a secondary electron

(SE) as well as a back-scattered electron (BE) detector for detection.

Temperature-programmed oxidation (TPO)

Temperature-programmed oxidation (TPO) is a common analytical method used for quantifying

different catalyst deposits (e.g. [136], [137], [138], [139]). By heating a sample in an oxidizing

atmosphere the deposits oxidize, which results in the sample losing weight relative to the amount of

oxidized products. Additionally, it is possible to measure the composition of effluent gas after the

TPO. To quantify carbon deposits it is possible to calculate the amount of oxidized carbon from the

CO2 content in the off-gas of the TPO.

To analyze the deposited catalyst samples a Linseis STA PT 1750 thermo-gravimetric analyzer (TGA)

coupled with an ABB Uras 26 continuous non-dispersive infrared sensor (NDIR) photometer and an

ABB O2 analyzer were used. Figure 4.9 depicts the flow sheet for the TPO setup. The TGA is equipped

with a gas box for a controlled dispensing of two different gases (argon and oxygen); it is directly

connected with the TGA cell. The catalyst sample is filled in a slotted ceramic crucible, placed in the

TGA cell. The gas leaves the TGA cell at the top and flows directly through the gas analyzer, which

measures the CO2 and O2 concentration.

Figure 4.9: Flow sheet of the TPO setup to determine carbon deposits

500 nm 500 nm 500 nma) b) c)

Ar

F

MFC Ar: max. 400 ml/min

O2

F

MFC O2: max. 300 ml/min

Gas analyzer O2

Gas analyzer CO2

Thermogravimetric analyzer(TGA)

Filter

Exhaust


46

The TPO method used (table 4.3) allows a quantitative as well as qualitative analysis of carbon

deposits on catalysts. In steps 1 and 2 the catalyst sample is heated to 100°C in an argon atmosphere

and is thereby dried – although it is usually already completely dry at that stage and would need no

further drying. Subsequently the catalyst is heated to 700°C at a rate of 3°C/min and a constant flow

of 10 vol. % O2 in Ar (step 3). Argon is used because it has a similar density as CO2; which is why the

two mix easily. This should reduce the influence of buoyancy variations on the balance. In step 4 a

reduction of the heating rate to 2°C/min at temperatures between 700°C and 850°C increases the

precision and lowers the maximum CO2 concentration.

Table 4.3: TPO method for a quantitative and qualitative analysis of carbon deposits

Step Temperature

[°C] Heating rate

[°C/min] Time [min]

Argon [ml/min]

Oxygen [ml/min]

1 100 3 - 380 0

2 100 - 20 380 0

3 700 3 - 380 42

4 850 2 - 380 42

5 850 - 30 380 42

6 20 30 - 100 0

Figure 4.10 shows the results of a TPO analysis of a methanation catalyst used under operating

conditions without carbon deposition. It plots the temperature profile of the TPO method, the mass

change of the sample and the CO2 content formed. The catalyst mass increases in the first

120 minutes due to the oxidation of nickel. For methanation, it is necessary to reduce the fresh

catalyst to metallic nickel since the fresh catalyst mainly contains nickel oxide. To allow easy handling

of the catalyst it is slightly oxidized (only on the surface) after methanation as metallic nickel is not

stable at oxidizing atmospheres. As a result the catalyst sample still consists mainly of nickel rather

than of nickel oxide. After 120 minutes the catalyst is almost completely oxidized and the mass stops

to increase. With increasing temperature graphitic carbon, a component of the catalyst, starts to

oxidize and the mass decreases as a result of the formation of CO2.

It is possible to distinguish carbon species according to their reactivity with oxygen [139]. Reactive

carbon oxidizes mainly below 300°C. Polymeric carbon (Cβ) typically reacts at temperatures between

450 and 600°C. The less reactive graphitic carbon (CC) needs temperatures above 600°C for oxidation.


47

Figure 4.10: Results of TPO analysis of a methanation catalyst used under operating conditions without

carbon deposits (reference sample)

Figure 4.11 shows the results of a TPO analysis carried out on a methanation catalyst which

contained the highest amount of deposited carbon encountered in the course of this work. The CO2

trend shows the different types of carbon deposits measureable with the method used. The first

peak (shoulder) at 260°C results from more reactive carbon. The next peak at 350°C cannot be clearly

assigned to any particular carbon species. It can be assumed that it is a type of less reactive

polymeric carbon. This unknown carbon type, as well as reactive carbon, was only measured on

catalysts with severe carbon deposits that were operated with real synthesis gas. Typically, only

polymeric carbon with its peaks at about 530°C and 570°C is measurable. Due to the already high

amount of graphite in the catalyst, it is not possible to use this method for detecting graphitic carbon

deposits. However, graphite is not of particular interest as it only forms in significant amounts at

temperatures higher than the ones used in methanation [113].

Figure 4.11: Results of TPO analysis of a methanation catalyst used under operating conditions with severe

carbon deposits (sample with maximum amount of carbon)

-0.5

0

0.5

1

1.5

2

2.5

3

3.5

4

4.5

-100

0

100

200

300

400

500

600

700

800

900

0 100 200 300 400

CO

2co

nte

nt

[vo

l. %

]

Ch

ange

in m

ass

[mg]

Te

mp

era

ture

[°C

]

Runtime [min]

Temperature profile

CO2 content

Change in mass

Graphitic carbon

-0.5

0

0.5

1

1.5

2

2.5

3

3.5

4

4.5

-100

0

100

200

300

400

500

600

700

800

900

0 100 200 300 400

CO

2co

nte

nt

[vo

l. %

]

Ch

ange

in m

ass

[mg]

Te

mp

era

ture

[°C

]

Runtime [min]

Temperature profile

CO2 content

Change in mass

Polymericcarbon

Graphitic carbon

Reactivecarbon


48

Due to the influence of nickel oxidation, it is not possible to use mass change for calculations up to

temperatures of about 500°C. The results gained through mass change and CO2 formation only

coincide in a small temperature range of about 500°C to 600°C. At higher temperatures the results

from mass change measurements are much higher, which is due to reaction/decomposition of other,

unknown catalyst components (e.g. sulfur). Hence, the amount of deposited carbon was calculated

only from the amount of CO2 that had formed.

To calculate the mass of carbon on the catalyst sample, the CO2 profile of the carbon-free reference

sample (figure 4.11) is subtracted from the CO2 profile of the analyzed sample. From the CO2

concentration and the total gas flow the amount of released carbon can be determined, which, when

put in relation to the mass of catalyst allows calculating the specific carbon content of the

catalyst (mgCarbon/gCatalyst).

Error analysis

The lowest carbon content measureable was determined with < 0.1 mgCarbon/gCatalyst. Different

handling and measuring errors may influence the results. Measuring errors may be due to varying gas

supply (volume flow), the gas analyzer (measured concentration) and the manual processing of the

results. Errors of the mass flow controllers result in error rates of < ± 1.3 %, weighing errors in error

rates < ± 0.4 % of the TPO value measured. Gas analyzer errors may lead to deviations of < ± 1.4 % of

the measured value. The main error caused by the GA, the offset error, is compensated for during

the processing of the data. Manual data processing can result in an additional error of < ± 2 %. The

total measuring error of TPO analysis is < ± 5.1 % of the measured value.

However, an even higher error percentage may come from the sample itself and the way it is

handled. For TPO analysis, a sample of 5 g is taken from the total sample amount of typically 30 g.

Coking occurs only on few catalyst pellets, especially at low coking rates. Therefore, variations in the

quantity of coked catalyst pellets that make up a TPO sample lead to variations of the measured

values. However, these variations can be described by means of statistical methods and reduced

through multiple testing. The variations of the measured values of different samples are < 23 % for a

confidence interval of 80 % and < 30 % for a confidence interval of 95 %, the maximum variation

being 48 %. The variation related to a medium value out of 2-4 samples is < 11 % for a confidence

interval of 80 % of the samples and < 25 % for a confidence interval of 95 %. The maximum variation

between these medium values is 36 %. Variations are generally higher in samples with lower carbon

content and lower in samples with higher carbon content.

Based on this statistical error analysis, the overall error of TPO method used in this investigation is

50 % for a measured carbon content below 0.5 mg/g, 25 % for a carbon content between 0.5-4 mg/g

and 20 % for a carbon content > 4 mg/g.

Although, the used TPO analysis is not a precise instrument for quantifying the amount of carbon

deposited on a catalyst, it is nonetheless a useful and valuable tool as this type of investigation does

not require great precision.


49

4.3. Poisoning

Poisoning is the loss of catalytic activity due to strong chemisorption of contaminates on active sites

[106]. A poison can affect catalytic activity via different mechanisms [105]:

1. A strongly adsorbed atom of poison physically blocks several adsorption/reaction sites and

topside sites on the metal surface of the catalyst.

2. It modifies the adsorption/dissociation ability of neighboring atoms by virtue of the strong

chemical bond.

3. It can restructure the surface of the catalyst, which can lead to dramatic changes in catalytic

properties.

4. It can block the access of adsorbed reactants to each other.

5. It may lower or prevent the surface diffusion of adsorbed reactants.

The major poisons for Ni catalysts present in synthesis gas from thermal gasification are sulfur and

chlorine components [106]. Besides H2S, as main sulfur component, also COS, CS2, thiophene, thiole

and other organic sulfur species act as catalyst poisons.

4.3.1. Poisoning by sulfur

Due to its high industrial relevance, poisoning by H2S has been well researched, e. g. [140], [141],

[142]. These studies show that H2S adsorbs rapidly, strongly and dissociatively on nickel surfaces and

indicate that the adsorption of both H2 and CO is poisoned by sulfur [105], resulting in a significant

loss of methanation activity, even at low concentrations between 15-100 ppb. In this context it has to

be added that most of these fundamental studies were performed several decades ago and a lot of

progress has been made since then in the development of more sulfur-tolerant catalysts (e.g. [143],

[144]). Typical limits for sulfur contaminations for commercially available Ni-containing methanation

catalysts are in an order of 200 ppb. However, the rate of sulfur poisoning strongly depends on the

operating conditions and the composition of the catalyst. Addition of additives such as Mo and B,

which selectively adsorb sulfur species, significantly increases the sulfur tolerance [105].

Figure 4.12: Displacement of the reactor temperature profile due to selective deactivation at the entrance of

a polytropic cooled fixed bed reactor

200

300

400

500

600

0 0.2 0.4 0.6 0.8 1

Normalized reactor length

Tem

per

atu

re [

°C]

Deactivated reactor area after txh

t0htxh


50

Since sulfur is adsorbed very fast and selectively at the entrance of a packed bed [105], it results in

displacement of the exothermic-reaction-related temperature peak due to catalyst deactivation

(figure 4.12).

4.3.2. Regeneration of sulfur-poisoned catalysts

One critical question in the context of poisoning is whether it is reversible or not and if the catalyst

can therefore be re-used after regeneration. From the point of view of thermodynamics, a

regeneration of sulfur by oxidation should be possible [145]. However, due to kinetic limitations

temperatures above 500°C are necessary for oxidation of sulfur [146]. Another aspect to consider is

the formation of nickel sulfates. To prevent the formation of sulfates the temperature should not be

below 500-600°C [145]. Catalysts to which different promoters have been added can require higher

temperatures for a proper regeneration [147]. Higher temperatures can, however, also lead to the

destruction of the catalyst surface.

Different experimental researches have investigated the regeneration of sulfur by means of H2 and

H2O. It has, for example, been reported that a removal of sulfur from unpromoted Ni catalysts with

H2O at temperatures above 600-650°C is possible [147]. The regeneration with H2 is more difficult

because it is slow, even at high temperatures [148]. Several investigations deal with promising

attempts to regenerate sulfur-poisoned reforming catalysts, but due to the high temperatures of

700-900°C, this is not convenient for methanation catalysts (e. g. [149], [150]).

Regeneration of sulfur-poisoned methanation catalysts is complicated and therefore not common.

Even if regeneration were possible, the activity would decrease with every regeneration cycle.

Therefore, all state-of-the-art methanation processes require the removal of sulfur contaminations

to avoid poisoning.

4.4. Thermal degradation

Thermal degradation of the catalyst can result in a loss of catalytic surface due to crystallite growth

and pore collapse, loss of support area due to support collapse and chemical transformation of the

catalytic phases to non-catalytic phases [105]. All thermal degradation and sintering mechanisms

cause a reduction in size of the catalytic surface and of the number of active sites.

Sintering of nickel-containing catalysts is affected by different operating conditions and catalyst

properties. Sintering rates increase exponentially with increasing temperatures [105]. Sintering in O2

is faster than in H2 [105]; H2O also increases the sintering rate [151]. The main influences on the

sintering rate are the type of catalyst support and the promoters used. Generally speaking, Al2O3 is

more stable than SiO2, but also the interactions between the catalyst and its support have to be

considered [105]. This is why the only commercially available methanation catalyst for high-

temperature methanation, Haldor Topsøe’s MCR-2X, which is designed for operating temperatures

of up to 750°C, uses an Al2O3 support. Other methanation catalysts for lower temperatures can also

contain mixtures of different support materials, like MgO and SiO2 for the BASF G1-80 catalyst

(operating temperature < 650°C) [152] and Al2O3 and SiO2 for the Südchemie ActiSorb S7 (operating

temperature < 550°C) [153]. It was found that promoters like potassium and contaminates like sulfur

significantly increase the sintering rate at high pressures. At low pressures no influence of promoters

and contaminates on the sintering rate could be observed [151].

However, for the methanation concepts investigated within this work, sintering and thermal

degradation are insignificant as the maximum methanation temperature is below 550°C.


51

4.5. Evaporation – nickel tetracarbonyl

By reaction of gaseous CO with solid Ni (equation 4.4) poisonous, gaseous nickel tetracarbonyl is

formed. Low reaction temperatures and high partial pressure of CO facilitate Ni(CO)4 formation

[105]. Figure 4.13 shows the equilibrium Ni(CO)4 concentration for different concentrations of CO in

synthesis gas. The line for 10.9 vol. % CO represents the CO concentration in the standard synthesis

gas used for the methanation tests.

( ) 4.4

Deactivation due to formation of nickel tetracarbonyl is mainly a problem at the inlet of the

methanation reactor, where the temperature is low enough and the CO partial pressure high enough

to allow this poisonous gas to form [105]. After it has formed, nickel tetracarbonyl is transported

downstream the reactor. If the CO partial pressure is much lower and the temperature is higher,

Ni(CO)4 can react reverse to CO and Ni, which results in Ni deposition downstream in the

methanation reactor.

Apart from the deactivation effect, another important aspect to consider is the fact that Ni(CO)4 is

extremely toxic. The median lethal concentration (LC50) for 30-minute exposure lies around 3 ppm

[154]. Therefore exposure even to low concentrations of Ni(CO)4 has to be avoided.

The most common method of preventing the formation of significant amounts of Ni(CO)4 is to ensure

the inlet temperature of the synthesis gas respectively the inlet catalyst temperature are high

enough. Generally speaking, an inlet temperature of > 250°C is sufficient for the methanation

concept and conditions introduced in this work. Lower temperatures at the reactor outlet are no

problem as the CO concentration is low, too.

Figure 4.13: Equilibrium concentration for Ni(CO)4 for different CO concentrations in dependency of the

temperature, calculated with FactSage

0.001

0.01

0.1

1

10

100

100 150 200 250 300 350 400

Ni(

CO

) 4co

nce

ntr

atio

n [

pp

m]

Ni catalyst temperature [°C]


52

Bench-Scale Methanation Tests with Clean Syngas

53

Chapter 5

5. Bench-Scale Methanation Tests with Clean Syngas –

Polytropic Reactor Concept

The main objectives of bench-scale methanation tests are to prove the proposed polytropic reactor

concept for methanation and to screen different methanation catalysts. All tests were performed

with bottle-mixed synthesis gas of the standard gas composition (table 3.1). To conduct the different

tests, a suitable test rig was constructed. Results from these tests are conversions expressed by

achieved gas compositions, and the long-term performance of the catalyst.

5.1. Experimental setup

The test rig (figure 5.1) consists of a gas mixing station for providing artificial, bottle-mixed syngas, a

bench-scale methanation reactor and a gas analysis unit.

Figure 5.1: Simplified flow sheet of the bench-scale methanation test rig

T

PD

Flare

Natural gas

16 thermocouples (side-mounted and center-mounted)

Dif

fern

enti

al p

ress

ure

sen

sor

Methanation reactor

Rea

cto

r co

olin

g

3 z

on

e h

eati

ng

Water saturator

Trace heated

lines (2

00

°C)

H2

F

MFC H2: max. 21 l/min

CO2

F

MFC CO2: max. 15 l/min

CO

F

MFC CO: max. 10 l/min

CH4

F

MFC CH4: max. 1.6 l/min

F

MFC N2: max. 55 l/minN2

L

P

Pressure sensor

Gas analyzer H2, O2

Gas analyzer CO, CO2, CH4 Condenser

unit

Pump unit

Pressurized air

Gas

sam

plin

g o

ver

reac

tor

len

gth

Air (for catalyst oxidation)

WaterDemineralization

Exhaust


54

The gas mixing station consists of mass flow controllers and a water saturator. The mass flow

controllers (for H2, CO, CO2, CH4, N2) provide a dry synthesis gas with a total flow rate of 3-50 l/min.

For addition of water the dry syngas flows through a temperature-controlled bubble column filled

with water to be saturated. The water content can be between 0-60 vol. %. An automatic water-

refilling system ensures a constant water level in the bubbler. Overheating of the gas at the outlet

prevents condensation.

An ABB AO2000 system (chapter 6.2.3) analyzes the permanent gas composition. The product gas

leaving the reactor is burned in a natural gas flare.

To allow performing methanation tests, a tube reactor was constructed. For the layout high

representativity for large-scale concepts was a primary consideration. Therefore, and also to allow

the usage of commercial catalysts in their original shape, a certain minimum size was required.

The inner reactor diameter is 27.6 mm to fulfill the requirement of the dR/dP ratio being > 8

(chapter 3.3.4) for catalyst particles with a particle size of 3 mm. The length was set in order to

achieve a catalyst bed height of 600 mm. This results in a reactor (catalyst) volume of 350 cm³.

Figure 5.2: 3D drawing of the tube reactor and sketch with positions of thermocouples and gas sample ports

The reactor is divided into three separate heating and cooling zones with a length of 10, 20 and

30 cm. Electrical heating cords heat the reactor until the methanation releases enough exothermic

heat. Excess heat is removed through controllable air cooling.

Inlet

Outlet

Air cooled mantle

T1: 0 cmT2: 1 cmT3: 2 cmT4: 3 cmT5: 5 cmT6: 7 cm

T7: 10 cm

T8: 15 cm

T9: 25 cm

T10: 35 cm

T11: 45 cm

T12: 57 cm

TM 1-5

G1

G2

G3

G4

G5

DR,i: 27.6mm

VR: 350 cm³


55

Numerous thermocouples measure the temperatures at various points of the reactor (figure 5.2),

thus provinding information about variations of the temperature profile. The inlet zone (first 10 cm)

is equipped with more thermocouples due to the higher temperature gradients in this area.

Additionally, a tube for axially displaceable thermocouples is placed in the center of the reactor. Five

modified supports for thermocouples enable additional gas sampling at different points of the

reactor.

A differential pressure sensor between reactor inlet and outlet measures the pressure increases

resulting from blockages caused by catalyst deposits.

The reactor operates at atmospheric pressure although pressurization would be possible, too.

5.2. Catalysts for methanation

For the methanation of hydrogen- and water-rich synthesis gases, several catalysts from commercial

manufacturers seemed promising. For the first screening, five different catalysts (table 5.1) were

chosen. In order to fulfill the non-disclosure agreements with the catalyst manufacturers, it is

necessary to use synonyms for the different catalysts.

Table 5.1: Overview of the catalysts used for the methanation tests

EVT01 EVT02 EVT03 EVT04 EVT05

63 % Ni on SiO2/MgO

56 % Ni on SiO2/Al2O3

Ni-based Ni (> 50 wt. %) Ni (> 50 wt. %) on

Al2O3/SiO2

3 x 3 mm tabs Extrudates,

1.6 mm Extrudates,

1.6 mm 3 x 3 mm tabs 1.9 x 3.5mm tabs

< 650°C < 550°C < 550°C < 500°C < 550°C

Reforming catalyst

Sulfur sorbent Methanation

catalyst Methanation

catalyst Methanation

catalyst

EVT01

EVT01 has a thermal stability up to 650°C and a good resistance to degradation in water steam.

Typical applications are the reforming of natural gas at low steam-to-carbon (S/C) ratios and the pre-

reforming of hydrocarbons from natural gas to naphtha.

EVT02

EVT02 is a sorbent for sulfur removal in hydrocarbon streams. Due to its high nickel content, it is

promising for catalytic methanation. It was already tested at the Institute of Thermal Engineering at

the Graz University of Technology in previous research projects [155]. The operation temperature

should be limited to 550°C.


56

EVT03

EVT03 is an experimental catalyst with properties similar to those of EVT02, but was especially

developed and tested for methanation. The operation temperature should be limited to 550°C.

EVT04

EVT04 has similar properties as EVT03, but with addition of promoters to improve the resistance to

coking. It has also been used for the reformation of naphtha with up to 20 vol. % benzene at a

temperature of 500°C.

EVT05

This new semi-commercial experimental catalyst, which was specially developed for methanation,

should provide a higher activity than both EVT03 and EVT04.

5.3. Test procedure

To allow good comparability all the tests were performed according to the same procedure. The

catalyst was filled into the reactor until the first thermocouple (figure 5.2, T1) was slightly covered,

resulting in a catalyst volume of around 350 cm³.

Before the application of synthesis gas it is necessary to reduce the catalyst since the fresh catalyst is

in an oxidized or partially oxidized state. Table 5.2 shows the standard reducing procedure.

The fully reduced catalyst is highly pyrophoric. Therefore, it is necessary to oxidize it before removing

it from the reactor. For mild oxidation, a mixture of 5 % O2 in N2 was used.

Table 5.2: Standard reducing procedure

Step Temperature [°C] Heating rate [°C/h] Time [h] Gas

1 200 100 (2) 100 vol. % N2

2 500 50 (6) 50/50 vol. % H2/N2

3 500 - 3 50/50 vol. % H2/N2

4 350 50 (3) 50/50 vol. % H2/N2

Evaluation

The evaluation of the tests is based on the measured gas composition. The aim was to investigate the

relation between gas compositions measured and gas compositions calculated according to the

thermodynamic equilibrium. The results indicate the activity of the used catalysts. Variable

parameters for these investigations were the GHSV, the synthesis gas water content and the reactor

outlet temperature.

Catalytic activity is generally expressed by the turnover frequency (TOF), which expresses the rate of

formation in relation to the catalyst concentration. In this study catalytic activity was determined by

comparing hydrogen conversion rates respectively the amounts of un-converted hydrogen present,

which corresponds to the methane formation rate. One advantage of this approach lies in the fact

that the hydrogen concentration in the product gas is much lower and more volatile as the methane

concentration and can therefore be measured more accurately. Additionally, hydrogen is a critical


57

parameter for a feed-in into the gas grid and therefore important for the evaluation of appropriate

catalysts. All gas compositions given in this work are on a dry basis if not otherwise stated.

Another indicator for catalytic activity is the temperature distribution along the catalyst bed.

Temperature profiles can be useful for comparing the activity of different catalysts. Higher

temperature gradients at the inlet indicate higher catalytic activity. However, heat transfer

properties of the catalyst must be considered too. Temperature profiles are particularly useful for

evaluating the long-term activity of catalysts. If a catalyst is becoming less activate, the temperature

distribution changes and in case of poisoning, a displacement of the temperature peak occurs (figure

4.12). Deactivation of the catalyst leads to a general decrease in temperatures.

5.4. Methanation tests with different catalysts

5.4.1. Basic performance screening

The aim of the basic performance tests was both to evaluate the polytropic reactor concept and to

choose promising catalysts for detailed investigations. The question for the polytropic reactor

concept was if sufficient cooling is feasible. Figure 5.3 shows the axial temperature profiles of the

reactor of the five tested catalyst. The inlet zone (scaled reactor length 0-0.17) was not cooled,

whereas the cooling of the middle and outlet zone was set to reach 265°C at the outlet. As it can be

seen in figure 5.3 the temperature peak occurs at the end of the inlet zone, with peak temperatures

between 490-520°C. The lower peak temperature of EVT03 points to lower activity of this catalyst.

However, since the temperature gradients of the different catalysts are quite similar, and considering

the different catalyst shapes and their effect on heat transfer properties, the temperature profiles

measured do not allow making any reliable assumption about how active the different catalysts are.

Figure 5.3: Temperature profiles of the tested catalysts at a GHSV of 4000h

-1 and an H2O content of 40 vol. %

The temperature distribution in the reactor signifies that the majority of the reaction heat is released

in the inlet zone of the reactor, which caused the high temperature increase. This also indicates a

high conversion rate within this inlet zone. Gas composition measurements taken at various points of

200

250

300

350

400

450

500

550

0.0 0.2 0.4 0.6 0.8 1.0

Tem

pe

ratu

re [

°C]

Scaled reactor length [-]

EVT01EVT02

EVT03

EVT04

EVT05


58

the reactor confirm this assumption. Figure 5.4 shows the gas composition at various points of the

reactor compared to the temperature-related equilibrium compositions (dotted lines). The cooling

conditions and the resulting temperature profile were similar to the results shown in figure 5.3. The

CO conversion (XCO) is already 60 % after 0.1 of the scaled reactor length, which corresponds well

with the high amount of released reaction heat. CO conversion and CO2 formation are in equilibrium

after 0.1 of the scaled reactor length. This implies that only the temperature respectively heat

removal from the reactor limits the further reaction of CO and CO2. Contrary to that, H2 and CH4 need

the whole reactor length to reach equilibrium. Therefore those gases are the limiting components

that need to be considered for activity analysis. Tests showed that additional cooling of the inlet zone

leads to reduced conversion, especially of H2. A reduction of the overall reactor temperature also

slows down the kinetics. Strong cooling, or isothermal operation, would therefore significantly

increase the reactor volume needed. This makes the polytropic reactor concept a good alternative to

state-of-the-art concepts as it combines a simple design with lower catalyst volumes.

Figure 5.4: Gas composition measured at various points of the reactor compared to temperature-related

equilibrium gas compositions for EVT05 at a GHSV of 1500 h-1

, 30 vol. % H2O

To get a first impression of the activity of the catalysts, methanation tests performed with a GHSV

high enough to exclude any possibility of the catalyst reaching the equilibrium for H2. Additionally,

the influence of the H2O content of the synthesis gas was analyzed. Figure 5.5 shows the measured

H2 content of the product gas for the different catalysts. The results confirm that no catalyst reached

equilibrium with a GHSV of 4000 h-1 and that the H2O content has a significant effect on

H2 conversion.

In the tests EVT01 performed best, followed by EVT05 and EVT02. The two specially developed

methanation catalysts, EVT03 and EVT04, showed the lowest activity. Due to their good

performance, EVT01 and EVT05 were chosen for detailed investigations. Although the activity of

EVT02 was similar to that of EVT05, it was rejected as it is a commercial sulfur sorbent for which the

catalyst manufacturer cannot guarantee a constant activity across all batches. Considering the first

results and the catalyst specifications, EVT05 looks the most promising: it shows good activity and it

was specially developed with a view to high coking resistance.

0

10

20

30

40

50

60

0 0.2 0.4 0.6 0.8 1

Gas

co

mp

osi

tio

n [

vol.

%]


H2

CH4

CO2

CO


59

Figure 5.5: H2 content in the product gas for different catalysts at varying synthesis gas H2O contents

at a GHSV of 4000 h-1

and 265°C reactor outlet temperature

5.4.2. Detailed catalyst screening

In the course of detailed catalyst screening, the performances of EVT01 and EVT05 were analyzed

under typical operating conditions. The parameters used for this purpose are the reactor outlet

temperature (varied between 220-280°C), the H2O content of the synthesis gas (varied between

30-40 vol. %) and the GHSV (varied between 1000-3000 h-1). The results obtained are not only useful

for comparing catalysts, but also show a general behavior of Ni-based methanation catalysts.

Figure 5.6 and figure 5.7 depict the influence of the reactor outlet temperature and the H2O content

of the synthesis gas on the H2 content after methanation. With decreasing temperature, the H2

content reaches a minimum until it starts to rise again. This increase at lower temperatures is due to

the lower catalytic activity and therefore slower kinetics.

Figure 5.6: H2 content in the product gas in dependency of the reactor outlet temperature and the water

content with EVT01 at a GHSV of 1500h-1

0

2

4

6

8

10

12

14

16

18

20

20 25 30 35 40

H2

con

ten

t [

vol.

%]

H2O content [vol. %]

EVT01

EVT02

EVT03

EVT04

EVT05

Equilibrium 265°C

0

1

2

3

4

5

6

7

220 230 240 250 260 270 280

H2

con

ten

t [v

ol.

%]

Reactor outlet temperature [°C]

30% H2O

33% H2O

35% H2O

37% H2O

40% H2O

Equilibrium 30 vol. % H2O


30% H2O

33% H2O

35% H2O

37% H2O

40% H2O

5 K


60

The achievable H2 minimum depends significantly on the H2O content of the synthesis gas. An

increase of the H2O content shifts the minimum to higher temperatures. This shift also depends on

the type of catalyst that is used. An increase in the H2O content from 30 to 40 vol. %, for example,

causes a shift of the H2 minimum of 5 K for EVT01, whereas for EVT05 the same increase results in a

shift of 25 K.

The H2O content influences the equilibrium composition only slightly. The equilibrium H2 content for

40 vol. % H2O is only 0.5 vol. % higher than for 30 vol. % H2O at 280°C. Since the difference is even

smaller at lower temperatures, the great influence of the H2O content on the H2 conversion must

result from kinetic limitations. Studies of the methanation kinetics show that due to adsorption

effects higher H2O concentrations limit both methanation and the WGS reaction [38], [156], [157].

Kopyscinski [38] reported that the effects of higher H2O concentrations were outbalanced by the

WGS reaction, which leads to higher levels of H2 in the product gas and lower CH4 yields.

Figure 5.7: H2 content in the product gas in dependency of the reactor outlet temperature and the water

content with EVT05 at a GHSV of 1500h-1

0

1

2

3

4

5

6

7

8

220 230 240 250 260 270 280

H2

con

ten

t [v

ol.

%]

Reactor outlet temperature [°C]

30 vol. % H2O

40 vol. % H2O



25 K


61

Figure 5.8 and figure 5.9 show the influence of the GHSV on the H2 content analogous to the previous

diagrams. Since the H2 content at the reactor outlet depends on the reactor outlet temperature

(figure 5.6 and figure 5.7), it was set so as to reach the lowest possible H2 content at each operating

point. The GHSV influences H2 content in a rather linear relationship, whereby the inclination of the

curve rises with increasing H2O content. Higher GHSVs and higher H2O contents result in higher

H2 contents in the product gas.

Figure 5.8: H2 content in the product gas in dependency of the GHSV and the water content with EVT01

The results show that a high H2 conversion requires low GHSVs. To fulfill the requirements for feed-in

into the gas grid, such as the DVGW G260 [158] and G262 [159] regulations, an H2 content in the SNG

below 5 vol. % is required. This means that the H2 content of the raw-SNG, as presented above, must

not exceed around 2.6 vol. % before CO2 removal. This value can only be achieved with low GHSVs,

e.g. 1500 h-1 at 30 vol. % H2O. However, other requirements, which may allow higher H2 contents or

additionally restrict the amount of H2 in SNG, also have to be taken into account.

The comparison of the two catalysts tested shows that EVT05 is more sensitive to higher GHSV and

higher H2O contents.

Figure 5.9: H2 content in the product gas in dependency of the GHSV and the water content with EVT05

0

1

2

3

4

5

6

7

1000 1500 2000 2500 3000

H2

con

ten

t [v

ol.

%]

GHSV [1/h]

30% H2O

33% H2O

35% H2O

37% H2O

40% H2O

290

280

270

260

250

240

220

Ave

rage

eq

uili

bri

um

te

mp

era

ture

[°C

]

30% H2O

33% H2O

35% H2O

37% H2O

40% H2O

0

1

2

3

4

5

6

7

8

9

10

1000 1500 2000 2500 3000

H2

con

ten

t [v

ol.

%]

GHSV [1/h]

320

300

280

260

240

220

Ave

rage

eq

uili

bri

um

tem

pe

ratu

re[°

C]

30 vol. % H2O

40 vol. % H2O


62

5.4.3. Long-term performance of catalysts

To prove the long-term performance of the catalysts, several tests over a longer period were

performed. Figure 5.10 shows the temperature trend of a long-term test with EVT01 at a GHSV of

1500 h-1 and with varying H2O contents of 35-40 vol. %. The temperature trends show temperatures

in the inlet respectively main reaction zone, in which a possible deactivation should be recognizable

first. The position of the different temperatures and a drawing of the reactor can be found in figure

5.2.

Figure 5.10: Reactor temperatures for a long-term test with water content between 35-40 vol. % with EVT01,

GHSV 1500h-1

As depicted in figure 5.10, the temperature decreases by around 10°C in the first 80 hours. After that

the overall temperature trend remains constant. The large and fast fluctuations of the trends are due

to short interruptions of the gas supply, unsteady operation of the water saturator and changes of

the H2O content. The decrease of the temperature in the first 80 hours is an indicator for minor

deactivation. However, due to the stabilization of the temperatures in the following period, this

minor deactivation can be accepted. Other tests also confirmed this effect. A possible explanation for

this initial deactivation is minor re-oxidation of the nickel catalyst due to the water contained in the

synthesis gas, which leads to a loss of active surface. After a certain time a state of equilibrium

between the reducing influence of H2 and the oxidizing influence of H2O is reached and deactivation

stops. The temperature profiles and the trend of the gas compositions (figure 5.11) show no

indication of deactivation of the catalyst. The variations of the gas compositions mainly result from

changes in the H2O content.

450

470

490

510

530

550

570

0 100 200 300 400

Tem

pe

ratu

re [

°C]

Runtime [h]

TM2: 5cm

TM3: 8.5cm

T4: 3cm

T7: 10cm


63

Figure 5.11: Gas composition for a long-term test with a water content of 35-40 vol. % with EVT01,

a reactor outlet temperature of 240-260°C and a GHSV 1500h-1

5.5. Conclusion bench-scale methanation tests

The results show that the dry raw-SNG mainly contains CH4 and CO2 in about the same quantities as

well as certain amounts of H2 (figure 5.4 or figure 5.11).

The tested catalysts, especially EVT01 and EVT05, have a good activity for methanation under the

test operating conditions. Catalysts EVT01 and EVT05 are active down to around 230°C, but strongly

dependent on the H2O content. Higher amounts of H2O result in a reduction of H2 conversion and a

lower CH4 yield. Low GHSVs are required to reach equilibrium for H2 and CH4 at the reactor outlet.

For high GHSVs and higher H2O contents the highest possible H2 content could be too high to meet

the requirements. This problem, however, can be easily dealt with other technical solutions, e.g. by

having a second reactor with previous water condensation. CO conversion is already in equilibrium

after the gas has passed the first section of the reactor, also at high GHSVs.

Long-term tests showed no indication of deactivation apart from some minor initial deactivation.

The axial temperature profile in the reactor is shaped as desired, with a temperature peak directly

behind the inlet and a long cooling zone. The polytropic reactor concept therefore constitutes a good

alternative to state-of-the-art reactor concepts as it combines a simple design with lower catalyst

volumes.

The tests with clean, bottle-mixed synthesis gas produced no evidence against the concept proposed

in this work. However, to allow studying catalyst behavior under more realistic conditions, testing

also needed to be done using contaminated synthesis gas (chapters 6 to 8).

0

2

4

6

0 100 200 300 400

Runtime [h]

46

48

50

Gas

co

mp

osi

tio

n [

vol.

%]

H2

CH4

CO2


64

Methanation Tests with Contaminated Syngas - Setup

65

Chapter 6

6. Experimental Investigations with Bottle-Mixed

Contaminated Syngas – Experimental Setup

The investigations with bottle-mixed contaminated synthesis gas focused on several specific

questions arising from the proposed methanation concept. The goal of the tests was to gain a better

understanding of the methanation process with hydrocarbon-contaminated synthesis gas and the

resulting interactions, in particular the formation of carbon deposits, the influence of operating

conditions on carbon formation and the effects of carbon deposition on methanation. Related to the

problem of carbon deposition is the question whether it is possible to convert higher hydrocarbons

directly during methanation without adversely affecting the methanation process and the catalyst in

particular.

In order to answer these questions, methanation tests with representative artificial synthesis gases

were performed. For this purpose a suitable test rig that allowed the mixing of synthesis gas from

individual gas bottles as well as the addition of different synthesis gas contaminations such as

ethylene, tars and sulfur species was constructed, as well as a reactor test rig to enable conducting a

large number of methanation tests.

6.1. Investigation focus and program

The investigation program addresses several questions concerning the process of methanation with

hydrocarbon-loaded synthesis gas and the resulting problems:

Which contaminates lead to carbon formation on Ni catalysts?

What are these carbon deposits like and how do they behave?

What is the influence of the operating conditions, and especially the temperature, on the

amount of deposited carbon?

Is it possible to convert higher hydrocarbons directly during methanation without any

negative impact on the methanation process and the catalyst in particular?

Is it possible to reduce or prevent carbon formation?

6.1.1. Definition of investigation parameters

Numerous parameters influence the methanation process as well as the performance and lifetime of

the catalysts (figure 6.1). Apart from influencing methanation, many of these parameters also

interact with each other. Due to the limited amount of time and resources available a variation of all

the parameters was not possible; this, however, is not necessary as many parameters are already

fixed by the process concept and only few parameters can be directly defined during methanation.


66

Since the focus of the investigations was on the influence of synthesis gas contaminations, and

higher hydrocarbons in particular, this is one of the main parameters to vary. Ethylene was chosen as

representative aliphatic hydrocarbon as it is the major C2-C4 hydrocarbon in synthesis gas formed

during gasification; furthermore, it is one of the synthesis gas components that promote carbon

deposition most heavily. Tars were represented by a mixture of benzene, toluene, phenol and

naphthalene, which are the main tar species found in the synthesis gas produced through

allothermal biomass gasification.

The composition of the synthesis gas is mainly defined by the gasification process. Therefore all the

experiments were conducted in accordance with the fixed standard gas composition shown in table

3.1. The layout of the whole process, as well as the reactor type, defines parameters such as the

reaction pressure and the residence time. Pressure influences methanation, and the conversion in

particular. However, as this influence is minor and pressurization significantly increases test rig

complexity, all methanation tests were performed at atmospheric pressure.

Figure 6.1: Parameters influencing methanation

The bench-scale methanation tests with clean synthesis gas (previous chapter) already confirmed the

suitability of the polytropic reactor concept and allowed determining suitable space velocities. The

bench-scale methanation tests and other tests also showed that the main conversion of CO and

higher hydrocarbons occurs within the first few centimeters of the reactor and that, if carbon is

deposited, this always happens directly after the reactor inlet (for fresh catalyst with full activity).

The part after the inlet zone is only necessary for reaching a high methane yield according to the

thermodynamic equilibrium. The bench-scale methanation tests already proved the possibility of full

conversion of synthesis gas. In this chapter the focus of investigations is on the influence of higher

hydrocarbons. For that purpose it is sufficient to consider just the reactor inlet zone, which was done

by downscaling and shortening the reactor, but maintaining the same axial velocities as in the bench-

scale reactor.

Due to its good results during the bench-scale tests and the promising properties, the catalyst EVT05

was chosen for the methanation tests.

Synthesis gas contaminations

Reactor type

Residence time

Permanent gas composition

Reactortemperature

Reactionpressure

Catalyst

Water contentsyngas

MethanationCarbon deposition

ConversionActivity


67

The amount of water in the syngas is one major factor influencing carbon deposition. For most tests

the water content was set to 40 vol. %, to be outside the thermodynamic equilibrium for carbon

deposition (figure 4.4).

As a result, the only parameter left to vary for methanation was the temperature, which has

significant influence on the reaction kinetics and therefore on the conversion as well as formation of

carbon (chapter 4.2.2). The dew point of tars is the factor limiting the lower temperature, whereas

catalyst properties limit the maximum temperature allowed.

Table 6.1 summarizes the parameters used and their variations for the methanation tests. Unless

otherwise stated, standard conditions were used for the tests, which will be presented in the next

chapter.

Table 6.1: Overview of parameters for the methanation tests

Parameter Variations

Synthesis gas composition H2: 52.6 vol. %, CO: 18.2 vol. %, CO2: 23.3 vol. %; CH4: 6.9 vol. %

Water content 30-40* vol. %

Contaminates C2H4: 0-1 vol. %, Tars: 0-12 g/Nm³, H2S: 0-1 ppm

Reactor temperature oven: 300-550°C, inlet: 280-460°C, peak: 455-530°C

Reactor pressure atmospheric

Catalyst EVT05, Ni-based

Residence time / gas flow 3500 ml/min, axial velocity 0.15 m²/s, GHSV of ≈10000 h-1

*standard conditions

6.1.2. Test program and procedure

The basic idea for the test procedure was to run numerous short-term and long-term methanation

tests with varying operating conditions and varying addition of contaminates. After the tests, the

amount of carbon deposited on the catalyst was analyzed quantitatively by means of the TPO

method. The temperature profiles recorded during the methanation tests also provided an indication

for possible catalyst deactivation.

The tests conducted can be classified in four groups: tests with clean synthesis gas, tests with

ethylene, tests with tars, and tests to reduce deposition of carbon. The tests with clean synthesis gas

were the basis and reference for the further investigations. The tests with ethylene showed the

influence of an aliphatic hydrocarbon on the formation of carbon deposits and analyzed its behavior.

The third series of tests investigated the influence of tars on methanation, while in the last test series

different ways of preventing or minimizing carbon deposition were analyzed.


68

6.2. Test rig assembly

The test rig (figure 6.2) consists of a gas mixing station with tar conditioning unit (figure 6.3), a

methanation reactor test rig (figure 6.5) and the gas analyzing unit (figure 6.7).

Figure 6.2: Photo of the test rig for tests with bottle-mixed, contaminated synthesis gases

6.2.1. Gas mixing station with tar conditioning unit

The gas mixing station (figure 6.3) consists of mass flow controllers (MFC), a water saturator and tar

saturators to allow conditioning a realistic artificial synthesis gas.

The MFCs provide a dry gas mixture of H2, CO, CO2, CH4 and N2. Additionally, two MFCs allow the

addition of different gaseous contaminates, e.g. C2H4, H2S, COS. The total dry gas flow is in a range of

around 500 to 5500 ml/min, depending on the composition. The dry gas mixture passes a

temperature-controlled bubble column with water in order to saturate the gas stream up to

60 vol. % H2O. An automatic refilling system ensures a constant water level in the bubbler. It consists

of a floating level switch inside a communicating vessel and a liquid mass flow controller for refilling.

If only dry gas is required, a solenoid valve allows bypassing the saturator. Overheating of the

saturated gas as well as heating of all downstream lines prevents condensation.

Gas analyzing unit

Methanation reactor

Gas mixing station with tar saturators


69

Figure 6.3: Flow sheet of the gas mixing station with tar conditioning unit

Tar saturators

Four independent tar saturators allow the addition of different higher hydrocarbons in gaseous state,

even if they are liquid or solid under ambient conditions. The tar saturators are bubble columns,

similar to the water saturator. MFCs dose carrier gas (8-50 ml/min N2), which passes the bubble

columns containing the liquid tar species.

The bubblers have a high length-to-diameter ratio (length: 400 mm, diameter: 66 mm) to ensure

good saturation. The filling level is at around 250 mm (850 ml); the high volume enables a long

operation without refilling.

The tar saturator is heated and temperature-controlled, because the vapor pressure, which is

relevant for saturation, depends on the temperature. To achieve high isothermality, a thigh-fitted

10 mm aluminum shell surrounds the bubble column, which is made of stainless steel. A heating cord

Demineralization

H2

F

MFC H2: max. 2800 ml/min

CO2

F

MFC CO2: max. 1000 ml/min

CO

F

MFC CO: max. 1000 ml/min

CH4

F

MFC CH4: max. 350 ml/min

F

MFC N2: max. 3500 ml/min

F


N2

TG1

F

MFC TG1: max. 380 mlN2/min

TG2

F

MFC TG2: max. 1200 mlN2/min

L

F

F

F

F

N2

Tar saturator 1(Benzene)

Tar saturator 2(Phenol)

Tar saturator 3(Naphthalene)

Tar saturator 4(Toluene)

F

Water

Water saturator with automatic refilling

Trace heated lines (200°C)

P

Pressure sensor

Overpressure valve

Outlet / to reactor

Static mixer


70

wrapped around the aluminum shell heats the whole bubbler. Only the upper part of the column is

outside the shell and additionally heated with a heating sleeve to overheat the saturated stream. The

control system automatically calculates the necessary temperatures from the preset tar

concentrations. To calculate the temperatures (T), the partial pressure (pi) for each tar species is

calculated via the relation of the mole contents (ni) (equation 6.1). The saturation temperature

required is determined by solving Antoine equations (equation 6.2) for the different tar species. The

constants for the Antoine equations are according to Landolt-Börnstein [160]. Table 6.2 shows

Antoine constants for commonly used tar species.

[ ] 6.1

( )

[ ]

6.2

For calculations that are more precise the standard Antoine equation can be extended according to

equation 6.3. The variable χ (equation 6.4) contains the temperature T for which the vapor pressure

should be calculated, the temperature of the lower boundary T0 and the critical temperature TC

(temperature of the upper boundary of the temperature range).

( )

[ ]

6.3

[ ] 6.4

Table 6.2: Constants for Antoine equations of different tar species, according to [160]

Temp. Range [K]

A B C n E F

Benzene 279-376 5.98523 1184.24 -55.623 2.3835 12.283 664.01

Toluene 281-393 6.05043 1327.62 -55.525 2.38083 50.777 -877.95

Phenol 315-351 6.7074 1633.05 -98.55 360-480 6.296 1523.42 -97.75

Naphthalene 300-353 8.70592 2619.91 -52.5 354-420 6.13555 1733.71 -71.291

o-Cresol 245-296 11.9247 3979.5 -0.15 0.43429 463.53 -36925 308-356 4.4627 782.97 -170.05 0.43429 463.53 -36925 356-493 6.1834 1534.54 -96.85 0.43429 463.53 -36925

o-Xylene 248-301 7.5862 2277.61 -0 2.3586 75.45 -880.27 301-445 6.09789 1458.706 -60.109 2.3586 75.45 -880.27

Indene 290-460 6.34410 1749.215 -52.375

Acenaphthene 290-310 4.32951 1266.801 -136.33 335-367 9.20403 3076.294 -56.060 367-415 6.36589 2089.345 -71.070

Acenaphthylene 280-325 9.70593 3781.506 -1.688

Anthracene 299-430 10.5899 4903.3 -1.58 504-615 7.47799 3612.44 44.91


71

Because already small temperature variations have a great influence on tar concentration and

saturation can only be achieved in theory, it is necessary to additionally measure and monitor the tar

concentration (Micro-GC, FID, SPA-method). Although the deviations from the calculated values are

small, they nevertheless have to be taken into account for precise measurements.

When the tar saturators are not used, a manual cone valve seals the lines. Before leaving the gas

mixing station, the streams pass a static mixer. The standard operating pressure of the test rig is that

of ambient or near ambient conditions (200 mbar overpressure at most). However, the design allows

even higher pressure and all parts are leak-tested to cope with an overpressure of up to 3 bars. For

safety reasons, a mechanical overpressure valve with a release pressure of 6 bars is installed at the

outlet.

More details about the construction and layout of the gas mixing station can be found in [161].

Control system

The whole test rig is monitored and controlled by an industrial control system with hardware by B&R

Automation. A high level of automation, including safety precautions, allows long-term testing

without manual monitoring. Several automatized routines, such as an automatized catalyst reduction

cycle, increase the repeatability of the tests.

Figure 6.4: User interface of the control system

6.2.2. Methanation reactor test rig

The reactor test rig (figure 6.5) is directly attached to the gas mixing station and integrated into its

control system. It consists of a reactor oven for two reactors, the reactors, a gas supply for reducing

the catalyst and a flare.

The methanation reactor is placed in the reactor oven and is directly connected to the outlet of the

gas mixing station. A natural-gas-supported flare burns the gas leaving the reactor. Sample ports at

the in- and outlet enable gas and tar analysis.

The second reactor oven allows simultaneous reduction of the catalyst. By replacing the used reactor

with a freshly reduced one, the time intervals between the methanation tests can be shortened


72

considerably. If the reduced reactor is not used immediately after reduction, it can be kept heated in

the reduction part by purging it with N2.

Figure 6.5: Flow sheet of the methanation reactor test rig

The reactor oven (figure 6.6) consists of two steel tubes heated with electrical heating cords. The

heating cords are covered with high temperature insulation to reduce heat losses and operate up to

750°C with separate control of each part.

Figure 6.6: 3D-drawing of the reactor oven with the reactors; right side: methanation, left side: reduction

Reactor

The fixed bed reactor represents only the inlet zone of large polytropic reactors. Thus, the L/dR-ratio

is low. The reactor is made from high-temperature-resistant stainless steel (1.4841 or 1.4541) and

has an inner diameter of 22 mm and a useable length of around 140 mm. However, the standard

catalyst filling height is 50-60 mm to ensure the gas is heated sufficiently as it passes the upper part

PD

T

H2

F

MFC H2: max. 2500 ml/min

N2

F


Reactor oven

Reactor in methanation

test part

Reactor in reduction part

Dif

fere

nti

al

pre

ssu

re s

enso

r

SPE

Port for SPE-sample

5 thermo-couples

To FID

To gas analyzing

unit

To FID

Trace heated lines

From gas mixing station

Flare

Natural gas

SPE

Port for SPE-sample


73

of the reactor. Using some inert bed material above the catalyst leads to better heating of the gas

and ensures the formation of a steady plug flow.

To measure the reactor temperatures over the whole length, a 4 x 0.5 mm tube with five axially

moveable thermocouples inside it is placed in the center of the reactor. Compression fittings connect

the reactors with the gas supply unit and the outlet, to allow easy unplugging, replacing and reuse of

the reactor.

6.2.3. Gas and tar analysis and measurement techniques

To enable proper validation of the results, a complex online analysis system (figure 6.7) was set up

allowing the use of a variety of methods. Additionally, different offline measurement methods were

used to support the analytical process. The main parts of the online analysis system are the

permanent gas analyzer (GA), a micro gas chromatograph (µ-GC), a flame ionization detector (FID)

and a UV-adsorptive tar analysis system. The offline methods comprise contamination

measurements with adsorption tubes (Dräger tubes) as well as tar sampling and analysis by means of

the solid-phase adsorption (SPA) method.

Figure 6.7: Flow sheet of gas analyzing unit

H2

CG

UV-adsorptive tar analysis

Ejector pump with flow regulation

Condenser

N2 (zero gas)Pressurized air

Flame ionization detector

Calibration gas (1 vol. % C3H8 in N2)

Pressurized air

Silica gel Activated carbon

Filter (1µm)

Activated carbon (heated, 90°C)

H2, combustion gas

Trace heated lines (200°C)

Micro-GC

Gas analyzer H2, O2

Gas analyzer CO, CO2, CH4

Activated carbon

Condenser unit

Pump unit

Exhaust

He

From reactor

Exhaust

From reactor inlet

From reactor outlet


74

Permanent gas analyzer (GA)

The GA measures the permanent gas composition (H2, CO, CO2, CH4, O2) either at the reactor inlet or

outlet. Different valves (shown in the flow sheet in figure 6.5) allow switching to the desired

measuring point.

The used ABB AO2000 gas analyzer system [162] consists of a Uras26, Caldos25 and a Magnos206

module as well as a pump and condenser unit.

The Uras26 analyzer is a non-dispersive infrared (NDIR) photometer, which measures the CH4, CO2

and CO content. The measurement range is between 0-60 vol. % for CH4, between 0-100 vol. % for

CO2 and between 0-25 vol. % for CO. It works on the principle that certain gases absorb infrared

radiation in relation to their concentration at a specific wavelength. Non-dispersive means that the

full spectrum of the infrared light source passes the sampling chamber and is filtered just before the

detector. For accurate measuring results it is important to consider that components of the gas

mixture have different absorption wavelengths and therefore do not cross-influences each other.

The Caldos25 analyzer measures the thermal conductivity of gases and, in doing so, the H2 content of

the gas mixture in a range of 0-100 vol. %. The thermal conductivity of a gas mixture depends on the

concentrations of specific gas species present in it. The analyzer module contains a chamber fitted

with thermostatically controlled resistors. The gas to be measured flows via a membrane into the

chamber and cools the resistors. The temperature drop thus created is compensated for by an

increase in electrical current flowing through the resistors. This electrical current relates to the gas

concentration. The thermal conductivity detector has high cross-sensitivity with other gas species.

This influence is computationally corrected via the mainboard by using the gas concentrations

measured in the other detector modules.

The oxygen analyzer Mangos206 uses the paramagnetic behavior of oxygen (oxygen molecules are

attracted by a magnetic field) and the magneto-mechanical measuring principle. The sensor

measures oxygen up to 25 vol. %. Since oxygen is not a synthesis gas component, it is not present in

any of the gas mixtures used in this investigation; however, the sensor allows monitoring of possible

leakages.

The three analyzer modules are mounted inside two 19” housings. One housing contains the

mainboard, where all signals are brought together. The readings are displayed on the user interface

of the GA, but are also integrated into the control system of the test rig via analog signals. This

enables combined recording of test rig data with the related values of gas composition.

As the GA requires dry, dust-free and non-condensing gas, the gas passes a condenser unit, which

cools the gas to around 2°C, before entering the analyzer. Furthermore, an activated carbon filter

removes contaminations, like tars and sulfur species. The typical gas flow through the GA is in a

range of 20-40 l/h.

Micro gas chromatograph (µ-GC)

Gas chromatography is a standard method for qualitative as well as quantitative analysis of gas

mixtures. A µ-GC uses miniaturized components, including an injector, a column and a detector,

inside fully configured modules. This allows faster, quasi-online measuring. The used Agilent 490

µ-GC [163] is a quad-channel version equipped with three different GC modules. Table 6.3 gives an

overview of the GC modules used. The µ-GC has a heated sample line and heated injectors up to

110°C, which also allows measuring the amount of condensing hydrocarbons such as BTX.

Typical runtimes of one sample are between 30 and 180 s (max. 600 s). The µ-GC measures either a

defined number of samples or continuously (e.g. one sample every 120 seconds). The results, i.e. the


75

concentrations in vol. % or ppm of the different compounds, are automatically displayed in a

spreadsheet.

An internal pump provides the µ-GC with the sample gas, which has to be dry or almost dry

(condensation at ambient temperature is sufficient). Online measurements at the methanation test

rig were therefore taken after the condenser (figure 6.7), although measuring before the condenser

is possible too.

A common method for offline gas analysis is the use of gas sampling bags, which has the advantage

that the gas does not cool down to below-ambient temperature and components such as H2S or BTX

remain in the gas almost completely while the water content is sufficiently low.

Table 6.3: Overview of used µ-GC-modules

Module Column type Column data Detectable species

MS5A 10m Molecular sieve 5Å (zeolite type A) PLOT-column with pre-column PoraBOND Q

L = 10 m, D = 0.25 mm, Tmax = 180°C

H2, O2, N2, CH4, CO, NO, Ar, He, Ne

PPQ 10m PLOT-column with polystyrene-divinylbenzene

L = 10 m, D = 0.25 mm, Tmax = 180°C

C1-C5, CO2, H2S, COS, SO2

5CB 8m WCOT-column with 100% dimethylpolysiloxane

L = 8 m, D = 0.15 mm, Tmax = 180°C

C3-C10, CS2, H2S, other S-compounds

Several GC methods have been developed for analyzing the different compounds and

contaminations of synthesis gas. Due to the large variations in concentrations and to realize low

runtimes no method allows complete analysis of all the species. One method enables measuring

almost all major components (table 6.4): C2-C4 hydrocarbons, BTX and the main sulfur species.

Permanent gases are not measured as they are covered by the GA. The only limitation of this

standard method is the fact that it does not allow separating C2H2 and C2H4.

More details on different methods, the calibration and the application of the used µ-GC can be found

in [164].

Table 6.4: Parameters for the standard µ-GC method for C2-C4 hydrocarbons, sulfur compounds and BTX

Parameters PPQ 10m 5CB 8m

Injector temperature 60°C 80°C

Column temperature 60°C 80°C

Injection time 50 ms 40 ms

Column pressure 200 kPa 350 kPa

Runtime 110 s 110 s

Carrier gas He He

Calibrated components C2H2, C2H4, C2H6, H2S, COS, C3H8, C3H6

H2S, C3H8, C3H6, C4H10, CS2, benzene, toluene, ethylbenzene*, o-,p-xylene*

*Runtime 220 seconds


76

Flame ionization detector (FID)

A flame ionization detector (FID) measures the concentration of organic compounds in a gas stream.

A compound needs to have C-H or C-C bonds to be measurable with a FID. The gas sample, which is

burnt in a hydrogen flame, contains organic species, which, when burnt, lead to formation of ions.

The ions released during combustion are proportional to the concentrations of organic compounds.

The signal intensity is generally equal to the number of carbon atoms in the molecule. However,

response factors, depending on the device and on the measured species, lead to deviations from this.

The output signal as well as the response factors are always related to a reference gas.

The FID used within this work is an ABB AO MultiFID 14 analyzer [165]. It has a heated sample gas

port and uses C3H8 as reference gas. An air-supplied ejector pump draws the sample gas into the

combustion chamber. Typical sample gas flows are in a range of 35-60 l/h and are set by varying the

air pressure for the ejector pump. H2 is used as combustion gas, and conditioned pressurized air (dry,

hydrocarbon- and dust-free) as combustion air.

Heated pipes connect the FID with the reactor inlet or outlet (figure 6.7), depending on the position

of the 3-port valve. By switching on a bypass, the whole gas stream passes a heated activated carbon

filter prior to passing through the FID. The activated carbon is supposed to remove all tars from the

sample gas without affecting the permanent gas composition. By comparing the original signal with

the signal of the filtered gas sample the total amount of tars present in the sample can be

determined. This differential measurement works well with tar-loaded N2, but is difficult with real

synthesis gas. Apart from tars also CO2 and H2O adsorb on the activated carbon. The resulting

increase in the concentration of FID-detectable gases - mainly CH4 - leads to higher signal readings.

Once the adsorption of H2O and CO2 has reached a state of equilibrium, which is the case after a

certain time, the gas composition of the sample gas is not influenced any more. However,

fluctuations of the gas composition, the limited tar adsorption capacity of activated carbon and the

different magnitudes of CH4 and tars make the measurement inaccurate. A good alternative for dry

gases would be the usage of a tar condenser (e.g. cooled silica wool) instead of a tar adsorber as it

does not influence the gas composition.

Mörsch [166] put a lot of effort in developing a FID-based online tar analyzer, working on a similar

principle as the system described above.

Due to the above-mentioned difficulties, the FID was mainly used for the calibration and monitoring

of the tar saturators and not for online measuring of hydrocarbon conversion. More details on the

operation and calibration of the FID can be found in [167].

Optical tar analysis with UV absorption

Optical methods are promising alternatives for online analysis of tar in different gas mixtures. An

often proposed option is the usage of fluorescence spectroscopy for quantitative and semi-

qualitative analysis of tar produced in biomass gasification [168], [169], [170]. Most aromatic

hydrocarbons show the ability to fluoresce after absorbing of UV radiation; the intensity of the

fluorescence signal corresponding with the concentration of tar molecules.

A measurement setup for fluorescence spectroscopy consists mainly of a heated measurement cell

with optical ports, a UV-light source (laser or LED) and a spectrometer. To measure low tar

concentrations, e.g. after tar reforming, expensive lasers and highly sensitive detectors are required.

An alternative to fluorescence spectroscopy is measuring the amount of absorption, which has the

main advantage that due to the higher signal intensity low-power light sources, such as UV LEDs, and

conventional photodiodes can be used.


77

The Beer-Lamberts law (equation 6.5) describes the relation of absorbance (A) to the properties of

the light-absorbing material [171]. Absorbance is the logarithmic quotient of the intensity of the

incident light (I0) and the intensity of the transmitted light (I1). It is further described as product of

the molar absorption coefficient (ε), the molar concentration (c) and the optical path length (b). The

Beer-Lamberts law is only valid for low concentrations, which are typical for tars in synthesis gas. The

molar absorption coefficient is an intrinsic property and depends on the species and the wavelength,

e.g. for naphthalene ε286nm = 9300 [l mol-1 cm-1], for toluene ε261nm = 300 [l mol-1 cm-1].

( ) [ ]

6.5

Equation 6.5 shows that the transmitted intensity (I1) depends exponentially on the optical path

length (b). Therefore the simplest way of measuring low concentrations is to increase the length of

the measurement cell. The maximum amount of absorbance of a species can only be reached at a

particular wavelength, e.g. for naphthalene at 286 nm.

The disadvantage of absorption measurements is that if more than one absorptive species is present,

cross-influences reduce accuracy. Systems with multiple light sources of different wavelengths can

reduce or overcome this drawback.

Since in this investigation the focus of the measurement setup was on measuring the conversion of

naphthalene during the methanation process, a 285 nm LED was chosen as a light source. The

detector used was a standard, amplified photodiode for detection of UV and visible light and the

measuring cell (figure 6.8) a steel tube 22 mm in diameter and with an optical length of 300 mm.

Light was coupled-in or out by means of 6 mm silica glass rods, which allowed heating of the cell

while enabling cooling of the electronic parts. The measurement cell can be heated up to 300°C, to

prevent condensation of tars.

Figure 6.8: UV absorption tar measuring cell

The measuring cell can take measurements either at the reactor inlet or outlet. An ejector pump

after the measurement cell with flow regulation ensures the sample is moved continuously through

the cell, the standard flow rate being between 0.1-0.3 l/min. Before they reach the pump, water and

tar are removed by a condenser. To measure the initial intensity of the light the measuring cell is

purged with N2. The control system of the test rig records and processes the amplified signal sent by

the photodiode. The tar concentration is automatically calculated using calibration factors.

UV-LED with mountingPhotodiode with

mounting Glass rod

Heated measuring cell


78

Detector tubes (Dräger tubes)

One of the classical techniques of fast gas analysis is the use of detector tubes, where a defined

amount of sample gas is pumped through a tube which contains chemicals that react with the

substance to be measured by changing color. Typically, the length of the discoloration represents the

concentration of the measured species. A printed scale allows direct reading of the measured value.

Standard deviations are in a range of ± 5 to 25 % [172]. Condensation of water and interferences with

other species of the sample have to be considered and reduce accuracy. A good method of

preventing condensation and reducing the interference of water is to collect the gas in a gas

sampling bag. If a specific humidity of the gas is needed, the gas bag may be additionally cooled.

Although their precision is lower, detector tubes are a valuable tool for measuring species where no

other technique is available or where the existing techniques do not have the right measuring range,

e.g. H2S < 1 ppm. Table 6.5 gives an overview of the most commonly used detector tubes for

measuring contaminates during this work.

Table 6.5: Overview of the most commonly used detector tubes for measuring contaminates in synthesis gas

Species Tube type Range Interferences

NH3 Gastec Ammonia 3La 2.5-200 ppm CO2 (corr. Factor)

H2S Gastec Hydrogen sulfide 4M 12.5-500 ppm -

H2S Gastec Hydrogen sulfide 4LT 0.1-4 ppm Mercaptans

H2S Dräger Hydrogen sulfide 0.2/b 0.2-6 ppm Mercaptans

HCl Dräger Hydrochloric acid 1/a 1-10 ppm -

HF Gastec Hydrogen fluoride 17 0.25-100 ppm HCl

Tar sampling and analysis according to tar protocol

The European tar-measurement standard CEN/TS 15439 [173], so called ‘tar protocol’, is an attempt

to standardize the different tar and dust analysis methods for biomass gasification; it regulates the

sampling as well as the analyzing process. The sampling is based on collecting dust in a heated filter

and in sampling tars by dissolving them in isopropyl alcohol (IPA). Since the sampling in this study

was done after the particle filter of the gasifier, the filter, as specified by the tar protocol, was not

used. The heated sample port of the gasifier was directly connected to the impinger bottles (figure

8.4) and a membrane pump was used to draw the gas through the washing bottles, the silica gel

adsorber and the gas meter. The gas flow was in a range of 150-250 l/h and sampling durations

between 30 to 180 minutes. A deviation from the standard was that the first impinger bottle was left

empty instead of being filled with IPA. This was necessary due to the high water content of the gas,

which condenses in the first bottle and would overflow it if it had been filled. Additionally, an internal

standard for GC analysis was added to the solvent in the second bottle. All other sampling

procedures were according to the standard CEN/TS 15439 [173].

After combining the different solutions from the washing bottles, a small sample was analyzed with

an Agilent GC 7890A with CP-Sil 8 CB 25 x 0.25 column with retention gap. To improve accuracy, the

phenolic fraction was analyzed using a different method. The major tar species, as presented for

example in figure 8.8, are calibrated by means of different standards.


79

Solid-phase adsorption (SPA) method

Based on solid-phase extraction (SPE), another common method of tar sampling is to draw a tar-

loaded gas stream through a sorbent cartridge, where the tars are adsorbed before being extracted

and analyzed in a laboratory. Brage [174] first introduced this method of sampling of tars produced in

biomass gasifiers.

The method used in course this work [175] is a modified version of the original method, the two main

differences being the use of an octadecyl-phase (C18) adsorbent in place of an aminopropyl-phase

(NH2) and the use of isopropyl alcohol as the only solvent for extraction. The C18 column showed the

same sampling performance as the NH2 column. Due to lower initial contamination the baseline and

the separation performance of GC analysis were better for samples taken with C18. A comparison of

the extraction performance of different solvents (IPA, acetonitrile, acetone, dichloromethane)

showed no significant benefits of the other solvents compared to IPA. IPA has the advantage of being

easy to handle and of allowing the use of the same GC methods for analyzing the SPA samples and

the tar protocol samples.

Figure 6.9: Configuration for SPA sampling

The small reactor test rig (figure 6.5) as well as the gas cleaning and methanation test rig for real

gases (figure 8.3) are equipped with several sample ports (via septum). Heating the sample ports to

300°C prevents the condensation of tars. The design of the ports ensures that the tip of the

0.9x70 mm needle is directly exposed to the gas stream (figure 6.9). A 100 ml glass syringe draws the

sample over the Chromabond C18ec SPE column within a typical sampling time of 1 minute. To relate

the concentration to standard conditions, it is necessary to know the gas temperature in the syringe.

For that purpose the syringe temperature is either measured or the syringe is kept at a constant

temperature, e.g. by means of a bag filled with an ice-water mixture that is wrapped around the

syringe. After sampling, the column is immediately closed with a silicone plug and stored in the

refrigerator for later analysis.

SPE-column

100 ml syringe

Septum retainer

Gas stream


80

Methanation Tests with Contaminated Syngas - Results

81

Chapter 7

7. Experimental Investigations with Bottle-Mixed

Contaminated Syngas – Results

This chapter presents the results of the methanation tests carried out with contaminated, bottle-

mixed synthesis gas. First, tests with non-contaminated synthesis gas were performed as a reference

(chapter 7.1), before in further tests, ethylene (chapter 7.2), tar mixtures (chapter 7.3) and hydrogen

sulfide (chapter 7.4) were added to the synthesis gas.

7.1. Parameter variations with non-contaminated synthesis gas

These tests showed the relation of the different temperatures to each other (reactor oven

temperature, inlet temperature and peak temperature) and their distribution in the reactor.

Furthermore, the tests proved that contamination-free synthesis gas does not cause carbon

deposition.

Reactor temperatures

Adequate measuring of the different reactor temperatures is crucial for analyzing and comparing the

results of methanation tests. An error analysis that was carried out shows the influences of errors on

the measurements. Errors may stem from the influence of the protective steel tube, the

thermocouple itself and the evaluation unit. Since in the used setup the temperature in the center of

the reactor is measured via thermocouples fitted inside a steel tube (chapter 6.2.2), thermal

conduction along the tube may influence the measured temperature. A computational analysis

showed that this influence is < ± 1.5°C. The maximum error range of the thermocouples themselves

is < ± 2°C and that of the evaluation unit < ± 1.5°C. Thus, the overall maximum error range in

measuring reactor temperature is < ± 5°C.

Figure 7.1 shows the profiles of the measured reactor temperatures for standard methanation

conditions in dependency of the reactor oven temperature. The reactor temperatures are controlled

and influenced by the reactor oven temperature. Since the lab-scale reactor represents the inlet zone

of the bench-scale reactor the scaled reactor length is related to the length of the bench-scale

reactor.

Figure 7.2 shows the relation of reactor temperatures to the reactor oven temperature. As can be

seen, there is a clear, linear relation between all measurable temperatures. The inlet temperature is

the temperature of the gas directly before contacting the catalyst. The medium inlet zone

temperature is the medium temperature measured within the first centimeter after the inlet. It is the

result of the inlet temperature and the release of exothermic heat of reaction and is important as it


82

influences carbon deposition, which affected this zone most. The peak temperature is the maximum

temperature measured; it is primarily influenced by the heat of reaction, and only secondarily by the

reactor oven temperature. The heat of reaction released decreases with increasing temperature due

to the lower conversion of synthesis gas. Therefore, the lower amount of released heat compensates

for some of the increase in reactor oven temperature.

Figure 7.1: Measured axial temperature profiles over the reactor at different reactor oven temperatures

The further results presented here are mainly based on the reactor oven temperature as this was the

set temperature for all experiments; however, all other temperatures directly relate to it, as shown

in figure 7.2.

Figure 7.2: Resulting reactor temperatures in dependency of the reactor oven temperatures

-0.05 0 0.05 0.1

200

250

300

350

400

450

500

550

-2 -1 0 1 2 3 4 5

Scaled reactor lenght [-]

Re

acto

r te

mp

era

ture

[°C

]

Reactor length [cm]

500°C450°C370°C320°C300°C550°C

Catalyst bed

250

300

350

400

450

500

550

300 350 400 450 500 550

Co

rre

spo

nd

ing

tem

pe

ratu

re [

°C]

Temperature reactor oven [°C]

Peak

Medium inlet zone

Inlet


83

Results of tests with non-contaminated synthesis gas

To prove the influence of synthesis gas without contaminations, several tests were performed under

standard conditions (table 6.1), with reactor oven temperatures between 320-450°C, runtimes

between 2-72 h and a water content between 20-40 vol. %.

In all the tests the measured catalyst carbon content after the test was below the detection limit of

0.1 mgC/gCatalyst in the TPO. Also SEM and EDX analysis did not show evidence for carbon deposits.

Some of the analyzed catalyst pellets showed cracks and minor spallings on their surface. The most

likely explanation for this is the occurrence of thermal stresses resulting from the sudden increase in

heat at the beginning of each test. Due to the exothermic heat released after the introduction of the

syngas, the temperature in the main reaction zone rises sharply within seconds. An increase in the

number of cracks with ongoing runtime was not observed. However, a negative impact of the cracks

on the tests is not assumed and therefore this was not further investigated. For applications on a

larger-scale, slower heating, e.g. by dilution of the synthesis gas, should be considered.

7.2. Parameter variations with aliphatic hydrocarbons – Ethylene

Ethylene is known to be one of the major promoters of carbon deposition. Its concentration in typical

producer gas from thermal gasification is the highest of any C2-C5 species present in it. Unfortunately,

removal of C2H4 from synthesis gas requires some effort and is not possible with the proposed hot

gas cleaning concept.

7.2.1. Behavior of carbon on the catalyst

To allow conducting a large number of experiments it would be helpful to reduce the runtime. To

make this possible it is, however, necessary to know how carbon deposits on the catalyst develops

with time. If the amount of formed carbon correlates with the runtime, extrapolation of short-term

tests could replace, or at least partially replace, long-term tests.

To analyze the behavior of carbon on the catalyst numerous tests were carried out under operating

conditions where coking was expected. Previous tests as well as investigations by the catalyst

manufacturer had shown that the presence of ethylene in amounts as small as > 0.3 vol. % already

promotes carbon formation. Therefore tests with different runtimes were conducted in which

0.5 vol. % and 0.7 vol. % of C2H4 were added to the standard synthesis gas. Additionally, the inlet

temperature was varied (280°C, 300°C, 330°C) in order to investigate the influence of temperature.

After each test run the amount of deposited carbon was determined by means of the TPO method.

Figure 7.3, figure 7.4 and figure 7.5 show the results of these investigations. The results show clearly

that if coking occurs, the amount of deposited carbon increases linearly with the runtime.

Furthermore, it can be seen that the gradients of the graphs increase significantly with higher

C2H4 contents. The temperature has an influence on the gradient as well, as will be discussed in

chapter 7.2.3. The error bars are based on a statistical analysis of errors occurring when using the

TPO method (chapter 4.2.4).


84

Figure 7.3: Carbon deposition on the catalyst at 300°C reactor oven temperature (≈280°C inlet) using different

C2H4 contents and runtimes, GHSV 10000 h-1



Since the methanation conditions are outside the thermodynamic area for carbon deposition, due to

a water content of 40 vol. %, deposition must result from kinetic effects. Carbon deposition caused

by ethylene is, according to figure 4.2, either the result of its decomposition to Cα or its

polymerization to coke. The favored route is not determinable.

Carbon can only accumulate if the kinetics for the formation of carbon is faster than any reactions

removing carbon deposits. The kinetics depends on the operating parameters and catalytic activity.

Since the operating parameters were the same for all runtimes, the kinetics should not be

influenced, either. If, additionally, catalytic activity remains constant, the kinetics can also be

0.5 vol. % C2H4

0.7 vol. % C2H4

Car

bo

nco

nte

nt

[mg C

arb

on/g

Cat

alys

t]

0

2

4

6

8

10

12

14

0 50 100 150 200 250

Runtime [h]

0

2

4

6

8

10

12

14

0 50 100 150 200 250

Runtime [h]

0.5 vol. % C2H4

0.7 vol. % C2H4

Car

bo

nco

nte

nt

[mg C

arb

on/g

Cat

alys

t]


85

assumed to be constant. Therefore the measured linear correlation between amount of deposited

carbon and runtime can also be explained by constant kinetic parameters over the whole runtime.

This further indicates that the amount of deposited carbon does not influence catalytic activity; at

least not with the amounts that had been occurred in course of this work. The temperature profiles

obtained during testing with different runtimes and high coking confirm that the kinetic behavior

during methanation also remain constant (figure 7.6). However, these results do not necessarily

mean that carbon deposition does not influence catalytic activity at all.



Figure 7.6: Temperature profiles of a test with high carbon deposition; reactor oven temperature 370°C,

0.7 vol. % C2H4, runtime 142 h

0

2

4

6

8

10

12

14

0 50 100 150 200 250

Runtime [h]

0.5 vol. % C2H4

0.7 vol. % C2H4

Car

bo

nco

nte

nt

[mg C

arb

on/g

Cat

alys

t]

300

350

400

450

500

0 1 2 3 4 5 6

Re

acto

r te

mp

era

ture

[°C

]

Reactor length [cm]

10 h

140 h

50 h

100 h


86

A more severe problem is the blockage of reactor voids by carbon deposits and the resulting increase

in differential pressure. Blockage due to agglomeration of catalyst pellets and coke occurs directly

after the reactor inlet in the main reaction zone, where most of the conversion happens. Figure 7.7,

for example, shows the differential pressure measured in the experiment above (illustrated in figure

7.6): it increased from around 10 mbars at the beginning to 150 mbars after 140 hours, which was

the upper limit and meant the end of the experiment.

Figure 7.7: Development of differential pressure across the reactor during a test with high carbon deposition;

reactor oven temperature 370°C, 0.7 vol. % C2H4, runtime 142 h

7.2.2. Definition of a critical/acceptable carbon content

The fact that carbon accumulates linearly with time raises the question how much carbon is

acceptable. The critical point is reached when coking leads to blockage of the reactor and

subsequent increase in differential pressure in the reactor. Deactivation of the catalyst by carbon

was not considered for definition of a critical carbon content due to the fact that it was not

noticeable under the conditions applied in this investigation.

Figure 7.8 shows the amount of carbon deposits related to the maximum differential pressure. There

is, of course, a tendency towards higher differential pressures with increased carbon deposition, but

there is no clear connection; some points with similar amounts of carbon have differential pressures

that are 2 to 5 times higher (e.g. 300°C/0.5 vol. %/187 h and 320°C/0.7 vol. %/70 h).

There may be several reasons for this. Besides the amount of carbon, its dispersal also influences

differential pressure. Most of the carbon is usually deposited along the first centimeter of the

catalyst bed, which is between 5.5 to 6 centimeters long. The amount of carbon is determined on the

basis of the total amount of catalyst material; its location and dispersal are not considered. If the

same amount of carbon is deposited in a smaller area, this leads to a greater blockage and thus to a

higher differential pressure than in the case of more widespread coking.


87

If coking is accompanied by catalyst poisoning, this can increase the amount of deposited carbon

before the blocking effect becomes problematic. As already been mentioned, poisoning leads to

deactivation and subsequent continuous shifting of the main reaction and the coking zone, causing

carbon deposition to be more dispersed. In tests with simultaneously poisoning carbon amounts of

15 mg/g and above did not lead to an increase in differential pressure (tests with real synthesis gas,

chapter 8).

When comparing results, such as the amount of deposited carbon, in particular, it is important to

consider the fact that the measured amounts are always a mixture of coked and uncoked catalyst

and that it is therefore important to use the same flows and geometric parameters (same axial

velocity, same L/dR ratio and, ideally, same dR/dP ratio) .

Another factor potentially influencing the relation between differential pressure and the amount of

carbon deposition is the filling procedure used. Variations in reactor bed density and in the flatness

of the surface of the catalyst bed can always occur; to minimize this influence, a standardized filling

procedure was used throughout this investigation.

Figure 7.8: Relation of differential pressure and the amount of carbon deposited in the reactor

Considering all these different issues it becomes clear that it is difficult to define a critical/acceptable

level of carbon content.

Figure 7.9 shows catalyst samples with different amounts of deposited carbon. At low carbon

contents (< 1 mg/g) only few catalyst pellets are partially covered with a thin layer of carbon. Since

the differential pressure is not increased, either, this amount of carbon can be assumed to be

uncritical. Samples with a content of up to 5 mg/g contain a larger number of catalyst pellets with

carbon deposits. They are also mainly partially covered, but have a thicker layer of carbon. These

layers lead to a diameter increase of up to 30 % and agglomerations to an increase in differential

pressure. Although high values were measured too, the average differential pressure lies in the

medium range, where, depending on the application, it is not problematic.

Carbon contents of around 10 mg/g lead to the formation of a large number of coked catalyst pellets.

They appear as both fully and partially covered pellets. The carbon layers are thick, and strong

0

40

80

120

160

200

0 2 4 6 8 10 12 14

Re

acto

r d

iffe

rnti

al p

ress

ure

[m

bar

]

Deposited carbon [mgC/gCatalyst]

Initial pressure320/0.5/70

320/0.5/210

320/0.7/70

320/0.7/142

320/0.7/46

300/0.5/70370/0.5/66

300/0.5/187

370/0.7/70

Temp./C2H4/Runtime [°C/vol.%/h]


88

agglomerations between the coked pellets lead to a high differential pressure across the reactor. A

carbon content above 10 mg/g could be problematic for many applications. Therefore, this work

assumes 10 mgC/gCatalyst to be the maximum carbon content acceptable (the critical catalyst carbon

content). However, it has to be mentioned that this value is based only on the results of the concept

proposed and the catalyst used in this investigation and might not be directly transferable to other

setups and applications.

A carbon content of 15 mg/g leads to the formation of mainly fully covered, agglomerated catalyst

pellets with a thick carbon layer and an increase in the differential pressure to around 20 times the

normal level. Such a high amount of carbon (> 15 mg/g) – which was never found in methanation

tests with bottle-mixed synthesis gas – is probably too much for continued operation of a reactor.

The sample with the almost fully coked catalyst containing 35 mg/g of carbon was taken from the

main reaction zone after a methanation test using real synthesis gas from biomass gasification.

Figure 7.9: Photographs of catalyst samples with different amounts of deposited carbon

If 10 mg/g is considered the maximum carbon content acceptable and the threshold value at which a

catalyst needs to be replaced, and assuming a linear coking rate, it is possible to calculate an

acceptable carbon content for any particular moment of the process. This, in turn, allows short-term

testing, e.g. over 22 hours, to replace long-term tests.

The permissible carbon content is also an important consideration when it comes to making

decisions about the required lifetime of the catalyst and the costs associated with it. Considering the

catalyst as a consumable, as the proposed concept does, means incurring additional costs.

Equation 7.1 allows the calculation of the necessary catalyst runtime tOp. [h] in dependency of

catalyst mass for one replacement mCat.[g], the catalyst costs CCat. [€/kg], the required specific catalyst

costs cCat. [€ct./kWhSyngas] and the synthesis gas power PSG [kW]. The main variable in this equation is

the specific catalyst costs, which represent the additional costs for the production of SNG due to

catalyst replacement. Transforming this equation makes it possible to calculate the specific cost of a

catalyst from a given runtime.

35 mg/g 15 mg/g 10 mg/g

5 mg/g 1 mg/g 0 mg/g


89

[ ] [ ]

[

] [ ] [ ]

7.1

The acceptable carbon content after a certain time CC(t) (equation 7.2) depends on the maximum

carbon content acceptable CCmax [mg/g], the time [h]and the required runtime tOp. [h]. The examples

with values are just intended to show of which orders the values typically are.

( )

[ ] [ ]

[ ] [

]

7.2

Figure 7.10 shows the specific amounts of catalyst consumption and the related specific catalyst

costs in dependency of different maximum carbon contents and the amount of deposited carbon per

hour. The costs are based on catalyst costs of 70 €/kg.

Figure 7.10: Specific amounts of catalyst consumption and cost in dependency of different carbon contents

In summary it can be said that it is quite difficult to define a value for an acceptable amount of

carbon. While carbon contents < 10 µg/gCatalyst·h are low and can be considered as unproblematic for

the proposed applications, contents > 70 µg/gCatalyst·h are considered as too high in terms of catalyst

consumption costs. In figure 7.11 the different carbon contents are rated on a colored scale.

7.2.3. Influence on ethylene-promoted carbon deposition

The main variable influences on ethylene-promoted coking are temperature, water content, and, of

course, the ethylene concentration itself. To ensure that no carbon formation results from the

thermodynamic equilibrium, the water content of the synthesis gas was fixed with 40 vol. % to

ensure being outside of the thermodynamic equilibrium for carbon formation.

0

0.5

1

1.5

2

2.5

3

3.5

0

0.1

0.2

0.3

0.4

0.5

0 50 100 150 200 250

Cat

alys

t co

sts

[€ct

/kW

hSy

nga

s]

Cat

alys

t co

nsu

mp

tio

n [

g/kW

hSy

nga

s]

Carbon content [µg/g·h]

30 mg/gMax. acceptable carbon content

20 mg/g

10 mg/g

5 mg/g


90

Figure 7.11: Influence of temperature on the amount of deposited carbon,

based on a runtime of 22 h, GHSV 10000 h-1

The amount of deposited carbon depends heavily on temperature (figure 7.11 and figure 7.12).

Starting at a reactor inlet temperature of 300°C respectively a peak temperature of 460°C, carbon

formation increases and reaches its maximum at an inlet temperature of 330°C (475°C peak

temperature). Further rises in temperature up to 420°C (515°C peak temperature), however, lead to

a decrease in carbon deposition. There is confirmation of this also in the literature.

Bartholomew [113], for example, reported that the formation of polymeric carbon from ethylene

reaches its maximum at 430°C because the coking rate is a matter of the kinetics between formation

and removal of carbon. Unfortunately, a direct comparison with temperature values found in the

literature is not possible as those values are always based on isothermal conditions and the use of

different catalysts. Furthermore, the high temperature gradients in a polytropic reactor make it very

difficult to define a representative temperature for the kinetics of carbon formation. However, since

this work focuses on application-oriented rather than on fundamental investigations, variation of the

inlet temperature is a very convenient way of controlling temperature as the inlet temperature is

also the one variable parameter in large-scale applications.

Below 300°C inlet temperature the coking rate starts to increase again. This effect, which has not

been reported in the literature so far, might be due to the condensation of ethylene polymerization

products. However, since operation beyond an inlet temperature of 300°C is not intended, this effect

was not considered further in this study.

Figure 7.12 shows all points of the parameter study for ethylene-promoted carbon deposition.

Besides temperature the amount of ethylene present in the synthesis gas has a strong influence on

the amount of deposited carbon. No carbon deposition occurred beyond a C2H4 content of

0.3-0.35 vol. %. This corresponds to the experience of the catalyst manufacturer according to which

ethylene-induced coking starts only at a certain C2H4 level.

C2H4 contents above 0.7 vol. % were not investigated as the amount of carbon at 0.7 vol. % is already

too high for long-term operation of a reactor. High coking also occurred at lower contents of

280 300 320 340 360 380 400 420 440

Car

bo

nco

nte

nt

[µg C

arb

on/g

Cat

alys

t · h

]

0.5 vol. % C2H4

0.7 vol. % C2H4

Reactor inlet temperature [°C]

0

20

40

60

80

100

120

140

160

0

20

40

60

80

100

120

140

160

250 300 350 400 450 500 550

Reactor oven temperature [°C]

Low

Med

ium

Hig

h


91

0.5 vol. %. However, there are certain inlet temperature ‘windows’ – 300°C and above 400°C – at

which the amount of deposited carbon could be low enough to allow longer operation.

In all the tests ethylene was fully converted and no other C2-C4 hydrocarbons were detected at the

outlet of the methanation reactor.

Figure 7.12: Amount of deposited carbon in dependency of the temperature and the C2H4 content,


In summary it can be said that already relatively low amounts of ethylene (> 0.5 vol. %) lead to severe

coking; unfortunately, C2H4 contents in synthesis gas are usually above this level. The formation of

carbon can only be prevented by keeping the ethylene content sufficiently low and ensuring that

methanation happens at a convenient temperature. Otherwise alternatives need to be found that

will help to reduce or prevent coking (chapter 7.4) or will make it possible to regenerate the catalyst.

7.3. Parameter variations with representative tar mixtures

The investigations with tars were performed in the same way as the tests with ethylene. Numerous

short-term tests (22 hours) with varying parameters showed the various influences on catalyst coking

that occurs due to the presence of tars in the synthesis gas. A mixture of four different tar species –

benzene, toluene, phenol and naphthalene – was chosen to represent tar contaminations in

synthesis gas because they constitute the main tar components produced in gasification and because

they represent different properties of tars. These tars, which are also easy to dose, are often referred

to as representative tar species in the literature, e.g. [99], [96], [95], [52].

The standard tar mixture used consisted of 3.5 g/Nm³ of benzene, 1 g/Nm³ of toluene, 1 g/Nm³ of

naphthalene and 0.5 g/Nm³ of phenol, which adds up to 6 g/Nm³ in total. The ratio between the

components remained constant even if a higher total concentration was used. The parameters varied

in these investigations are temperature, the tar concentration and, contrary to the tests with

ethylene, also the water content. Water is one of the main influences on the reforming and

conversion of higher hydrocarbons. An increased water content enhances the reforming of

hydrocarbons and should reduce the propensity for coking.

0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1

250 300 350 400 450 500 550

Reactor inlet temperature [°C]C

2H

4co

nte

nt

[vo

l. %

]


280 300 320 340 360 380 400 420 440

60

133 54

26 16

0012

3

3910

20

8989

0 0 0

20 µgC/gCatalyst·h


92

Figure 7.13 shows amounts for carbon deposition based on 22 hour methanation tests in which the

6 g/Nm³ standard tar mixture was added and both synthesis gas water contents (between

30-40 vol. %) and reactor temperatures were varied.

It can be seen that a higher amount of water leads to tendentially lower amounts of deposited

carbon. Furthermore, higher temperatures result in greater coking of the catalyst. The only point not

showing this trend is the one at a water content of 40 vol. % and a reactor oven temperature of

320°C. The reason for that is probably the condensation and further polymerization of tar

compounds as this operating point was at the lowest inlet-zone temperature. Besides the inlet and

oven temperature, shown in the diagram, the reactor peak temperature and the resulting

temperature of the inlet-zone (according to figure 7.2) also have an impact on the conversion of

higher hydrocarbons. For the points at 30 and 35 vol. % H2O, the inlet-zone temperature was slightly

higher than shown in the diagram as lower amounts of water lead to higher reaction temperatures.

Figure 7.13: Amount of deposited carbon in dependency of the temperature and the H2O content,

syngas with standard tar concentration (6 g/Nm³), based on a runtime of 22 h, GHSV 10000 h-1

Figure 7.14 shows the amount of deposited carbon when using the standard tar concentration

(6 g/Nm³) and twice that concentration (12 g/Nm³). As expected, higher tar concentrations led to

greater coking of the catalyst. Furthermore, these results indicate more clearly the influence of the

operating temperature. Operation with inlet temperatures below 300°C result in higher amounts of

carbon being deposited on the catalyst, probably for the above reasons. Inlet temperatures above

400°C also increase coking.

However, by comparing these results with those obtained with ethylene, it can be seen that far less

carbon was deposited in the tests with tars than with ethylene. If suitable operating conditions are

chosen, low coking rates can be expected, which should allow long-term operation with tar-

contaminated synthesis gas.

25

30

35

40

45

300 350 400 450 500 550


H2O

co

nte

nt

[vo

l. %

]


300 320 340 360 380 400 420 440

17 4.5 5 6.5

1 4.5 9.5 11

1.5 6.5 12 13.5



93

Figure 7.14: Amount of deposited carbon in dependency of the temperature and the tar concentration,

syngas with H2O content of 40 vol. %, based on a runtime of 22 h, GHSV 10000 h-1

What happens with tars during methanation?

As the test have shown, tar contaminations of synthesis gas lead only to minor catalyst coking during

methanation. It is, however, also important to know if tars are really converted during methanation

or if they just pass through the reactor. As can be seen in figure 7.15, which shows tar conversions

for the standard methanation configuration based on tar concentrations measured using the

SPA method, tars were fully converted under typical methanation conditions. Minor amounts of

benzene and toluene were detectable at the reactor outlet only at higher temperatures.

Figure 7.15: Tar conversion during a methanation test with the standard catalyst filling (30g) and standard

tar concentration in dependency of the reactor temperature, GHSV 10000 h-1

0

2

4

6

8

10

12

14

16

300 350 400 450 500 550


Tar

con

cen

trat

ion

[g/

Nm

³]


300 320 340 360 380 400 420 440

17 4.5 5 6.5

25 13.5 12.5 24.5


320 370 450 550 700


0.90

0.91

0.92

0.93

0.94

0.95

0.96

0.97

0.98

0.99

1.00

450 500 550 600 650

Tar

con

vers

ion

[-]

Reactor peak temperature [°C]

Benzene

Toluene

Naphthalene

Phenol

Total


94

A test with a reduced amount of catalyst and increased tar concentration (figure 7.16) shows the

temperature influence on tar conversion even more clearly. The diagram indicates that the

methanation temperature has only a minor influence on the conversion of tars as the overall

conversion rate is between 96.5-98 % for all tested temperatures. Due to the low amount of catalyst

material used, the total catalyst bed height was only 1.5 cm, which, however, was enough to convert

> 96.5 % of the tars. Therefore these results represent the conversion that took place directly after

the inlet of the reactor.

Figure 7.16: Tar conversion during a methanation test with reduced catalyst filling (9 g) and 10 g/Nm³ of tar,

GHSV 33000 h-1

This high and fast conversion of tars is somehow contrary to what the literature says on tar

reforming. Many authors report that high temperatures – of up to 900°C – are needed for sufficient

conversion of higher hydrocarbons (chapter 3.3.3). Industrial steam reforming applications also

typically operate at outlet temperatures of up to 900°C [88]. Although the majority of the literature

claims that full reforming of tars at typical methanation temperatures is not possible, practice shows

that it is, at least with tars formed under steam gasification conditions.

Previous investigations [176] already analyzed the influence of methanation conditions on the

conversion of higher hydrocarbons (figure 7.17). In three test runs, performed for this study, 3 g/Nm³

of toluene were added to different gas compositions. The maximum temperature was constant in all

three tests. In the first test an H2/H2O mixture was used to represent reforming conditions; in the

second and third test run 6 vol. % and 18 vol. % of CO were added respectively to simulate

methanation conditions. Figure 7.17 clearly shows that the tar conversion rate is much higher when

CO is added rather than a mixture of H2 and H2O is used only. Similar results can be found in

publications by Vosecký [99], however, without any explanation of the reasons for this behavior.

The main difference between the tests with and without CO lies in the amount of reaction partners

that are available. Therefore the additional CO must somehow influence the kinetics of tar

conversion. One hypothesis is that the additional carbon or oxygen enhances the kinetics of tar

conversion, e.g. by faster dehydrogenation. However, which detailed mechanism is really at work

here cannot be clarified by this investigation as this would require additional, more detailed kinetic

studies.

0.90

0.91

0.92

0.93

0.94

0.95

0.96

0.97

0.98

0.99

1.00

450 500 550 600 650

Tar

con

vers

ion

[-]


Benzene

Toluene

Naphthalene

Phenol

Total

320 370 450 550 700



95

Figure 7.17: Influence of methanation conditions on the conversion of toluene; adapted from [176],

standard reactor setup with EVT01 catalyst, GHSV 10000 h-1

Methanation with simultaneously addition of C2H4 and tars

Since ethylene and tars are both present in gasification-derived synthesis gas, interactions between

them are likely. To prove these interactions methanation tests were carried out in which tars and

ethylene were added simultaneously. Figure 7.18 shows the effects this had on carbon deposition.

The determination of the conversion rate of higher hydrocarbons confirmed previous results, which

had shown that higher hydrocarbons are completely converted under the chosen operating

conditions. However, since the amount of deposited carbon resulting from simultaneous conversion

is not a simple addition of the amounts of separate conversions, an interaction between the different

contaminates is obvious.

Figure 7.18: Amount of deposited carbon resulting from methanation with simultaneous addition of C2H4 and

tars compared to separate addition, based on a runtime of 22 h, GHSV 10000 h-1

0.80

0.85

0.90

0.95

1.00

350 400 450 500 550 600

Tolu

en

e c

on

vers

ion

[-]


54/6/4042/18/40

60/0/40

H2/CO/H2O [vol. %]

0

20

40

60

80

100

120

140

160

250 300 350 400 450 500 550


0.5 vol. % C2H4

0.7 vol. % C2H4

280 300 320 340 360 380 400 420 440

0.7 vol. % C2H4 + Tar

6 g/Nm³ Tar

0.5 vol. % C2H4+Tar


condensation/polycondensation expected

Car

bo

nco

nte

nt

[µg C

arb

on/g

Cat

alys

t · h

]


96

At lower reactor inlet temperatures (≈330°C) the measured amount of deposited carbon for

coincident conversion of tars and C2H4 was lower than when C2H4 was converted alone. At higher

temperatures the opposite effect occurred.

Raising the C2H4 content from 0.5 to 0.7 vol. % caused the amount of carbon deposited from tars and

C2H4 to rise by the same amount as when C2H4 was converted separately. It therefore seems that the

magnitude of coking is determined mainly by the ethylene content, whereas the trend of increased

coking with higher temperatures results from the influence of tars.

Which mechanism causes these interactions was not investigated in the course of this work. Jess [96]

reported that during conversion of hydrocarbon mixtures some species hinder or reduce the

conversion of others. A similar effect may also reduce or increase the amount of carbon formed

during the conversion of a hydrocarbon mixture and, in doing so, produce the results described

above.

Although simultaneous conversion of C2H4 and tars can be beneficial for lower coking than in

methanation with conversion of C2H4 only, the coking rates obtained are probably still too high for

commercial applications. Further investigations into ways of reducing coking are therefore necessary.


97

7.4. Reduction of carbon deposition by addition of sulfur species

An observation made during bench-scale methanation tests with real synthesis gas led to the

assumption that sulfur somehow influences the formation of carbon on the catalyst. After having

made improvements on the desulfurization unit, severe coking occurred on the catalyst although

methanation tests under similar operating conditions had not shown this problem. Kienberger

et al. [100] reported that no coking occurred during methanation of H2S-loaded synthesis gas

although the syngas contained ethylene and tars in a concentration at which, according to the results

presented earlier in this thesis, carbon deposition was to be expected.

The coke-reducing effect of sulfur during steam reforming is well reported by Rostrup-Nielsen [129].

Sulfur chemisorbs and deactivates the catalyst. At low H2S concentrations delineated active zones

remain, which enable the reforming reaction but inhibit carbon formation reactions (chapter 4.2.2).

The same effect may lead to the coke-reducing behavior of sulfur observed during methanation.

However, despite the fact that sulfur may prevent carbon deposition, it is nevertheless a strong

poison for nickel, which is why even minor amounts of sulfur will deactivate the catalyst. Therefore

the use of sulfur to reduce coking makes only sense if the degree of catalyst degradation due to

sulfur is smaller than the amount of coking or if the gas cleaning effort can be reduced.

Catalyst deactivation by sulfur

To determine catalyst degradation from sulfur, methanation tests with addition of H2S were

performed. Figure 7.19 shows the degree of catalyst deactivation and the amount of adsorbed sulfur

that was measured in these tests. The active catalyst area was determined by measuring the axial

temperature profile along the whole length of the reactor. Due to deactivation the temperature of

the reactor inlet zone decreases (figure 4.12). Therefore, the area beyond the temperature graph

reduces as well with ongoing deactivation. If the temperature graph equals the graph of the inert

temperature, a state of full deactivation has been reached.

The H2S concentration was 14 ppm for the first 52 hours, 130 ppm between hour 53 and 116, and

14 ppm again after that. It can be seen that deactivation at 130 ppm is faster than at 14 ppm, but

that H2S is not adsorbed in the same ratios. This leads to the assumption that the specific amount of

catalyst deactivation, e.g. gCatalyst/gH2S, is also a matter of overall sulfur concentration. Which

mechanism causes these results cannot be clarified. One possible explanation is that the diffusional

restrictions inside the catalyst pellets slow down the adsorption rate [177] so that at higher

concentrations sulfur adsorbs in greater amounts on the surface of the catalyst pellets and thus does

not deactivate the inner part of the catalyst. The results also indicate that the deactivation rate

decreases with high amounts of sulfur and high deactivation of the catalyst. This may be due to sulfur

adsorption being spread more widely across the reactor, which leads to a smaller temperature

decrease. However, this effect will not be pursued any further as such high degrees of catalyst

deactivation are not suitable for methanation anyway.


98

Figure 7.19: Measured catalyst degradation from poisoning with H2S (14 and 130 ppm) for EVT05

From the measured deactivation rates it is possible to calculate the specific amount of catalyst

consumption and, consequently, the specific catalyst cost, which is the additional cost of catalyst

replacement due to deactivation of the catalyst. Figure 7.20 shows the amount of catalyst

consumption and the specific cost of catalyst degradation based on the results from methanation

tests with poisoning with 14 ppm H2S. Due to the above-mentioned influence of the sulfur

concentration on the deactivation rate, the values may differ somewhat for lower H2S

concentrations. However, the magnitude of catalyst consumption should be represented well.

Figure 7.20: Specific catalyst consumption and related specific catalyst cost due to poisoning with H2S,

determined with 14 ppm H2S for EVT05; estimated catalyst cost of 70 €/kg

If the H2S concentration is low, the resulting catalyst consumption and cost might well be in an

acceptable range for medium and small-scale applications. If the H2S concentration in the feed is

0.5 ppm, the deactivation of the catalyst results in a catalyst consumption in a range of

0.015-0.02 g/kWhSyngas.

0

0.2

0.4

0.6

0.8

1

1.2

1.4

1.6

0.0

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

0.9

1.0

0 50 100 150 200

Ad

sorb

ed

H2S

[g]

No

rmal

ize

d a

ctiv

e c

atal

yst

are

a [-

]

Runtime [h]

130 ppm 14 ppm 14 ppm H2S

0.0

0.5

1.0

1.5

2.0

2.5

0

0.1

0.2

0.3

0.4

0 1 2 3 4 5 6 7 8 9 10

Cat

alys

t co

sts

[ct/

kWh

Syn

gas]

Cat

alys

t co

nsu

mp

tio

n [

g/kW

hSy

nga

s]

H2S concentration [ppm]


99

Influence of sulfur on carbon deposition

To investigate the influence of sulfur in the feed, the same procedure as in the methanation tests

with C2H4 and tars was applied; additionally, however, 0.25 to 1 ppm of H2S was added to the

synthesis gas. The tests were performed at a reactor oven temperature of 370°C, where the highest

coking had occurred during the tests with C2H4. This operating point was therefore of particular

interest concerning the impact of sulfur. As can be seen in figure 7.21, which shows the effects of

adding C2H4 and H2S, the addition of as little as 0.25 ppm of H2S already leads to a significant

reduction in the amount of deposited carbon, and even with 1 vol. % of C2H4 almost-carbon-free

methanation was possible by addition of 1 ppm of H2S.

Figure 7.21: Influence of C2H4 and H2S on the amount of deposited carbon at 370°C reactor oven temperature,


Since C2H4 and tars are normally present simultaneously in gasification-derived synthesis gas, the

influence the addition of H2S has on such gas compositions was also investigated. For that purpose

the standard tar mixture with a total of 6 g/Nm³ of tar was added, along with C2H4, to the syngas.

Figure 7.22: Influence of a C2H4, tars and H2S on the amount of deposited carbon at 370°C reactor oven

temperature, based on a runtime of 22 h, GHSV 10000 h-1

0.2

0.4

0.6

0.8

1.0

1.2

-0.2 0 0.2 0.4 0.6 0.8 1 1.2

C2H

4co

nte

nt

[vo

l. %

]


605 3.5

133

17 8.5 9

2.5


0.2

0.4

0.6

0.8

1.0

1.2

-0.2 0 0.2 0.4 0.6 0.8 1 1.2

C2H

4co

nte

nt

[vo

l. %

]


19

99 62

29 8

506 187 61

C2H4 + 6 g/Nm³ Tar 20 µgC/gCatalyst·h


100

Figure 7.22 shows the measured carbon contents for methanation tests with different C2H4 and H2S

concentrations. These results also confirm the coking-reducing property of sulfur. However,

compared to the results with C2H4/H2S only (figure 7.21), this effect is less strong. At a C2H4 content

of 0.7 vol. %, 6 g/Nm³ tars and an H2S concentration of 0.5 ppm the carbon deposition was 8 µg/g·h,

which could be sufficiently low to allow long-term operation.

To also prove the influence of H2S over a longer runtime, long-term tests were performed as well. In

one such test with 0.7 vol. % of C2H4, 6 g/Nm³ of tar and addition of 0.5 ppm of H2S over a runtime of

270 hours, 2.19 mg/g of carbon was deposited. In a test without addition of H2S, this amount was

already reached after 22 hours. The differential pressure across the reactor was constant for the

whole runtime. Due to the addition of sulfur 10 % of the catalyst was deactivated, which corresponds

to a specific catalyst consumption of around 0.03 g/kWhSyngas. This degree of deactivation is higher

than can be explained by the effect of sulfur alone (figure 7.20). However, the results show the great

potential that lies in the addition of sulfur to the synthesis gas for the reduction or even prevention

of coking.

7.5. Visual evaluation of carbon deposits

Besides quantitative evaluation, a visual evaluation provides additional information on the type and

consistency of carbon deposits. Therefore selected samples were analyzed by means of microscopy.

Figure 7.23: States of polymeric carbon coverage on a catalyst pellet

Figure 7.23 shows the different states of polymeric carbon coverage on a catalyst pellet. Depending

on the intensity of coking, the pellet can be covered only partially, or fully, or fully covered with

agglomeration between the individual catalyst pellets.

Figure 7.24: SEM photos of polymeric carbon deposits on the catalyst resulting from C2H4 after (a) 22 h,

(b) 70 h runtime; operating conditions: 320°C reactor oven temperature, 40 vol. % H2O, 0.5 vol. % C2H4

Clean Ni-catalyst Partial, even carbon deposition

Full, evencarbon deposition

Full, agglomerated carbon deposition

a) b)200 µm 100 µm


101

Due to the temperatures of the catalyst occurred in the tested methanation concept only polymeric

carbon deposits were expected (chapter 4.2.1). This assumption was confirmed by the samples

analyzed, in which no graphitic carbon was found. Figure 7.24 shows SEM photos of carbon deposits

formed from ethylene after different runtimes. After shorter tests (runtime of 22 h), layers of carbon

can be found on the surface of the catalyst. Photos taken with a higher resolution showed that these

deposits consist mainly of carbon filaments (figure 7.25) and minor amounts of carbon layers (figure

7.26). After a longer runtime (70 h), areas with pitting/erosion of the surface were found; these were

covered with polymeric carbon deposits, which indicates that in these areas catalytic material was

removed by filamentous carbon deposits. Pitting could be found only on one sample. However, all

analyzed samples with carbon deposits promoted by C2H4 and C2H4/tar mixtures contain mainly

filamentous carbon. It can therefore be assumed that other catalyst samples are also affected by

catalyst destruction, which is important if regeneration of a catalyst becomes an issue.

The samples which had H2S added to synthesis gas containing C2H4 and tars showed lower numbers

of carbon filaments but more carbon layers and other unstructured, amorphous deposits.

Since only a small number of samples were analyzed by means of SEM, it is important to note that

the results presented here are not necessarily representative for all samples.

Figure 7.25: SEM photos of polymeric carbon filaments resulting from C2H4

Figure 7.26: SEM photos of polymeric carbon layers resulting from C2H4

10 µm 400 nm

2 µm 400 nm


102

7.6. Summary and conclusion bottle-mixed syngas tests

Carbon deposition, which is promoted by different contaminations like ethylene and tars, can cause

severe problems during methanation. While the main focus of the tests was on a quantitative

analysis of carbon deposition, catalytic activity and conversion of the different contaminations were

also considered.

The results of the investigation into coking behavior show that coke accumulates linearly with the

runtime under the specified operating conditions, which makes it possible to partially replace long-

term tests with extrapolation of shorter tests.

The addition of C2H4 in amounts higher 0.3-0.35 vol. % to clean synthesis gas leads to severe

formation of carbon on the catalyst. The amount of deposited carbon depends on the reactor

temperatures and the ethylene content; it increases strongly with the amount of C2H4 added, and

increases and then decreases with rising temperature.

The addition of a representative tar mixture to the clean synthesis gas also leads to carbon

formation. Higher syngas water contents and lower temperatures reduce the coking rate. Tars cause

considerably less coking than ethylene. Therefore, the direct conversion of tars during methanation

can be assumed as being less problematic than the conversion of ethylene. In all methanation tests

performed, ethylene and tars were fully converted under typical methanation conditions.

Simultaneous conversion of C2H4 and tars showed that the magnitude of carbon deposition is

determined by the C2H4 content, while tars give the trend to stronger coking at higher temperatures.

Since typical concentrations of ethylene and tars in synthesis gas are in a range where severe coking

can be expected, ways to reduce coking need to be found. One such option could be the addition of

certain sulfur species, like H2S, to the synthesis gas, which was successfully tried in several

methanation tests carried out in the course of this investigation. The results clearly show that minor

amounts of H2S added to the syngas can significantly reduce coking, thus proving the great potential

of this method.

Bench-Scale Tests with Real Syngas from Gasification

103

Chapter 8

8. Bench-Scale Tests with Real Synthesis Gas Produced in

Allothermal Gasification

8.1. Investigation focus and program

The aim of the bench-scale tests was to apply the proposed concept of methanation using real

synthesis gas produced in thermal gasification. The focus of the tests was on the performance of the

catalyst under the realistic conditions of usage of contaminated synthesis gas. In the method

proposed some of the contaminations, such as particles, alkalis and sulfur species, are removed by

the hot gas cleaning unit prior to methanation whereas higher hydrocarbons remain in the syngas

and are converted during methanation.

Before five long-term tests with runtimes of up to 200 hours were performed, an existing, indirectly

heated gasifier was modified and connected with a newly built gas cleaning unit and a bench-scale

methanation reactor. During the tests gas compositions and contaminations were measured at all

stages of the process. The main test results are catalyst degradation rates, which were evaluated on

the basis of the temperatures measured in the reactor. Further important results gained are

conversion rates of higher hydrocarbons during methanation and removal efficiency of the hot gas

cleaning unit. After each test the amount of carbon deposited on the catalyst was measured using

the TPO method. The results thus obtained can serve as a basis for further process improvements

and for the design of large-scale concepts.

8.2. Test rig assembly and setup

8.2.1. Test rig assembly

The test rig assembly for the bench-scale tests consists of an indirectly heated, fluidized bed gasifier,

a hot gas cleaning unit and a methanation reactor. An additional gas mixing station provides bottle-

mixed gases for reducing of the catalysts (figure 8.1). Different analysis systems measure gas

compositions and contaminations at different points of the process.


104

Figure 8.1: Photo of the bench-scale test rig for SNG production with real synthesis gas from gasification

Lab-scale gasifier

The indirectly heat fluidized bed gasifier (figure 8.2) is used to produce a realistic synthesis gas from

lignite and biomass. It was constructed and modified in the course of various previous studies [10],

[178]. The nominal fuel power is 5 kW, but it typically operates at 1-2 kW. The whole system is

designed for pressures up to 4 bars. To overcome the pressure losses of the downstream parts,

operation at an overpressure of 0.5-1 bar is sufficient.

The gasification reactor consists of a bubbling fluidized bed, which has an inner diameter of 60 mm

and a length of around 150 mm. A well-dimensioned freeboard prevents the excessive discharge of

bed material while providing enough resistance time to allow a high amount of coke to be converted,

which is important for long-term operations. The main bed material is olivine (Magnolithe, Austria)

with a medium grain size of 250 µm. Water steam is used for fluidization of the bed and as a

gasification medium; the steam is generated by a commercial steam generator as used in

conventional steam stations for irons. The steam flow is measured and regulated by means of an

orifice flow measurement and a proportional valve. Additionally, it is also possible to use N2 for

fluidization. An electrical tube furnace heats the reactor up to 850°C to provide the heat required for

the endothermic gasification reactions.

Gasifier

Methanation andgas-cleaning test rig

Gas mixing station


105

Figure 8.2: Flow sheet of the indirectly heated, fluidized bed gasifier

The used fuel-feed system was newly designed for the tests within this work and replaces the

previously used fuel input described in [10]. The new system consists of a screw conveyor, which

doses the fuel from a fuel reservoir, and a feed-lock system to pressurize the fuel and feed it into the

fluidized bed. The feed-lock system also prevents air from getting into the reactor and blocks the

release of gasification gases from the reactor by purging with N2. It consists of three pneumatic ball

valves which alternately open and close to transport the fuel into the reactor. Purging and

pressurization takes place between the two ball valves fitted close to the reactor. Due to the

programmed sequence, fuel drops into the reactor every 15 seconds. For safety reasons, the lowest

ball valve only opens if the pressure above it is slightly higher than the reactor pressure and if the

two upper ball valves are closed.

A cyclone filter and a sinter metal filter remove particles from the gas. Since the sinter metal filter

operates at around 350°C, alkalis also condensate on the filter cake. A jet-pulse filter-cleaning system

removes the filter cake when the differential pressure of the filter becomes too high. A pressure-

retaining valve at the outlet of the gasification system keeps the system pressure constant. The

synthesis gas leaves the reactor with a temperature of around 350-400°C.

A large number of pressure sensors and thermocouples enable the monitoring of all important

operating parameters. The use of a Bernecker&Rainer (B&R) industrial control system makes the

whole assembly fully automatized and consequently allows unsupervised operation, which is

important for long-term testing.

PD

F


PD

T

PD

Cyclone filter

Sinter metal filter

Automatic filter cleaning

Differential pressure sensor (filter)

Feed-lock system

N2-purging and pressurizing

Differential pressure sensor (reactor)

Orifice flow measurement

Steam regulation

valve

Steam generator

Pressure retaining

valve

Demineralization

Reactor oven with fluidized bed reactor

P

Pressure sensor (steam)

P Pressure sensor (reactor)

P Pressure sensor (N2)

P

Pressure sensor

Fuel reservoir

Screw conveyor

Purging

N2

N2

5 thermo-couples

SPE

Port for SPE-sample

Water

Exhaust

Pressurized air

Exhaust

Pressurized air

Outlet / to methanation

test rig


106

Gas cleaning and methanation unit

Trace-heated lines connect the gasifier with the gas cleaning and methanation unit (figure 8.3). The

adsorptive hot gas cleaning unit consists of two tubular fixed bed reactors. The larger reactor 1 has

an inner diameter of 54 mm, a length of 800 mm and a total useable volume of around 1.7 liters.

Reactor 2 has an inner diameter of 34 mm, a length of 600 mm and a volume of around 0.5 liters.

Both reactors are placed in reactor ovens which can be heated up to 700°C (larger reactor) and 450°C

(smaller reactor). A 6 mm tube in the center of the reactors supports axially displaceable

thermocouples for measuring various reactor temperatures.

Both the methanation reactor and the gas mixing station are the same as those used in the bench-

scale methanation tests with clean synthesis gas (chapter 5.1).

Figure 8.3: Flow sheet of the bench-scale hot gas cleaning and methanation unit

A large number of valves enable a variable interconnection of the different reactors. This is necessary

especially for the start-up procedure or when one of the gas cleaning reactors needs to be replaced

during operation. A gas meter is used to measure and control the volume flow of the methanation

reactor. Since the gas contains a large amount of water, a condenser and a silica gel adsorber remove

it before it reaches the gas meter. After each reactor, sample ports allow the taking of gas samples

for the micro-GC and SPA samples; these ports are also directly connected with the permanent gas

analyzer. Therefore the full spectrum of measuring techniques can be applied at all stages of the

process.

T T T

PD

Flare

Natural gas

SPE

Port for SPE-

sample

SPE

Port for SPE-

sample

SPE

Port for SPE-

sample

16 thermo-couples

Dif

fern

enti

al p

ress

ure

sen

sor

2 thermo-couples

2 thermo-couples

Hot gas cleaning 2

Methanation reactor

Hot gas cleaning 1

Pre

ssu

rize

d a

ir

(rea

cto

r co

olin

g)

To GA / gas sample bag

Condenser

Silica gelGas meter

From gasifier

Trace heated lines

3 z

on

e h

eati

ng

Volumeflow measurement unit



From gas mixing station


107

Gas analysis unit

The gas analysis unit for the bench-scale methanation tests with real synthesis gas (figure 8.4)

consists of a permanent gas analyzer and a tar-sampling unit constructed according to the tar

protocol. Before the gas reaches the gas analyzer, washing bottles with IPA remove tars from the

synthesis gas as they would damage the gas analyzer. Solenoid valves enable switching between

different sample ports (gasifier, gas cleaning and methanation).

The tar-sampling unit takes a slip stream coming from the gasifier, but it is also possible to connect it

with the sample ports of the different reactors.

Further analysis equipment and techniques used are the micro-GC (for analysis of contaminations),

detector tubes, and the SPA method (for tar sampling). Chapter 6.2.3 provides a detailed description

of the different analysis methods.

Figure 8.4: Flow sheet of the gas analysis unit for methanation and gasification tests

8.2.2. Test setup and operating conditions

Fuel

The gasification was performed with biomass and lignite. Standard wood pellets according to the

ENplus-A1 standard were used as a biogenic fuel. The lignite was of the high-quality type

RWE PowerSPLIT, which is normally used in fluidized bed applications. After crushing, the grain size

was between 2-4 mm, which is ideal for the feed-system of the gasifier. Table 8.1 shows the main

parameters of the two fuels. Tests 1-4 of the long-term tests presented in the next section were

performed with lignite, biomass was used for test 5.

Lignite was used for two reasons: first, because this research was initially part of a European coal

research project CO2freeSNG [26] and funded by it, and, secondly, because lignite represents a kind

of ‘worst-case biomass scenario’. The gasification properties as well as the gas qualities that can be

reached using the chosen type of lignite are similar to those of biomass, the main difference being

the higher amount of sulfur contaminations caused by lignite. Thus, if the applied gas cleaning

concept works with lignite-derived synthesis gas, it should work with all kinds of biomass-derived

gases.

Activated carbon

Condenser unit

Pump unit

From gasifier

Exhaust Tar sampling according tar protocol

Washing bottlesVolume-sampling module

From methanation

From gas cleaning 1

From gas cleaning 2

-20°C +20°C

-20°C +20°C

Gas analyzer H2, O2

Gas analyzer CO, CO2, CH4


108

Table 8.1: Fuel parameters for the used lignite and biomass

Wood pellets (ENplus-A1)

RWE PowerSPLIT [179]

C [wt. %] 47.6 53.6 H [wt. %] 5.8 3.9 O [wt. %] 39.0 19.2 N [wt. %] < 0.2 0.6 S [wt. %] 0.04 0.35 H2O [wt. %] 6.9 19.0

Ash [wt. %] 0.47 3.5 Fixed carbon [wt. %] 17.4 [180] 35.5 Volatiles [wt. %] 74.2 [180] 42 LHV [kJ/kg] 18100 19800

xH2O,min [kgH2O/kgFuel,Wet] 0.208 0.431

Gasifier operating conditions

The main purpose of the gasifier is to provide a constant and representative flow of synthesis gas.

Therefore, the operating conditions were set accordingly (table 8.2). The main requirement for the

synthesis gas is a water content in the range of 40 vol. % and a constant gas flow. Operation at low

fuel inputs ensures long-term operation of the gasifier due to high carbon conversion rates.

Furthermore, the GHSV of the methanation reactor can be kept lower at reduced fuel inputs as the

methanation reactor is connected in-line with the gasifier.

Table 8.2: Operating parameters for the real gas methanation tests

Lignite Biomass Test 1 Test 2 Test 3 Test 4 Test 5

Bed Temp. [°C] 770-820 770-820 770-820 770-820 790-810 Fuel input [kW] 0.7-1 0.7-1 0.7-1 0.7-1 1-1.3 Pressure [bar] 0.5 0.5 0.5 0.5 0.5 Steam flow [kg/h] 0.35-0.4 0.34-0.38 0.33-0.37 0.33-0.37 0.22-0.28

Sorbent Desulf. 1 ZnO ZnO ZnO ZnO ZnO Temp. Desulf. 1 ≈300°C ≈300°C ≈300°C ≈300°C ≈300°C Sorbent Desulf. 2 AC AC GS6+AC GS6 - Temp. Desulf. 2 ≈300°C ≈300°C ≈300°C ≈300°C -

Catalyst Meth. EVT01 EVT05 EVT05 EVT05 EVT05 Inlet-Temp. Meth. 275-300°C 275-300°C 275-300°C 275-300°C 275-300°C GHSV Meth. [h-1] ≈3000 ≈2700 ≈2500 ≈2500 ≈2500

Gas cleaning operating conditions

The two hot gas cleaning reactors can be filled with different types of sorbents. Zinc oxide is the most

common adsorbent for hot removal of H2S. The commercial ZnO sorbent Clariant/Südchemie

ActiSorb S2 had performed well in previous desulfurization tests and was therefore used in the first

reactor. The temperature of the ZnO adsorber was around 300°C. At that temperature the

equilibrium H2S concentration is below 0.1 ppm.


109

The second reactor was used for testing various other sorbent types. ZnO is probably not able to

remove organic sulfur species completely. Promising alternatives to it are different impregnated

activated carbons and a sorbent on basis of copper-/manganese oxide. Since those materials also

showed hydrodesulphurization activity, a short ZnO bed was placed after them. The filling of the

second reactor for the five long-term tests presented in the next section was chosen accordingly:

Test 1: ROZ3 (AC), Test 2: ROZ3 (AC), Test 3: FCDS-GS6 (CuO/MnO) + ROZ3 (AC), Test 4: FCDS-GS6

(CuO/MnO), Test 5: empty. Chapter 3.2 provides more information about the sorbent materials used.

To prevent condensation of tars the operating temperature was set to around 300°C.

Methanation operating conditions

Due to its promising results during the bench-scale methanation tests with clean synthesis gas, the

catalyst EVT05 was also used for the tests with real gas, except for test 1, in which EVT01 was used.

The inlet temperature for the methanation tests was between 275 and 300°C, which is a few degrees

lower than the results of chapter 7 would suggest. Therefore, future tests will operate with inlet

temperatures of 300-330°C. Since the focus of the tests was on the degradation behavior of the

catalyst, the reactor was not actively cooled. Active cooling would influence the temperature

profiles, which are the basis for the evaluation of the degradation. The typical outlet temperatures

were in the range 350-420°C. The methanation reactor was typically operated with a GHSV between

2500-3000 h-1.


110

8.3. Results

The following section presents the results of gasification, hot gas cleaning and methanation obtained

in five long-term tests. Tests 1-4 were performed with lignite, whereas wood pellets were used in

test 5.

8.3.1. Gasification

Since the main purpose of gasification was to produce a representative synthesis gas, the main

results are gas compositions and contaminations of the synthesis gas. Parameter studies on

gasification were not performed within this thesis, but a large number of parameter variations were

carried out in previous works [181], [155], [182] [178].

Figure 8.5 shows average permanent gas compositions on a dry basis and without N2 for synthesis

gas produced by gasification of wood pellets and lignite. In practice gasifiers always have to deal with

fluctuating gas compositions. The gasifier used in the tests of this investigation also showed certain

fluctuations in gas compositions, which were mainly due to a non-steady fuel input and variations of

the bed temperature. Therefore, typical ranges of gas compositions are shown additionally.

However, some fluctuations may even exceed those ranges.

Figure 8.5: Mean permanent gas composition from gasification of woody biomass and lignite (dry, N2-free),

gasification temperature ≈800°C, σ ≈6 (biomass), σ ≈4 (lignite)

The permanent gas compositions of biomass and lignite are fairly similar. Due to the higher amount

of volatiles in biomass, its gasification produces a higher amount of CH4 and therefore a lower

amount of H2. The mean gas composition of biomass gasification corresponds well to the standard

synthesis gas composition defined and used for the methanation tests described in previous

chapters. Since this standard syngas gas composition is based on equilibrium calculations, the

measured gas composition from gasification is in or close to equilibrium (except CH4). This

equilibrium-like gas composition results from operation at low fuel power and therefore long times

of residence of the gas in the gasification reactor. Average gas residence times in the fluidized bed

are around 1.6 s, whereas average residence times in the whole reactor are around 30 s, which is due

to the large freeboard of the gasifier.

0

10

20

30

40

50

60

H2 CO CO2 CH4

Gas

co

mp

osi

tio

n [

vol.

%]

Lignite (left bar)

Woody biomass (right bar)

H2 CO CO2 CH4

Typical range


111

Figure 8.6 presents the synthesis gas compositions on a wet basis including N2. Nitrogen is produced

by purging of the fuel input. During the tests with biomass modifications of the feed-lock system

allowed a reduction of N2 purging so that lower amounts were present in the syngas. The medium

water content of around 42 vol. %, was slightly higher than the water content of the standard syngas.

Figure 8.6: Mean permanent gas composition from gasification of woody biomass and lignite,


Besides the desired syngas components, synthesis gas also contains certain amounts of higher

hydrocarbons, the main species found being C2H4 (around 0.8 vol. % for biomass and 0.35 vol. % for

lignite – figure 8.7). The higher amount of volatiles in biomass led to the formation of higher amounts

of hydrocarbons in biomass gasification, in analogy to the higher CH4 content.

Figure 8.7: Mean C2-C4 content from gasification of woody biomass and lignite,


A comparison of the C2-C4 amounts with results from the Agnion HPR pilot plant shows that they

match well for biomass [7] as well as for gasification of lignite [26]. Compared to the Güssing

FICFB gasifier [9], the measured C2-C4 amounts are 3-4 times lower. This could be due to the longer

0

10

20

30

40

50

H2 CO CO2 CH4 N2 H2O

Gas

co

mp

osi

tio

n [

vol.

%]

Lignite (left bar)


H2 CO CO2 CH4 N2 H2O

Typical range

0.001

0.01

0.1

1

C2H4 [vol. %]C2H6 [vol. %]C3H6 [vol. %]C3H8 [vol. %]C4H10 [vol. %]

Gas

co

mp

osi

tio

n [

vol.

%]

Lignite (left bar)


C2H4 C2H6 C3H6 C3H8 C4H10

Typical range


112

residence time of gas in the lab-scale gasifier and different temperatures compared to the Güssing

plant.

Apart from non-condensable hydrocarbons certain amounts of condensable hydrocarbons (tars), are

also produced during gasification. Figure 8.8 shows the main tar components and their average

concentrations measured in the syngas. The values are based on averaging of results determined by

means of the SPA method and the tar protocol. In addition, BTX concentrations were measured in

gaseous state with the micro-GC.

The total average tar concentrations were 10.8 g/Nm³ for biomass and 5.4 g/Nm³ for lignite. This

total tar concentration includes BTX concentrations of 5.9 g/Nm³ for biomass and 3.8 g/Nm³ for

lignite. The lower tar amount of lignite was expected as it contains a lower amount of volatiles. The

measured tar concentrations compare well with the results of large-scale allothermal fluidized bed

gasifiers, like the Güssing gasifier [9] or the HPR plant [7].

Figure 8.8: Mean tar concentrations from gasification of woody biomass and lignite,


Figure 8.9 shows the mean contaminations measured in synthesis gas produced by biomass and

lignite gasification. Due to the higher sulfur content of lignite its conversion leads to the production

of much higher amounts of gaseous sulfur species than when using biomass. The main sulfur species

formed from biomass, H2S, has an average content of around 12 ppm. This is several times lower

than in the Güssing gasifier and results from the usage of wood pellets, which contain low amounts

of sulfur, instead of wood chips. Similar H2S contents, in the range of 10-23 ppm, were reported for

the gasification of wood pellets in the HPR plant [7]. However, since tests were also performed with

lignite, the whole concept was also tested with syngas with a high sulfur content.

The only measurable organic sulfur species were COS and CS2. Unfortunately, it was not possible to

measure other organic sulfur species, like thiols or tiophene.

0.01

0.1

1

10

Be

nze

ne

Tolu

ene

Xyl

en

e (

m, p

, o)

Ind

ane

Ind

ene

Nap

hth

ale

ne

2-M

eth

yln

aph

thal

en

e

1-M

eth

yln

aph

thal

en

e

Bip

he

nyl

Ace

nap

hth

ylen

e

Ace

nap

hth

ene

Flu

ore

ne

Ph

enan

thre

ne

An

thra

cen

e

Flu

ora

nth

ren

e

Pyr

ene

Ph

eno

l

Cre

sol (

m, p

, o)

Xyl

en

ol

Tar

con

cen

trat

ion

[g/

Nm

³]

Lignite (left bar)



113

Figure 8.9: Mean contaminations from gasification of woody biomass and lignite,


8.3.2. Adsorptive hot gas cleaning

The main purpose of adsorptive hot gas cleaning was the removal of sulfur contaminations without

influencing the other synthesis gas components. Contaminations were measured after the first

reactor, which contained ZnO only and after the second reactor, which contained, depending on the

test, ZnO, activated carbons or CuO/MnO sorbents. The comparison of measured contaminations

after gas cleaning of fully contaminated syngas (figure 8.10) shows that all measurable sulfur species

were below the detection limit of around 0.1 ppm for H2S and 0.2 ppm for COS. The complete

removal of sulfur was both independent of the type of fuel used and of whether ZnO was used with

other sorbents or alone. It can therefore be claimed that under the applied conditions ZnO

(ActiSorb S2) allows the removal of H2S and COS to an extent which is sufficient for catalytic

applications. The permanent gas compositions, as well as other contaminations and tars, were not

measurably influenced by hot gas cleaning.

Figure 8.10: Comparison of the mean contaminations resulting from gasification of lignite before and after

hot gas desulfurization with ZnO at ≈300°C, GHSV gas cleaning ≈ 500 h-1

0.1

1

10

100

1000

H2S [ppm] COS [ppm] CS2 [ppm] NH3 [ppm] HCl

Gas

co

mp

osi

tio

n [

pp

m]

Lignite (left bar)


H2S COS CS2 NH3 HCl

Typical range

0.1

1

10

100

1000

10000

H2S [ppm] COS [ppm] NH3 [ppm] BTX C2-C4

Gas

co

mp

osi

tio

n [

pp

m]

Gasification Gas Cleaning

H2S COS NH3 BTX C2-C4


114

8.3.3. Methanation

The evaluation of the methanation tests is shown on basis of the results of test 5, which represents

the results of the other tests very well.

Figure 8.11 shows the trend of the permanent gas composition on a dry basis at the outlet of the

methanation reactor. The fluctuations of the trend are due to fluctuations in the gas production of

the gasifier. Greater variations or interruptions result from sampling of gas or tar. The raw-SNG

contains a high amount of unconverted H2, which would make it unsuitable for feed-in into the gas

grid. This high H2 content results from the high outlet temperatures of around 400°C, which are due

to the fact that the reactor was not cooled during these tests. Comparing the measured gas

composition with the equilibrium composition related to the outlet temperature, full

conversion/yield is reached at the reactor outlet.

Figure 8.11: Trend of the permanent gas composition after methanation for test 5,

reactor outlet temperature ≈400°C, GHSV ≈2500 h-1

One of the main results obtained is information about the amount of higher hydrocarbons present

after methanation. Since no C2-C4 hydrocarbons were detected at the outlet of the methanation

reactor (the detection limit for these species being around 10 ppm), it can be assumed that they

were fully converted. This confirms the results of the tests performed with bottle-mixed synthesis

gas, in which there was full conversion of ethylene within the first 1-2 centimeters of the reactor.

Figure 8.12 shows measured tar concentrations after methanation as well as the related tar

conversions for the main tar species. The amounts of tars after methanation are at the limit for

detection. Benzene and toluene were not detected after methanation. However, as the detection

limit for these species is 40 mg/Nm³ for benzene and 15 mg/Nm³ for toluene, the BTX conversion

rate is at least 99 %, perhaps even as high as 99.95 % might be possible. The conversion of heavier

tars (tars without BTX) is > 99.4 %. This leads to a total conversion rate of at least 99.2 % up to

99.8 %. Kienberger [10] measured a total tar conversion during methanation of 97.9 %, which is in

the same order as the results found in this investigation.


115

Figure 8.12: Measured tar concentration after methanation and the related tar conversion for test 5,

reactor outlet temperature ≈400°C, GHSV ≈2500 h-1

During methanation the trend of the differential pressure of the reactor (figure 8.13) provides

important information: an increased value indicates blockage of the reactor due to severe coking.

Neither in test 5 nor in the other tests was such an increase observed.

The fluctuations in differential pressure are due to gas flow fluctuations. Since the correlation

between differential pressure and gas flow is quadratic, even minor gas flow variations result in

considerable variations of the pressure.

Figure 8.13: Trend of the differential pressure across the methanation reactor over the runtime for test 5

0

10

20

30

40

50

60

Be

nze

ne

Tolu

ene

Xyl

en

e

Ind

ane

Ind

ene

Nap

hth

ale

ne

2-M

eth

yln

aph

t.

1-M

eth

yln

aph

t.

Bip

he

nyl

Ace

nap

hty

len

e

Ace

nap

hte

ne

Flu

ore

ne

Ph

enan

thre

ne

An

thra

cen

e

Flu

ora

nth

ren

e

Pyr

ene

Ph

eno

l

Cre

sol

Xyl

en

ol

0.80

0.85

0.90

0.95

1.00

Tar

con

cen

trat

ion

[m

g/N

m³]

Tar

con

vers

ion

[-]


116

Coking during the test with real synthesis gas

The differential pressure trend indicates that the coking which occurred is not severe enough to

cause a blockage of the reactor voids. To quantify the amount of deposited carbon, catalyst samples

from different points of the reactor were analyzed by means of the TPO method. The results (shown

in figure 8.14) confirm previous findings of this investigation: most of the coking occurred in the inlet

zone of the reactor, and no carbon deposits were found on the catalyst far after the inlet zone; i.e.

most of the reactor was free of carbon deposits.

The amount of carbon deposited in the inlet zone is high compared to the tests with bottle-mixed

synthesis gases, where carbon contents above 5 mg/g had led to a significant rise of the differential

pressure, whereas in the real-gas tests carbon contents of 35 mg/g did not cause an increase. The

main difference between the tests with bottle-mixed syngas and the bench-scale tests with real

synthesis gas is that despite gas cleaning, the latter, probably contains certain amounts of catalyst

poisons, such as organic sulfur. By slowly deactivating the catalyst these poisons shift the main

reaction zone, thus causing carbon deposits to become more dispersed. No coking occurs after the

main reaction zone as all the hydrocarbons are converted in the inlet zone.

Figure 8.14: Measured catalyst carbon contents at different points of the methanation reactor after test 5

As in the tests with bottle-mixed synthesis gas, the majority of carbon deposits were of a filamentous

kind (figure 8.15); in addition, minor amounts of polymeric carbon layers were visible under the

electron microscope. What was different from the tests with bottle-mixed syngas was the presence

of laminar (probably graphitic) carbon deposits (figure 8.16). Since the operating conditions of the

lab-scale and the bench-scale methanation reactor were quite similar, the carbon deposits occurred

might be influenced by the gas compositions or the contaminations (or they were just not detected

on the samples of the tests with bottle-mixed gases).

Unlike in the tests with bottle-mixed syngas, the tests with real gas led to the formation of a high

amount of broken catalyst pellets (figure 8.17) in the inlet zone. Although thermal stressing cannot

be ruled out as a possible cause, the destructive behavior of filamentous carbon is a much more

likely explanation [109]. The fact that the broken catalyst pellets were observed only in areas with

severe coking also points to that.

250

300

350

400

450

500

550

0

5

10

15

20

25

30

35

40

0.0 0.2 0.4 0.6 0.8 1.0

Re

acto

r te

mp

era

ture

[°C

]

Car

bo

n c

on

ten

t [m

g Car

bo

n/g

Cat

alys

t]


35.7

8.7

0.69 0.10 0.11

Temperature


117

Figure 8.15: SEM-photos of polymeric carbon filaments on a catalyst sample taken after

longtime real gas tests

Figure 8.16: SEM-photos of laminar (graphitic) carbon deposits on a catalyst sample taken after

longtime real gas tests

Figure 8.17: SEM-photo of cracks on a catalyst tab after 200 h runtime with real synthesis gas

10 µm 1 µm

10 µm 400 nm

200 µm1 mm


118

Catalyst deactivation

In all tests with real synthesis gas, deactivation of the catalyst occurred; the displacement of the

temperature profile measured in the reactor is a clear indication of that. Figure 8.18 depicts such

temperature profiles for different runtimes. This deactivation is most probably caused by organic

sulfur species which could not be removed during hot gas cleaning; unfortunately, quantification of

these sulfur species was not possible. Although the tar species used in the tests with bottle-mixed

syngas did not cause any deactivation of the catalyst, deactivation due to other tar species cannot be

excluded. Further investigations are necessary to better understand the reasons for this deactivation.

Figure 8.18: Axial temperature trends in the methanation reactor for different runtimes for test 5

It is possible to quantify the deactivation by determining the integral area under the curve of the

reactor temperature. The area represents the released heat and it is therefore an indicator for the

activity of the catalyst. The amount of deactivation can be calculated as follows: To get the active

temperature profile, the inert temperature profile (Tinert) is subtracted from the temperature profile

of a particular runtime (Tt). By integrating the active temperature profile over the length (l) one gets

the active-catalyst-area aCatalyst (equation 8.1).

( ) ∫ (

) 8.1

( )

8.2

Normalizing the active-catalyst-area after a certain runtime to the active-catalyst-area at the

beginning, one gets the normalized-active-catalyst-area an,Catalyst (equation 8.2). Figure 8.19 shows the

trend of the normalized-active-catalyst-area for the five long-term tests with real synthesis gas.

Due to variation of the synthesis gas composition, the synthesis gas flow and the resulting gas

compositions after methanation, a useful comparison is only possible by relating the active-catalyst-

area to a representative value, such as synthesis gas power or SNG power.

250

300

350

400

450

500

550

0 0.2 0.4 0.6 0.8 1

Tem

pe

ratu

re [

°C]


0 h10 h

40 h

90 h

150 h

190 h


119

Since the decrease of the normalized-active-catalyst-area corresponds to the loss of catalyst active, it

becomes possible to calculate the specific amount of catalyst consumption. Figure 8.20 shows the

calculated specific catalyst consumption values of the five bench-scale test runs. The high variations

are mainly due to fluctuations in the supply of syngas and inconsistent determination of values such

as syngas power.

Figure 8.19: Measured catalyst degradation for tests 1-5 expressed by the normalized-active-catalyst-area

The tests with lignite (tests 1-4) show a linear decrease of the catalyst-area (catalyst consumption),

whereas the amount of degradation in test 5 (using biomass) decreases with runtime. One

explanation for this reduction is that the sulfur concentration of the feed might have fallen for some

unknown reason.

The average catalyst consumption values measured are between 0.13-0.34 g/kWhSyngas. According to

the catalyst consumption value determined for synthesis gas containing H2S (figure 7.20) these

measured consumption values would be equivalent to around 4-10.5 ppm of H2S.

In test 1 a different catalyst and a slightly higher H2O content of synthesis gas was used than in the

other tests. Therefore, a direct comparison of test 1 with tests 2-5 is not possible. The results of tests

with lignite (2-4) show a greater extent of catalyst consumption than those in which biomass was

used (5). This might be explained by the greater amounts of sulfur contaminations found in lignite-

derived synthesis gas. Although H2S was below 0.2 ppm in the feed stream for methanation, a certain

amount of organic sulfur was probably still present in the syngas.

However, due to certain unknown parameters, such as organic sulfur contents of synthesis gas, slight

variations between the different tests, and multiple determinations not having been carried out, a

clear explanation of the differences in the extent of catalyst consumption in the different tests is not

possible. Nevertheless, the values obtained provide useful clues as to the extent of catalyst

consumption.

0.80

0.85

0.90

0.95

1.00

0 50 100 150 200

No

rmal

ize

d a

ctiv

e c

atal

yst

are

a [-

]

Runtime [h]

Test 1

Test 2

Test 3

Test 4

Test 5


120

Figure 8.20: Measured specific catalyst consumptions for tests 1-5

When assessing the results, the cost factor also has to be taken into account. Assuming a catalyst

costs of 70 €/kg, additional costs of 0.9-2.4 ct/kWhSyngas would be incurred due to the consumption of

the catalyst. Considering the low feed-in price for natural gas of, currently, around 2.7 ct/kWh

(November 2013), such a high cost for a catalyst would clearly be uneconomical for lignite-to-SNG

systems.

However, further research should make it possible to reduce catalyst consumption. If that, together

with rising gas prices or the availability of additional funding, leads to an increase in revenues, the

concept proposed here can be a real alternative for the future as it will allow economical operation

of small-scale, decentralized biomass-to-SNG plants.

Test 1

Test 5

Test 2

Test 4

Test 3

0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

4.0

0

0.1

0.2

0.3

0.4

0.5

0.6

0 50 100 150 200

Runtime [h]

Cat

alys

t co

nsu

mp

tio

n [

g/kW

hSy

nga

s]

Cat

alys

t co

sts

[€ct

/kW

hSy

nga

s]

Conclusion

121

Chapter 9

9. Conclusion

This thesis makes a contribution to the development of a methanation process that allows the

production of substitute natural gas in small-scale, decentralized facilities. Smaller plant sizes require

a reduction of plant complexity, which can be achieved by introducing a reduced form of gas cleaning

and a simplified methanation process. With reduced gas cleaning certain contaminations remain in

the synthesis gas. These contaminations may harm the methanation catalyst; however, a certain

extent of deactivation of the catalyst can be accepted if it helps to reduce plant complexity. The

maximum degree of catalyst deactivation acceptable (extent of catalyst consumption) is a matter of

economics. The achievable amount depends on parameters such as the type of catalyst used, the

operating conditions and the gas compositions and contaminations. This thesis therefore seeks to

investigate the methanation process itself and, in particular, the influences of different

contaminations present in synthesis gas.

The results of bench-scale methanation tests show that the polytropic reactor concept proposed in

this thesis is a good alternative to existing concepts as it combines a simple design with lower

catalyst volumes. It allows clean synthesis gas to be fully converted up to around 230°C without

noticeable deactivation of the catalyst, and the SNG thus produced is, after conditioning, suitable for

feed-in into the gas grid.

Figure 9.1: Influence of contaminations on the specific amount of catalyst consumption;

parameters: 330°C reactor inlet temperature, tar mixture with 6 g/Nm³, H2S = 0.5 ppm, 40 vol. % H2O

Synthesis gas complexity

Syngas Syngas+ C2H4

Syngas+ C2H4

+ Tar

Syngas+ C2H4

+ Tar+ H2S

Realsyngas

0.7 vol. % C2H4

0.5 vol. % C2H4

Lignite

Biomass

PoisoningCoking

0.0

0.5

1.0

1.5

2.0

2.5

0.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

Cat

alys

tco

nsu

mp

tio

n[g

/kW

hSy

nga

s]

Cat

alys

tco

sts

[€ct

/kW

hSy

nga

s]

Conclusion

122

The situation is even more complex if different types of contaminations are present in synthesis gas.

Hydrocarbon-based contaminations are directly converted in the inlet zone of the methanation

reactor. However, certain hydrocarbon types and concentrations can cause coking of the catalyst.

Figure 9.1 summarizes the resulting extent of catalyst deactivation, expressed as specific catalyst

consumption values, for methanation with different contaminations.

The addition of 0.3-0.35 vol. % of ethylene to clean synthesis gas results in coking of the methanation

catalyst. C2H4 contents of above 0.5 vol. %, as are typical of biomass-derived synthesis gas, lead to

severe coking and therefore high specific catalyst consumption values. If, in addition to C2H4, tars are

also present in the syngas, coking and consequently the specific degree of catalyst consumption are

reduced. Further addition of minor amounts (< 1 ppm) of H2S allows methanation without deposition

of carbon despite the presence of C2H4 and tars. Since H2S is a strong catalyst poison, it slightly

deactivates the catalyst, but with low (acceptable) catalyst consumption.

These results were obtained under controlled and well reproducible lab-scale conditions.

Unfortunately, bench-scale methanation tests with real synthesis gas from thermal gasification

caused a much higher deactivation rate although the operating conditions of methanation and the

amounts of C2H4 and tars were the same order, the main difference being probably the larger

amount of organic sulfur contaminations in the feed, which could not be removed completely by the

form of gas cleaning applied. Coking also occurred in the bench-scale methanation test. Since it was

sufficiently low, it was, however, not the reason for the high degree of catalyst deactivation.

Figure 9.2, which shows the various degrees of catalyst deactivation in relation to sulfur

concentrations, provides a different view of the results. It illustrates clearly that methanation

without consumption of the catalyst is not possible with the proposed concept, but that the extent of

deactivation is fairly low for sulfur concentrations of 0.7 to 1 ppm.

Figure 9.2: Influence of sulfur concentration and ethylene content on specific catalyst consumption

0.0

0.5

1.0

1.5

2.0

2.5

0

0.1

0.2

0.3

0.4

0 1 2 3 4 5

Cat

alys

t co

nsu

mp

tio

n [

g/kW

hSy

nga

s]

Sulfur concentration [ppm]

1 vol. % C2H4

0.7 %

0.5 %

Syngas

Real syngas

Biomass

Lignite

Syngas + HC

Cat

alys

tco

sts

[€ct

/kW

hSy

nga

s]

Conclusion

123

In summary it can be said that a complete conversion of higher hydrocarbons is possible. The

deposition of carbon, which occurs in the process, can be prevented or minimized by adding traces of

hydrogen sulfide. Although further investigations are necessary, the results show the great potential

of the proposed concept for the production of SNG and contribute to a better understanding of the

various factors influencing methanation.

Further investigations will have to look at ways of reducing catalyst deactivation resulting from

methanation with real synthesis gas. Since catalyst degradation is probably caused by insufficient

removal of sulfur contaminations and not by the methanation process itself, the optimization of the

gas cleaning process seems to be the next logical step to take. Rather than aiming at complete

removal of sulfur, which may result in coking, this thesis advocates the addition of traces of H2S as a

simple and economical solution.

References

124

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