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
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.
III
IV Abstract
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.
V Kurzfassung
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.
VI Content
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
VII Content
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
VIII Content
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
IX Content
X Figures
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
XI Figures
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 7.4: Carbon deposition on the catalyst at 320°C using different C2H4 contents ...... 84
Figure 7.5: Carbon deposition on the catalyst at 370°C using different C2H4 contents ...... 85 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
XII Figures
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
XIII Tables
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
XIV Nomenclature
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
XV Nomenclature
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] -1 -1 ελ Molar absorption coefficient [l mol m ]
ηMeth Chemical efficiency for methanation [-] λ Air ratio [-] σ Excess steam ratio [-]
XVI Introduction
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.
Fuel H2O H2O/CO2
Thermal Gas Gas cleaning Methanation gasification conditioning
Heat/O2 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.
1 Introduction
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.
Hot gas filter Methanation Gas conditioning
SNG Fuel Ash Gasifier Heat
Steam Adsorptive gas cleaning H2O CO2
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.
2 Introduction
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.
3 Introduction
4 State-of-the-Art
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
5 State-of-the-Art
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.
adiabateAdiabatic Festbettreaktoren fixed bed reactors Beispiel: Beispiel:Cooled reactors adiabateadiabateadiabate Festbettreaktoren Festbettreaktoren Festbettreaktoren Beispiel:Beispiel:gekühltegekühltegekühlte gekühlteReaktoren Reaktoren Reaktoren Reaktoren Beispiel:Beispiel:Beispiel:Beispiel: Lurgi Methanation e.g. Lurgi methanation process LurgiLurgi LurgiMethanation Methanation Methanation e.g. SynthaneSynthanehotSynthaneSynthane tubeHotSynthane Tube reactorHot Hot Reactor Tube Hot Tube TubeReactor Reactor Reactor gestufte gestufte ZwischenkühlungZwischenkühlung gestuftegestufte methanreiches ZwischenkühlungZwischenkühlungEdukt- Edukt- methanreichesmethanreiches intercooling Edukt-Edukt-staged feed Produktgasmethanreiches zugabe zugabe ProduktgasProduktgas zugabezugabeinjection Produktgasraw-SNG up to bis 600°Cbis 600°Cbis 600°C 600bis- 700600°C°C Beispiel: Beispiel: heat-transfer- Beispiel:Beispiel: Synthane Hot Tube Reactor adiabate adiabate fluid cooling SynthaneSynthane Hot Tube Hot Reactor Tube Reactor adiabateadiabate Thermoöl- Thermoöl- Synthane Hot Tube Reactor Festbett- Festbett- Thermoöl-Thermoöl- Festbett-Festbett- Kühlung Kühlung katalysatoren katalysatoren KühlungKühlung katalysatorenkatalysatoren methanreiches methanreiches adiabatic methanreichesmethanreiches ProduktgasProduktgasrawProduktgas-SNG fixed beds ca. 400°C ca. 400°C Produktgas ca.ca. 400°Cca. 400°C 400°C
Synthesegas Synthesegas SynthesegasSynthesegasca. ca. ZwischenkühlungZwischenkühlungZwischenkühlung ca.ca. intercoolingZwischenkühlung (CO, H2)(CO, H2)(CO,300-400°C H2) 300-400°C SynthesegasSynthesegasSynthesegas (CO,syngas H2) 300-400°C300-400°C300-400°C gestufte gestufte Synthesegas gestuftegestuftestaged (CO, H2)(CO,syngas H2)(CO, H2) Produktgas- Produktgas- Edukt- Edukt- (CO, H2) Produktgas-Produktgas- Edukt-Edukt-feed rezirkulationrecyclerezirkulation compressor zugabe zugabe rezirkulationrezirkulation zugabezugabeinjection
Flüssigphasen-MethanierungFlüssigphasen-Methanierung Wirbelschicht-MethanierungWirbelschicht-Methanierung Flüssigphasen-MethanierungFlüssigphasen-MethanierungThree-phase methanation FluidizedWirbelschicht-MethanierungWirbelschicht-Methanierungbed methanation Beispiel: Beispiel: Beispiel:Beispiel:Beispiel:Beispiel:e.g. Chem systems liquid phase methanation Beispiel:Beispiel:e.g. Thyssengas Comflux process Chem ChemSystemsChem Systems Systems ThyssenThyssen ThyssenComflux-Thyssen Comflux- Comflux- Comflux- Chem Systems methanreiches methanreiches methanreiches methanreiches Liquid PhaseLiquid Phase heat exchanger methanreichesmethanreiches Verfahren Verfahren methanreichesmethanreiches LiquidLiquid Phase Phase Zwischenkühlung Zwischenkühlung Produktgas Produktgas VerfahrenVerfahren Produktgasraw-SNG Produktgas ZwischenkühlungZwischenkühlung rawProduktgasProduktgas-SNG ProduktgasProduktgas MethanationMethanationMethanationMethanation Synthesegas/ Synthesegas/ Synthesegas/Synthesegas/ 3-Phasen- 3-Phasen- ProduktgasProduktgasProduktgas druckaufgeladenedruckaufgeladenedruckaufgeladene Synthesegas/ Synthesegas/ 3-Phasen-3-Phasen- reactantsProduktgas Separator Separator pressurizeddruckaufgeladene reactantsSynthesegas/Synthesegas/ Wirbelschicht Wirbelschicht (gasförmig) (gasförmig) separatorSeparator Separator WirbelschichtWirbelschichtWirbelschicht Produktgas Produktgas WirbelschichtWirbelschicht (gaseous)(gasförmig)(gasförmig) fluidizedWirbelschicht bed (gaseous)ProduktgasProduktgas three-phase bis ca. 70 bar bis ca. 70 bar (gasförmig) (gasförmig) ca. 70 barca. 70 barca. 70 bar inertes inertes up to 70 bisbar bisca. ca. 70 70bar bar (gasförmig)(gasförmig) fluidizedca.bed 70 bar inertesinertes Trägerölinert fluidTrägeröl TrägerölTrägeröl Kühlung Kühlung (flüssig)(liquid)(flüssig) KühlungKühlung (flüssig)(flüssig) cooling catalyst Katalysator-catalyst Katalysator- Katalysator- Katalysator- Katalysator-Katalysator- particles Katalysator-Katalysator- PartikelparticlesPartikel Partikel Partikel PartikelPartikel (solid) PartikelPartikel Synthesegas Synthesegas (fest) (solid) (fest) Synthesegas Synthesegas (fest) (fest) SynthesegasSynthesegas (fest)(fest) Trägeröl- Trägeröl- Synthesegas (fest)(fest) (CO, H2) (CO, H2) Trägeröl- Trägeröl- syngas Synthesegas (CO,syngas(CO, H2) H2) Rezirkulation Rezirkulation (CO, H2)(CO, H2)(CO, H2) RezirkulationRezirkulation (CO, H2)
Figure 2.2: Different reactor concepts and processes for methanation of synthesis gas, according to [12]
6 State-of-the-Art
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].
7 State-of-the-Art
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].
8 State-of-the-Art
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.
Coal Gas Shift Rectisol Fixed bed Syngas to cooling conversion unit methanation ammonia plant Coal lock
SNG Steam Lurgi gasifiers
O2 Condenser Air ASU Ash lock Gravity Water separation Naphta N2, Xe, Ash Kr CO2 for enhanced oil recovery Tar oil By-product Phenol, Cresol, processing Ammonium sulfate
Figure 2.3: Simplified flow sheet of the DGC Great Plains synfuels plant, adapted from [3]
9 State-of-the-Art
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
10 State-of-the-Art
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.
Filter RME Bulk H2S Fine H2S Fluidized bed NH3/H2O CO2/H2 Flue scrubber removal removal methanation removal removal gas CO Pre- 2 RME heating Biomass FICFB Ash SNG gasifier Water H2 Air Tars/ Steam To gas engine tiophene
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.
11 State-of-the-Art
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.
Steam Fuel Syngas
Flue gas
Reformer
Heat pipes
Combustion chamber
Figure 2.5: a. Conceptual design of the HPR, b. idea for decentralized SNG production, according to [41]
12 State-of-the-Art
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.
13 State-of-the-Art
14 Theoretical Background
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
15 Theoretical Background
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
16 Theoretical Background
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. 0.6
H 0.5 2,dry H2Owet
0.4 CO
[mole fraction][mole dry 0.3 CO2,dry 0.2
0.1 CH4,dry Gas Gas composition
0.0 2 4 6 8 10 Excess steam ratio σ [-]
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
17 Theoretical Background
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. GC 7-rings and 1 Very heavy tars, gravimetric tar undetectable higher Cyclic hydrocarbons with heteroatoms, water phenol, cresol, 2 Heterocyclic soluble pyridine Compounds that usually do not pose problems xylene, styrene, 3 Light aromatic regarding condensation or water solubility toluene naphthalene, Light 2/3-ring compounds that condense at intermediate 4 acenaphthene, polyaromatic temperatures at higher concentrations anthracene Heavy 4-to-6-ring compounds that condense at high fluoranthene, 5 polyaromatic temperatures at low concentrations 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
18 Theoretical Background 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
19 Theoretical Background
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.
Woody Biomass Lignite (higher quality) Measured contaminates lab scale gasifier 10000
1000
100
10
1
Concentration [ppm][g/Nm³]Concentration BTX/Tarfor or 0.1 H2S COS CS2 org. S* HCl NH3 BTX Tar**
*excluding COS and CS2; **Total tar amount including BTX;
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
20 Theoretical Background
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
100 CaO Fe3O4
10 [ppm] 1 MnO Ni 0.1
concentration ZnO
S S 2 H 0.01
0.001 200 300 400 500 600 700 800 900 Temperature [°C]
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.
21 Theoretical Background
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]
3.16
3.17
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.
22 Theoretical Background
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 Sour gases (H S, HCl, HF, SO , NO ), Carbotech 10-20 wt. % 2 2 2 K2CO3 CS2 D47/3-KC10
Iron oxide 10 wt. % H2S, mercaptans, COS - H S, PH , Hg, AsH , radioactive Potassium iodide KI 1-5 wt. % 2 3 3 Norit ROZ 3 gases, mercaptans Potassium 5 wt. % H2S (without O2), aldehydes - permanganate KMnO4 Potassium hydroxide Sour gases (H S, HBr, HCl, HF, SO , Donau Carbon 10 wt. % 2 2 KOH NO2), mercaptans Desorex K43Na Sodium hydroxide 10 wt. % H S, mercaptans - NaOH 2
23 Theoretical Background
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.
24 Theoretical Background
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.
0.8
0.7 H2 0.6
0.5 CH4 [mole fraction] [mole 0.4
H2O 0.3 CO 0.2
CO Gas composition Gas 0.1 2
0 0 200 400 600 800 1000 Reaction temperature [°C]
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.
25 Theoretical Background
0.5 CH4
0.4 H2O
0.3 [molefraction]
0.2
H2 0.1
Gascomposition CO2 0 CO 0.1 1 10 100 Reaction pressure [bar]
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).
0.7 0.9 ηMeth 0.8 0.6 0.7
0.5 Meth η H2O 0.6 0.4
[mole fraction][mole 0.5
0.3 0.4 CO 0.3 0.2 2
0.2 Chemical efficiency CH Gas Gas composition 0.1 4 H 0.1 2 CO 0 0 200 300 400 500 600 Reaction temperature [°C]
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
26 Theoretical Background
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 ̇
0.6
0.5 CO2
CH4
0.4 [molefraction] 0.3
0.2
0.1
H2 Dry gas gas Dry composition CO 0 0 0.1 0.2 0.3 0.4 0.5 0.6
H2O content of synthesis gas [mole fraction]
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].
27 Theoretical Background
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.
H H C O 1 2 3
O H H C C O
O H H C 4 H H C O
H H C O 5
H H H H H H H H H C H H C H C H C 6 7 H 8 H H
H H H H O H 9 H O 10 O
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).
28 Theoretical Background
H H C O O C O O C O H H 1 2 E-R 6 7 O O H H H C O C O H O C O H H O C O H H L-H 3 4 5
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.
29 Theoretical Background
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].
H H O H H 1 H H O O H H O H H 2 3
H H H C C H H H 4 H H H H H H H H C C H H H H C C H H 5
C O H H H 8 O H C O H H C O C 6 7
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].
30 Theoretical Background
Hydrogenation Isomerization Ring Opening Dealkylation
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].
31 Theoretical Background
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]
Benzene Naphthalene Toluene 100%
80% [%]
60%
40%
conversion
m
H n
C 20%
0% 200 250 300 350 400 450 500 Temperature [°C]
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]
32 Theoretical Background
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].
33 Theoretical Background
34 Catalyst Deactivation and Carbon Deposition
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 Strong chemisorption of species on catalytic sites, Poisoning Chemical thereby blocking sites for catalytic reaction Physical deposition of species from fluid phase Fouling Mechanical onto the catalytic surface and in catalyst pores Thermally induced loss of catalytic surface area Thermal degradation Thermal (sintering) Reaction of gas with catalyst phase to produce Vapor formation Chemical volatile compounds, e.g. Ni(CO)4 Vapor-solid and solid-solid Reaction of fluid, support or promoter with Chemical reactions catalytic phase to produce inactive phase Loss of catalytic material due to abrasion or Attrition/Crushing Mechanical crushing of catalyst particles
35 Catalyst Deactivation and Carbon Deposition
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. Carbon (a) (c) (b)
Nickel
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.
36 Catalyst Deactivation and Carbon Deposition
Table 4.2: Carbon species formed on Ni catalyst, adapted from [113] and [112]
Formation Formation Peak Carbon species temperature from temperature from temperature for
CO decomposition C2H4 decomposition reaction with H2 Adsorbed, atomic 100-330°C C 200-400°C 200°C (surface carbide) (Peak 220°C) Polymeric, amorphous 330-620°C C 250-500°C 400°C films or filaments (Peak 430°C) Vermicular (polymeric,
CV amorphous) filaments, 300-1000°C - 400-600°C fibres or whiskers 230-330°C C Nickel carbide 150-250°C 275°C (Peak 270°C) Graphitic (crystalline) C 500-550°C - 550-850°C C platelets or films
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).
CxHy
via CHn
6 7 8 CH4 3 +H2O -H2 via H2/CO2 via CHn CxHy 5 CH4 +H2 +H2O via H /CO 4 2 +/- H2 9 2 2
CO 1 C 10 C 12 13 11
fast reaction C Cv Cc
kinetically limited reaction Gaseous Adsorbed Solid
Figure 4.2: Reaction paths for formation, gasification and transformation of coke and carbons, adapted from [113] and [108]
37 Catalyst Deactivation and Carbon Deposition
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.
Amorphous Ni Ni carbon CO / CxHy Support Support 1 2
C C Ni
4 5 Support 3
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.
38 Catalyst Deactivation and Carbon Deposition
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.
Carbon C 0.0 1.0 0.1 0.9 0.2 0.8 0.3 0.7 0.4 0.6 0.5 Lignite 0.5 CH O 0.6 0.87 0.27 0.4 Biomass 600°C 0.7 CH O 500°C 1.36 0.61 0.3 Phase 0.8 equilibrium Syngas dry 400°C lines for 0.2 300°C 200°C solid carbon 0.9 Syngas with 0.1
40 vol. % H2O 1.0 H O 2 0.0 0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 Hydrogen H Oxygen O
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,
39 Catalyst Deactivation and Carbon Deposition 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
40 Catalyst Deactivation and Carbon Deposition
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.
Temperature [°C] 7001 600 500 4001.5 300 2002 150 2.5 10
CO C+O
) 8 x 10³ x
(N 6 ln
4
C+HCH
2 C+H CH
C C
Rateof formation 0
-2 1 1.5 2 2.5 Reciprocal Temperature 1/T [10-3 K]
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].
41 Catalyst Deactivation and Carbon Deposition
Temperature T [°C] 700 600 500 400
Propene 100
10
1-Butene Rateof deposition [µg/min cm²]
1 0.9 1 1.1 1.2 1.3 1.4 1.5 Reciprocal Temperature 1/T [10-3 K]
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)
42 Catalyst Deactivation and Carbon Deposition 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.
43 Catalyst Deactivation and Carbon Deposition
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.
100
80
60
40
20
Reactor differntial pressure[mbar] No or low coking
0 0 20 40 60 80 100 Runtime [h]
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.
44 Catalyst Deactivation and Carbon Deposition
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.
500 nm 500 nm 500 nm a) b) c)
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.
Exhaust
Gas analyzer F O2 MFC Ar: max. 400 ml/min Ar
Gas analyzer F CO2 MFC O2: max. 300 ml/min O2 Filter Thermogravimetric analyzer (TGA)
Figure 4.9: Flow sheet of the TPO setup to determine carbon deposits
45 Catalyst Deactivation and Carbon Deposition
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
Temperature Heating rate Time Argon Oxygen Step [°C] [°C/min] [min] [ml/min] [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.
46 Catalyst Deactivation and Carbon Deposition
900 4.5 800 Temperature profile 4 700 3.5
600 3
C] ° [mg] Change in mass 500 2.5 Graphitic carbon [vol. %] 400 2
300 1.5 content
CO2 content 2 Temperature[
200 1 CO Changein mass 100 0.5 0 0 -100 -0.5 0 100 200 300 400 Runtime [min]
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].
900 4.5 800 Temperature profile 4 700 3.5
600 3
C]
° [mg]
500 Change in mass 2.5 [vol. %] Graphitic carbon 400 2
300 1.5 content
Polymeric 2 Temperature[ 200 Reactive CO content 1 CO
Changein mass carbon 2 carbon 100 0.5 0 0 -100 -0.5 0 100 200 300 400 Runtime [min]
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)
47 Catalyst Deactivation and Carbon Deposition
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.
48 Catalyst Deactivation and Carbon Deposition
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].
600
t0h txh
500
C] °
400 Temperature [ Temperature 300
Deactivated reactor area after txh 200 0 0.2 0.4 0.6 0.8 1 Normalized reactor length
Figure 4.12: Displacement of the reactor temperature profile due to selective deactivation at the entrance of a polytropic cooled fixed bed reactor
49 Catalyst Deactivation and Carbon Deposition
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.
50 Catalyst Deactivation and Carbon Deposition
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.
100
10 [ppm] 1
0.1
concentration 4
0.01 Ni(CO)
0.001 100 150 200 250 300 350 400 Ni catalyst temperature [°C]
Figure 4.13: Equilibrium concentration for Ni(CO)4 for different CO concentrations in dependency of the temperature, calculated with FactSage
51 Catalyst Deactivation and Carbon Deposition
52 Bench-Scale Methanation Tests with Clean Syngas
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.
16 thermocouples Air (for catalyst oxidation) (side-mounted and Pressure T center-mounted) P sensor Methanation reactor
F
H2 MFC H2: max. 21 l/min
r
o s n
F e s
g e n r CO2 MFC CO2: max. 15 l/min i l L u o s s o
c e
PD r F r p o
l t a c
CO MFC CO: max. 10 l/min i a t Water g e n n
i R e
saturator t r n a r e
F e e v h h f
t
o f
i g CH e 4 MFC CH4: max. 1.6 l/min g n D n n e i o l
l
z
r p 3 o m F t c a s a
e N2 MFC N2: max. 55 l/min s r a G T r a Pressurized air c
Water e
Demineralization h e a t
Exhaust Pump unit e d
l i Gas analyzer n e s
H2, O2 ( 2
0 Natural
0 Flare ° gas C
Gas analyzer ) CO, CO2, CH4 Condenser unit
Figure 5.1: Simplified flow sheet of the bench-scale methanation test rig
53 Bench-Scale Methanation Tests with Clean Syngas
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.
TM 1-5
G1 T1: 0 cm T2: 1 cm Inlet G2 T3: 2 cm T4: 3 cm G3 T5: 5 cm T6: 7 cm
T7: 10 cm
G4 T8: 15 cm
T9: 25 cm
G T : 35 cm Air cooled mantle 5 10
T11: 45 cm
T12: 57 cm
Outlet DR,i: 27.6mm
VR: 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.
54 Bench-Scale Methanation Tests with Clean Syngas
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 56 % Ni on Ni (> 50 wt. %) on Ni-based Ni (> 50 wt. %) SiO2/MgO SiO2/Al2O3 Al2O3/SiO2 Extrudates, Extrudates, 3 x 3 mm tabs 3 x 3 mm tabs 1.9 x 3.5mm tabs 1.6 mm 1.6 mm < 650°C < 550°C < 550°C < 500°C < 550°C Reforming Methanation Methanation Methanation Sulfur sorbent catalyst catalyst catalyst 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.
55 Bench-Scale Methanation Tests with Clean Syngas
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
56 Bench-Scale Methanation Tests with Clean Syngas 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.
550 EVT02 EVT01 500 EVT04 450
C] EVT05 ° 400 EVT03
350
Temperature[ 300
250
200 0.0 0.2 0.4 0.6 0.8 1.0 Scaled reactor length [-]
-1 Figure 5.3: Temperature profiles of the tested catalysts at a GHSV of 4000h 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
57 Bench-Scale Methanation Tests with Clean Syngas 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.
60
CO 50 2
40 CH4 30
20
H2 Gascomposition [vol. %] 10 CO 0 0 0.2 0.4 0.6 0.8 1 Scaled reactor length [-]
Figure 5.4: Gas composition measured at various points of the reactor compared to temperature-related -1 equilibrium gas compositions for EVT05 at a GHSV of 1500 h , 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 -1 equilibrium with a GHSV of 4000 h 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.
58 Bench-Scale Methanation Tests with Clean Syngas
20 18 16
14 EVT03
12 EVT02 [vol. [vol. %] 10 EVT04
8 EVT05
content content EVT01
2 6 H 4 Equilibrium 265°C 2 0 20 25 30 35 40
H2O content [vol. %]
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.
7
6 5 K
5 . . %]
4 Equilibrium [vol 40 vol. % H2O 3 30% H2O
content 30% H2O 2
H 2 33%33% H2OH2O Equilibrium 35%35% H2OH2O 30 vol. % H O 1 2 37%37% H2OH2O 40%40% H2OH2O 0 220 230 240 250 260 270 280 Reactor outlet temperature [°C]
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
59 Bench-Scale Methanation Tests with Clean Syngas
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.
8
7 25 K
6 40 vol. % H2O
. . %] 5 Equilibrium
[vol 4 40 vol. % H2O 30 vol. % H2O
3
content 2 H 2 Equilibrium 30 vol. % H O 1 2
0 220 230 240 250 260 270 280 Reactor outlet temperature [°C]
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
60 Bench-Scale Methanation Tests with Clean Syngas
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.
7 C]
290 ° 6 280
5 270 . . %] 4 260 [vol 250 3 240
30%30% H2OH2O
content 2
2 33%33% H2OH2O 220 equilibrium [ temperature H 35%35% H2OH2O
1 37%37% H2OH2O Average 40%40% H2OH2O 0 1000 1500 2000 2500 3000 GHSV [1/h]
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, -1 e.g. 1500 h 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.
10 320 C]
9 ° [ 8 300
7 40 vol. % H2O . . %]
6 280
temperature [vol 5
4 30 vol. % H2O 260
content
2 H 3 240 equilibrium 2 220
1 Average 0 1000 1500 2000 2500 3000 GHSV [1/h]
Figure 5.9: H2 content in the product gas in dependency of the GHSV and the water content with EVT05
61 Bench-Scale Methanation Tests with Clean Syngas
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 -1 1500 h 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.
570
550 C]
° 530 TM2: 5cm T : 8.5cm 510 M3 T4: 3cm
Temperature[ 490 T7: 10cm
470
450 0 100 200 300 400 Runtime [h]
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.
62 Bench-Scale Methanation Tests with Clean Syngas
50
CH4 48
CO2 46
6
H2
4 Gas composition Gas [vol. %] 2
0 0 100 200 300 400 Runtime [h]
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).
63 Bench-Scale Methanation Tests with Clean Syngas
64 Methanation Tests with Contaminated Syngas - Setup
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.
65 Methanation Tests with Contaminated Syngas - Setup
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.
Catalyst
Water content Reactor syngas temperature
Activity Conversion Synthesis gas Permanent gas Methanation contaminations composition Carbon deposition
Reaction pressure Residence time
Reactor type
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.
66 Methanation Tests with Contaminated Syngas - Setup
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.
67 Methanation Tests with Contaminated Syngas - Setup
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).
Methanation reactor
Gas analyzing unit Gas mixing station with tar saturators
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.
68 Methanation Tests with Contaminated Syngas - Setup
Trace heated lines (200°C) F
H2 MFC H2: max. 2800 ml/min
F
CO2 MFC CO2: max. 1000 ml/min Water saturator with L automatic refilling F
MFC CO: max. 1000 ml/min CO F
F
CH4 MFC CH4: max. 350 ml/min
F
MFC N2: max. 3500 ml/min Water
N2 F Demineralization
MFC N2: max. 350 ml/min
F
TG1 MFC TG1: max. 380 mlN2/min
F
TG2 MFC TG2: max. 1200 mlN2/min
N2 Static
F F F F mixer
Pressure sensor P
Overpressure valve Tar saturator 1 Tar saturator 2 Tar saturator 3 Tar saturator 4 Outlet / to (Benzene) (Phenol) (Naphthalene) (Toluene) reactor
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
69 Methanation Tests with Contaminated Syngas - Setup 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. A B C n E F Range [K] 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
70 Methanation Tests with Contaminated Syngas - Setup
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
71 Methanation Tests with Contaminated Syngas - Setup 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.
From gas
S F mixing P Port for E station SPE-sample MFC H2: max. 2500 ml/min H2 To FID 5 thermo- T couples F MFC N2: max. 25000 ml/min N2 r o
l s
a Reactor n i t e s n oven
e Natural e r r PD Flare e u f gas s f i s D e r p Reactor in Reactor in methanation reduction part test part S
P Port for To gas E analyzing SPE-sample unit Trace heated lines
To FID
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
72 Methanation Tests with Contaminated Syngas - Setup 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.
UV-adsorptive tar N (zero gas) Pressurized air analysis 2
Exhaust
Trace heated lines (200°C) Ejector pump with Condenser flow regulation
Flame ionization detector H2, combustion gas
H2 From Calibration gas (1 vol. % C3H8 in N2) reactor inlet
CG Pressurized air
Silica gel Activated Filter Activated carbon carbon (1µm) (heated, 90°C)
Exhaust Micro-GC
Gas analyzer From reactor H2, O2 outlet
Gas analyzer CO, CO2, CH4 He
Activated carbon From reactor
Pump Condenser unit unit
Figure 6.7: Flow sheet of gas analyzing unit
73 Methanation Tests with Contaminated Syngas - Setup
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
74 Methanation Tests with Contaminated Syngas - Setup 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 Molecular sieve 5Å (zeolite type L = 10 m, D = 0.25 mm, H , O , N , CH , CO, MS5A 10m A) PLOT-column with pre-column 2 2 2 4 T = 180°C NO, Ar, He, Ne PoraBOND Q max PLOT-column with polystyrene- L = 10 m, D = 0.25 mm, C -C , CO , H S, PPQ 10m 1 5 2 2 divinylbenzene Tmax = 180°C COS, SO2 WCOT-column with 100% L = 8 m, D = 0.15 mm, C -C , CS , H S, 5CB 8m 3 10 2 2 dimethylpolysiloxane Tmax = 180°C 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 H S, C H , C H , C H , CS , C H , C H , C H , H S, 2 3 8 3 6 4 10 2 Calibrated components 2 2 2 4 2 6 2 benzene, toluene, COS, C H , C H 3 8 3 6 ethylbenzene*, o-,p-xylene* *Runtime 220 seconds
75 Methanation Tests with Contaminated Syngas - Setup
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.
76 Methanation Tests with Contaminated Syngas - Setup
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, -1 -1 -1 -1 e.g. for naphthalene ε286nm = 9300 [l mol cm ], for toluene ε261nm = 300 [l mol cm ].
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.
UV-LED Photodiode with with mounting mounting Glass rod
Heated measuring cell
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.
77 Methanation Tests with Contaminated Syngas - Setup
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.
78 Methanation Tests with Contaminated Syngas - Setup
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.
100 ml syringe
Gas stream SPE-column
Septum retainer
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.
79 Methanation Tests with Contaminated Syngas - Setup
80 Methanation Tests with Contaminated Syngas - Results
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
81 Methanation Tests with Contaminated Syngas - Results 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.
Scaled reactor lenght [-] -0.05 0 0.05 0.1 550
500 C] ° 450
400
350 500°C 450°C 300 370°C
Reactortemperature [ 320°C 250 Catalyst bed 300°C 550°C 200 -2 -1 0 1 2 3 4 5 Reactor length [cm]
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.
550 C] ° 500 Peak
450 Medium inlet zone
400
350 Inlet
Correspondingtemperature[ 300
250 300 350 400 450 500 550 Temperature reactor oven [°C]
Figure 7.2: Resulting reactor temperatures in dependency of the reactor oven temperatures
82 Methanation Tests with Contaminated Syngas - Results
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).
83 Methanation Tests with Contaminated Syngas - Results
14
] 12
Catalyst g
/ 10
Carbon 8 mg [ 0.5 vol. % C2H4 6
content 0.7 vol. % C H 4 2 4
Carbon 2
0 0 50 100 150 200 250 Runtime [h]
Figure 7.3: Carbon deposition on the catalyst at 300°C reactor oven temperature (≈280°C inlet) using different -1 C2H4 contents and runtimes, GHSV 10000 h
14
] 12
Catalyst g
/ 10 0.7 vol. % C2H4
Carbon 8
mg [ 6
content 4
0.5 vol. % C2H4
Carbon 2
0 0 50 100 150 200 250 Runtime [h]
Figure 7.4: Carbon deposition on the catalyst at 320°C reactor oven temperature (≈300°C inlet) using different -1 C2H4 contents and runtimes, GHSV 10000 h
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
84 Methanation Tests with Contaminated Syngas - Results 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.
14 ]
12 0.7 vol. % C2H4
Catalyst g
/ 10
Carbon 8
mg [
6 content 4 0.5 vol. % C2H4
Carbon 2
0 0 50 100 150 200 250 Runtime [h]
Figure 7.5: Carbon deposition on the catalyst at 370°C reactor oven temperature (≈330°C inlet) using different -1 C2H4 contents and runtimes, GHSV 10000 h
500
140 h
C] 10 h
° 450 100 h
50 h 400
350 Reactortemperature [
300 0 1 2 3 4 5 6 Reactor length [cm]
Figure 7.6: Temperature profiles of a test with high carbon deposition; reactor oven temperature 370°C,
0.7 vol. % C2H4, runtime 142 h
85 Methanation Tests with Contaminated Syngas - Results
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.
86 Methanation Tests with Contaminated Syngas - Results
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.
200 Temp./C2H4/Runtime [°C/vol.%/h] 300/0.5/187
160 320/0.7/142
120 320/0.5/210 370/0.7/70 80 320/0.7/70
40 370/0.5/66 Reactor differntial pressure[mbar] 300/0.5/70 Initial pressure 0 320/0.5/70 320/0.7/46 0 2 4 6 8 10 12 14
Deposited carbon [mgC/gCatalyst]
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
87 Methanation Tests with Contaminated Syngas - Results 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.
35 mg/g 15 mg/g 10 mg/g
5 mg/g 1 mg/g 0 mg/g
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.
88 Methanation Tests with Contaminated Syngas - Results
[ ] [ ] 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.
0.5 3.5 ]
Max. acceptable 30 mg/g 3 ]
Syngas 0.4 carbon content
2.5 Syngas kWh
20 mg/g kWh
0.3 /
2
ct € 1.5 0.2 10 mg/g 1 0.1 0.5 Catalyst costs[
Catalyst Catalyst consumption[g/ 5 mg/g 0 0 0 50 100 150 200 250 Carbon content [µg/g·h]
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.
89 Methanation Tests with Contaminated Syngas - Results
Reactor inlet temperature [°C] 280 300 320 340 360 380 400 420 440 160 160
140 140 · h] ·
120 120
Catalyst Catalyst High /g 100 100
0.7 vol. % C2H4 Carbon
g 80 80 [µ
60 60
content 40 40 0.5 vol. % C2H4
20 20 Carbon
0 0 Low Medium 250 300 350 400 450 500 550 Reactor oven temperature [°C]
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
90 Methanation Tests with Contaminated Syngas - Results
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.
Reactor inlet temperature [°C] 280 300 320 340 360 380 400 420 440 1 20 µg /g ·h 0.9 C Catalyst 0.8
0.7 89 89 133 54
0.6 20 [vol. [vol. %] 10 26 16 0.5 39 60 3
content 0.4 4
H 12 0 0
2 0.3 C 0.2 0.1 0 0 0 0 250 300 350 400 450 500 550 Reactor oven temperature [°C]
Figure 7.12: Amount of deposited carbon in dependency of the temperature and the C2H4 content, based on a runtime of 22 h, GHSV 10000 h-1
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.
91 Methanation Tests with Contaminated Syngas - Results
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.
Reactor inlet temperature [°C] 300 320 340 360 380 400 420 440 45 20 µgC/gCatalyst·h
17 4.5 5 6.5 40
[vol. %] [vol. 1 4.5 9.5 11
35 O O content 2 1.5 6.5 12 13.5 H 30
25 300 350 400 450 500 550 Reactor oven temperature [°C]
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.
92 Methanation Tests with Contaminated Syngas - Results
Reactor inlet temperature [°C] 300 320 340 360 380 400 420 440 16
14 25 13.5 12.5 24.5 12
10
8 17 4.5 5 6.5 6
4 Tarconcentration [g/Nm³] 2 20 µgC/gCatalyst·h 0 300 350 400 450 500 550 Reactor oven temperature [°C]
Figure 7.14: Amount of deposited carbon in dependency of the temperature and the tar concentration, -1 syngas with H2O content of 40 vol. %, based on a runtime of 22 h, GHSV 10000 h
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.
Reactor oven temperature [°C] 320 370 450 550 700 1.00 0.99 0.98
] 0.97 - 0.96 0.95
0.94 Benzene
Tarconversion [ 0.93 Toluene Naphthalene 0.92 Phenol 0.91 Total 0.90 450 500 550 600 650 Reactor peak temperature [°C]
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
93 Methanation Tests with Contaminated Syngas - Results
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. Reactor oven temperature [°C] 320 370 450 550 700 1.00 0.99 0.98
0.97
] - 0.96 0.95 0.94 Benzene Toluene 0.93 Tarconversion [ Naphthalene 0.92 Phenol 0.91 Total 0.90 450 500 550 600 650 Reactor peak temperature [°C]
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.
94 Methanation Tests with Contaminated Syngas - Results
1.00 42/18/40 54/6/40
] 0.95 -
60/0/40 0.90
Tolueneconversion [ 0.85
H2/CO/H2O [vol. %] 0.80 350 400 450 500 550 600 Reactor peak temperature [°C]
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. Reactor inlet temperature [°C] 280 300 320 340 360 380 400 420 440 160 condensation/ 0.7 vol. % C H + Tar 140 polycondensation 2 4 · h] · expected
120
Catalyst Catalyst /g
100 0.7 vol. % C2H4 Carbon
g 80 [µ
60
content 40 0.5 vol. % C2H4+Tar 20 0.5 vol. % C2H4 Carbon 6 g/Nm³ Tar 0 250 300 350 400 450 500 550 Reactor oven temperature [°C]
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
95 Methanation Tests with Contaminated Syngas - Results
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.
96 Methanation Tests with Contaminated Syngas - Results
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.
97 Methanation Tests with Contaminated Syngas - Results
1.0 1.6 14 ppm 130 ppm 14 ppm H2S
] 0.9 - 1.4 0.8 1.2 0.7
1 S S [g]
0.6 2 0.5 0.8
0.4 0.6
0.3 AdsorbedH 0.4 0.2
Normalizedactive catalyst area[ 0.1 0.2 0.0 0 0 50 100 150 200 Runtime [h]
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.
0.4
] 2.5
] Syngas 0.3
2.0 Syngas
kWh
kWh /
1.5 ct 0.2
1.0 0.1
0.5 Catalyst costs [ Catalyst Catalyst consumption[g/
0 0.0 0 1 2 3 4 5 6 7 8 9 10 H S concentration [ppm] 2 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.
98 Methanation Tests with Contaminated Syngas - Results
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. 1.2 20 µgC/gCatalyst·h 2.5 1.0
0.8 17 8.5 9 133
content [vol. [vol. content%] 0.6
4 5 3.5 H
2 60 C 0.4
0.2 -0.2 0 0.2 0.4 0.6 0.8 1 1.2
H2S concentration [ppm] Figure 7.21: Influence of C2H4 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
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. 1.2
1.0 506 187 61
0.8 29 8 99 62
content [vol. [vol. content%] 0.6
4 19
H
2 C 0.4
C2H4 + 6 g/Nm³ Tar 20 µgC/gCatalyst·h 0.2 -0.2 0 0.2 0.4 0.6 0.8 1 1.2 H S concentration [ppm] 2 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
99 Methanation Tests with Contaminated Syngas - Results
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.
Clean Ni-catalyst Partial, even Full, even Full, agglomerated carbon deposition carbon deposition carbon deposition
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.
a) 200 µm b) 100 µm
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
100 Methanation Tests with Contaminated Syngas - Results
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.
10 µm 400 nm
Figure 7.25: SEM photos of polymeric carbon filaments resulting from C2H4
2 µm 400 nm
Figure 7.26: SEM photos of polymeric carbon layers resulting from C2H4
101 Methanation Tests with Contaminated Syngas - Results
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.
102 Bench-Scale Tests with Real Syngas from Gasification
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.
103 Bench-Scale Tests with Real Syngas from Gasification
Methanation and gas-cleaning test rig Gasifier
Gas mixing station
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.
104 Bench-Scale Tests with Real Syngas from Gasification
Exhaust Fuel Purging Pressurized air reservoir Feed-lock system N2
P Pressure
Screw conveyor sensor (N ) Pressure Automatic 2 filter cleaning sensor Pressurized air P Port for N2-purging and pressurizing Sinter Pressure SPE-sample Exhaust metal filter retaining SPE
valve
Outlet / to 5 thermo- methanation T P couples D test rig Cyclone Differential pressure filter sensor (filter)
Reactor oven P Pressure sensor (reactor) with fluidized bed reactor Differential pressure PD sensor (reactor) Steam P D regulation
valve Pressure F sensor
(steam) MFC N2: max. 3500 ml/min N2 Orifice flow P measurement Demineralization Steam Water generator
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.
105 Bench-Scale Tests with Real Syngas from Gasification
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).
From gas Natural mixing station Flare gas Trace heated lines
From gasifier 2 thermo- 2 thermo- 16 thermo- couples couples E E E P P P S couples T S T S T Port for Port for Port for SPE- SPE- SPE-
r
sample o sample sample s
n e s )
r g e i g r n a i n
u l i s t d o s a e o e z e c
i r
h r r
p PD
u o l e s t a n s i c t o e a z r n
e P r e 3 ( n r
e Hot gas Hot gas f f i
D cleaning 2 cleaning 1 Methanation reactor To GA / gas To GA / gas sample bag sample bag To GA / gas sample bag Volumeflow measurement unit
Condenser Gas meter Silica gel
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.
106 Bench-Scale Tests with Real Syngas from Gasification
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.
From Exhaust Tar sampling according tar protocol gasifier
Gas analyzer H2, O2 -20°C +20°C Volume-sampling Washing bottles Gas analyzer module CO, CO2, CH4 From gas cleaning 2
Activated From gas cleaning 1 carbon From methanation
-20°C +20°C Pump Condenser unit unit
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.
107 Bench-Scale Tests with Real Syngas from Gasification
Table 8.1: Fuel parameters for the used lignite and biomass
Wood pellets RWE PowerSPLIT (ENplus-A1) [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.
108 Bench-Scale Tests with Real Syngas from Gasification
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.
109 Bench-Scale Tests with Real Syngas from Gasification
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.
60 Lignite (left bar) 50 Woody biomass (right bar) Typical range 40
30
20 Gas Gas composition%] [vol. 10
0
H2H2 COCO CO2 CO2 CH4CH4
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.
110 Bench-Scale Tests with Real Syngas from Gasification
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.
50 Lignite (left bar) 40 Woody biomass (right bar) Typical range
30
20
Gas Gas composition%] [vol. 10
0
H2H2 COCO CO2 CO2 CH4CH4 N2N2 H2OH2O
Figure 8.6: Mean permanent gas composition from gasification of woody biomass and lignite, gasification temperature ≈800°C, σ ≈6 (biomass), σ ≈4 (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.
1 Lignite (left bar) Woody biomass (right bar) Typical range 0.1
0.01 Gas Gas composition%] [vol.
0.001
C2H4C 2[vol.H4 %]C2H6C 2[vol.H6 %]C3H6C [vol.3H6 %]C3H8C [vol.3H8 %]C4H10C4 H[vol.10 %]
Figure 8.7: Mean C2-C4 content from gasification of woody biomass and lignite, gasification temperature ≈800°C, σ ≈6 (biomass), σ ≈4 (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
111 Bench-Scale Tests with Real Syngas from Gasification 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].
10 Lignite (left bar) Woody biomass (right bar)
1
0.1 Tarconcentration [g/Nm³]
0.01
Indane
Phenol
Indene
Pyrene
Xylenol
Toluene
Benzene
Biphenyl
Fluorene
Anthracene
Naphthalene
Phenanthrene
Fluoranthrene
Acenaphthene
Cresol(m,o) p,
Xylene (m, p,o)
Acenaphthylene 1-Methylnaphthalene 2-Methylnaphthalene Figure 8.8: Mean tar concentrations from gasification of woody biomass and lignite, gasification temperature ≈800°C, σ ≈6 (biomass), σ ≈4 (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.
112 Bench-Scale Tests with Real Syngas from Gasification
1000 Lignite (left bar) Woody biomass (right bar) 100 Typical range
10
Gas composition Gas [ppm] 1
0.1
H2S H[ppm]2S COS COS [ppm] CS2 [ppm]CS2 NH3NH [ppm]3 HCl
Figure 8.9: Mean contaminations from gasification of woody biomass and lignite, gasification temperature ≈800°C, σ ≈6 (biomass), σ ≈4 (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.
10000 Gasification Gas Cleaning
1000
100
10
Gas Gas composition[ppm] 1
0.1 H2SH [ppm]S COS COS [ppm] NH3 NH [ppm] BTXBTX C2-C4C -C 2 3 2 4 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
113 Bench-Scale Tests with Real Syngas from Gasification
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.
114 Bench-Scale Tests with Real Syngas from Gasification
60 1.00
50
0.95 ] 40 -
30 0.90
20 Tarconversion [ 0.85 Tarconcentration [mg/Nm³] 10
0 0.80
Cresol
Xylene
Indane
Indene Phenol
Pyrene
Xylenol
Toluene
Benzene
Biphenyl
Fluorene
Anthracene
Naphthalene
Acenaphtene
Phenanthrene
Fluoranthrene
Acenaphtylene 1-Methylnapht. 2-Methylnapht. 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
115 Bench-Scale Tests with Real Syngas from Gasification
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.
40 550 35.7
] 35
500
C] Catalyst
30 °
/g Temperature 450
25 Carbon
mg 20 400
15 350
10 8.7 Reactortemperature[ 300
Carboncontent [ 5 0.69 0.10 0.11 0 250 0.0 0.2 0.4 0.6 0.8 1.0 Scaled reactor length [-]
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.
116 Bench-Scale Tests with Real Syngas from Gasification
10 µm 1 µm
Figure 8.15: SEM-photos of polymeric carbon filaments on a catalyst sample taken after longtime real gas tests
10 µm 400 nm
Figure 8.16: SEM-photos of laminar (graphitic) carbon deposits on a catalyst sample taken after longtime real gas tests
1 mm 200 µm
Figure 8.17: SEM-photo of cracks on a catalyst tab after 200 h runtime with real synthesis gas
117 Bench-Scale Tests with Real Syngas from Gasification
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.
550 10 h 0 h 500 40 h
C] 450 ° 90 h 150 h 400 190 h
350 Temperature[
300
250 0 0.2 0.4 0.6 0.8 1 Scaled reactor length [-]
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.
118 Bench-Scale Tests with Real Syngas from Gasification
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.
1.00
] -
0.95 Test 5 Test 1
0.90 Test 2 Test 3
0.85
Normalizedactive catalyst area[ Test 4 0.80 0 50 100 150 200 Runtime [h]
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.
119 Bench-Scale Tests with Real Syngas from Gasification
0.6
4.0 ]
0.5 3.5 ] Syngas
3.0 Syngas
kWh 0.4 kWh
Test 4 2.5 / ct 0.3 Test 3 2.0 € Test 2 1.5 0.2 Test 5 1.0 Test 1 0.1 Catalyst costs[
0.5 Catalyst Catalyst consumption[g/
0 0.0 0 50 100 150 200 Runtime [h]
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.
120 Conclusion
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.
Synthesis gas complexity 0.35 2.5
] Lignite
0.30 ]
Syngas 2.0 Coking Poisoning
0.25 Syngas
kWh
[g/ kWh
1.5 /
0.20 Biomass
ct
€ [ 0.15
0.7 vol. % C2H4 1.0 costs
consumption 0.10 0.5 0.05 Catalyst
Catalyst 0.5 vol. % C2H4 0.00 0.0 Syngas Syngas Syngas Syngas Real
+ C2H4 + C2H4 + C2H4 syngas + Tar + Tar + H S 2 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
121 Conclusion
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.
0.4
] 2.5
1 vol. % C2H4 ]
Syngas 0.3 Lignite
2.0 Syngas
kWh kWh 0.5 % /
Real syngas ct
1.5 € 0.2 Biomass [
Syngas + HC 1.0 costs
0.1 Syngas 0.7 %
0.5 Catalyst Catalyst Catalyst consumption[g/ 0 0.0 0 1 2 3 4 5 Sulfur concentration [ppm]
Figure 9.2: Influence of sulfur concentration and ethylene content on specific catalyst consumption
122 Conclusion
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.
123 References
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