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GASIFICATION OF BIOMASS IN SUPERCRITICAL WATER
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Promotion committee:
Prof.dr. W.H.M. Zijm Chairman University of Twente Prof.dr.ir. A. Bliek Secretary University of Twente Prof.dr.ir. W.P.M. van Swaaij Promoter University of Twente Dr.ir. W. Prins Assistant-promoter University of Twente/BTG BV Dr. S.R.A.Kersten Assistant-promoter University of Twente Dr.ir. L. van de Beld BTG B.V. Prof.dr.ir. J.A.M. Kuipers University of Twente Prof.dr.ir. M. Wessling University of Twente Prof.dr. E. Dinjus Forschungszentrum Karlsruhe
The research reported in this thesis was executed under:
1. A grant of the Netherlands Organization for Scientific Research – Chemical Sciences
(NWO-CW) in the framework of the research program “Towards Sustainable
Technologies”, subproject BIOCON with the financial contributions from Shell Global
Solutions International B.V. and the Dutch Ministries of Economic Affairs
(EZ/SenterNovem) and Environmental Affairs (VROM).
2. A grant of the European Commission in the framework of the “Super Hydrogen”
program EC-contract: ENK5-CT2001-00555.
3. A grant of the New Energy and Industrial Technology Development Organization
(NEDO) from Japan in the framework of the research program “Technical Feasibility
of Biomass Gasification in a Fluidized Bed with Supercritical Water”.
Cover designer: Vesna Smiljanić
Publisher: Wöhrmann Print Service, Zutphen, The Netherlands © Biljana Potic, Enschede, The Netherlands, 2006
No part of this work may be reproduced by print, photocopy or any other means without the permission in writing from the author.
ISBN 90-365-2367-2
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GASIFICATION OF BIOMASS IN SUPERCRITICAL WATER
Proefschrift
ter verkrijging van de graad van doctor aan de Universiteit Twente,
op gezag van de rector magnificus, prof. dr. W.H.M. Zijm,
volgens besluit van het College voor Promoties in het openbaar te verdedigen
op vrijdag 12 mei 2006 om 13.15 uur
door
Biljana Potic
geboren op 20 september 1970 te Novi Sad, Servië en Montenegro
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Dit proefschrift is goedgekeurd door de promotor: Prof.dr.ir. W.P.M. van Swaaij en de assistent-promotoren: Dr.ir. W. Prins Dr. S.R.A. Kersten
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Luna, Mila and Goran, you are my life and my orientation. I love you!
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Contents
Summary
1
Samenvatting
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Chapter 1 Introduction
11
Chapter 2 A High-throughput Screening Technique for Conversion in
Hot Compressed Water
37
Chapter 3 Gasification of Model Compounds and Wood in Hot
Compressed Water
53
Chapter 4 Pilot Plant
79
Chapter 5 SCWG Experiments in a Micro Continuous Flow Reactor
97
Chapter 6 Fluidization with Supercritical Water in Microreactors
111
Chapter 7 Reactor Design Considerations for Biomass Gasification in
Hot Compressed Water
131
Acknowledgements
141
List of Publications
145
Curriculum Vitae
147
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1
SSSuuummmmmmaaarrryyy
Application of biomass and waste, as a renewable and possibly sustainable energy source, has gained
an important role in the world’s future energy policy. The Dutch government, for instance, has set a
target for 2020, which states that 5% of the total primary energy production must be based on biomass
and waste input. For the year 2040 this could rise to 30%. The European Committee has put similar
directives forward.
An important fraction of the available biomass and waste streams has a high moisture content (say
more than 70 wt % water). Wet streams cannot be converted economically by thermal conversion
techniques like combustion, pyrolysis and gasification because of the large amount of energy required
for evaporation of water. Partial conversion of wet biomass by anaerobic biological processes,
producing convenient energy carriers like methane or ethanol, is possible and practiced for suitable
feedstock materials. Over the last decades research activities worldwide have been devoted towards
the development of new thermochemical processes, which can convert wet biomass efficiently and
economically. One of the novel technologies for conversion of wet biomass and waste streams is
gasification in supercritical water (SCWG: T > 3750C, P > 221 bar). Hydrogen and/or methane can be
produced by SCWG with selectivities that can be controlled by the process conditions and catalysis.
The research described in this thesis is dealing with the SCWG process. Due to the severe process
conditions (typically: T = 600°C, P = 300 bar and a corrosive environment), experimental
investigation on SCWG is expensive and time consuming. Despite these challenges, experiments were
conducted by different groups in the world revealing the influence of process conditions (temperature,
pressure, residence time, concentration of the organics, catalysis) on the yields and the selectivity of
the desirable gas products. All laboratory scale experiments reported in the past were conducted in
metal reactors. This was demonstrated to cause a difficult to quantify catalytic effect of the reactor
wall that makes interpretation and comparison of the available data uncertain. Therefore, part of the
present work is focusing on gasification experiments not influenced by catalytic effects.
For this non-catalytic gasification, a high throughput technique for mapping of the reaction space has
been developed. In this technique, quartz capillaries of 1 mm inside diameter are used as batch
reactors. SCWG experiments could then be performed in a safe, cheap, and quick manner; one
capillary measurement takes about 5 min. compared to several hours for earlier batch and continuous
methods. Via validation measurements with formic acid and glucose solutions it has been
2
demonstrated that the technique is sufficiently reliable for screening purposes including trend
detection.
The quartz capillary batch reactor technique has been subsequently used to study the non-catalytic
gasification of glycerol, glucose and pinewood in supercritical water. Mapping of the reaction space
has been done by performing over 700 experiments in which T, P, the heating time, the reaction time
and the concentration of the feedstock were varied. The most important observations were that the
pressure turned out to have no effect on the conversion rate to gas-phase products and product yields,
and that, without catalytic effects, complete conversion to the gas-phase is only possible for very
diluted feedstock solutions (<2 wt %). Compared with previous work in metal reactors, conversion
rates and final conversion to gas phase products was lower while the product gas composition was also
different. By adding Ruthenium on a TiO2 carrier to the capillaries, the potential of heterogeneous
catalysis has been demonstrated. When adding this catalyst to the capillaries, glucose solutions in the
range of 1 to 17 wt % could be gasified completely already at 600oC.
A first generation pilot plant (intake 30 l feed/h) for the supercritical water gasification process,
designed to gain operating experience and to identify problems related to the overall process from feed
preparation to product gas analysis, is described and operated. Non-charring feedstocks such as
glycerol and methanol solutions could be successfully gasified with this plant. However, charring
feedstocks like starch caused fouling and blockage in the reactor and heat-exchanger demonstrating
that already during heating up problems occur. Results have been compared with those of the micro
reactor systems used in this work. The observed differences have been explained by the presence and
extend of the catalytic activity of metal reactor walls, or absence of such activity.
As the operation of this large plant was expensive and cumbersome, a new micro continuous reactor
system was set-up. Unfortunately the use of quartz in this system led to too frequent failure of the
reactor and connections. Therefore, the reactors used in the system were tubular reactors of 1 mm
internal diameter made of different metals namely Inconel 625 and stainless steel (SS-316). In this
system, the influence of the process conditions and reactor construction materials on feedstock
conversion to different gas phase products has been examined. In a second series of experiments, the
empty tube reactor is placed in series with a packed bed of catalyst particles to study its influence on
conversion and selectivity. Catalysts used are 3 wt % Ru on TiO2 and charcoal of beech wood. After
the gasification, burn-off of carbonaceous deposits was performed and the CO2 production was
measured. The tested Ru catalyst increased strongly the carbon conversion to gas phase products,
probably by converting intermediates products that without catalyst would slip trough the reactor and
end-up as liquid dissolved products. Glucose produced carbonaceous by-products in the form of
deposits in the SCWG reactor. In a few tests, these deposits were combusted with air under controlled
3
process conditions. This burn-of or regeneration may be an important aspect of an industrial SCWG
reactor system, where the catalyst will most likely be in a fluidized state. As ultimately this convenient
continuous reactor set-up would be used for catalytic SCWG with a highly active catalyst compared to
the moderate catalytic influence of the reactor wall, the catalytic activity of the wall was accepted.
To pave the way for introducing a fluidized catalyst bed in a SCWG reactor and especially in a
continuous micro plant, a new micro-fluid bed technique for investigation of supercritical water
fluidization was designed and operated. A cylindrical quartz reactor of only 1 mm internal diameter
with a quartz ball as distributor was used for process conditions up to 5000C and 244 bar. Properties of
the fluid bed like: minimum fluidization velocity (Umf), minimum bubbling velocity (Umb), bed
expansion, and the identification of the fluidization regime were investigated by visual observation.
Dedicated 2D and 3D Discrete Particle Models (DPM) models were used to simulate the micro-fluid
beds for gas-solid fluidization (fluid density range from 16 to 230 kg/m3). It has been found that the
results for Umb from the simulations are in accordance with experimental data and somewhat higher
than predictions from existing empirical relations. Experiments in three different scale cylindrical
reactors namely: 26 mm, 12 mm and 1 mm inside diameter, showed that to mimic large-scale
homogeneous fluidized bed systems in the 1 mm reactor only small particles (d < 100 micron) can be
used.
Finally, based on the present work and literature results, the different requirements for a SCWG
system are listed and discussed. A preliminary conceptual design is proposed including
fluidized/circulating bed systems, heat-exchangers and a regeneration section. Such a system could
possibly cope with the operational problems expected for a pilot plant or commercial SCWG system
processing real biomass, but much more research and development work is required before it can be
realized.
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5
SSSaaammmeeennnvvvaaattttttiiinnnggg
Het gebruik van biomassa en afval als een hernieuwbare, en mogelijk duurzame energiebron, wordt
wereldwijd een steeds belangrijker onderdeel van het energie beleid. De Nederlandse regering
bijvoorbeeld, heeft in haar doelstelling voor 2020 bepaald dat 5% van de totale primaire
energieproductie gebaseerd moet zijn op biomassa en afval als grondstof. De Europese Commissie
heeft vergelijkbare richtlijnen voorgesteld.
Een belangrijk deel van de beschikbare biomassa- en afvalstromen heeft echter een hoog vochtgehalte
(meer dan 70 gewichtsprocent water). En zulke natte stromen kunnen niet op economisch
verantwoorde wijze worden omgezet met de traditionele thermische conversietechnieken
(verbranding, pyrolyse en vergassing) vanwege de grote hoeveelheid energie die er nodig is voor de
verdamping van water. Gedeeltelijke omzetting van natte biomassa met behulp van biologische
processen waarbij nuttige energiedragers zoals methaan of ethanol worden gevormd, is mogelijk en
wordt in de praktijk veelvuldig toegepast voor bepaalde typen biomassa.
De afgelopen twintig jaar is er op een aantal plaatsen in de wereld ook onderzoek verricht m.b.t. de
ontwikkeling van nieuwe thermochemische processen om natte biomassa efficiënt en economisch
verantwoord om te zetten in nuttige producten. Eén van de nieuwe technologieën voor de verwerking
van natte biomassa en afval is de vergassing in superkritiek water, in het Engels aangeduid met de
afkorting SCWG (Super Critical Water Gasification). Bij temperaturen boven 375oC en drukken boven
221 bar (het superkritieke punt van water) kan er waterstof en/of methaan gemaakt worden. De
productselectiviteit wordt geregeld met de keuze van procescondities en de toepassing van
katalysatoren.
Het onderzoek dat in dit proefschrift is beschreven gaat over het SCWG-proces. Gewoonlijk zijn de
experimenten tijdrovend en kostbaar vanwege de strenge procescondities (reactie in een corrosief
medium, typisch bij T = 600oC en P = 300 bar). Ondanks deze bezwaren zijn er door verschillende
onderzoeksgroepen in de wereld experimenten uitgevoerd waarbij duidelijk is geworden wat de
invloed is van de procescondities (temperatuur, druk, verblijftijd, biomassaconcentratie en katalyse) op
de opbrengst en de verdeling van de gewenste producten. Alle laboratoriumexperimenten waarover in
het verleden is gerapporteerd zijn uitgevoerd in reactoren van metaal. Inmiddels is ook aangetoond dat
dit de interpretatie en vergelijking van resultaten bemoeilijkt omdat de metaalwand een onbekend
6
katalytisch effect veroorzaakt. Daarom was een deel van het onderzoek voor dit proefschrift gericht op
vergassingsproeven waarbij zulke ongewenste katalytische effecten zijn uitgesloten. Er is een techniek
ontwikkeld om snel, en veel experimenten te doen (high-throughput screening) ten einde de mogelijke
reactiecondities in kaart te brengen. Hierbij zijn kwartsglazen capillairen van 1 mm inwendige
diameter gebruikt als “batch” reactor. De SCWG experimenten konden op deze wijze veilig, goedkoop
en snel worden uitgevoerd; een experiment met zo’n capillair duurt maar vijf minuten terwijl in het
verleden een experiment vele uren in beslag nam.
De techniek is eerst gevalideerd, door de omzetting van mierenzuur en glucose te bestuderen, en bleek
voldoende betrouwbaar voor snelle verkennende proeven en het vastleggen van trends. Vervolgens is
de capillaire techniek gebruikt om de niet-katalytische vergassing van glycerol, glucose en dennenhout
in superkritiek water te onderzoeken. Het operatiegebied voor de vergassingsreactie is in kaart
gebracht door 700 experimenten uit te voeren waarbij de temperatuur, de druk, de opwarmsnelheid, en
de concentratie van de organische stof in het water zijn gevarieerd. De belangrijkste waarnemingen
waren: a) dat de druk geen invloed bleek te hebben op de snelheid van de omzetting naar gasfase
producten en de productopbrengsten, en b) dat een volledige omzetting naar gasfase producten zonder
toepassing van katalyse alleen maar mogelijk is voor verdunde oplossingen (minder dan 2
gewichtsprocent organische stof in water). In vergelijking met het eerdere werk uitgevoerd in metalen
reactoren, bleek de omzettingssnelheid en de uiteindelijke omzetting naar gasfase producten lager te
zijn terwijl ook de productsamenstelling afweek van eerder gepubliceerde resultaten. Door ruthenium
op een TiO2-drager als katalysator toe te voegen aan de capillairen, kunnen glucoseoplossingen met
een concentratie variërend van 1 tot 17 gew. % al bij 600oC volledig vergast worden, waarmee de
potentie van katalyse voor SCWG is aangetoond.
In dit proefschrift wordt ook het ontwerp, de bouw en het testen van een volledige pilot-plant
installatie beschreven. Deze SCWG pilot-plant heeft een capaciteit van 30 liter vloeistof per uur, en is
bedoeld om procesgerelateerde problemen van allerlei aard te identificeren en op te lossen, en ervaring
te krijgen met continue operatie. Voedingen die geen coke vormen, zoals glucose en methanol, konden
in deze pilot-plant goed worden vergast. Maar de vergassing van zetmeel, leidde tot vervuiling en
tenslotte tot verstopping van de warmtewisselaar en de reactor, waaruit valt af te leiden dat er al
problemen ontstaan tijdens de opwarming van de voeding naar de gewenste reactietemperatuur. De
resultaten van succesvolle pilot-plant zijn vergeleken met die van de kleinschalige
laboratoriumproeven en de waargenomen verschillen zijn verklaard met de katalytische activiteit van
de metalen wand van de pilot-plant reactor.
7
Het bedrijven van de pilot-plant is duur en lastig gebleken. Daarom is er een veel kleinere opstelling
gebouwd voor continue SCWG-vergassingsproeven. Helaas bleek het niet goed mogelijk om in deze
micro-opstelling een kwartsglazen reactor te gebruiken. Met name de verbinding met staal leidde vaak
tot breuk tijdens het op druk en temperatuur brengen van de reactor. Er is tenslotte besloten om met
een metalen buisreactor te werken van 1 mm inwendige diameter, gemaakt van Inconel 625 of
roestvrij staal (SS-316). De opstelling is gebruikt om de invloed van de procescondities, en het type
metaal van de reactorwand, op de conversie van de voeding naar de verschillend productgassen te
onderzoeken. En in een tweede serie experimenten is er een gepakt bed van katalysatorpoeder in serie
geplaatst met een lege buisreactor om de invloed van de katalysator op de conversie en selectiviteit te
bestuderen. De beproefde katalysatormaterialen zijn 3 gew.% ruthenium op TiO2 en kool van
beukenhout. Na de vergassing werd de geaccumuleerde kool in de reactoren afgebrand met lucht, en
de hoeveelheid ervan bepaald door de CO2 productie te meten. Er bleek dat de ruthenium katalysator
de koolstofconversie naar gasfaseproducten sterk verbetert, waarschijnlijk doordat intermediare
producten, die zonder de aanwezigheid van een katalysator door de reactor slippen en als opgeloste
componenten in de waterfase terechtkomen, in dit geval wel worden omgezet. Het gebruik van glucose
als voeding veroorzaakte koolachtige bijproducten in de vorm van waarneembare afzettingen in de
reactor. In enkele proeven zijn deze koolafzettingen verbrand met lucht onder gecontroleerde
omstandigheden. Het afbranden, c.q. uitbranden van koolafzettingen zal voor een grootschalig SCWG-
proces erg belangrijk zijn om het systeem te reinigen en de katalysator (mogelijk in een fluid bed
reactor) te regenereren.
De kleine continue opstelling is gemakkelijk te bedrijven en zal in de toekomst verder worden
gebruikt voor het testen van katalysatoren voor SCWG. In dat verband is een beperkte katalytische
activiteit van de wand acceptabel omdat deze wegvalt tegen de veel grotere activiteit van de
toegepaste katalysatordeeltjes.
Om de introductie van een fluid bed als katalytische reactor in SCWG voor te bereiden, en met name
ook voor toepassing van fluïdisatie in kleine laboratoriumopstellingen, (microreactor), is er een
nieuwe microtechniek ontwikkeld en toegepast om de fluïdisatie van katalysatordeeltjes in
superkritiek water te onderzoeken. Een cilindrische kwartsbuis van slechts 1 mm inwendige diameter,
met een kwartskogeltje onderin om het fluïdisatie medium te verdelen, is gebruikt voor proeven bij
temperaturen tot 500oC en drukken tot 244 bar. Eigenschappen van het gefluïdiseerde bed, zoals de
minimum fluïdisatiesnelheid (Umf), de minimale belsnelheid (Umb), de bedexpansie, en het
fluïdisatieregiem, zijn onderzocht aan de hand van visuele waarnemingen. Bovendien zijn er
geavanceerde twee- en driedimensionale modellen toegepast (DPM: discrete particle models) om gas-
vast fluïdisatie te simuleren voor zulke micro fluid bedden met dichtheden variërend van 16 tot 230
kg/m3. De uitkomsten van de simulaties zijn vergelijkbaar met de experimentele resultaten, maar zijn
8
iets hoger dan wat de gangbare empirische correlaties voorspellen. Experimenten waarbij de
reactordiameter is gevarieerd (26, 12, en 1 mm inwendige diameter) hebben laten zien dat het gedrag
van grootschalige homogeen gefluïdiseerde bedden goed kan worden nagebootst, mits er in de 1mm
reactordeeltjes worden gebruikt die kleiner zijn dan 100 µm.
Tenslotte is er, op basis van dit onderzoek en de literatuur, een overzicht gemaakt van de verschillende
eisen die aan een SCWG-systeem gesteld moeten worden. Er is een voorlopig ontwerp gepresenteerd
waarin gefluïdiseerde bedden en vaste stof circulatie zijn opgenomen, om warmtewisseling tussen
voeding en reactor effluent alsmede katalysatorregeneratie, mogelijk te maken. Zo’n systeem zou
mogelijk de operationele problemen die te verwachten zijn van een pilot-plant of een commerciële
SCWG installatie voor echte biomassa, kunnen ondervangen. Maar er is meer onderzoek en
ontwikkelingswerk nodig om dit allemaal te realiseren.
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CCChhhaaapppttteeerrr 111 IIInnntttrrroooddduuuccctttiiiooonnn
Biomass as sustainable energy source
During the last three decades, concerns on the use of fossil sources have increased enormously. Care
about the environment and uncertainties about the security of energy supply are the main reasons for
these concerns.
It is a well-established fact that the concentration of carbon dioxide in the atmosphere has increased
notably over the past 150 years (United Nations Environment Programme, 2005). There is also
considerable evidence that the increase of the carbon dioxide concentration in the atmosphere since the
industrial revolution is primarily caused by human activity (United Nations Environment Programme,
2005). How the combustion of fossil fuels and the resulting unbalanced global CO2 flows will affect
life on earth on the long term is strictly speaking not known. The fact that the consequences are
unknown should however be enough to start considering sustainable energy systems. Our measures to
counteract the increasing CO2 level should be guided by the precaution principle. It is by far the safest
strategy to assume that mankind’s increasing fossil CO2 emissions turns out to do irreversible harm to
the well-being and prosperity of generations to come. At present, we are able to develop the required
technologies for the conversion of sustainable sources into heat, power, fuels, and chemicals. The most
important hurdles for rapid development of such technologies are economic and mainly related to the
still low price of crude oil, natural gas, and coal.
Besides these concerns on global warming, the world’s current political situation also gives rise for
reconsideration of energy strategies. The majority of the crude oil reservoirs and reserves (90%,
(World Energy Council, 2004)) are located in potentially political unstable countries in the Middle
East, which makes availability principally unsure. This motivates several countries to invest in the
development of alternatives from locally available fossil resources such coal, shale and natural gas, but
also increasingly from renewable sources like hydro, solar, wind, and biomass (e.g. State of the Union
12
Speech of US president Bush, January 2006), (Bush, 2006, The National Renewable Energy
Laboratory's, 2006).
Moreover, it is projected by most scenarios that the world’s oil and gas reserves are limited to ca. 60
years (e.g. BP statistical review (BP, 2005)). Obviously, this will have far reaching consequences for
the security of supply in a fossil dominated energy system. For the distant future it is evident that we
have to rely on sustainable sources, no matter if the fossil reserves turn out to sufficient for 200, 500 or
1000 years.
Because biomass is built-up from carbon extracted from the atmosphere in the form of CO2, which, in
a short cycle, is returned again to the atmosphere after decay or combustion, the utilization of biomass
does not overall influence the atmospheric CO2 concentration. While considering alternatives for fossil
feedstocks, it becomes clear that biomass is the only sustainable source containing carbon and thus
allows direct matching with existing fossil derivatives like transportation fuels and chemicals.
Integrating and partnering with the existing fossil-based industries and logistic facilities is of utmost
importance in the transition towards a completely sustainable society, because it lowers the capital
investments and offers guaranteed markets for the products (Van Swaaij et al., 2004). Without this
initial partnering with the fossil industry it is to be doubted if the very large capital investments
required for achieving the directives on sustainable energy of the Dutch government and the European
Union will ever be brought up. Moreover, the total amount of biomass in the form of forestry,
agricultural and plantation residues is huge (say the equivalent of half the world’s crude oil
production, (Groeneveld, 2000). Moreover, even possible to grow biomass especially for energy
production (e.g. seeds for bio-diesel, and corn or sugar cane for bio-ethanol), also as part of
agricultural policies. When combining biomass-based processes with CO2 sequestration, the CO2
concentration in the atmosphere can even be reduced. Biomass as a renewable source should of course
be used in a sustainable fashion. Several factors should be considered like competition with the food
chain, energy for production and transportation of biomass, ecological aspects, bio-diversity soil
exhaustion, etc. As an example it should be realized that biomass always contains minerals, which
should be carefully considered in relation to a sustainable and ecological responsible use of biomass.
Biomass conversion
For the conversion of biomass to energy or energy carriers, the moisture content of the biomass is of
crucial importance in view of the net lower heating value and the flame temperature (see Figure 1.1).
The net lower heating value is defined here as:
LHVnet = (1 – Xmoisture) x LHV – Xmoistue x ∆Hvap
in which Xmoisture is the mass fraction of moisture in the biomass, LHV is the lower heating value of the
dry fraction and ∆Hvap the enthalpy of vaporization of water. This definition of the net lower heating
value thus includes the energy required for vaporization of the moist.
Biomass with a moisture content of more than 60% on dry weight basis cannot be combusted
efficiently. Drying of the biomass is expensive and will consume rapidly a large fraction of the energy
available in the wet biomass unless special measures are taken. Water can be removed by methods like
sun drying, mechanical pressing, utilization of hot flue gas, multiple effect evaporation, vapor
recompression, and superheated steam drying. Sun drying in the open air is the cheapest but not
always possible. Other drying techniques are costly and generally not yet optimized for bulky biomass.
That is why conversion technologies for wet biomass streams, which do not request the initial
evaporation of the water associated with the biomass, can be very attractive. Examples of wet biomass
streams are by-products from food processing and certain streams in future biorefineries.
0.0 0.2 0.4 0.6 0.8 1.00
5
10
15
20
25
moisture content [kg/lg]
Flame temperature [oC/100]
LHVnet [MJ/kg]
Figure 1.1 Influence of the moisture content on the net heating value of the feedstock and the
adiabatic flame temperature in air.
In this introduction we will consider briefly the technologies that utilize dry or dried biomass, after
which the conversion technologies for wet biomass will be discussed. 13
Dry biomass conversion
Examples of dry biomass are: rice husk, wood chips, straw, saw dust, cotton stalk, nut shell, etc. Dry
biomass can be used as a solid fuel in direct combustion or be converted thermochemically for the
production of various secondary energy carriers. Figure 1.2 gives an overview.
Apart from the classical combustion in household stoves, grate boilers etc., also more advanced boilers
like fluidized bed and circulating fluidized bed boilers are used now to generate energy from e.g. wood
and straw (Quaak et al., 1999). A recent development is large-scale combustion of biomass in coal-
based power stations, by co-firing up to a level of 10 wt % (IEA, 1998). This is indeed a quick and
efficient way to reduce fossil carbon-based CO2 emissions and produce “green electricity”, for which
reason it is stimulated in some countries by fiscal measures and subsidies. The combustion with
energy generation of organic fractions in waste incineration is also counted sometimes as a
contribution to the reduction of CO2 emissions. Here, of course, only organic fractions from non-fossil
sources should be counted, e.g. thus excluding most plastics.
Dry Biomass
directcombustion
flash pyrolysis gasification oil
pressing
ash ash ash
organic liquid(+ water)
fuel gas orsynthesis gas oil
liquid fuelshydrocarbonsmethanol etc.
Dry Biomass
directcombustion
flash pyrolysis gasification oil
pressing
ash ash ash
organic liquid(+ water)
fuel gas orsynthesis gas oil
liquid fuelshydrocarbonsmethanol etc.
Figure 1.2 A simplified overview of the conversions of dry biomass into secondary energy carriers.
Flash pyrolysis (Bridgwater et al., 1999; Wang et al., 2005) is a process in which dry biomass is
heated rapidly to produce vapors that can be condensed to a liquid product that is usually indicated as
bio-oil. If carried out properly, about 70 - 80% of the product can be converted into bio-oil. Some char
and incondensable gas is produced, which can be used to generate heat and power for the process
itself, or for the local market. Conversion to bio-oil increases the bulk density of the material from
applied directly in engines, furnaces, turbines etc. After gasification in existing oil gasifiers (from 200-14
15
300 kg/m3 to 1300 kg/m3 allowing for long distance transport and easy storage. The oil can be SHELL,
LURGI or TEXACO), the produced bio-oil syngas can be used for the production of chemicals,
Fischer Tropsch diesel, or as a fuel-cell feedstock. Biomass can also be pyrolyzed slowly, that is in a
more traditional way, to produce e.g. charcoal from wood as a secondary solid fuel. But that will not
be discussed here.
Dry biomass can be gasified (Beenackers and Maniatis, 1997; Kersten, 2002; Kersten et al., 2003) by
partial oxidation with air or oxygen/steam mixtures, to transfer the combustion value of the biomass to
fuel gas or syngas respectively. A wide range of processes to achieve this gasification has been
developed. Over the last decades, the research group in Twente has analyzed the performance and
made experimentally verified reactor models of various gasifier types, such as the co-current moving
bed (Groeneveld, 1980), fluidized bed (Van den Aarsen, 1985), and circulating fluid bed (Kersten,
2002). For large-scale operation entrained flow gasifiers are attractive (Higman and Van der Burgt,
2003). Product gas can be fired in boilers, turbines and engines or shifted to hydrogen for application
in fuel cells. It can be converted to liquid fuels like methanol or hydrocarbons via the Fischer-Tropsch
synthesis.
If biomass consists of oilseeds, oil can also be extracted and converted to “green diesel” but this will
not be discussed here (Boerrigter et al., 2003).
Wet biomass conversion
Examples of wet biomass streams are vegetable, fruit and garden waste, waste streams from
agricultural, food and beverage industries, manure, sewage sludge, and some household wastes.
Seaweed and micro-algae are examples of cultivated wet biomass crops. According to estimates of
ECN (Hemmes, 2004), between 5.3 and 12 million ton (dry matter) of wet streams are annually
available in The Netherlands. Besides, these streams that are already available it is expected that
within a future biorefinery also diluted side streams (by-products) are present. A bio-refinery is an
integrated concept aiming at full utilization of biomass, in which different fractions of biomass are
converted in large-scale plants (economy of scale) or in standardized small-scale units (economy of
numbers) in an economically optimal product slate.
This thesis deals with the conversion of wet biomass streams. An overview of these processes is given
in Figure 1.3. By biological processes methane can be produced which will easily separate from the
bioreactor fluid, leaving a non-fermentable organic fraction (Hill and Bolte, 2000).
“Wet” Biomass
Pressure cooker
P=300 bar, T=600oC
Pressure cooker
P=150 bar, T=300oC
Phase separation
Water + Ash Methane, alcohols etc.
Phase separation
Water + Ash Hydrophobic oil
Phase separation
Water + Ash Gaseous fuel (H2)
HTU SCWG
Figure 1.3: Simplified overview of conversion processes for wet biomass.
Bio-ethanol can be produced also from various feedstock materials, especially from starch and sugars.
It can also be recovered from the bioreactor by refined distillation or extraction processes (Reith et al.,
2002). Biological processes with their special feedstock and operation conditions differ considerably
from the thermochemical processes and will not be further discussed here. Biological and
thermochemical conversion process will both be needed to develop a biorefinery that utilizes the
whole lignocellulosic feedstock. In this refinery a multitude of conversion steps will be present; by-
products from biological conversions will be used as feedstock for thermochemical conversions and
vise versa.
Two other conversion processes indicated in Figure 1.3 are thermochemical, viz. HTU® and
supercritical water gasification (SCWG). These processes are operated at elevated temperatures, and
under high pressure to avoid the evaporation step of water. They both require extensive heat exchange
between the feed inlet and product outlet streams in counter current operation to become energy
efficient.
HTU® or HydroThermal Upgrading is a thermochemical process in subcritical water in which
biomass is converted to liquid bio-crude with a reduced oxygen content (Goudriaan and Peferoen,
1990). In some hydrothermal processes by application of a catalyst the production of methane is aimed
for (Elliott et al., 2004).
Gasification of biomass in supercritical water for the production of hydrogen or methane rich gas is
the topic of this thesis. 16
Gasification of wet biomass feedstocks in supercritical water
As stated above, the high moisture content of so-called wet biomass streams makes conventional
thermochemical technologies inefficient due to the high-energy requirement for water evaporation (2.4
MJ/kg at atmospheric conditions). Although wet biomass (moisture > 60 wt %) has a very low overall
heating value, products with a high heating value can still be extracted from it by applying advanced
conversion processes. Biomass gasification in hot compressed water (SCWG, 600°C, 300 bar) is
considered as a promising technique to convert such wet streams into a gas that is rich in either
hydrogen or methane depending on the operating conditions and applied catalysis. In hot compressed
water (P > 200 bar), the heat effect associated with water evaporation is marginal compared to that at
ambient conditions (∆Hvap becomes zero at Pc). Therefore, by practicing counter-current heat
exchange between the feed stream and the reactor effluent, high thermal efficiencies can be reached
despite the low dry matter content of the feedstock. In Figure 1.4 a conceptual flow sheet of the
SCWG process is given.
Figure 1.4 Conceptual flow sheet of the SCWG process.
It is therefore crucial for the process that the sensible heat content of the reactor effluent is utilized as
far as possible to pre-heat the feedstock stream (mainly water) to reaction conditions (see Figure 1.4).
The efficiency of the heat exchange in relation to the applied pressure can be calculated from the heat
balance for a counter-current shell and tube heat-exchanger. The result is presented in Figure 1.5, in
which the heat-exchanger efficiency is plotted as a function of the operating pressure and the available
17
area per unit throughput in kilogram per second. In case of an infinitely long heat-exchanger, the steep
asymptotic approach to 100% efficiency above 200 bar is a result of the sharp decrease of ∆Hvap to
zero beyond that pressure. For heat-exchangers with finite surface area, the effect of the operating
pressure is less pronounced. In practice, a hundred percent transfer of the available heat in the reactor
effluent to the feedstock stream is impossible. In fact, efficiencies of ca. 75% are typical for liquid-
liquid shell and tube heat-exchangers (Kersten et al., 2004). For such efficiency the operating pressure
should be ca. 200 bar in case of 50 m2 per kg/s throughput or 300 bar in case of 25 m2 per kg/s
throughput (see Figure 1.5).
0 50 100 150 200 250 300 3500.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9
1.0
cold,out cold,inHE
hot,in cold,in
H -Hη =
H -H
effic
ienc
y (η
HE) [
-]
pressure [bar]
AHE = 10
2550
100∞
Supercritical Pressure
Figure 1.5 Calculated efficiencies of a water-water counter-current heat-exchanger plotted versus the
operating pressure for different surface areas, AHE = area (m2) per unit throughput (kg/s). The flow
rates (kg/s) on both sides were assumed to be equal. U = 1000 W/(m2.K), Inlet conditions: Thot,in =
600oC, Tcold,in = 25oC.
Promises of biomass gasification in supercritical water are that:
● The technology is suitable for efficient processing of wet feedstock (> 70 wt % moisture);
evaporation of water is avoided and feedstock/product heat exchange is quite well possible.
● Contrary to anaerobic digestion and fermentation processes, the technology allows in principle
for complete feedstock conversion.
● The product gas is made available at high pressure (> 250 bar) and, for its application,
expensive gas compression can be avoided.
18
19
● The product gas is clean; minerals, metals, and the undesired gases like CO2, H2S and NH3
(which have a high solubility in compressed water) remain in the water phase and can thus be
recovered.
● The product gas is not diluted with inert gas.
● The selectivity towards either methane or hydrogen can be controlled with temperature,
pressure and the application of catalysts.
● Sequestration of (pure) CO2 seems convenient.
These promises however go together with a series of problems that need to be solved in the process
development. Pumping of biomass slurries to pressures of up to 300 bar is a challenge. The high
temperatures and pressures involved put serious demands on the construction materials to be used,
especially because corrosion problems are expected. Heat exchange between the reactor feed and
effluent is required to make the process efficient, but heating of a biomass slurry is likely to cause
fouling and plugging in the heat-exchanger as the biomass starts to decompose already around 250oC.
Ash deposits may cause similar problems, and an effective ash removal system must therefore be part
of the reactor/process. Although carbon formation can be suppressed by applying high temperatures or
using a catalyst, the process should include an option to burn this carbon off the catalyst (when
applied) and the reactor walls.
When the future process development is successful in solving the problems of heat exchange,
corrosion, and fouling, interesting applications are foreseen.
Compressed methane can be used for blending with synthetic natural gas and compressed natural gas
for motorcars. In The Netherlands CH4/H2 mixtures could be fed to the natural gas pipelines.
Compressed hydrogen is an attractive feedstock for upgrading processes within the bio refinery
concept and for storage of hydrogen for mobile and stationary fuel cells. CO2 sequestration is
especially interesting because it would, together with the use of biomass as a feedstock material, create
the opportunity to reduce the amount of CO2 in the atmosphere (carbon sink).
Relevant properties of supercritical water
The vapor pressure curve in the P-T diagram of water, which indicates the co-existence of liquid and
vapor, ends in the critical point at Tc = 373.95 and Pc = 220.64 bar.
Figure 1.6 Phase diagram of water.
At the critical point, the distinction between vapor and liquid disappears. The right upper quadrant
enclosed by the critical isotherm Tc and the critical isobar P = Pc is the area where the fluid is
supercritical. The properties of supercritical water are quite different from those of the normal liquid
or steam at atmospheric pressure. For instance, at the critical point the density ρ is around 0.3 g/cm3
(versus 1 g/cm3 for liquid water at ambient conditions), the dielectric constant e is 5 (versus 80 for
liquid water at ambient conditions) and the ion product Kw = [H+][OH-] is 10-11 (versus 10-14 for
liquid water at ambient conditions). The most striking feature of supercritical water is the possibility to
manipulate and control its properties around the critical point by tuning the temperature and pressure.
For instance, by raising the pressure from 10 to 50 MPa at a temperature of 400oC, the dielectric
constant is increased from 2 to 14, the ion product from 10-28 about to 10-12 , and the density from less
than 0.1 to about 0.5 g/cm3. As an important consequence of the change in the dielectric constant,
supercritical water behaves like a non-polar solvent and exhibits a high solubility towards non-polar
organic compounds like benzene. Also gases like oxygen, nitrogen, carbon dioxide and methane are
completely miscible in supercritical water. Contrary, the solubility of inorganic salts like NaCl is
decreased to very low values. All relevant properties of (supercritical) water are known accurately (see
e.g. NIST).
20
21
Chemistry and thermodynamics
Biomass gasification in supercritical water is the result of thermal decomposition reactions (catalyzed
or not), followed by a multitude of reactions between intermediates and end-products (Kruse et al.,
2002; Kruse and Gawlik, 2003). In the vicinity of the critical point where the ion product is high (10-
11), the H+ concentration is about 30 times higher than at ambient conditions offering increased
opportunities for acid catalyzed reactions (Buhler et al., 2002; Kabyemela et al., 1997; Kabyemela et
al., 1997; Kabyemela et al., 1998; Kabyemela et al., 1999). Presumably, at the higher temperatures
where hydrolysis as a result of the very low ion product Kw is impossible, radical reactions (pyrolysis
and cracking) will control the chemistry (Antal et al., 1993; Antal et al., 1999; Antal et al., 2000; Hao
et al. 2003; Lee et al., 2002; Potic et al., 2002; Potic et al., 2004; Potic et al., 2005; Schmieder, 1999).
Steam reforming and methanization of the biomass cover the extremes of possible stoichiometric
equations:
C6H10O5 + 7H2O → 6CO2 + 12H2 (1)
C6H10O5 + 1H2O → 3CO2 + 3CH4 (2)
Thermal decomposition of biomass and the subsequent reforming of intermediate fragment molecules
at higher temperatures could profit from the supercritical conditions, because of the good miscibility
of organic compounds and gases in supercritical water. It offers an opportunity to conduct chemistry in
a single phase that otherwise would have to occur in a multiphase system. As a consequence, there are
no interphase mass transfer limitations reducing the reaction rates, and higher concentrations of
reactants and intermediate products can be maintained. There is an extensive literature (Kabyemela et
al., 1997; Kabyemela et al., 1997; Kabyemela et al., 1997; Kabyemela et al., 1998; Kabyemela et al.,
1999; Kruse et al., 2002; Kruse and Gawlik, 2003; Kruse et al., 2003; Minowa et al., 1998; Mok and
Antal, 1992; Saka and Ueno, 1999; Sinag et al., 2003; Sinag et al., 2004; Stein et al., 1983) available
on the elucidation of the chemical pathways of biomass conversion in hot compressed water. Although
there seems to agreement on certain elements of the whole degradation chain, an overall model
suitable for reactor design is not present yet. As there is no mechanistic reaction path model available,
equilibrium calculations could be used to produce indicative results for comparison with experimental
results.
The thermodynamic calculations for supercritical gasification of C6H10O5 (cellulose), of which the
results are presented in Figure 1.7, are obtained with Gibbs free energy minimization model (Kyle,
1999) using the predictive Soave Redlich Kwong equation of state to calculate the required fugacity
22
coefficients (Bertucco et al., 1995; Soave, 1972; Soave et al., 1993). Such calculations have a limited
quantitative value in case the reactions involved are too slow to reach equilibrium, but they may be
useful in predicting trends and the results desired upon application of an appropriate catalyst. In Figure
1.7 yields are plotted that are defined as the moles produced of a certain component out of one mole of
C6H10O5. If water is consumed in the reaction the corresponding water yield is negative.
Complete gasification of wet feedstock is thermodynamically possible. Actually, at 600°C and 250 bar
dry matter concentrations of up to 50 wt % do not have thermodynamic or stoichiometric limitations
(see Figure 1.7a) regarding the full conversion the gas-phase. At a lower temperature of 350°C,
complete conversion is only possible for feedstocks with less than 30 wt % organics (see Figure 1.7a).
In the thermodynamic calculations the non gasified part feedstock remains as solid carbon.
Thermodynamics predict that high temperature gasification would produce a hydrogen-rich gas (at
least for dry matter contents of less than 10 wt %) while at low temperature a methane-rich gas is
produced (see Figures 1.7b, 1.7c and 1.7d). Complete reforming of the feed (Eq. 1) to hydrogen is
thermodynamically possible at temperatures above 600°C and concentrations of organics lower than 1
wt %. Below 400°C while using feedstocks with more than 15 wt % organics, complete conversion of
the feed to methane (Eq 2) is favored by the Second Law. Between 350°C and 600°C there is a large
difference in composition of the produced gas (compare Figures 1.7b and 1.7c), while increasing the
temperature from 600°C to 700°C results in only minor changes (compare Figures 1.7c and 1.7d). It is
also interesting to note that according to thermodynamics water is net consumed in the decomposition
reactions under all relevant conditions (see Figures 1.7b, 1.7c, 1.7d, 1.7e and 1.7f).
For low temperature gasification, the content of dry matter in the feed does not influence the product
distribution to a large extent; the yields are almost unaffected (see Figure 1.7b). On the contrary, at
higher temperature there is a continuous varying product distribution ranging from nearly pure
hydrogen for very low weight percentages of dry matter, to a mixture hydrogen methane for high
organic fractions in the feed (see Figures 1.7c and 1.7d). Below 200 bar, the pressure has a profound
influence on the product distribution (see Figure 1.7e). Once above 200 bar, the operating pressure
does not influence the product distribution to a large extent (see Figure 1.7e). It is worthwhile to note
that, from a thermodynamic point of view, high temperature gasification should be carried out at the
lowest possible pressure to achieve maximal hydrogen yields (see Figure 1.7e). Figure 1.7f shows
predictions for typical SCWG conditions, viz. 250 bar and a feed stream that contains 10 wt %
organics. For these conditions, according to thermodynamics, there is strong shift from methane
towards hydrogen and carbon monoxide upon increasing the temperature. Methane-rich gas can be
produced up to temperatures of approx. 500oC, higher temperatures favor the production of hydrogen.
23
100 20 40 60 80 00.0
0.2
0.4
0.6
0.8
1.0600 oC & 250 bar
350 oC & 250 bar
fract
ion
carb
on g
asifi
ed
wt.% organics in feed0 5 10 15 20 25 30
-6
-3
0
3
6
9
12 CO CO2
CH4
H2
H2O
350 oC & 250 bar
Yie
lds,
mol
e pr
oduc
ed/m
ole
C6H
10O
5
wt.% organics in feed (a) (b)
0 5 10 15 20 25 30
-6
-3
0
3
6
9
12 CO CO
2
CH4
H2
H2O
600 oC & 250 bar
Yi
elds
, mol
e pr
oduc
ed/m
ole
C6H
10O
5
wt.% organics in feed 0 5 10 15 20 25 30
-6
-3
0
3
6
9
12 CO CO2
CH4
H2
H2O
700 oC & 250 bar
Yi
elds
, mol
e pr
oduc
ed/m
ole
C6H
10O
5
wt.% organics in feed (c) (d)
0 100 200 300 400
-6
-3
0
3
6
9
12 CO CO
2
CH4
H2
H2O
Yie
lds,
mol
e pr
oduc
ed/m
ole
C6H
10O
5
Pressure, bar
10 wt% organics in feed, 600 oC
300 400 500 600 700 800
-6
-3
0
3
6
9
12
CO CO
2
CH4
H2
H2O
10 wt% organics in feed, 250 bar
Yi
elds
, mol
e pr
oduc
ed/m
ole
C6H
10O
5
T, oC (e) (f)
Figure 1.7 Results of equilibrium calculations of cellulose (C6H10O5) gasification in hot compressed
water.
24
Literature review
Here a short literature review is presented dealing with those reported results that are relevant for
process development. For a more complete and detailed review the reader is referred to a joint paper of
the SCWG community (Matsumura et al., 2005), which was co-authored by the author of this thesis.
The reported experimental results concerning the effects of the operating conditions and applied
catalysis are analyzed from a reactor engineering point of view. There is also significant information
available concerning the chemical pathways of model compounds near the critical point (see e.g.
Buhler et al., 2002; Kabyemela et al., 1998; Kruse and Gawlik, 2003; Savage, 1999). These results are
however less relevant for the severe conditions of SCWG. Listing and discussing this information is
therefore beyond the scope of this work.
Historical background
In the mid seventies of the last century, researchers at MIT (Amin, 1975; Woerner, 1976) discovered
that biomass could be liquefied without producing char by processing it in supercritical water. In 1985,
Modell (Modell, 1985) reported these experiments. It concerned experiments in which glucose,
cellulose, hexanoic acid, polyethylene and maple wood sawdust samples were quickly immersed in
sub- en supercritical water. Most experiments were performed with glucose or maple wood sawdust,
of which 10 and 2 g were injected respectively in ca. 300 g of water. Char was collected as reaction
product when operating below the critical point of water, while just above the critical point char was
not observed. Other interesting observations were that polyethylene produced char under sub- and
supercritical conditions and that using 30 g of glucose instead of 10 g resulted in significant char
formation. Also in 1985, Elliot and Sealock (Elliott and Sealock, 1985) presented work that showed
that the combined advantage of a high-pressure water environment and a metal catalyst (nickel) could
compensate in the slow reaction kinetics typical for gasification at lower temperature. In supercritical
water of 450°C, 80% of the carbon present in the feedstock (10 wt % wood flour in water) could be
converted to gases. In hindsight, the results presented in these papers should be regarded as the
pioneering work that set off the research on gasification of biomass in hot compressed water. It is
interesting to note that the motivation for this early research was not to develop a process for the
effective conversion of wet biomass streams, but to minimize char production and to optimize steam
reforming of biomass. At that time, the carbonaceous by-product of gasification and liquefaction
processes was considered a serious problem. It was reasoned that char decreased the yield of the fluid
products and could cause technical processing difficulties. Already in 1978 biomass was proposed as a
potential feedstock for hydrogen production via steam reforming (Antal, 1978) and studies were
reported of the reaction kinetics of cellulose pyrolysis in steam. It was found that biomass did not react
directly with steam at atmospheric pressure to produce the desired products. Significant amounts of tar
25
and char were formed, and the gas contained higher hydrocarbons in addition to the desired light gases
(Antal, 1985; Antal et al., 1987; Antal and Mok, 1988; Mozaffarian et al., 2004; Stein et al., 1983).
Since the mid eighties, roughly speaking, two approaches to gasification in hot compressed water have
been investigated in terms of reaction temperature ranges. In low-temperature catalytic gasification at
350 to 500oC, the feedstock is gasified with the help of a catalyst into a methane-rich gas. High
temperature supercritical gasification is carried out in the range of 500 to 750oC, with or without
catalysis, and aims at producing primarily hydrogen.
The carbon efficiency is often used in literature to characterize the gasification process; it is defined as
fraction of carbon in the feedstock that has been converted to permanent gases.
Low temperature gasification
Looking at the reported results, the catalytic low-temperature process (350 - 500°C) must be regarded
as very promising. It provides the opportunity to produce a pressurized methane-rich gas at relatively
low temperatures. Even for feedstocks with a dry matter concentration of up to 10 wt %, carbon
efficiencies of 90% and up were achieved in batch and continuous bench-scale facilities, while without
catalyst the carbon efficiency is limited to ca. 20% (Elliott et al., 1993; Elliott et al., 1994; Elliott et al.,
2004). However, near 100 percent carbon efficiency was never reached yet; typically the carbon
efficiency was in the range of 94 to 98% when using the optimal catalyst at 400°C. The comparison of
several feedstocks using Ru and Ni catalyst showed the highest reactivity with manure, lignocellulosic
feedstocks showed lower activity as a group. The observed effect of temperature was obvious (higher
carbon efficiency at higher temperature), but there was no dramatic effect noticeable passing the
supercritical point of water (374oC). Only a limited range of catalytic metals and supports can be used
in the process because of the oxidation and degradation problems in the hot-water environment
(Matsumura et al., 2005). New catalyst formulations for low-temperature gasification include
combinations of metals stable under the applied conditions, such as ruthenium or nickel bimetallics
and stable supports, such as certain titania, zirconia, or carbon. For example, the ruthenium on rutile
titania extrudate is particularly effective in this process (Matsumura et al., 2005).
High temperature gasification
All reported laboratory results on high temperature gasification (500 - 750 °C) were obtained in metal
reactors. Stainless steel, Inconel, Hastelloy and corroded Hastelloy were used as construction
materials. Both empty tubular reactors and stirred cells were used in the laboratory work that was
carried out so far. Results obtained by processing model compounds and various wet biomass species
in tubular reactors were reported mainly by Antal and co-workers (Antal et al., 1993; Antal et al.,
2000). Hao et al. (Hao et al., 2003) and Lee et al. (Lee et al., 2002) reported experiments in tubular
26
reactors with glucose solutions as feedstock while varying the operating conditions. The diameter of
the tubular reactors was in the range of 1.44 mm to 6.22 mm with a corresponding specific possible
catalytic wall area of 2778 m2 to 643 m2 per unit volume reactor. Sinag et al. (Sinag et al., 2004)
conducted SCWG experiments in a 190 mL metal autoclave equipped with a stirrer. In their
experiments a cold feed stream was continuously injected into the hot autoclave. Literature data
indicate that the reactor material has a significant effect on the gasification process. Yu et al. (Yu et
al., 1993), for instance, reported results of gasification of glucose in the same experimental set-up
under identical conditions, but using other metals for the reactor. They found large differences
between experiments carried out in Inconel, Hastelloy and corroded Hastelloy reactors. For a test
series using 0.6 M acetic acid they observed 14% carbon efficiency in the Inconel reactor whereas in
the corroded Hastelloy reactor the carbon efficiency was as high as 53%.
In a nutshell the results reported for non-catalytic (not taking into account the catalytic effect of the
reactor wall) SCWG in the range of 500 - 750°C can be summarized as follows:
● The metal reactor wall itself and deposits on it are catalytically active influencing both the
carbon efficiency and the product distribution. Due to this catalytic effect of the wall, which
may also be influenced by corrosion, published data are difficult to compare and interpret.
Besides, the catalytic activity of small laboratory equipment cannot be translated to large-scale
reactors, which have a much smaller wall area over volume ratio.
● Temperature is the most important process parameter; at 500°C without catalyst hardly any gas
is produced, while at 600°C using a low concentrated feedstock complete gasification can be
achieved. It is important to note here that results reported by Antal and co-workers before
report the temperature in the last section of the tubular reactor while it turned out later that in
the entrance (mixing) region peak temperatures were present that were more than 100°C higher.
● Complete conversion of the feedstock to the gas-phase is achieved only for very diluted
solutions of glucose and glycerol (< 5 wt %). Higher concentrations (> 5 wt %) of model
compounds and more complicated feedstocks (e.g. wood, lignin, starch) cannot be gasified
completely. The carbon conversion is typically between 75 and 90% for solutions with more
than 5 wt % organics in it.
● The overall reaction is reasonably fast reaching the maximum conversion at 600°C within one
minute.
● Processing complicated feedstocks such as wood and starches that contain more than 10 wt %
organics cause severe plugging problems in the reactor due to deposits of carbonaceous and
mineral material.
27
● The catalytic activity of the reactor wall suggested that heterogeneous catalysis might be
employed to increase the extent of gasification for concentrated feeds. Antal and co-workers
employed carbonaceous catalysts to investigate this possibility (Antal et al., 2000; Xu et al.,
1996; Xu and Antal, 1998) and concluded that the available surface area of the carbon does not
significantly influence the results. Of the tested carbons, coconut shell activated carbon was the
most active one. They obtained complete gasification of a 22 wt % glucose solution at 600°C
(peak temperature not given). Besides glucose also other model compounds (e.g. ethylene
glycol, phenol) and various realistic biomass feedstocks such as sewage sludge, wood sawdust,
and starches were tested. Although complete gasification was not always possible,
carbonaceous catalysts were found effective for all the compounds tested. To achieve
reasonable carbon efficiencies (> 85%) for feedstocks such as potato starch catalyst bed
temperatures of up to 750°C had to be applied.
High temperature gasification pilot plants
SCWG is in an early stage of development. Due to its promises with respect to possible conversion of
waste materials to a valuable gas, the laboratory research is developing rapidly. However, large-scale
commercial installations do not yet exist. The gap between small-scale testing in laboratories to
practical demonstration of a new process is bridged by experimentation with pilot plants. At present
there are two pilot plants being operated in the world. The largest plant, in operation since the
beginning of 2003, is the one of Forschungszentrum Karlsruhe (FzK) in Germany (Boukis et al.,
2002). It has a design capacity of 100 l/h, and was built to demonstrate supercritical gasification of wet
residues from wine production. EU subsidies plus a grant awarded by the Japanese NEDO enabled the
construction of a well-equipped process development unit (PDU) in Enschede, The Netherlands, with
a maximum throughput capacity of 30 l/h (Potic et al., 2002, Potic et al., 2004, Van de Beld et al.,
2003). BTG Biomass Technology Group B.V. has been responsible for the technical realization and
start-up of this small pilot plant.
28
Outline of this thesis
The further development of the SCWG process requires input from various disciplines. On the
feedstock side, pre-treatment and pumping problems have to be tackled. Due to the severe operating
conditions, studying of the long-term performance of selected construction materials is of utmost
importance. Detailed chemical studies that provide more insight in the reaction pathways and kinetics
at the severe SCWG conditions would help to select the optimal process conditions.
This thesis focuses on the reactor engineering aspects of SCWG. Four main research lines (goals) will
be addressed:
● Development of high-throughput experimental techniques.
● Mapping of the operating window under catalytically inert conditions.
● Investigating the potential of heterogeneous catalysis.
● Development of elements of a SCWG that can deal with heterogeneous catalysis under fouling
conditions.
Below the relevance of these research lines will be elucidated.
1. Development of high-throughput experimental techniques. Due to the severe operating
conditions of SCWG, conventional laboratory equipment (e.g. autoclaves with significant
volumes) requires extensive safety measures, and is costly and time intensive. In such
systems, typically, 1 to 2 experiments can be performed per day, while 10 to 20 would be
desirable to speedup the pace of the research. In this thesis, a high-throughput batch screening
technique using quartz capillaries of only 1 mm internal diameter (Chapter 2) has been
developed. With this method 20 cheap and safe tests per day have been performed. For
catalyst screening and to study possible ways to recover or to avoid carbonaceous deposits
when processing realistic feedstock materials, a continuous flow micro system (1 mm i.d.) is
designed and applied (Chapter 5). In the same system, fluidization with sub- and supercritical
water has been studied (Chapter 6).
2. Mapping of the operating window under catalytically inert conditions. It is know that metal
reactor walls are catalytically active and influence the gasification process. Due to this
catalytic effect of the wall published data are difficult to compare and interpret. Besides, the
catalytic activity of small laboratory equipment cannot be translated to large-scale reactors,
which have a much smaller wall area over volume ratio. In the present work, catalytically inert
quartz capillaries have been used for the investigation of gasification in hot compressed water
(Chapters 2 and 3). These results are supposed to represent the plain performance of the
process and can provide reliable information on the chemistry and kinetics of biomass and
waste gasification in high temperature and high-pressure water.
29
3. Investigating the potential of heterogeneous catalysis. Results of previous research already
made it clear that with catalysis only very diluted feed streams (say < 5 wt % organics) can be
gasified completely. When processing feedstock with higher concentrations, carbonaceous
deposits (not always e.g. not for glycerol) and water soluble organics are produced next to
gases. Carbonaceous deposits cause fouling and blocking problems within the process. The
organics containing effluent water requires an additional waste water treatment system, which
obviously affects the economics of the process negatively. In this thesis, a catalyst developed
for subcritical conversion has been tested under SCWG conditions in the microreactors
(Chapters 3 and 5). Both the gasification performance and the burn-off characteristics of this
catalyst have been investigated. The latter aspect is of importance for the development of the
SCWG process.
4. Development of elements of a SCWG process that can deal with heterogeneous catalysis
under fouling conditions. Like mentioned above, catalysis is needed when dealing with
concentrated feedstocks. Besides, when processing charring feedstock fouling will occur.
These two items define the severe boundary conditions for process development; viz. the
SCWG should be able to deal with catalysis under fouling conditions. In Chapter 7, based on
results gathered in Chapters 2 to 6, an integrated heat-exchanger/reactor system based on fluid
beds and circulating solids is proposed.
Notation
AHE Area per unit throughput, m2/(kg/s)
H Enthalpy, J/kg
∆Hvap Enthalpy of vaporization of water, MJ/kg
Kw Ion product
LHV Lower heating value of the dry fraction, MJ/kg
LHVnet Net lower heating value, MJ/kg
Pc Critical pressure, bar
Tc Critical temperature, oC
Thot,in Inlet temperature of the hot stream, oC
Tcold,in Inlet temperature of the cold stream, oC
U Overall heat transfer coefficient, W/(m2.K)
Xmoisture Mass fraction of moisture in the biomass
ηHE Heat-exchanger efficiency
30
References
Amin, S.I., MSc Thesis, MIT, 1975. Antal, M.J., Energy from Biomass and Waste. In Class, D.L.; Chicago, 1978. Antal, M.J., Mathematical Modelling of Biomass Pyrolysis Phenomena. Fuel 1985, 64 1483. Antal, M.J.; Brittain, A.; DeAlmeida, C.; Ramayya, S.; Roy, J.C., Heterolysis and Homolysis in Supercritical Water. In Supercritical Fluids: Chemical and Engineering Principles and Applications, Squires, T.G.; Paulaitis, M.E., Eds.; American Chemical Society, Washington, DC, 1987. Antal, M.J.; Mok, W.S., A Study of the Acid Catalyzed Dehydratation of Fructose in Near-Critical Water. In Research in Thermochemical Biomass Conversion, Bridgwater, A.V.; Kuester, L.J., Eds.; Elsevier, New York, 1988. Antal, M.J.; Manarungson, S.; Mok, W.S., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. In Advances in Thermochemical Biomass Conversion; Bridgwater, A.V., Ed.; Blackie Academic and Professional: London, 1993, 1367. Antal, M.J.; Allen, S.; Lichwa, J.; Schulman, D.; Xu, X., Hydrogen Production from High-Moisture Content Biomass in Supercritical Water. In U.S. DOE Hydrogen Program Review, 1999. Antal, M.J.; Allen, S.G.; Schulman, D.; Xu, X.; Divilio, R.J., Biomass Gasification in Supercritical Water. Ind. Eng. Chem. Res. 2000, 39, 4040. Beenackers, A.A.C.M.; Maniatis, K., Gasification Technologies for Heat and Power from Biomass. In Biomass Gasification and Pyrolysis: State of Art and Future Prospects; Kaltschmitt, M.; Bridgwater, A.V., Eds.;CLP Press, UK, 1997. Bertucco, A.; Barolo, M.; Soave, G., Estimation of Chemical Equilibria in High-Pressure Gaseous Systems by a Modified Redlich-Kwong-Soave Equation of State. Ind. Eng. Chem. Res. 1995, 34, 3159. Boerrigter, H.; Uil, H.D.; Calis, H.P., Green Diesel from Biomass by Fischer-Tropsch Synthesis: New Insights in Gas Cleaning and Process Design. ECN-RX--03-047; ECN: Petten, The Netherlands, 2003. Boukis, N.; Diem, V.; Dinjus E.; Galla U.; Kruse A., Advances with the Process of Biomass Gasification in Supercritical Water. Presented at the 4th International Symposium on High Pressure Technology and Chemical Engineering, Venice: Italy, 2002. BP Statistical Review of World Energy 2005, 2005. http://www.bp.com/genericsection.do?categoryId=92&contentId=7005893 Bridgwater, A.V.; Meuer, D.; Radlein, D., An Overview of Fast Pyrolysis of Biomass. Org. Geochem. 1999, 30, 1479. Buhler, W.; Dinjus, E.; Ederer, H.J.; Kruse, A.; Mas, C., Ionic Reactions and Pyrolysis of Glycerol as Competing Reaction Pathways in Near- and Supercritical Water. J. Supercrit. Fluids 2002, 22, 37.
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Bush, J.W., State of the Union Speech; 2006. http://www.whitehouse.gov/stateoftheunion/2006/index.html Elliott, D.C.; Sealock, L.J., Low Temperature Gasification of Biomass Under Pressure. In Fundamentals of Thermochemical Biomass Conversion, Overend, R.P.; Milne, T.A.; Mudge, L.K., Eds.; Elsevier Applied Science Publishers: London, 1985. Elliott, D.C.; Sealock, L.J.; Baker, E.G., Chemical Processing in High-Pressure Aqueous Environment: 2. Development of Catalyst for Gasification. Ind. Eng. Chem. Res. 1993, 32, 1542. Elliott, D.C.; Sealock, L.J., Jr.; Baker, E.G., Chemical Processing in High-Pressure Aqueous Environment. 3. Batch Reactor Process Development Experiments for Organics Destruction. Ind. Eng. Chem. Res. 1994, 33, 558. Elliott, D.C.; Neuenschwander, G.G.; Hart, T.R.; Butner, R.S., Low-Temperature Catalytic Gasification of Wet Biomass Residues. In Proceedings of the 2nd World Biomass Conference, Biomass for Energy, Industry, and Climate Protection; van Swaaij, W. P. M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004, p. 746. Goudriaan, F.; Peferoen, D.G.R., Liquid Fuels from Biomass via a Hydrothermal Process. Chem. Eng. Sci. 1990, 45, 2729. Groeneveld, M.J., Personal communication, 2000. Groeneveld, M.J., The co-current moving bed gasifier. Technical High School of Twente, Enschede, 1980. Hao, X.H.; Guo, L.J.; Mao, X.; Zhang, X.M.; Chen, X.J., Hydrogen Production from Glucose used as a Model Compound of Biomass Gasified in Supercritical Water. Hydrogen Energy 2003, 28, 55. Hemmes, K. Vergassing van natte biomassa/reststromen in superkritiek water (SG), voor de productie van groen gas (SNG), SNG/H2-mengsels, basis chemicaliën en puur H2. ECN: 2004. Higman, C.; van der Burgt, M., Gasification. Elsevier 2003. Hill, D.T.; Bolte, J.P., Methane Production from Low Solid Concentration Liquid Swine Waste Using Conventional Anaerobic Fermentation. Bioresource Technology 2000, 74, 241. IEA World Energy Outlook 1998; IEA/OECD; IEA: Paris, France, 1998. Kabyemela, B.M.; Adschiri, T.; Malaluan, R.; Arai, K., Degradation Kinetics of Dihydroxyacetone and Glyceraldehyde in Subcritical and Supercritical Water. Ind. Eng. Chem. Res. 1997, 36, 2025. Kabyemela, B.M.; Adschiri, T.; Malaluan, R.M.; Arai, K., Kinetics of Glucose Epimerization and Decomposition in Subcritical and Supercritical Water. Ind. Eng. Chem. Res. 1997, 36, 1552. Kabyemela, B.M.; Adschiri, T.; Malaluan, R.M.; Arai, K.; Ohzeki, H., Rapid and Selective Conversion of Glucose to Erythrose in Supercritical Water. Ind. Eng. Chem. Res. 1997, 36, 5063. Kabyemela, B.M.; Takigawa, M.; Adschiri, T.; Malaluan, R.M.; Arai, K., Mechanism and Kinetics of Cellobiose Decomposition in Sub- and Supercritical Water. Ind. Eng. Chem. Res. 1998, 37, 357.
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Kabyemela, B.M.; Adschiri, T.; Malaluan, R.M.; Arai, K. Glucose and Fructose Decomposition in Subcritical and Supercritical Water: Detailed Reaction Pathway, Mechanisms, and Kinetics. Ind. Eng. Chem. Res. 1999, 38, 2888. Kersten, S.R.A., Biomass Gasification in Circulating Fluidized Beds. University of Twente: Enschede, 2002. Kersten, S.R.A.; Prins, W.; van der Drift, B.; van Swaaij, W.P.M., Principles of a Novel Multistage Circulating Fluidized Bed Reactor for Biomass Gasification. Chem. Eng. Sci. 2003, 58, 725. Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Reactor Design Considerations for Biomass Gasification in Hot Compressed Water. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 777. Kruse, A.; Gawlik, A.; Henningsen, T., Biomass Liquefaction and Gasification in Near- and Supercritical Water: Key Compounds as a Tool to Understand Chemistry. Presented at the 4th International Symposium on High-Pressure Technology and Chemical Engineering, Venice: Italy, 2002. Kruse, A.; Gawlik, A., Biomass Conversion in Water at 330 - 410oC and 30 - 50 MPa. Identification of Key Compounds for Indicating Different Chemical Reaction Pathways. Ind. Eng. Chem. Res. 2003, 42, 267. Kruse, A.; Henningsen, T.; Sinag, A.; Pfeiffer, J., Biomass Gasification in Supercritical Water: Influence of the Dry Matter Content and the formation of Phenols. Ind. Eng. Chem. Res. 2003, 42, 3711. Kyle, B.G., Chemical and Process Thermodynamics. Prentice Hall PTR, New Jersey, 1999. Lee, I.; Kim, M.-S.; Ihm, S.-K., Gasification of Glucose in Supercritical Water. Ind. Eng. Chem. Res. 2002, 41, 1182. Matsumura, Y.; Minowa, T.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; van de Beld, L.; Elliott, D.C.; Neuenschwander, G.G.; Kruse, A.; Antal, M.J., Biomass Gasification in Near- and Supercritical Water: Status and Prospects. Biomass Bioenergy 2005, 29, 269. Minowa, T.; Zhen, F.; Ogi, T., Cellulose Decomposition in Hot-Compressed Water with Alkali or Nickel Catalyst. J. Supercrit. Fluids 1998, 13, 253. Modell, M., Gasification and Liquefaction of Forest Products in Supercritical Water. In Fundamentals of Thermochemical Biomass, Overend, R.P.; Milne, T.A.; Mudge, L.K., Elsevier Applied Science Publishers Ltd.: London, 1985, p. 95. Mok, W.S.L.; Antal, M.J., Uncatalyzed Solvolysis of Whole Biomass Hemicellulose by Hot Compressed Liquid Water. Ind. Eng. Chem. Res. 1992, 31, 1157. Mozaffarian, M.; Deurwaarder, E.P.; Kersten, S.R.A., Green Gas (SNG) Production by Supercritical Gasification of Biomass; ECN-C-03-066; ECN: Petten, The Netherlands, 2004. National Institute of Standards and Technology, (NIST), Gaithersburg, MD.
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Potic, B.; van de Beld, L.; Assink, D.; Prins, W.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water. In Proceedings of the 2nd European Conference and Exhibition on Biomass for Energy, Industry and Climate Protection, Palz, W.; Spitzer, J.; Maniatis, K.; Kwant, K.; Helm, P.; Grassi, A., Eds.; ETA Florence, WIP Munich: Amsterdam, 2002, p. 777. Potic, B.; Kersten, S.R.A.; Prins, W.; Assink, D.; van de Beld, L.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water: Results of Micro and Pilot Scale Experiments. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W. P. M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004, p. 742. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M. Gasification of Model Compounds and Wood in Hot Compressed Water. Submitted to Ind. Eng. Chem. Res. 2005. Quaak, P.; Knoef, H.; Stassen, H.M., Energy from Biomass: A Review of Combustion and Gasification Technologies. World Bank Publications, 1999. Reith, J.H.; Uil, H. de; Veen, H. van; Laat, W.T.A.M. de.; Niessen, J.J.; Jong, E. de.; Elbersen, H.W.; Weusthuis, R.; Dijken, J.P. van; Raamsdonk, L., Co-production of Bio-ethanol, Electricity and Heat from Biomass Residues. ECN-RX-02-030; Petten: The Netherlands, 2002. Saka, S.; Ueno, T., Chemical Conversion of Various Celluloses to Glucose and its Derivates in Supercritical Water. Cellulose 1999, 6, 177. Savage, P. E., Organic Chemical Reactions in Supercritical water. Chem. Rev. 1999, 99, 603. Schmieder, H.A.J.; Boukis, N.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht, M., Hydrothermal Gasification of Biomass and Organic Wastes. Proceedings of the 5th Conference on Supercritical Fluids and their Applications: Garda, Italy, 1999, p. 347. Sinag, A.; Kruse, A.; Schwarzkopf, V., Key Compounds of the Hydropyrolysis of Glucose in Supercritical Water in the Presence of K2CO3 . Ind. Eng. Chem. Res. 2003, 42, 3516. Sinag, A.; Kruse, A.; Rathert, J., Influence of the Heating Rate and the Type of Catalyst on the Formation of Key Intermediates and on the Generation of Gases During Hydropyrolysis of Glucose in Supercritical Water in a Batch Reactor. Ind. Eng. Chem. Res. 2004, 43, 502. Soave, G., Equilibrium Constants from a Modified Redlich-Kwong Equation of State. Chem. Eng. Sci. 1972, 27, 1197. Soave, G.; Barolo, M.; Bertucco, A., Estimation of High-Pressure Fugacity Coefficient of Pure Gaseous Fluids by Modified SRK Equation of State. Fluid Phase Equilib. 1993, 91, 87. Stein, Y.S.; Antal, M.J.; Jones, M., A Study of the Gas-Phase Pyrolysis of Glycerol. J. Anal. Applied Pyrolysis 1983, 4, 283-296. The National Renewable Energy Laboratory's, Energy Efficiency and Renewable Energy Technology Development in China; 2006. United Nations Environment Programme, Introduction to Climate Change; 2005.
34
Van de Beld, L.; Wagenaar, B.M.; Assink, D.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; Penninger, J.M.L., Biomass and Waste Conversion in Supercritical Water for the Production of Renewable Hydrogen. Presented at the 1st European Hydrogen Energy Conference, Grenoble: France, 2003. Van den Aarsen, F.G., Fluidised Bed Wood Gasifier Performance and Modelling. Technical High School of Twente, Enschede, The Netherlands, 1985. Van Swaaij, W.P.M.; Prins, W.; Kersten, S.R.A., Strategies for the Future of Biomass for Energy, Industry and Climate Protection. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004. Wang, X.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Experimental Validation of Model Results. Ind. Eng. Chem. Res. 2005, 8767. Woerner, G. A., MS Thesis, MIT, 1976. World Energy Council, Survey of Energy Resources 2004. Elsevier, 2004. Xu, X.; Matsumura, Y.; Stenberg, J.; Antal, M.J., Carbon-Catalyzed Gasification of Organic Feedstocks in Supercritical Water. Ind. Eng. Chem. Res. 1996, 35, 2522. Xu, X.; Antal, M.J., Jr., Gasification of Sewage Sludge and Other Biomass for Hydrogen Production in Supercritical Water. Enviromental Progress 1998, 17, 215. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels 1993, 7, 574.
35
36
37
u
CCChhhaaapppttteeerrr 222 AAA HHHiiiggghhh---TTThhhrrrooouuuggghhhpppuuttt SSScccrrreeeeeennniiinnnggg TTTeeeccchhhnnniiiqqquuueee fffooorrr CCCooonnnvvveeerrrsssiiiooonnn iiinnn HHHooottt CCCooommmppprrreeesssssseeeddd WWWaaattteeerrr
Abstract
Conversion in hot compressed water (e.g. 600oC and 300 bar) is considered to be a promising
technique to treat very wet biomass or waste streams. In this chapter, a new experimental method is
described that can be used to screen the operating window in a safe, cheap, and quick manner (one
measurement takes about 5 min). Small sealed quartz capillaries (i.d. = 1 mm) filled with biomass or
model compounds in water are heated rapidly in a fluidized bed to the desired reaction temperature.
The reaction pressure can be controlled accurately by the initial amount of solution in the capillary.
After a certain contact time, the capillaries are lifted out of the fluidized bed, rapidly quenched, and
destroyed to collect the produced gases for GC analysis.
Results of measurements for formic acid and glucose solutions have shown that the technique is
reliable enough for screening purposes including trend detection. For conversions above 30%, three
identical measurements are sufficient to produce reasonably accurate average values with a
confidence level of 95%.
This chapter has been published in Ind. Eng. Chem. Res, 43 (2004), p. 4580
38
Introduction
Process intensification is a novel design approach in which reduction of equipment size leads to less
energy consumption, improved safety, lower capital costs, and less pollution (Burns and Ramshaw,
1999; Dummann et al., 2003; Rebrov et al., 2003). In this chapter, the intensification principle has
been applied to a screening technique for the conversion of biomass in hot compressed water, in
particular gasification of biomass/waste in supercritical water (SCWG). Very wet biomass (moisture
content > 70 wt %) cannot be converted economically by traditional techniques such as combustion
and gasification due to the energy required for water evaporation (2.4 MJ/kg). In SCWG, water
evaporation is avoided and intensive countercurrent heat exchange is practiced. Therefore, SCWG is
considered as a promising technique to convert wet streams into a hydrogen-rich gas (Antal et al.,
1993; Holgate et al., 1995; Modell, 1985). Biomass is converted in the presence of water, e.g. via:
C6H12O6 + 6H2O 6CO2 + 12H2
The above stoichiometric equation is highly idealized; in practice, also CO, CH4, C2-3-components,
liquids (including H2O), and polymers are formed next to CO2 and H2 (Schmieder et al., 2000). The
product distribution appears to be a strong function of the reactor temperature and the weight
percentage of organic material in the feedstock (Kruse et al., 1999). Antal and co-workers (Yu et al.,
1993) found that the wall material (Hastelloy, Inconel) of their bench-scale continuous flow reactor
had a large effect on the obtained results. This finding points to an important role of catalysis in
SCWG. Consequently, data obtained in small-scale metal reactors are obscured by undefined catalytic
effects, and therefore difficult to interpret and extrapolate. The novel method uses quartz capillaries of
1 mm i.d., 2 mm o.d., and 150 mm length as batch reactors, which have no catalytic activity.
Unfortunately, the number of experimental SCWG data published is small, while at the same time the
range of conditions explored is narrow. This is partly due to the severe operating conditions (300 bar
and 600oC), which make laboratory testing at bench-scale (1 - 100 g/h, 10 mL < Vreactor < 1 L)
problematic, expensive, and time-consuming. The proposed technique will allow mapping of the
complete operating area and trend detection of the process at moderate costs in just a short period of
time. Because of the small diameter and the properties of quartz, these microreactors can withstand
conditions up to 900oC and 600 bar. Apart from SCWG of biomass, this technique can be used also to
study other thermochemical conversion processes of wet feedstock, such as hydrothermal liquefaction,
oxidation in supercritical water (SCWO), and possibly many other high pressure processes. Hereafter,
the novel technique will be explained and validated in detail.
Experimental
Methodology
The method uses quartz capillaries of 1 mm i.d., 2 mm o.d., and 150 mm length as batch reactors (see
Figure 2.1). An experiment is started by charging a capillary with a known amount of aqueous solution
(biomass or a model compound dissolved in water). Hereafter the capillary is sealed and rapidly
heated in a high-temperature fluidized bed to the desired reaction temperature. In the heating
trajectory, the pressure increases as a result of the evaporating water. By varying the amount of
solution in the capillary, the final pressure is known and can be adjusted quite accurately, as the water
vapor pressure predominantly determines the total pressure. If a more precise estimate of the pressure
is required, thermodynamic calculations can be applied on the basis of the initial amount of water and
the measured amount of product gases. After a certain contact time at a specific reaction temperature
and pressure, the capillary is quenched to ambient conditions, which ensures that all reactions are
stopped. Finally, the gas phase inside the capillary is analyzed with respect to the absolute amount and
the composition.
Figure 2.1 Photograph of the capillaries together with 1 euro coin.
Advantages and Restrictions
This new experimental method has several advantages:
● conducting an experiment is fast, cheap, and safe (the reactor volume is 0.12 mL),
● the quartz capillaries are strong enough to withstand extremely high pressures (600 bar) and
high temperatures (900oC),
● quartz has no or hardly any catalytic activity,
● quartz is resistant to corrosion,
● catalysts can be easily added , and only very small amounts are required,
● the reactor content can be inspected visually (char, water-soluble organic compounds, and tars
can be identified in case of SCWG of biomass) and,
● the quartz reactors can be heated and cooled rapidly, leading to more precise results.
39
40
A disadvantage of the technique is that analysis of the liquid phase is extremely difficult, due to very
small sample volume, and has not been practiced up to now. Hence, when products are present in the
liquid phase, complete mass balance closure cannot yet be obtained. Depending on the
development/availability of liquid analysis methods with respect to the minimum required sample
volume, this disadvantage can be overcome. Small (diameter) reactors are required in order to hold the
large pressure built up and to ensure fast heat transfer. Fast heat transfer will guarantee accurate
knowledge concerning the reaction temperature, which is necessary for correct data interpretation. It is
common knowledge that in fluidized beds high heat transfer coefficients can be achieved, especially
when the bed particles and the immersed objects are small (Parmar and Hayhurst, 2002; Prins et al.,
1986; Prins et al., 1989). This has been checked by submerging a silver cylinder with the same
dimensions as the used capillaries in a fluidized bed (dp = 250 µm, u/umf = 2). The external heat
transfer coefficient derived from the heating curve appeared to be ca. 1000 W/(m2.K), which is in good
agreement with predictions from literature correlations (Parmar and Hayhurst, 2002, Prins et al.,
1989). Inside the quartz wall of the capillaries, no important temperature gradients will be present
because of the small dimensions (Dwall = 0.5 x 10-3 m) and the relatively high thermal conductivity (λ
= 1.8 W/(m.K)). The heating time of the used capillaries in the FB was determined experimentally to
be approximately 5 s at a reactor temperature of 600oC. For this measurement, a thermocouple was
welded into a water-filled capillary. The measured heating time (5 s) is in good agreement with the
calculated one (4.5 s), which is based upon the measured external heat transfer coefficient, the known
properties of the quartz wall, and minimum heat transfer conditions from the wall to the solution
inside.
The pressure-temperature and density trajectories of the feed material during heating to the reaction
conditions differ from those of continuous reactors. Continuous reactors are operated essentially
isobaric, whereas the capillaries, because of the fixed volume and initial mass, are isochoric reactors.
This could, in principle, be a cause for deviations in results between continuous reactors and the batch
capillaries. However, from a few special SCWG tests in which the heating trajectory of the capillaries
was varied, it appeared that the product composition is determined almost entirely by the final process
pressure and temperature.
Experimental Setup
As mentioned before, the capillaries were heated in a fluidized bed (sand particles of dp = 250 µm,
u/umf = 2), which was placed in a temperature-controlled oven. For each measurement two capillaries
were placed on a holder especially designed to allow quick and safe manual operation. The
temperature inside the FB was measured with thermocouples. After a certain contact time at the
desired reaction temperature and pressure, the capillaries were quenched in a water bath to ambient
41
conditions, to ensure that no further reactions proceeded. The produced gases were analyzed by the
following method:
● a quenched capillary was put into a stainless steel chamber of 50 mL,
● then this chamber was flushed with helium to remove traces of product gas from a previous
run,
● hereafter, the chamber was closed (with valves) and the capillary was crushed completely with
a hammer mechanism,
● finally, a gas sample was taken from the chamber and led to a gas chromatograph.
It is essential to crush the capillaries completely. Otherwise gas gaps, containing products, could be
captured between liquid phases, and excluded from the gas analysis.
Gaseous products were analyzed with a gas chromatograph using TCD cells (Varian Micro GC CP-
2003). Two columns were used: a Molsieve 5A, operated at 90ºC and 155 mbar, and a Porapak Q
operated at 75ºC and 155 mbar. Helium was used as a carrier gas. The chromatograms were obtained
and interpreted with Varian CP-Maître Elite software. Figure 2.2 is a schematic representation of the
setup.
N2
He
out
in
GC
out
in
Analyticalsection
2
3
4
5
5
71
18
7
9
TI
TI
PI
PI
6
TC
10
Figure 2.2 Schematic representation of the capillary reactor technique: (1, capillary; 2, nitrogen
supply for the fluidized bed; 3, oven; 4, fluidized bed; 5, thermocouples; 6, capillary holder; 7,
sampling chamber; 8, pressure indicator; 9, gas chromatograph; 10, temperature controller for the
oven).
Data Interpretation and Possible Errors
The degree of conversion of carbon in biomass to permanent gases (carbon efficiency, Xc) is chosen
here as the main process parameter for SCWG, because it indicates the distribution of carbon over the
desired products (permanent gases) and the undesired product (liquids and polymers). In practice, the
product composition can be modified catalytically, viz., by water-gas shift to produce CO2 and H2, or
by methanation to produce CH4.
Xc is defined by:
( )( ) ( )2 4 2 4
,
,
2 ...c permanent gas chamberc CO CO
solution cc solution
PVN RTXm f nN CH C H
M
ω ω ω ω−= = + + + +
42
Like every other experimental technique, also the newly developed capillary reactor method contains
certain error margins in the experimental parameters (see Table 2.1).
Table 2.1 Possible maximum experimental errors estimated for an experiment with a 17 wt % glucose
solution in water.
Quantity Value (x) Error (∆x) Units ∆x/x (%)
( )chamberRT
PV
3 x 10-3 3 x 10-5 Mol 1
( )solution cm f nM
9 x 10-5 3.6 x 10-6 Mol 4
∑ω Varies Varies - 2
TFB 1
Maximal possible error 8
As the main output of the method is the carbon efficiency, the error analysis will be focused on this
parameter. The number of moles present in the sampling chamber is known accurately by precise
measurement of T (within 1oC), P (within 500 Pa), and V (within 0.1 mL). The reaction solution was
prepared in batches of 50 g using high-precision balances (accuracy ± 0.001 g). For the mass of
solution actually put into the capillary, the error is the highest (ca 4%). This error is introduced mainly
because the capillaries are sealed rapidly and carefully in a hydrogen flame, but, nevertheless, some
water will be evaporated. Furthermore, it has been found that small variations in the fluid bed
temperature may cause differences in carbon efficiency ( ± 1% every 3oC for a 17 wt % glucose
solution at Xc < 50%; see Figure 2.4). In the low conversion regime this error is significant, whereas
near the maximum conversion, obviously, the effect of the exact process conditions disappears.
Overall, the maximal possible experimental error in the carbon efficiency was estimated to be
approximately 8% in a typical case (see Table 2.1), which is accurate enough for trend detection.
43
Results and Discussion
Visualization
As mentioned before, some products, such as tars and solid carbon (char), for SCWG, are clearly
visible after the reaction has been carried out in the quartz capillaries. In practice, char formation in the
heat-exchanger (the feed-site exit temperature is typically 450oC) will be one of the major bottlenecks
of the SCWG process (Potic et al., 2002). Figure 2.3 shows capillaries with high and low levels of
char/tar formation, indicating that possible problems can indeed be identified visually.
(a)
(b)
Figure 2.3 Capillaries with a different level of char/tar formation after SCWG (a) 17 wt % glucose in
water at 600oC, (b) 10 wt % glucose in water at 720oC.
Effects of the Heating Trajectory
To understand the effect of the heating trajectory of the capillaries on the carbon efficiency, special
tests have been performed. The following results were compared:
● capillaries with a short residence time in the FB (columns 1-3 of Table 2.2),
● capillaries which were quenched after a short residence time (the same as in columns 1-3 of
Table 2.2) and then reheated up to a 1 min total residence time in the FB (column 4 of Table
2.2),
● capillaries with a 60 s residence time without interruption (column 5 of Table 2.2).
44
Table 2.2 Carbon efficiency, Xc (%), measured for a 10 wt % glucose solution in water. The final
conditions are 600oC and 300 bar. Illustration of the effect of heating interruption after 1, 2, or 3 s.
Heating time and sequence
1 s 2 s 3 s
60 s residence time
after the interruption
60 s residence time
without any interruption
Xc 0.08 0.69 0.69
Xc 0.08 0.76 0.72
Xc 0.1 0.69 0.73
Xc 0.15 0.71 0.73
From Table 2.2 it can be concluded that the heating trajectory has no notable effect on the final carbon
efficiency. Continuation of interrupted experiments results approximately in the same final carbon-to-
gas conversion degrees as those of noninterrupted experiments. The differences between the figures in
the last two columns of Table 2.2 are all well below the experimental error. It should be realized that
the results of Table 2.2 have been derived for isochoric conditions. Whether or not the corresponding
conclusion is also valid for continuous tubular reactors should be proven in subsequent experimental
work.
Validation of the Technique
Mass balance closure was investigated by experiments with components which, according to the
literature (Yu and Savage, 1998), should be converted completely to gas-phase components at
temperatures above 700oC and low feed concentrations. Additional measurements were carried out
under conditions that should give a much lower carbon conversion. They were used to evaluate the
reproducibility of the method for SCWG in case a significant amount of carbon remains in the liquid
or solid phase. Formic acid and glucose were selected as model compounds. On basis of this analysis,
the number of identical experiments could be defined whose average should represent a good estimate
of the mean value. As mentioned before, the carbon efficiency was selected as an identifier for the
process. Table 2.3 shows series of measurements for which the average carbon efficiency and related
experimental error were calculated.
45
Table 2.3 Results of SCWG measurement series illustrating the reproducibility and accuracy of the
experimental technique (P > Pc).
Nr required for 95% CI Model
Compd
Concn.
wt % T(oC) Nr <Xc> σ
σ/<Xc>
(%) 5%
error
10%
error
Formic Acid 1.2 700 4 1.005 0.041 4.1 3 1
Glucose 0.8 700 3 1.003 0.041 4.1 3 1
Glucose 18 550 4 0.535 0.043 8.1 10 3
Glucose 17 460 8 0.203 0.033 16.2 41 10
The first two series (100% carbon efficiency, see Table 2.3) show that all produced gases can be
recovered with the novel method despite the very small volumes used. For all the series, the
experimental error found is acceptable and not in disagreement with the estimated one, except for the
low conversion range (< 25%). The reaction system is here possibly very sensitive to auto-catalytic
effects.
On the basis of the estimations of Table 2.1, in combination with statistical analysis (Devore, 2000) of
the first three series of Table 2.3, it was decided that three identical experiments is sufficient to
guarantee, for a confidence level of 95%, a mean value with an error width of 0.1 < Xc > (see Table
2.3, last column).
Clearly, the experimental technique described in this chapter can be used very well to estimate
reaction kinetics, determine trends, and explore the operating window for SCWG. An example of a
measured trend is presented in Figure 2.4.
46
400 500 600 700 8000
20
40
60
80
100
X c [%
]
T [oC]
Figure 2.4 Measured carbon efficiency of a 17 wt % glucose solution versus the reactor temperature.
τ = 60 s. The plotted line represents a trend line.
It has been found that, for a 17 wt % glucose solution in water above Pc with a contact time of 1 min,
the carbon efficiency increases asymptotically from 20% to 80% within a temperature range of 460 -
700oC.
47
48
Conclusion
A novel high-throughput screening technique has been developed which allows quick and safe
experimentation with thermochemical conversion reactions of wet feedstock at severe operating
conditions in a noncatalytic environment. This technique gives sufficient accuracy for all the
measuring conditions and is certainly satisfactory for the overall mapping of the experimental area and
trend detection. Advantages of the technique are the possibility of applying high temperatures and
pressures, the absence of undesired catalytic effects of the reactor wall, the absence of corrosion, rapid
heating and cooling, and easy visual inspection. Future developments may open the possibility of
liquid analysis, which has not been practiced up to now. The reactor system is isochoric, and the
heating trajectories of the reactants differ from those of a continuous isobaric system, but this does not
seem to affect the results in an important way. For this work, formic acid and glucose were used as
model compounds. Future experiments will include real biomass samples such as wood and
agricultural waste.
Notation
dp Particle diameter in the fluidized bed oven, µm
F Fraction of biomass in the solution, g/g
ID Inside diameter, mm
M Molar mass, g/mol
msolution Mass of the solution, g
nc Number of carbon atoms in one molecule of biomass
Nc,permanent gas Number of carbon moles in the permanent gas
Nr Number of samples
Nsolution Number of carbon moles in the solution
o.d. Outside diameter, mm
P Pressure, Pa
Pc Critical pressure, Pa
R Gas constant, J mol-1 K-1
T Temperature, K
TFB Temperature of the fluidized bed, K
U Velocity of the air in the fluidized bed oven, m/s
umf Minimum fluidization velocity, m/s
V Volume, mL
Vreactor Volume of the reactor, mL
49
Xc Carbon efficiency, %
Greek Symbols
σ Standard deviation
τ Reaction time, s
ω Mole fraction
Subcripts
c Carbon
Operators
< > Mean
References
Antal, M.J.; Manarungson, S.; Mok, W.S., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. In Advances in Thermochemical Biomass Conversion; Bridgwater, A.V., Ed.; Blackie Academic and Professional: London, 1993, 1367. Burns, J.R.; Ramshaw, C., Development of a Microreactor for Chemical Production. Trans. Inst. Chem. Eng. 1999, 77, 206. Devore, J.L., Probability and Statistics for Engineering and the Sciences; Duxbury: Pacific Grove, CA, 2000. Dummann, G.; Quittmann, U.; Groschel, L.; Agar, D.W.; Worz, O.; Morgenschweis, K., The Capillary-Microreactor: a New Reactor Concept for the Intensification of Heat and Mass Transfer in Liquid-Liquid Reactions. Catal. Today, 2003, 79-80, 433. Holgate, H.R.; Meyer, J.C.; Tester, W.J., Glucose Hydrolysis and Oxidation in Supercritical Water, AIChE J. 1995, 41, 637. Kruse, A.; Abeln, J.; Dinjus, E.; Kluth, M.; Petrich, G.; Schacht, M.; Sadri, H.; Schmieder, H., Gasification of Biomass and Model Compounds in Hot Compressed Water. Presented at the International Meeting of the GVC-Fachausschuβ "Hochdruckverfahrenstechnik", Karlsruhe, Germany, 1999; paper no. 107. Modell, M., Gasification and Liquefaction of Forest Products in Supercritical Water. In Fundamentals of Thermochemical Biomass, Overend, R.P.; Milne, T.A.; Mudge, L.K., Eds.; Elsevier Applied Science Publishers Ltd.: London, 1985, p. 95. Parmar, M.S.; Hayhurst, A.N., The Heat Transfer Coefficient for a Freely Moving Sphere in a Bubbling Fluidised Bed. Chem. Eng. Sci. 2002, 57, 3485.
50
Potic, B.; van de Beld, L.; Assink, D.; Prins, W.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water. In Proceedings of the 12th European Conference and Exhibition on Biomass for Energy, Industry and Climate Protection; Palz, W.; Spitzer, J.; Maniatis, K.; Kwant, K.; Helm, P.; Grassi, A., Eds.; ETA Florence, WIP Munich: Amsterdam, 2002; p. 777. Prins, W.; Draijer, W.; van Swaaij, W.P.M., Heat Transfer to Immersed Spheres Fixed or Freely Moving in a Gas-Fluidized Bed. In Heat and Mass Transfer in Fixed and Fluidized Beds. Van Swaaij, W.P.M.; Afgan, N.H., Eds.; Hemisphere Publishing Corp.: Washington, DC, 1986. Prins, W.; Harmsen, G.J.; De Jong, P.; van Swaaij, W.P.M., Heat Transfer from an Immersed Fixed Silver Sphere to a Gas Fluidized Bed of Very Small Particles. In Fluidization VI; Grace, J.R.; Shemilt, L.W.; Bergougnou, M.A., Eds.; Engineering Foundation: New York, 1989. Rebrov, E.V.; Duinkerke, S.A.; de Croon, H.M.J.M.; Schouten, J.C., Optimization of Heat Transfer Characteristic, Flow Distribution and Reaction Processing for a Microstructured Reactor/Heat-exchanger for Optimal Performance in Platinum Catalyzed Ammonia Oxidation. Chem. Eng. J. 2003, 93, 201. Schmieder, H.; Abeln, J.; Boukis, N.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht, E., Hydrothermal Gasification of Biomass and Organic Wastes. J. of Supercrit. Fluids 2000, 17, 145. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels 1993, 7, 574. Yu, J.; Savage, P.E., Decomposition of Formic Acid under Hydrothermal Conditions. Ind. Eng. Chem. Res. 1998, 37, 2.
51
52
53
C
CCChhhaaapppttteeerrr 333 GGGaaasssiiifffiiicccaaatttiiiooonnn ooofff MMMooodddeeelll CCooommmpppooouuunnndddsss aaannnddd
WWWooooooddd iiinnn HHHooottt CCCooommmppprrreeesssssseeeddd WWWaaattteeerrr
Abstract
In this chapter, an experimental investigation is presented that concerns the gasification of glycerol,
glucose, and pinewood in supercritical water. The batch experiments were performed in quartz
capillary reactors with an internal diameter of only 1 mm. Because these quartz reactors are
catalytically inert, the process could be studied in absence of the interfering catalytic influence of a
metal reactor wall, as used in all previous studies. The reaction space has been mapped by performing
over 700 experiments in which the temperature, pressure, reaction time, and concentration of the
feedstock were varied. The most important observations were that the pressure turned out to have no
effect on the conversion and product yields, and that, noncatalytically, complete conversion to the gas
phase is only possible for very diluted feedstock solutions (< 2 wt %). By adding ruthenium on a TiO2
carrier (Ru/TiO2) to the capillaries the potential of heterogeneous catalysis has been demonstrated.
When adding this catalyst to the capillaries, glucose solutions in the range of 1 - 17 wt % could be
gasified completely.
54
Introduction
Application of biomass and waste, as a renewable energy source, has gained an important role in the
world’s future energy policy. The Dutch government, for instance, has set a target for 2020, which
states that 5% of the total primary energy production must be based on biomass and waste input
(MEZ, 1995, 1997). Similar directives have been put forward by the European Committee in their
White Paper (European Commission, 2000).
Biomass and waste streams can be classified by their moisture content. Wet feedstock is defined here
as containing > 70 wt % moisture. Examples of wet waste streams include the following: vegetable,
fruit and garden waste; waste streams from agricultural, food and beverage industries; manure; sewage
sludge; and some household wastes. Seaweed and micro-algae are examples of cultivated wet biomass
crops. According to estimates of the Energy Research Centre of The Netherlands (ECN), (Hemmes,
2004), 5.3 - 12 million tons (dry matter) of wet streams are annually available in The Netherlands.
This corresponds to 70 - 160 PJ of energy (assuming 65% efficiency), which represents 2.3 - 5.3% of
the current Dutch annual energy consumption.
This chapter examines the conversion of wet streams into H2 and/or CH4 by means of gasification in
hot compressed water. Wet biomass cannot be converted economically by traditional techniques such
as pyrolysis, gasification, and combustion, because of the large amount of energy required for the
evaporation of water. Partial conversion by anaerobic digestion is possible and already practiced for
suitable feedstock materials (Sarada and Joseph, 1996). Gasification in supercritical water (SCWG) is
a novel technology for the conversion of wet biomass and waste streams to hydrogen or methane-rich
gas (Antal et al., 2000; Hao et al., 2003; Holgate et al., 1995; Lee et al., 2002; Manarungson et al.,
1993; Matsumura, 2002; Matsumura et al., 2005; Modell, 1985; Yu et al., 1993). Regarding the useful
products of SCWG, hydrogen, and/or methane, it is interesting that the selectivity can be controlled by
the process conditions and catalysis. Biomass-derived hydrogen could be applied in the future as a
renewable feedstock for fuel cells. Methane from biomass may be attractive as a renewable substitute
of natural gas (SNG). The product gases could be available at high pressure, which is required for
storage and transportation and in many end applications. The CO2 byproduct is almost pure and
suitable for sequestration or fertilization in greenhouses.
For proper process development of SCWG, data concerning the influence of the process conditions
(temperature (T), pressure (P), reaction time (t), concentration (c)) on the gas yields and byproduct
formation (char, liquid intermediates) are needed. Laboratory-scale results of SCWG experiments
have been reported by several research groups (Antal et al., 1993; Hao et al., 2003; Holgate et al.,
1995; Kabyemela et al., 1999; Kruse et al., 1999; Kruse et al., 2000; Kruse et al., 2002; Kruse and
Gawlik, 2003; Kruse et al., 2003; Modell, 1985; Potic et al., 2002; Schmieder, 1999; Sinag et al.,
2004; Xu et al., 1996; Xu and Antal, 1997; Yu et al., 1993). However, previously reported data are
55
obtained in small metal (e.g., Inconel, Hastelloy, stainless steel) reactors, which, to a variable extent,
exhibit catalytic activity (Diem et al., 2003; Sinag et al., 2004; Yu et al., 1993). Because of this ill-
defined catalytic activity, the reported data on SCWG are difficult to compare and interpret.
In this work, it is shown that Inconel indeed does have catalytic influence on SCWG reactions. Results
of noncatalytic SCWG experiments are presented. These results have been obtained by
experimentation with inert quartz microreactors (Potic et al., 2004) and are supposed to represent the
plain (noncatalytic) performance of the process. Results of model compounds and pinewood are
presented. The potential of heterogeneous catalysis for SCWG is discussed on the basis of experiments
in which Ru/TiO2 catalyst was added to the quartz reactors.
56
Experimental
The investigation presented here is based on more than 700 experiments that have been conducted in
the batch capillary reactors. In a previous paper (Potic et al., 2004), this measurement technique is
described in detail, including an error analysis. The technique uses quartz capillaries with an inner
diameter (ID) of 1 mm, an outer diameter (OD) of 2 mm, and a length of 150 mm as batch reactors.
An experiment starts by charging a capillary with a known amount of aqueous solution (biomass or a
model compound suspended or dissolved in water). Hereafter, the capillary is sealed and rapidly
heated in a high-temperature fluidized bed to the desired reaction temperature. The heating time of the
capillary in the fluid bed is 5 s. In the heating trajectory, the pressure increases as a result of the
evaporating water. By varying the amount of solution in the capillary, the final pressure is known and
can be adjusted quite accurately, because the water vapor pressure predominantly determines the total
pressure. The reaction time is defined as the time that the capillary has been immersed in the fluid bed,
minus the heating time (5 s). In a standard experiment, a reaction time of 60 s has been applied,
because this reaction time turned out to be sufficient to reach the maximum conversion at the lowest
typical SCWG temperature under noncatalytic conditions (600oC; see the section, “Effects of Reaction
Time”). After a certain fixed time at a specific reaction temperature and pressure, the capillary is
quenched to ambient conditions, which ensures that all reactions are stopped. Gas products are
released in a metal sampling chamber (50 mL) by smashing the capillary with a hammer mechanism.
Finally, the gas phase inside the capillary is analyzed to determine the absolute amount and the
composition. On the basis of the readings of a gas chromatograph, the total number of moles present in
the sampling chamber, and the amount of feedstock, the mass balance (yields and conversion) is
calculated.
The aim of the present investigation is mapping of the operating window of noncatalytic SCWG. The
following parameters were varied: temperature (400 - 800oC), pressure (50 - 450 bar) and the
concentration of organics in the feedstock (1 - 20 wt %). Different model compounds were used as
feedstock (viz., glucose, glycerol, and cellulose). Besides these model compounds, pinewood also was
used.
Hereafter, every experimental data point presented in the figures shown is an average value that is
based on at least six experimental data points (average standard deviation on the carbon conversion is
6%). The investigation is divided into two areas: noncatalytic gasification and catalytic gasification.
For the catalytic investigation, ruthenium on TiO2 (Ru/TiO2) has been applied. Moreover, Inconel 625
powder has been used as an additive for identification and quantification of the catalytic influence of
the metal wall of the reactors that have been reported in open literature.
The carbon conversion, the product yields, and the cold-gas efficiency are chosen here as the main
output parameters of the process.
The carbon conversion is defined as degree of conversion of carbon from biomass to permanent gases:
100,
,
xN
NX
feedc
ici
c
∑=
The product yield is defined as:
feed
ii N
NY =
In case of wood, the elemental composition has been normalized to C6HyOz (glucose C-basis).
The cold-gas efficiency (LHV) is defined as:
100xLHV
LHVY
feed
iii
LHV
∑=η
As described previously, the capillary is charged in air, resulting in a certain amount of oxygen (air)
being inevitably present in the capillary after the sealing. The amount of hydrogen lost due to
oxidation by slipped-in oxygen can be estimated from the experimental results of low-concentration
glucose solutions (1 wt %) processed at high temperature (800°C). Typical product gas yields per mol
of glucose of SCWG at 800oC were as follows: 5.5 for CO2, 0.5 for CH4, 0 for CO, 0 for C2- and C3-
components, and 6 for H2 (see Figure 3.4 which will be discussed later in this work). From a carbon
balance under these conditions, it follows that carbon gasification is complete for this experiment:
5.5C (CO2) + 0.5C (CH4) = 6C (in glucose).
If the reaction would proceed via the pathway
C6H12O6 + 5H2O → νΗ2 + 5.5CO2 + 0.5CH4
the value of ν would be 10. However, for these low-concentration solutions, ν has been measured to
be ± 6. This would mean that 4 mol of H2 per mol of glucose were oxidized by slipped-in O2. In
absolute numbers, this comes down to ca. 1 x 10-6 mol of H2. The maximum amount of O2 from air
present in the capillaries is 9 x 10-7 mol. If we may assume that the O2 inclusion is the same in absolute
value, the effect of slipped-in oxygen can be neglected for all other conditions investigated in this
work.
57
58
Results and discussion
Catalytic Effects of Metal Reactors
As mentioned earlier, all previously reported results on gasification in hot compressed water were
obtained in metal reactors. Stainless steel, Inconel, Hastelloy, and corroded Hastelloy were used as
construction materials. Both empty tubular reactors and stirred cells were used in the laboratory work
that has been performed thus far. Results obtained in tubular reactors were reported mainly by Antal
and co-workers (Antal et al., 2000; Xu et al., 1996; Yu et al., 1993). The diameter of the tubular
reactors was in the range of 1.44 mm to 6.22 mm with a corresponding specific possible catalytic wall
area of 2778 m2 per unit volume reactor to 643 m2 per unit volume reactor. Reynolds numbers were
mostly in the range for laminar flow. Nevertheless, the flow in these reactors presumably approached
plug flow. Because of the small reactor diameter, radial mixing by molecular diffusion was fast
enough to compensate for mixing effects caused by the laminar flow velocity profile (the Peclet
number always exceeded 50).
Sinag et al. (Sinag et al., 2003) conducted SCWG experiments in a 190 mL metal autoclave that was
equipped with a stirrer. In their experiments, a cold feed stream was continuously injected into the hot
autoclave.
Literature data already indicate that the reactor material has a significant effect on the gasification
process. Yu et al. (Yu et al., 1993), for instance, reported results of gasification of glucose in the same
experimental setup under identical conditions, but using other metals for the reactor. They found large
differences between experiments that were conducted in Inconel, Hastelloy, and corroded Hastelloy
reactors. For a test series using 0.6 M acetic acid, they observed 14% carbon conversion in the Inconel
reactor, whereas in the corroded Hastelloy reactor, the carbon conversion was as high as 53%.
In the present work, catalytically inert quartz capillaries have been used for the investigation of
gasification in hot compressed water. These results are supposed to represent the plain performance of
the process. Table 3.1 shows a comparison between glucose gasification tests in the quartz capillaries
and in metal reactors (data from literature) for almost identical conditions.
Clearly, the results obtained in the noncatalytic quartz tubes deviate from those obtained in metal
reactors (see Table 3.1) in cases where the feedstock contains only a low amount (1.8 wt %) of
glucose. Compared to Inconel reactors, the conversion of carbon from the feed to the gas phase is less
in the quartz capillaries. Apparently, nickel in the Inconel reactor wall (there is ca. 60 wt % nickel in
Inconel 625) catalyzes the glucose decomposition into gaseous components. Results that have been
obtained in the stainless steel reactor are similar to those obtained in quartz capillaries, with respect to
carbon conversion; however, the hydrogen production, as a result of the water-gas shift reaction, is
significantly higher. Al data obtained in the quartz capillaries were compared with (almost) identical
tests that have been reported in the literature. Based on this analysis, it can be concluded that the
carbon conversion in the quartz capillaries is lower than that in Inconel and Hastelloy reactors. The
difference in carbon conversion is high, up to 20%, for low-concentration feeds (for example, 3 wt %)
and reduces to zero for solutions that contain > 10 wt % of organics. Metal reactors always show more
water-gas shift activity than the quartz capillaries, irrespective of the feedstock concentration.
Table 3.1 Comparison between SCWG results obtained in inert quartz capillaries and metal reactors
(data from literature)
feed glucose, low concentration
1.8 wt % 1.8 wt % 1.8 wt %
glucose, high concentration
10 wt % 10 wt %
reference (Hao et al.,
2003) (Yu et al., 1993) this work (Lee et al., 2002) this work
reactor SS 316 tubular Inconel tubular Quartz batch Hastelloy tubular Quartz batch
T [oC] 600 600 600 600 600
P [bar] 250 345 300 280 300
Xc [-] 61 90 70 67.3 69
gas composition [mol %]
H2 31.2 61.6 13.3 43.1 11.7
CO2 41.7 29 20 28.2 8.9
CO 21.8 2 53 9.7 60.5
CH4 4.0 7.2 6 11.6 12.9
C2+ 1.3 not available 7.4 7.4 6.0
Generally, it can be concluded that, in metal reactors, carbon monoxide formation is low and H2
formation is high, compared to the quartz capillaries (see Table 3.1). This indicates that small-scale
metal reactors have a tendency to promote the water-gas shift reaction.
To show the catalytic effect of Inconel, Inconel powder (dp = 100 - 200 µm) was added to the inert
quartz capillaries. Tests were conducted at 600°C with a 5 wt % glycerol solution. The results are
presented in Figure 3.1.
59
0 2 4 60.0
0.2
0.4
0.6
0.8
1.0
1.2
1.4
H2 CO2
CO CH4
C2H4 C2H6
inconel powder [g Inconel/g solution]
yiel
d [m
ole
gas/
mol
e gl
ycer
ol]
0
20
40
60
80
100
carbon efficiency [%]carbon efficiency [%
]
Figure 3.1 Effect of Inconel 625 on the carbon efficiency and product-gas yields. Inconel was added
to the capillaries in powder form. A 5 wt % glycerol solution in water was used as feedstock. T = 600 oC, P = ± 300 bar, τ = 60 s.
Although the mixing of the powder in the capillary was not ideal, resulting in some scatter in the
results, the general trend of the carbon conversion and gas distribution can be observed. The carbon
conversion and the yields of CO, CO2, CH4, and H2 increase due to the addition of Inconel powder to
the capillary reactor (see Figure 3.1). Although it is impossible to relate the amount of Inconel added
to the capillaries to the catalytic activity of a metal reactor wall, it is obvious that Inconel has a
profound influence on the results of glucose gasification.
It is clear from the data presented above that the metal laboratory reactors used in previous
investigations exhibited a catalytic effect. Metal walls seem to catalyze the carbon conversion (total
gas production) and the water-gas shift equilibrium. Because of this catalytic effect of the wall, which
may also be influenced by corrosion, published data are difficult to compare and interpret. Besides, the
catalytic activity of small laboratory equipment cannot be translated to large-scale reactors, which
have a much smaller ratio of wall area to volume. As a consequence, reliable information on the
chemistry and kinetics of biomass and waste gasification in high-temperature and high-pressure water
is missing. The difficulties regarding the interpretation of laboratory-scale data become already clear
when comparing the available laboratory data with the few pilot-plant experiments that have been
conducted so far. In pilot-plant runs with the UT/BTG facility (Potic et al., 2004; Van de Beld et al.,
2003), less H2 and more CO production was observed, compared to the tests conducted in laboratory
reactors made from the same material and under similar conditions (T, P, c, τ). This can be ascribed to
the lower ratio of (catalytic) wall-area to volume of the pilot facilities. The tubular reactor in the pilot
facility of BTG/UT has a specific wall area of 286 m2 per m3 reactor, being ca. a factor of 7 lower than
for the majority of the laboratory reactors (see above). 60
Noncatalytic Gasification in Hot Compressed Water
It has been demonstrated that tests in small laboratory reactors are obscured by undefined catalytic
effects. In this section, results of gasification experiments in catalytically inert quartz capillaries that
are supposed to represent the plain performance of SCWG will be presented.
Effects of temperature and concentration
Decomposition in hot compressed water starts already at temperatures as low as 200oC, producing
mainly tarry products and char. Figure 3.2 shows a photograph of a capillary reactor after reaction.
The reaction medium of 10 wt % glucose solution was exposed to 400oC and 250 bar for a reaction
time of 60 s. The photograph shows that, at such low temperatures, liquid and solid products are
formed. Currently, the exact nature of these reaction products cannot be identified with the applied
capillary technique.
Figure 3.2 Photograph of the content of a capillary reactor after low-temperature gasification. A 10
wt % glucose solution in water was used as feedstock. (T = 400oC, P = ± 300 bar, t = 60 s).
Figure 3.3 shows the effect of the reactor temperature and the concentration of organics in the feed on
the carbon conversion. For the model compounds (glucose and glycerol), the conversion reaches a
constant level beyond 650°C. Further increasing the temperature up to 800oC does not change the
carbon conversion. Hence, the carbon conversion reaches a maximum value. This maximum
conversion is dependent on the concentration of organic material in the feed: 1 wt % glucose in water
can be converted completely, whereas the maximum carbon conversion of 17 wt % glucose is only
83%. This dependency of the conversion on the concentration was also found by Antal and co-workers
(Yu et al., 1993). Pinewood has a much lower maximum conversion (ca. 40%) than the model
compounds used. Below 650°C, the carbon conversion is a strong function of the temperature. For
glucose and glycerol, the difference between 600°C and 650°C is, more or less, equivalent to 20%
conversion. This result points toward the need for precise control and measurement of the reactor
temperature to obtain well-defined experimental results. In that respect, the temperature gradients and
peak temperatures, sometimes reported to be present in the reactors used to study SCWG (Antal et al.,
2000), may have influenced earlier results significantly. This has also been discussed by Antal et al.
(Antal et al., 2000). 61
400 500 600 700 8000
20
40
60
80
100
glycerol 19 wt % wood 15 wt% glucose 1 wt% glucose 10 wt% glucose 17 wt%
Xc
[%]
T [oC] (a)
0
20
40
60
80
100
700650600
gas
com
posi
tion
[%]
T [oC]
glycerol 19 wt% H2
CO2
CO CH4
5500
20
40
60
80
100
glucose 17 wt% H2
CO2
CO CH
4
(b)
Figure 3.3 (a) Influence of the temperature and the concentration on the carbon efficiency for SCWG
of glycerol, glucose, and pine wood. (b) Influence of the temperature on the gas production for SCWG
of high-concentration solutions of glucose and glycerol (C2 and C3 components are not shown. P = ±
300 bar. τ = 60 s.).
Glucose and glycerol show identical gasification behavior, with respect to the carbon conversion (see
Figure 3.3a). This is rather surprising, given the fact that they are different chemical species: glucose
is a sugar and glycerol is an alcohol. It may indicate that the main reactions of gasification in hot
compressed water are thermal cracking reactions, just like in dry gasification. Also, in dry gasification,
it has been found that the exact nature of the feedstock hardly influences the process performance
(Van der Drift et al. 2001). However, looking at the H2/CO ratio for high concentrations of glycerol
and glucose at temperatures of 550, 600, 650, and 700oC (see Figure 3.3b), a significant difference is
observed. The production of hydrogen, mainly as a result of the water-gas shift reaction, is at least 2
times higher in the case of glycerol. Apart from H2, CO, CO2, and CH4, some C2 (C2H4, C2H6) and C3
(C3H6, C3H8) components are produced (~10%), with ethane being a major component.
62
The carbon conversion of wood is very low, compared to that of glucose and glycerol, which can be
ascribed by the presence of lignin in wood. Lignin itself is difficult to gasify and it has been observed
that lignin blocks the conversion of wood’s other constituents: cellulose and hemi-cellulose (Yoshida
and Matsumura, 2001; Yoshida et al., 2004).
The effect of the reactor temperature on the product distribution is discussed based on the results of
glucose solutions processed at 300 bar and a reaction time of 60 s, with a low (1 wt %), medium (10
wt %), and high (17 wt %) concentration of glucose (see Figure 3.4). The main gas products of the
reaction are carbon monoxide, carbon dioxide, hydrogen, methane, ethane, and propane. Along with
the carbon conversion, the yield of CO2, H2, and CH4 increase as the temperature increases. Also, the
CO yield initially increases, but above 700°C, CO rapidly decreases at constant carbon conversion (see
Figure 3.4c), indicating that, above this temperature, the water-gas shift reaction starts to proceed. C3
components are not present in the product gas at temperatures of > 700°C and the yield of C2
components strongly decreases at > 700°C. This could point to enhanced reforming of C2 components
at > 700oC.
400 500 600 700 8000
2
4
6
8
10 1wt%
10 wt%
17 wt%
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
hydrogentheoretical value
400 500 600 700 800
0
1
2
3
4
5
6
1wt%
10 wt%
17 wt%
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
carbon dioxide
(a) (b)
400 500 600 700 8000.0
0.5
1.0
1.5
2.0
2.5
3.0 1wt%
10 wt%
17 wt%
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
carbon monoxide
400 500 600 700 800
0.0
0.5
1.0
1.5
1wt%
10 wt%
17 wt%
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
methane
(c) (d)
63
400 500 600 700 8000.0
0.1
0.2
0.3
0.4
1wt%
10 wt%
17 wt%
yi
eld
[mol
e ga
s/ m
ole
gluc
ose]
T [oC]
C2
400 500 600 700 8000.00
0.05
0.10
0.15
0.20
1wt%
10 wt%
17 wt%
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
C3
(e) (f)
Figure 3.4. Influence of the temperature and the concentration of the glucose on the gas yields of (a)
hydrogen, (b) carbon dioxide, (c) carbon monoxide, (d) methane, (e) C2 components, and (f) C3
components. (P = ± 300 bar, t = 60 s).
The effect of the concentration on the yields can be summarized as follows. H2 and CO2 increase with
decreasing concentration while CO, CH4, and C2 decrease. Figure 3.5 shows that, over the entire
temperature range, the measured yields do not resemble the yields predicted by chemical equilibrium.
At < 700oC, the deviation is very large. At 800oC and 1wt %, the experimentally found yields slowly
start to approach chemical equilibrium mainly due to an enhanced water-gas shift reaction (see Figure
3.5) The equilibrium calculations were performed according to the Gibbs minimization method (Kyle,
1999). Fugacity coefficients were calculated with the modified Soave-Redlich-Kwong equation of
state (Soave et al., 1993). The parameters for the modified Soave-Redlich-Kwong equation of state
were obtained from Bertucco and Soave (Bertucco et al., 1995, Soave, 1972). Thermodynamic
properties were taken from a NIST-JANAF source (NIST).
64
400 500 600 700 8000
1
2
3
4
5
6
7
CH4
CO
Capillary resultsH2CO2COCH4
gas
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
H2
CO2
400 500 600 700 8000
1
2
3
4
5
6
7
CH4
CO
Capillary resultsH2CO2COCH4
gas
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
T [oC]
H2
CO2
(a)
400 500 600 700 8000
2
4
6
8
10
12
CH4CO
CO2
H2Capillary results
H2 CO2 CO CH4
ga
s yi
eld
[mol
e ga
s/ m
ole
gluc
ose]
T [oC] (b)
Figure 3.5 Comparison between the measured and predicted chemical equilibrium product
distribution: (a) 17 wt % glucose and (b) 1 wt % glucose. (P = ± 300 bar, t = 60 s).
In Figure 3.6, the cold-gas efficiency, and the individual contribution of the components to it, are
plotted. In the same figure, the LHV of the product gas also is presented.
Clearly at low temperature (low carbon conversion), only a small amount of energy stored in the feed
can be transferred to the gas. For all three glucose solutions (1, 10, and 17 wt %) the maximum cold-
gas efficiency is attained near the temperature at which the carbon conversion is maximal (compare
Figure 3.3 and 3.6). Up to 700 oC, the product gas can be characterized as fuel gas: it consists of CO,
H2, CH4, and C2, all of which contribute significantly to the heating value. Above 700oC, the product
65
0
5
10
15
20
400 500 600 700 800 9000.0
0.2
0.4
0.6
0.8
1.0
1.2 H2
CO CH4
C2,3
LH
V g
as/ L
HV
gluc
ose
T [oC]
glucose 1 wt%
LH
V g
as, M
J/m
3
LHVgas
(a)
400 500 600 700 800 9000
5
10
15
20
400 500 600 700 800 9000.0
0.2
0.4
0.6
0.8
1.0
1.2
LHV
gas/
LH
V gl
ucos
e
T [oC]
glucose 10 wt%
LH
V ga
s [M
J/m
3 ]
(b)
400 500 600 700 800 9000
5
10
15
20
400 500 600 700 800 9000.0
0.2
0.4
0.6
0.8
1.0
1.2glucose 17 wt%
T [oC]
LH
V ga
s/ L
HV
glu
cose
LH
V g
as, M
J/m
3
(c)
Figure 3.6. Cold gas efficiency, the individual contribution of the components to it, and the LHV of
the produced gas versus the temperature: (a) 1 wt %, (b) 10 wt %, (c) 17 wt %. (P = ± 300 bar, t = 60
s).
66
67
gas is more pure: for the 1 wt % solution, the majority of the energy in the feed is transformed to
hydrogen; a 10 wt % glucose solution produces a H2/CH4 mixture with some CO and C2, and in the
case of 17 wt % glucose solution, methane represents most of the gas’ energy content.
When the gas mixture becomes less complex at higher temperatures, the cold-gas efficiency slightly
decreases. This effect can be ascribed to pre-reforming reactions of C2 components to methane, and to
hydrogen production via the water-gas shift reaction, which both are exothermic.
The volumetric heating value of the product gas goes through a maximum in the range of 600 - 700oC.
Hence, above 700oC, the produced gas is more pure, but the heating value of the gas is significantly
lower.
Effect of reaction time
Figure 3.7 shows that already after 5 s of reaction time at 600oC, some gaseous products are produced
from glucose SCWG. After 40 s of reaction time at 600oC, the carbon conversion and product
distribution reach their asymptotic value. Increasing the reaction time does not increase the carbon
conversion, nor does it change the product distribution. The concentration of the glucose does not
significantly influence the relative conversion rate, which suggests a possible first-order overall
reaction.
68
000 20 40 60 80 10
20
40
60
80
100
1.8 wt.% 20 wt.%
Xc [%
]
time [s]
glucose
(a)
0 20 40 60 80 1000
1
2
3
0 20 40 60 80 1000
1
2
3
20 wt.% H2
CO2
CO CH4
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
time [s]
1.8 wt.% H2
CO2
CO CH4
(b)
Figure 3.7. Influence of the reaction time on (a) the carbon conversion and (b) the gas yield for 1.8
wt % and 20 wt % glucose solution. (T = 600oC, P = ± 300 bar).
Another important finding regarding the reaction time is that the heating trajectory has no considerable
effect on the final carbon conversion. This result was obtained by comparing the carbon efficiency
from interrupted and noninterrupted experiments (Potic et al. 2004).
In a practical SCWG process, the reaction zone is spread out over the heat-exchanger and the reactor.
Although laboratory-scale tests showed that the heating trajectory has no effect on the final
conversion, it remains important that all products formed in the heat-exchanger (gas, liquid, and
solids) end up in the reactor and are not deposed on the heat-exchanger wall. In that way, the final
carbon efficiency will be the result of the process conditions as prevailing in the reactor. This may be
achieved by using a particle loaded scouring heat-exchanger (Kersten et al., 2004).
Effect of the pressure
Although it has been suggested that supercritical water properties would be beneficial for the
gasification reaction, experimental results from capillary reactors showed that pressure has no
influence on the carbon efficiency (see Figure 3.8a) and product gas yield (see Figure 3.8b). It is
crucial for the SCWG process that the heat content of the reactor effluent be utilized as much as
possible to preheat the feedstock stream (mainly water) to reaction conditions. However, heat
exchange between these streams is not practical at low pressure, because of the high heat of
evaporation under almost isothermal and isobaric conditions. Heating of the feedstock stream to the
desired gasification temperatures in a heat-exchanger without evaporation requires operation at high
pressures. This is the true incentive of the high pressures involved in wet gasification.
0 100 200 300 400 5000
20
40
60
80
100
0 100 200 300 400 5000
20
40
60
80
100
0 100 200 300 400 5000
20
40
60
80
100
Xc [%
]
P [bar]
glucose 10 wt.%
cellulose 10 wt.%glycerol 10 wt. %
(a)
H2 CO2 CO CH4 C2H4 C2H6 C3H6 C3H80
1
2
3
4
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
138 bar 162 bar 197 bar 255 bar 297 bar 418 bar
glucose 650 oC, 10 wt.%
(b)
Figure 3.8 Effect of pressure on (a) the carbon efficiency for glucose, glycerol, and cellulose, and (b)
the gas yield for the 10 wt % glucose solution. (T = 650 C, t = 60 s)o .
69
Catalytic SCWG
Addition of potassium and sodium cations
Many researchers (Kruse et al., 2000; Schmieder et al., 2000; Sinag et al., 2003; Sinag et al., 2004;
Van de Beld et al., 2003) have investigated the influence of K+ and Na+ cations on the SCWG process.
Kruse et al., for example, found that these alkali metals promote the water-gas shift reaction slightly,
but not the carbon conversion. We performed a few experiments with these alkali metals, which
confirmed their conclusions. The influence of NaOH on the gas distribution is presented in Table 3.2.
Table 3.2. Influence of the addition of NaOH to a 17 wt % glucose feed solution.
parameter no additives 0.5 wt % NaOH 1 wt % NaOH 3 wt % NaOH
temperature, T [oC] 600 600 600 600
pressure, P [bar] 300 300 300 300
Reaction time, t [s] 60 60 60 60
carbon conversion, Xc 60 55 52 60
gas composition
H2 9.9 21 17 17
CO2 8.2 47 45 47
CO 62.2 7 12 9
CH4 12.9 19 20 21
C2 4.5 4 4.2 4.2
C3 2.3 2 1.8 1.8
Use of 3 wt % ruthenium on TiO2
A compound comprised of 3 wt % ruthenium on TiO2 (Ru/TiO2) was kindly provided to us by Pacific
Northwest National Laboratory. This catalyst was used by Pacific Northwest National Laboratory for
the conversion of waste streams, in subcritical water (200 bar, 350°C), to methane (Elliott et al., 1993).
To test the catalyst, a 17 wt % glucose solution that cannot be converted completely without a catalyst
was used. Figure 3.9 shows the results of identical experiments with and without a catalyst.
70
400 500 600 700 8000
20
40
60
80
100
glucose 17 wt.% + Ru catalyst glucose 17 wt.%
Xc[%
]
T [oC]
Figure. 3.9 Influence of the temperature on the carbon efficiency of the 17 wt % glucose solution for
noncatalytic and catalytic processes. (T = 600oC, t = 60 s).
The ruthenium catalyst is able to increase the conversion from 70% to 100% at the commonly used
process temperature of 600oC (see Figure 3.9). This is an interesting and surprising result, because this
catalyst was not designed for SCWG conditions. The measured product distribution, as a function of
the temperature, is given in Figure 3.10. With a catalyst, the product distribution is clearly different
than without a catalyst (compare Figure 3.5 to Figure 3.10). In fact, except for hydrogen, the gas at
600oC resembles the equilibrium gas for those conditions rather well (compare Figure 3.5 and 3.10).
H2 CO2 CO CH40.0
0.5
1.0
1.5
2.0
2.5
3.0
3.5
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
400 oC 500 oC 600 oC
Glucose 17 wt.%
Figure 3.10 Effect of temperature on the gas production for the 17 wt % glucose solution with
ruthenium catalyst (P = ± 300 bar, t = 60 s).
Contrary to noncatalytic SCWG, in case of the Ru/TiO2 catalyst, the conversion is not a function of the
concentration of organics in the feed. For solutions of 1, 5, 12, and 17 wt % glucose in water, complete
71
conversion of carbon to the gas phase is achieved. However, the product distribution remains a
function of the concentration (see Figure 3.11). Ru/TiO2 is designed as a methanation catalyst. If
applied for SCWG, the methane production is strongly dependent on the concentration of the organics
in the feed (see Figure 3.11). The product gas for a 1 wt % glucose solution consists of only hydrogen
and carbon dioxide. With increasing concentration, the product-gas distribution shifts from hydrogen
to the methane-rich gas. With these results, the potential of heterogeneous catalysis is clearly
demonstrated. However, this catalyst may not be suitable for the actual process, because it was
developed for T < 400oC and lifetime at 600oC is never examined. Besides, the support (TiO2) of the
ruthenium catalyst may not be stable under the applied process conditions (Wang et al., 2004).
H2 CO2 CO CH40
2
4
6
yi
eld
[mol
e ga
s/ m
ole
gluc
ose] glucose 1 wt.% + Ru
glucose 5 wt% + Ru glucose 12 wt.% + Ru glucose 17 wt.% + Ru
Figure 3.11 Effect of the glucose concentration on the gas production for the catalytic gasification of
glucose. (T = 600oC, P = ± 300 bar, t = 60 s).
`
72
73
Conclusions
The conclusions of this work can be summarized as follows.
● It has been demonstrated in the present work that results obtained in small-scale (1 mm < d < 7
mm) metal reactors are obscured by undefined catalytic effects of the reactor wall, improving
the carbon conversion (except for the stainless steel reactor) and/or gas shift activity.
● The noncatalytic gasification in supercritical water (SCWG) has been studied in inert quartz
capillaries. From these tests, the following can be concluded:
● Below 650oC, the carbon conversion is a strong function of the temperature.
● Increasing the temperature from 650oC to 800oC hardly affects the carbon conversion. The
maximum conversion is a function of the concentration. Complete conversion of the feedstock
to the gas phase is achieved only for very diluted solutions of glucose and glycerol (1 wt %).
Higher concentrations (≥ 5 wt %) cannot be gasified completely. The carbon conversion is
typically in the range of 75% - 90% for solutions that contain > 5 wt % organics. A high
temperature of 800oC is needed for enhanced water-gas shift activity to produce hydrogen, but
equilibrium still is not reached.
● For glucose solutions with a concentration of 1 - 20 wt %, a reaction time of 40 s is sufficient to
reach the maximum conversion at 600oC. The relative conversion rates are hardly dependent on
the concentration of the feedstock solution, indicating an overall first-order reaction.
● Under identical conversion conditions, pinewood has a considerably lower maximal
conversion than glucose and glycerol (e.g., 45% vs 83% at 700oC).
● Pressures in the range of 50 - 500 bar have no influence on the carbon conversion and the
product distribution is not affected in the range of 50 - 450 bar.
● Up to 700oC, the product gas can be characterized as fuel gas that consists of CO, H2, CH4, and
C2+ components. Above 700oC, the product gas becomes less complex: at low glucose
concentration, an excess of H2 is produced; a medium concentration results in a mixture of H2
and CH4, and a high concentration produces mainly CH4.
● The addition of K+ or Na+ cations to the reaction mixtures promotes the water-gas shift
reaction, which leads to more H2 and less CO but does not increase the carbon conversion.
● By adding a 3 wt % ruthenium on TiO2 (Ru/TiO2) catalyst, complete conversion of solutions
with 1 to 17 wt % glucose is achieved. These experiments show the potential of heterogeneous
catalysis. However, a catalyst for the process is still missing and the development of
commercial catalysts should be considered as a necessary and essential step in the process
development.
Notation
feedLHV Lower heating value of feed dry and ash free, MJ/m3, MJ/mol
igasproductLHV ,− Lower heating value of product gas component i, MJ/m3, MJ/mol
icN , Number of moles carbon in product gas component I produced, mol
feedcN , Number of moles carbon in the feed, mol
igasproductN ,− Number of moles of product gas component i
feedN Number of moles of feed
P Reaction pressure, bar
T Reaction time, s
T Reaction temperature, oC
Xc Carbon conversion, dimensionless, or %
iY Gas yield, mol/mol
LHVη Cold gas efficiency, %
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Matsumura, Y.; Minowa, T.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; van de Beld, L.; Elliott, D.C.; Neuenschwander, G.G.; Kruse, A.; Antal, M.J., Biomass Gasification in Near- and Supercritical Water: Status and Prospects. Biomass Bioenergy 2005, 29, 269. Modell, M.; Gasification and Liquefaction of Forest Products in Supercritical Water. In Fundamentals of Thermochemical Biomass. Overend, R.P.; Milne, T.A.; Mudge, L.K., Eds.; Elsevier Applied Science Publishers Ltd.: London, 1985, p. 95. National Institute of Standards and Technology (NIST), Gaithersburg, MD. Potic, B.; van de Beld, L.; Assink, D.; Prins, W.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water. In Proceedings of the 12th European Conference and Exhibition on Biomass for Energy, Industry and Climate Protection; Palz, W.; Spitzer, J.; Maniatis, K.; Kwant, K.; Helm, P.; Grassi, A., Eds.; ETA Florence, WIP Munich: Amsterdam, 2002; p. 777. Potic, B.; Kersten, S.R.A.; Prins, W.; Assink, D.; van de Beld, L.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water: Results of Micro and Pilot Scale Experiments. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection, van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 742. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., A High-throughput Screening Technique for Conversion in Hot Compressed Water. Ind. Eng. Chem. Res. 2004, 43, 4580. Sarada, R.; Joseph, R., A Comparative Study of Single and Two Stage Processes for Methane Production from Tomato Processing Waste Process. Biochem. (Oxford, U.K.) 1996, 31, 337. Schmieder, H.; Abeln, J.; Boukis, N.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht, M., Hydrothermal Gasification of Biomass and Organic Wastes. In Proceedings of the 5th Conference on Supercritical Fluids and their Applications, Garda: Italy, 1999; p. 347. Schmieder, H.; Abeln, J.; Boukis, N.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht, E., Hydrothermal Gasification of Biomass and Organic Wastes. J. Supercrit. Fluids 2000, 17, 145. Sinag, A.; Kruse, A.; Schwarzkopf, V., Key Compounds of the Hydropyrolysis of Glucose in Supercritical Water in the Presence of K2CO3. Ind. Eng. Chem. Res., 2003, 42, 3516. Sinag, A.; Kruse, A.; Rathert, J., Influence of the Heating Rate and the Type of Catalyst on the Formation of Key Intermediates and on the Generation of Gases During Hydropyrolysis of Glucose in Supercritical Water in a Batch Reactor. Ind. Eng. Chem. Res. 2004, 43, 502. Soave, G., Equilibrium Constants from a Modified Redlich-Kwong Equation of State. Chem. Eng. Sci. 1972, 27, 1197. Soave, G.; Barolo, M.; Bertucco, A., Estimation of High-Pressure Fugacity Coefficient of Pure Gaseous Fluids by Modified SRK Equation of State. Fluid Phase Equilib. 1993, 91, 87. Van de Beld, L.; Wagenaar, B.M.; Assink, D.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; Penninger, J.M.L., Biomass and Waste Conversion in Supercritical Water for the Production of Renewable Hydrogen. Presented at the 1st European Hydrogen Energy Conference, Grenoble: France, 2003.
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Van der Drift, A.; van Doorn, J.; Vermeulen, J.W., Ten Residual Biomass Fuels for Circulating Fluidized-bed Gasification. Biomass Bioenergy 2001, 20, 45. Wang, J.A.; Cuan, A.; Salmones, J.; Nava, N.; Castillo, S.; Moran-Pineda, M.; Rojas, F., Studies of Sol-Gel TiO2 and Pt/TiO2 Catalysts for NO Reduction by CO in an Oxigen-rich Condition. Appl. Surf. Sci. 2004, 230, 94. White Paper on Environmental Liability. European Commission, 2000. Xu, X.; Matsumura, Y.; Stenberg, J.; Antal, M.J., Carbon-Catalyzed Gasification of Organic Feedstocks in Supercritical Water. Ind. Eng. Chem. Res. 1996, 35, 2522. Xu, X.; Antal, M.J. Jr., Gasification of Sewage Sludge and Organics in Supercritical Water. Presented at the 1997 AIChE Annual Meeting, 1997. Yoshida, T.; Matsumura, Y., Gasification of Cellulose, Xylan and Lignin Mixtures in Supercritical Water. Ind. Eng. Chem. Res. 2001, 40, 5469. Yoshida, T.; Oshima, Y.; Matsumura, Y., Gasification of Biomass Model Compounds and Real Biomass in Supercritical Water. Biomass Bioenergy 2004, 26, 71. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels 1993, 7, 574.
78
79
i l
CCChhhaaapppttteeerrr 444 PPPiilllooottt PPPllaaannnttt
Abstract
In this Chapter, a first generation pilot plant for supercritical water gasification with a capacity of 30
litres per hour wet intake is presented. Several experiments with non-fouling feedstock materials such
as glycerol and methanol are successfully performed and the results are compared with experiments
obtained in continuous and batch microreactors. Experiments with more complex feedstocks such as
starch revealed that charring feedstocks cannot be gasified in this pilot plant without operational
problems such as fouling and blockages of heat-exchanger and reactor.
80
Introduction
Research on SCWG was started in Twente in the year 2000 in an international program involving the
University of Twente, a private Dutch biomass enterprise, BTG B.V., and the Japanese partners NIRE
and the University of Hiroshima. At that time, this consortium obtained a research grant from the
Japanese NEDO organization to explore the opportunities of SCWG. Part of the NEDO-grant was
designated to design, construct and operate a pilot plant facility in Twente to gather process design
data. The potential of SCWG has recently been extensively discussed in the literature (Antal et al.,
2000; Boukis et al., 2002; Hao et al., 2003; Kruse et al., 2003; Lee et al., 2002; Matsumura et al.,
2005; Potic et al., 2002; Potic et al., 2005; Van de Beld et al., 2003) and in some earlier work (Antal et
al., 1999; Holgate et al., 1995; Kabyemela et al., 1998; Kabyemela et al., 1999; Manarungson et al.,
1993; Modell, 1985; Schmieder, 1999; Xu et al., 1996; Xu and Antal, 1998; Yu et al., 1993). On the
other hand, possible problems such as the feeding of slurries and fouling in the reactor and auxiliary
equipment are not yet fully identified. The team in Twente, however, anticipated some of the problems
and proposed a preliminary concept of a supercritical fluid bed system to prevent fouling and to allow
for the introduction of catalytic material and its possible regeneration. Unfortunately, in 2000 there
was no information available necessary for a design of a pilot plant based on a supercritical fluid bed.
Nevertheless, it was decided to make a start with the pilot plant research program. The strategy was to
start with a first generation plant in order to get continuous operation experience with process
pressures of ca. 250 bar and temperatures exceeding 600oC. This first generation plant included a shell
and tube heat-exchanger and a tubular reactor and was supposed to take in simple “non-fouling”
feedstocks. It was projected that after this initial phase more basic data on SCWG and supercritical
fluidization would have been collected, amongst others from the present thesis project. The shell and
tube heat-exchanger and the tubular reactor would subsequently be replaced by a fluid bed system to
develop the process for more complex feedstock materials. Several foreseen and unforeseen problems
such as legislation difficulties, lack of operational experience with some of the high-pressure
machinery and the extremely long delivery time of special parts and materials caused delays in the
program. In fact, the first satisfactorily runs with the first generation pilot plant could only be
performed as late as mid 2004. At the time of writing, the fluid bed system has not been installed in
the pilot plant yet.
In this chapter, the first generation SCWG pilot plant with a maximal intake of 30 kg wet biomass per
hour is presented. Several experiments with non-fouling model compounds are reported and compared
with results obtained in the microreactors (see Chapters 3 and 5). In addition, by analyzing the test
runs with more complicated feedstocks possible operational problems are identified.
81
The pilot plant
The pilot plant is designed and constructed by BTG B.V. as a part of the consortium activities and is
located in a concrete high-pressure safety box in the High-Pressure Laboratory of the University of
Twente. It is designed for operation at temperatures up to 650oC and pressures of 300 bar with a
maximum capacity of 30 kg/h wet feedstock. A schematic presentation of the plant is given in Figure
4.1. The plant is modular built and consists of a number of units: a pump, a shell and tube heat-
exchanger, a coiled tubular reactor positioned in an oven, a cooler, a gas-liquid separation section and
gas product analysis.
Figure 4.2 is a detailed flow sheet of the pilot plant. The feed storage system consists of two vessels
(S-002 and S-003, see Figure 4.2). One additional vessel for water (S-001) is installed to facilitate the
start-up and shut-down procedure. The effluent liquid stream from the process (mainly water) is
collected in a fourth vessel (S-006). To monitor the actual flows out of and into the vessels, weighing
balances are installed under each vessel. The feed is supplied to the system by a LEWA piston pump
(P-001) with maximum capacity of 30 kg/h and maximum operating pressure of 300 bar. A shell and
tube heat-exchanger (HE-001) with a length of 1.5 m is installed between the inlet and outlet streams
of the reactor. The feedstock flows through the outer tube (i.d. = 5/4" ~ 30 mm, Incoloy 825), and the
reactor effluent through the inner tube (i.d. = 1/2" ~ 1.3 mm, Incoloy 825). The reactor (R-001) is a
coiled tube (Incoloy 825) with a total length of 16 m and an inside diameter of 13.84 mm. The coil
itself has a diameter of 30 cm and a height of 60cm. It is placed in the center of a gas oven with a
maximum capacity of 20 kW. After passing through the heat-exchanger, the reactor effluent is cooled
with water in a double wall cooler (C-001). After this cooler, the two-phase product stream is
separated in the separation section, which consists of a high pressure (HP) separator (SEP-001 in
Figure 4.2), a pressure reducing valve system and two low-pressure (LP) separators (SEP-002 and
SEP-003 see Figure 4.2). The high-pressure separator is a vessel of stainless steel with a volume of
about 2 L and is operated at the prevailing process pressure. In this vessel the two-phase process
stream is separated into a liquid (mainly water) and a gas stream at a temperature of 50 to 100oC. The
liquid level in the vessel is controlled with an ultrasonic level indicator in combination with an electro-
pneumatic control valve. After the high pressure separator, the pressure is reduced to ambient
conditions by a pneumatically operated ball valve in the gas-phase line. An electronic backpressure
controller (PIC-001) placed after this ball valve is used to control the system pressure. Both the gas
and the liquid stream from the high pressure separator are again separated at ambient conditions in
SEP-002 and SEP-003, respectively. In the low pressure separator SEP-002 handling the gas stream
from the HP separator, water is removed down to a level corresponding to atmospheric pressure and
ca. 20oC. Dissolved gases in the liquid stream from the HP separator are removed from the liquid
stream in SEP-003. The product gas streams from the plant GAS 1 and GAS 2 are monitored
Figure 4.1 Schematic presentation of the pilot plant.
The pilot plant is started-up by pressuring it with nitrogen. As a result the gaseous product streams
GAS 1 and GAS 2 (see Figures 4.1 and 4.2) contain, during the first hours of operation, considerable
amounts of nitrogen. The mass balance was derived during stable operation when the nitrogen fraction
in the outlet streams was below 10 volume percent. To avoid extremely long run times requested to
purge out all the initial nitrogen and because the partial pressure of the water vapor is dominant while
the process was found to be relatively insensitive to the absolute pressure (see Chapter 3), for the
purpose of the present investigation this remaining nitrogen was accepted.
separately with respect to composition and flow rate. In the mass balance calculations the component
flows in kg/h of the streams are added and together define the gaseous product flow (see below).
R-001Reactor
(in gas heated oven)
To gas analysis section
To gas analysis section
FEED
GAS-LIQUID separation section
P-001Pump
HE-001Shell and tube heat exchanger
C-001Cooler
SEP-001HP separator
C-002
SEP-003
SEP-002
WATER 2
WATER 1
GAS 1
GAS 2
82
5
R-001E-1922
H-002
FC003
FC004
FI006
NATURAL GAS
AIR
TI014
TI005
TIC002
T
P-001 FC001
4
C-001D-1922
HE-001E-1922
TI003
TI007
8
6
TI002
HPS001
PI001
SV-001
TI008
9
TI015
FI007
TI016
TI009-1
10
S-001
LI001
LLA001
LI002
LLA002
LI003
LLA003
S-002
BV001
BV002
BV003
1 2 3
FI001
S-003
wate
r
wat
er /
etha
nol
wat
er /
biom
ass
FEEDING SECTION
HEAT EXCHANGER
coolant in
coolant out
FURNACE
HLA 002
LI005
BV005
S-006
Prod
uced
wat
er
FI005-2
Water II
Water II17
17
TI001
QI001
TI021
BV010 FS
001
QS001
QS002
QS003
TS021
TS014
TS016
003
SEP-001C-1922
LLA004
LI004
BV004
FC002
15
SEP-003
S-005
BV009
QI013
FI005 -1
TI013
17
17
17
16
GAS II
QI009
TIC007
TI012 -2
FI004
TI012 -1
S-004
P-002
TI011
14
14
14
13
GAS I
QI003
FI002
TI010 -2
TIC006
TI010-1
11
12 SEP-002
WATER II
WATER I
HP-G/L SEPARATOR
LP-G/L SEPARATOR
LP-G/L SEPARATOR
P-003
PIC002
NITROGEN
HLA001
BV007
FC006
TI009 -4
TI009-2
Coolant out
Coolant in
BV006
FC005
TI009 -7
TIC005-2
PIC001
TI009-3
TI017
Coolant in
TI018
Coolant out
BV008
LLA005
001
QI004
TI009-5
HLA004
PI003
002
QI010
LLA006
Coolant in TI
019
Coolant out
TI020
HLA003
C-002
TI009-6
TITmax
TITmax
TITmax
PI002
83
Figure 4.2 Detailed flow sheet of the plant.
Measurements during pilot plant operation
Temperatures, pressures, flow rates and concentrations are measured at several locations of the system
(see Figure 4.2). The run time starts from the moment that set process parameters (temperature and
pressure) are reached and feeding system is switched from water to the feedstock. All data are
collected and processed by an industrial process computer equipped with ADAM data acquisition
modules (Adam 5000E, supplier Advantech, GeniDaq software was used).
In Figure 4.3 direct read-outs from the weighing balances of the feedstock supply vessel (S002 and
S003) and the total liquid effluent collecting vessel (S006) are presented. The total liquid effluent of
the pilot plant is the sum of the liquid out-let streams of the low-pressure flash vessels SEP-001 and
SEP-002. It can be seen that the inlet stream has a constant rate (after more than 3 h run time), but that
the exit flow rate fluctuates requiring a considerable time of averaging. The exit flow rate fluctuates
predominantly as a result of control actions of the level controller in the high-pressure separator.
Figure 4.3 Readouts from the weighing balances of the feedstock and total liquid effluent. Liquid
mass in the vessels versus the run time are plotted.
Exit gas flow rates (GAS 1 and GAS 2) are measured with mass flow indicators. These flows are
fluctuating (see Figure 4.4), also because they are affected by the actions of controllers. The flow rate
of GAS 1, being the dried gas from the high-pressure separator, is directly influenced by the system
pressure controller. GAS 2 which, is composed of the gases dissolved in the high-pressure liquid
(water), is influenced by the level controller in the high-pressure separator. In the mass balance
calculations, the last 30 min of gas production are used, which corresponds to the lowest nitrogen
level.
84
3.33 3.55 3.77 3.99 4.21 4.44 4.66 4.860.0
0.2
0.4
0.6
0.8
1.0
1.2
1.4
corre
cted
gas
flow
rate
[m3 / h
]
time [h]
GAS 1 (rich in H2) GAS 2 (rich in CO2)
Figure 4.4 Produced gas flow rates versus the run time. Differential mass balances were made over
the boxed time span.
As mentioned before, the pilot plant is started up and initially pressurized with nitrogen and as a result
nitrogen is present in the exit gas streams during the whole test period. Besides, it has been observed
that the decay of the nitrogen concentration in the exit gas does not follow a PFR or a CISTR pattern.
In reality, the nitrogen concentration level first drops rapidly and then levels off slowly indicating
some dead-zones in the system, which makes mass balancing (using flow rates) over the whole run
time difficult. After careful consideration of all options it has been decided that mass balances during
the time-span of the experiment in which the nitrogen fraction in the exit gas is less than 10 volume
percent will give the most accurate information. The production of gas phase components is calculated
based on the actual gas flow rates and concentrations derived from the read-outs of the GCs. Averaged
values over a time interval of 30 min are taken. For the calculation of the total product yields and the
carbon conversion to gas-phase products, the values of GAS 1 and GAS 2 are added. In Figure 4.6,
typical examples are given of the measured gas compositions in both gas-phase flows over a stable run
period. On-line GCs are used to determine the product gas composition. Components that can be
detected are: H2, CO, CO2, CH4, ethylene, ethane, propylene and propane. Every 5 min, a gas sample
is analyzed by the GC and the analysis time is about 3 to 4 min. Typical yields and conversions to gas-
phase products are only based on the feedstock and the gaseous products. Analyses of the liquid
effluent were only performed incidentally. For these analyses a combined liquid sample of WATER 1
and WATER 2 is taken which is analyzed with respect to its C, H and N content. The elemental
composition of the liquid effluent samples is determined with an elemental analyzer (Flash EA 1112
Thermo Quest, Rodano, Italy).
85
4.4 4.48 4.57 4.67 4.75 4.840
10
20
30
40
50
gas
com
posi
tion
[%]
time [h]
H2 C2H6 N2 C3H6 CO C3H8 CH4 CO2 C2H4
GAS 1
4.36 4.45 4.55 4.65 4.75 4.850
20
40
60
80
100
GAS 2 H2 C2H6 N2 C3H6 CO C3H8 CH4 CO2 C2H4
gas
com
posi
tion
[%]
time [h]
Figure 4.5 Compositions GAS 1 and GAS 2 versus the run time.
The carbon conversion to the gas-phase and the product yields are chosen here as the main output
parameters of the process. The carbon conversion to the gas-phase is defined as degree of conversion
of carbon from biomass to permanent gases:
feedc,feed
compounds gaseous ofnr ic,i
c fx Θ
fx Θ X
∑= 4.1
The product yield is defined as:
feedfeed
iii /MΘ
/MΘY = 4.2
Like mentioned before, the used mass flow rates in these equations are average values of the last 30
min. of gas production and only experiments are used that have a nitrogen fraction of less than 10 vol
% in this period. During the test runs the total pressure in the reactor can be measured accurately by
the installed pressure transducers. Definition of the exact reactor temperature is on the other hand
extremely difficult. One problem is the presence of a heat up trajectory inside the reactor; the feed
stream exits the heat-exchanger at ca. 450°C and is then subsequently heated in the oven in the first
part of the reactor tube to ca. 600°C. A second problem is the temperature profile existing in the gas
oven. Scanning experiments have revealed that this temperature gradient is ca. 100°C from bottom to
the top, the bottom where the gas flame is located being hotter. By the complete lack of an alternative,
it is decided to use the temperature at the exit of the reactor as the characteristic process temperature.
We are aware of the fact that this imprecise temperature definition makes interpretation of the results
cumbersome, because especially in the temperature region below 650°C the conversion is affected to a
large extent by the temperature (see Chapter 3). On the other hand, it has been observed that it is the
86
final temperature that predominantly determines the conversion and product distribution (see Chapters
2 and 3).
Results
Apart from the limitations mentioned above, successful experiments showing no operational problems
have been carried out in the pilot plant with glycerol and methanol as feedstock. Test with starch (a
fouling feedstock) have shown severe operational problems in the form of blockages at several
locations, for example in the heat-exchanger and in the reactor. These blockages occur rapidly and
prevent operation with charring feedstocks.
Mass Balances
In Figure 4.6 the measured temperatures, pressures, flow rates and product gas compositions during a
gasification test with 9.1 wt % glycerol are presented as a typical example. The actual mass balance of
this presented experiment is listed in Table 4.1.
Figure 4.6 Graphical presentation of a typical experiment with glycerol. The presented gas
compositions are normalized leaving nitrogen out.
87
Table 4.1. Mass balance of a typical pilot plant experiment with glycerol. The presented gas
composition is normalized leaving nitrogen out.
Feed
T [oC]
P [bar]
Feed concentration [wt %]
Catalyst
Glycerol (C3H8O3)
602
270
9.1
None
Inlet
Feed flow total [kg/h] 8.07
Feed flow organics [kg/h] 0.734
Feed flow carbon [kg/h] 0.287
Outlet
Water flow [kg/h] 7.33
Carbon in water [kg/h] not determined
Total gas flow [kg/h] 0.498
Carbon in gas [kg/h] 0.221
Gas produced (GAS 1 + GAS 2)
H2 28.2 vol.% 0.0127 kg/h
CO 33 vol.% 0.2050 kg/h
CH4 14.5 vol.% 0.0516 kg/h
C2H4 0.6 vol.% 0.0034 kg/h
C2H6 6.7 vol.% 0.0454 kg/h
C3H6 0.0 vol.% 2.4E-05 kg/h
C3H8 0.0 vol.% 2.5E-05 kg/h
CO2 17.0 vol.% 0.167 kg/h
Carbon conversion to the gas-phase 0.74 (kg c in gas/kg c in feed)
If outlet temperature from the oven is accepted to be representative for the reaction temperature (ca.
602oC), the obtained 74% carbon conversion is in agreement with results obtained in the batch micro
reactor (see Chapter 3).
In Table 4.2 a more complete mass balance is presented for methanol gasification, including analyses
of the liquid effluent streams. The carbon recovery of this experiment with methanol is above 95%
(see Table 4.2). This result indicates that despite the many shortcomings of this first generation pilot
plant, reasonably good mass balances can be obtained when simple non-charring feedstocks are
processed.
88
Table 4.2. Mass balance of an experiment with methanol. The presented gas composition is
normalized leaving nitrogen out.
Feed
T [oC]
P [bar]
Feed concentration [wt %]
Catalyst
Methanol (CH3OH)
595
270
4.2
Na2CO3, 0.01 wt % added to the wet feed
Inlet
Feed flow [kg/h] 8.11
Feed flow organics [kg/h] 0.341
Feed flow carbon [kg/h] 0.128
Outlet
Water flow [kg/h] 7.70
Carbon in water [kg/h] 0.033
Total gas flow [kg/h] 0.336
Carbon in gas [kg/h] 0.092
Gas produced (GAS 1 + GAS 2)
H2 72.1 vol % 0.0368 kg/h
CO 2.8 vol % 0.0196 kg/h
CH4 0.4 vol % 0.0017 kg/h
C2H4 0.0 vol % 0.0 kg/h
C2H6 0.0 vol % 0.0 kg/h
C3H6 0.0 vol % 0.0 kg/h
C3H8 0.0 vol % 0.0 kg/h
CO2 24.8 vol % 0.2784 kg/h
Carbon conversion to the gas-phase 0.70 (kg c in gas/kg c in feed)
Carbon recovered in the liquid effluent 0.26 (kg c in water effluent/kg c in feed)
Total carbon recovery 0.96 (kg c in products/kg c in feed)
Addition of Na2CO3
During some runs, small amounts of Na2CO3 (0.01 wt %) were added to the feedstock solution. By
comparing the results with and without Na2CO3 (see Figure 4.7) it can be seen that it influences the
product distribution strongly. In fact, it is shown that already by addition of such small amounts of
89
Na+, the water-gas shift equilibrium is approached to a large extent resulting in a high hydrogen yield
and low CO levels. Apart from its affect on the product gas distribution, the results also show that the
carbon conversion increases when adding Na2CO3. The influence of Na2CO3 on the carbon conversion
was not observed in the batch microreactors. This difference between batch capillaries and the pilot
plant results cannot be explained from any existing theory, although it could in principle originate
from interaction of the metal wall with Na2CO3..The difference in the measured carbon conversions
with and without Na+ may well be ascribed by the imperfect control and definition of the reactor
temperature in the pilot plant. From measurements in the capillaries (see Chapter 3) it is known that
20°C change in the reactor temperature represents ca. 10% carbon conversion around 600°C. As stated
before, in the oven of the pilot plant the temperature gradient is 100°C.
0
20
40
60
80
100
H2 CO2 CO CH4 C2H6 -- Na2CO30.0
0.5
1.0
1.5
2.0
yiel
d [m
ole
gas/
mol
e gl
ycer
o] without additives 0.01 wt%, Na2CO3
glycerol, 10wt%, T = 602oC, 280 bar
Xc
[%]
Figure 4.7 Influence of Na2CO3 addition (0.01 wt %) on the product gas yields and carbon
conversion. A 10 wt % glycerol solution has been used as feedstock. T = 602°C, P = 280 bar.
Comparison of Pilot Plant Results with Data from Microreactors
In Figure 4.8 gasification of 10 wt % glycerol in the pilot plant is compared with results obtained in
the batch quartz capillaries (see Chapter 3) and the Inconel micro continuous reactor (see Chapter 5)
for the same process conditions.
90
0
20
40
60
80
100
H2 CO2 CO CH4 C2H4 C2H6 C3H6 C3H8 -- -- --0.0
0.2
0.4
0.6
0.8
1.0
1.2
1.4
1.6
yiel
d [m
ole
gas/
mol
e gl
ycer
ol] Pilot plant (incoloy 825)
micro continuous reactor (inconel 625) micro capillary reactor (quartz)
car
bon
effic
ienc
y [%
]
Figure 4.8 Gasification results for 10 wt % glycerol at 600oC and ± 280 bar. Results from the pilot
plant, the micro continuous reactor (see Chapter 5) and the capillary reactor (see Chapter 3) are
compared.
The highest carbon efficiency is observed in the micro continuous Inconel reactor, which has the
highest catalytic wall area (Inconel contains 60% Ni) due to its small internal diameter of only 1 mm.
The pilot plant reactor has also a catalytically active reactor wall, but a lower specific wall area (i.d. =
12 mm) resulting in a lower carbon conversion than the micro continuous reactor. Complete absence
of catalytic activity in the batch capillary reactors results in the lowest carbon conversion.
Influence of Feed Concentration
To investigate the influence of the concentration of the feedstock on the carbon efficiency and gas
production, gasification experiments with 4.8, 10 and 17wt % of glycerol including 0.01 wt % Na2CO3
were performed. The results are presented in Figure 4.9.
91
0
20
40
60
80
100
H2 CO2 CO CH4 C2H6 4.8% 10% 17%0
1
2
3
yiel
d [m
ole
gas/
mol
e gl
ycer
ol]
4.8 wt% 10 wt% 17 wt%
car
bon
effic
ienc
y [%
]
Figure 4.9. Influence of concentration of the glycerol solution on the product gas yields and the
carbon conversion. T = 600°C, P = 270 bar.
Figure 4.9 shows that, like in the capillary reactors (see Chapter 3), higher feed concentrations (>
10%) cannot be converted completely to the gas-phase. Moreover, the water gas shift reaction is not
reaching equilibrium, especially not at 17 wt % resulting in a significant CO concentration in the
product gas.
92
93
Conclusions and recommendations
The conclusions from this chapter can be summarized as follows:
● A first generation pilot plant for supercritical water gasification with a maximal throughput of
30 kg/h wet feedstock has been used for conversion tests. This plant is based on a shell and
tube heat-exchanger and a tubular reactor and is designed to process non-fouling feedstock
materials (in this case glycerol and methanol).
● Tests with glycerol and methanol as feedstock have shown that the pilot plant can be run
without operational problems. While built for other purposes, the pilot plant is however not
suited for very accurate measurement under well defined conditions like for the reactor
temperature.
● With glycerol several experiments have been performed and the results have been compared
with similar test conducted in the micro capillary reactors (see Chapter 3). It can be concluded
that the trends observed in both systems are identical, viz.:
● Complete conversion of the feedstock to the gas-phase is only possible for very diluted
feedstock solutions (< 5 wt %).
● Addition of Na2CO3 promotes the water gas shift reaction resulting in a high hydrogen yield
and low CO levels.
● On the other hand, in absolute sense there is a clear difference between the results obtained in
the pilot plant and in the capillaries. In the pilot plant, a higher carbon conversion from
feedstock to gas phase products and more H2 production is achieved under similar conditions.
This is ascribed to the catalytic activity of the wall of the Incoloy reactor of the pilot plant. The
pilot plant, however, shows a lower carbon to gas phase conversion in comparison with the
metal continuous micro system presented in Chapter 5 which can be explained by the lower
specific catalytic wall area of the pilot plant reactor compared to the continuous micro reactor
(i.d. pilot plant is 1.38 cm vs i.d. micro continuous is 1 mm)
● Test with starch (a fouling feedstock) showed severe operational problems in the form of
blockages at several locations.
● It has been decided to stop the pilot plant program for the time being. The focus is shifted
towards the development of a micro continuous plant that provides the essential information,
but is much cheaper and allows much faster experimentation (see Chapter 5). In similar micro
systems the research on fluidization under supercritical conditions is continued and this will
hopefully pave the way to efficient SCWG while allowing for fouling feedstocks.
Notation
fc,i Carbon fraction (mass basis) of component i
feedLHV Lower heating value of feed dry and ash free, MJ/m3
igasproductLHV ,− Lower heating value of permanent gas, MJ/m3
M Molar mass, kg/mole
P Reaction pressure, bar
T Reaction time, s
T Reaction temperature, oC
cX Carbon conversion, %
gasproductY − Gas yield
Θ feedMass flow rate of organics in the feed, kg/h
Θ iMass flow rate of component i, kg/h
References
Antal, M.J.; Allen, S.; Lichwa, J.; Schulman, D.; Xu, X., Hydrogen Production from High-Moisture Content Biomass in Supercritical Water. In U.S. DOE Hydrogen Program Review, 1999. Antal, M.J.; Allen, S.G.; Schulman, D.; Xu, X.; Divilio, R.J., Biomass Gasification in Supercritical Water. Ind. Eng. Chem. Res. 2000, 39, 4040. Boukis, N.; Diem, V.; Dinjus E.; Galla U.; Kruse A., Advances with the Process of Biomass Gasification in Supercritical Water. Presented at the 4th International Symposium on High Pressure Technology and Chemical Engineering, Venice: Italy, 2002. Hao, X.H.; Guo, L.J.; Mao, X.; Zhang, X.M.; Chen, X.J., Hydrogen Production from Glucose used as a Model Compound of Biomass Gasified in Supercritical Water. Hydrogen Energy 2003, 28, 55. Holgate, H.R.; Meyer, J.C.; Tester, W.J., Glucose Hydrolysis and Oxidation in Supercritical Water. AIChE J. 1995, 41, 637. Kabyemela, B.M.; Takigawa, M.; Adschiri, T.; Malaluan, R.M.; Arai, K., Mechanism and Kinetics of Cellobiose Decomposition in Sub- and Supercritical Water. Ind. Eng. Chem. Res. 1998, 37, 357. Kabyemela, B.M.; Adschiri, T.; Malaluan, R.M.; Arai, K., Glucose and Fructose Decomposition in Subcritical and Supercritical Water: Detailed Reaction Pathway, Mechanisms, and Kinetics. Ind. Eng. Chem. Res. 1999, 38, 2888. Kruse, A.; Henningsen, T.; Sinag, A.; Pfeiffer, J., Biomass Gasification in Supercritical Water: Influence of the Dry Matter Content and the Formation of Phenols. Ind. Eng. Chem. Res. 2003, 42, 3711. Lee, I.; Kim, M.-S.; Ihm, S.-K., Gasification of Glucose in Supercritical Water. Ind. Eng. Chem. Res. 2002, 41, 1182.
94
95
Manarungson, S.; Mok, W.S.L.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. In Advances in Thermochemical Biomass Conversion; Bridgwater, A.V., Ed.; Blackie Academic and Professional: London, U.K., 1993; p. 1367. Matsumura, Y.; Minowa, T.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; van de Beld, L.; Elliott, D.C.; Neuenschwander, G.G.; Kruse, A.; Antal, M.J., Biomass Gasification in Near- and Supercritical Water: Status and Prospects. Biomass Bioenergy 2005, 29, 269. Modell, M., Gasification and Liquefaction of Forest Products in Supercritical Water. In Fundamentals of Thermochemical Biomass, Overend, R.P.; Milne, T.A.; Mudge, L.K., Eds.; Elsevier Applied Science Publishers Ltd.: London, 1985, p. 95. Potic, B.; van de Beld, L.; Assink, D.; Prins, W.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water. In Proceedings of the 12th European Conference and Exhibition on Biomass for Energy, Industry and Climate Protection; Palz, W.; Spitzer, J.; Maniatis, K.; Kwant, K.; Helm, P.; Grassi, A., Eds.; ETA Florence, WIP Munich: Amsterdam, 2002; p. 777. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Gasification of Model Compounds and Wood in Hot Compressed Water. Submitted to Ind. Eng. Chem. Res., 2005 Schmieder, H. Abeln, J.; Boukis, N.; Dinjus, E.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht, M., Hydrothermal Gasification of Biomass and Organic Wastes. In Proceedings of the 5th Conference on Supercritical Fluids and their Applications, Garda, Italy, 1999; p. 347. Van de Beld, L.; Wagenaar, B.M.; Assink, D.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; Penninger, J.M.L., Biomass and Waste Conversion in Supercritical Water for the Production of Renewable Hydrogen. Presented at the 1st European Hydrogen Energy Conference, Grenoble, France, 2003. Xu, X.; Matsumura, Y.; Stenberg, J.; Antal, M.J., Carbon-Catalyzed Gasification of Organic Feedstocks in Supercritical Water. Ind. Eng. Chem. Res. 1996, 35, 2522. Xu, X.; Antal, M.J. Jr., Gasification of Sewage Sludge and Other Biomass for Hydrogen Production in Supercritical Water. Environmental Progress 1998, 17, 215. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels. 1993, 7, 574.
96
97
o t
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Abstract
This chapter is dealing with SCWG of glycerol and glucose in a continuous reactor. For this
investigation a continuous micro reactor system has been developed. The reactors used are stainless
steel 316 and Inconel capillaries with an internal diameter of 1 mm and a length of 1 meter. One goal
of this investigation is to show the catalytic effect of the reactor construction material and to show the
influence of the process conditions on the carbon conversion and the product gas distribution. In
addition, selected catalysts, namely 3 wt % Ru on TiO2 and charcoal of beech wood, are tested in a
fixed bed micro reactor system. This investigation has revealed that the carbon conversion and
product gas distribution strongly depend on the selected reactor construction material. Moreover, it is
confirmed that pressure has no effect on the carbon conversion and gas yield and that an increase of
concentration of organics in the feed results in a lower carbon conversion and less hydrogen
production. Addition of Ru catalyst provokes cracking/reforming of oxygenated hydrocarbons that
without catalyst would slip trough the reactor and end-up as liquid product. Demonstration of
combustion with air under realistic process conditions of produced carbonaceous deposits from
glucose gasification, represents an important aspect of an industrial SCWG reactor system.
98
Introduction
With a quartz batch capillary technique (Potic et al., 2004) more than 700 experiments have been
performed to investigate the gasification of organic compounds in supercritical water (Chapter 2 and
3). These data form the basis of a complete map of the SCWG operating window for glucose and
glycerin under non-catalytic conditions. However, for catalyst screening and to study possible ways to
recover or to avoid carbonaceous deposits when processing realistic feedstock materials, also tests
under continuous operating conditions are required. For these purposes, several tests have been
performed in a pilot-plant with a nominal throughput of 10 kg wet feedstock per hour (see Chapter 4).
Operation of this pilot-plant is labor intensive and very costly. Besides, the results are not yet accurate
enough to serve as a basis for process development. To enable quick and cheap experimentation under
supercritical conditions a continuous micro tubular reactor (1 mm internal diameter) has been
developed and tested. Both an empty tube and an empty tube in series with a packed bed of catalyst
particles have been used as reactor system. Experiments have been done to investigate the catalytic
effect of the construction material and the influence of the process conditions. Besides, some
preliminary tests have been performed with selected catalysts (3 wt % Ru on TiO2 and charcoal of
beech wood), including burn-off of carbonaceous deposits.
99
Experimental
The continuous micro reactor for SCWG has been developed to allow simple, safe and cheap
experimentation without using a special high-pressure infrastructure. Such a system should be flexible
with easy to replace parts allowing fast cleaning and short experimental runs at steady state. These
demands were met by miniaturization and by elimination of a complex multi-phase pressure reduction
valve.
The continuous SCWG reactor system consisted out of a HPLC piston pump, a tubular reactor, an
electrical oven, a cooler, a pressure vessel, a liquid collection vessel and a sampling bag for gases.
Figure 5.1 is a schematic representation of the experimental set-up. Stainless steel 316 and Inconel
capillaries with an internal diameter of 1 mm and a length of 1 meter were used as tubular reactors (V
= 0.785 mL). Quartz reactors were also tested but these were found to be too fragile for continuous
operation. The set-up could in principle be operated with a catalytically inert reactor like quartz but
this could not yet be technically solved.
The flow rate of the feedstock was controlled accurately with a HPLC pump (LabAlliance). This pump
could deliver flow rates of water in the range of 0.01 to 24 mL per minute at a maximum pressure of
500 bar. The pressure directly after the pump was measured and monitored with the in-built pressure
sensor of the pump. For the rapid cooling of the reactor effluent a cooler was placed just after the
oven. The cold liquid, gases and entrained solids (if any) were collected in a high pressure vessel (V =
30 mL), which was connected to a pressure sensor. Between the cooler and the pressure vessel a
simple valve was placed. This pressure vessel was also used to start-up the system in the following
way. While the valve was closed, the vessel was pressurized with helium and water was pumped into
the reactor and cooler until the pressures up and downstream the valve equalized. Then the valve was
opened and the helium in the vessel maintained the pressure in the system. Hereafter, the experiment
was started by switching from water to the feedstock solution. During a typical experiment the
pressure increased ca. 50 bar. The representative pressure of an experiment was defined as the average
of the start and end pressure of that tests (<p>). We allowed for this pressure increase, because in the
batch capillaries it was observed that the operating pressure did not effect the results in the range of 40
to 400 bar (see Chapter 3).
During operation, the temperature was measured with three thermocouples (tip = 0.5 mm) placed on
the outer wall of the capillary. The temperature difference between these thermocouples never
exceeded more than 3°C during operation. Additionally, the axial temperature profile inside the
reactor was initially measured by a movable thermocouple. By these measurements it was shown that
the first 11 cm of the reactor (total length is 1 meter) were needed to reach the final reactor
temperature for a typical flow rate of 0.04 mL/min. For all employed flow rates the length of the
heating zone was determined. In this way, the reaction time, defined as the time that the fluid was at
the final reactor temperature, could be calculated. For all investigated cases the heating time was about
the same as the reacting time.
After the experiments, the gas products were collected in a sampling bag, which was then connected to
a GC for the gas analysis. On the basis of the readings of the GC, the total number of moles present in
the pressure-vessel (calculated from the known T, P and V of this vessel), and the amount of
feedstock, the mass balance (yields and conversion) was calculated. The carbon conversion and the
product yields, are chosen here as the main process qualifiers. The carbon conversion is defined as
degree of conversion of carbon from biomass to permanent gases:
100,
,,
xN
NX
feedc
igasproductci
c
−∑=
The product yield is defined as:
feed
igasproductigasproduct N
NY ,
,−
− =
For the catalytic tests, the last 30 cm of the capillaries were filled with catalyst. The catalyst zone was
placed between two supports to ensure fixed bed operation. Ru (3 wt %) on TiO2 (dp = 150 micron)
and charcoal (dp = 150 micron) from beech wood were tested as catalysts. In this way, in first 70 cm of
the reactor heating non-catalytic gasification is taking place. The reaction time in this zone is about 45
s , which is, according to earlier results (Chapter 3), enough to reach the maximum carbon conversion
at temperatures of 600oC and higher. In catalytic zone, all primary reaction products (gas, liquid and
solid) are brought in contact with catalyst. The residence time in the catalytic bed was ca. 10 s.
100
Feed tank
TI
TI
TI
HPLC pump
Electrical oven
Reactor
Thermocouples
Pressure vessel
Pressure sensor
Cooler
Collectingvessel
Gas bag
Fixed bed
PI
Figure 5.1. Schematic representation of the experimental set-up.
Burn-off of carbonaceous deposits was done by leading pressurized air through the system. At the start
of these experiments, the system was at atmospheric pressure and the air-cylinder at 180 bar. The
burn-off tests were ended when the pressure in the air-cylinder and the pressure vessel equalized (ca.
40 bar).
101
102
Validation of the Technique
Mass balance closure was checked by experiments with formic acid, which according to the literature
(Yu and Savage, 1998) should be converted completely to gas-phase components at temperatures
above 650oC and low feed concentrations. Several runs were performed at 700oC and feedstock
concentrations below 2.5 wt %. The observed carbon conversion (fraction of carbon in the feedstock
converted to gas-phase components) varied between 98 and 102%.
As an example, Table 5.1 lists the results of two identical experiments. These results are typical for the
obtained reproducibility. It can be concluded that all produced gases can be recovered with the method
used and the reproducibility is satisfactorily. In the tables and figures of this Chapter individual data
points are reported, although always a duplo was performed. The relative deviation of the duplos was
never higher than 5%.
Table 5.1. Result of two identical tests with formic acid in the stainless steel continuous flow reactor.
Feedstock: 1.3 wt % formic acid in water, T = 700oC, <p> = 300 bar, and 60 s reaction time.
Formic acid 1.3 wt %
Run 1 2
T [oC] 700 700
P [bar] 300 300
t [s] 60 60
Gas composition [%]
H2 55.92 55.20
CO2 42.94 43.67
CO 1.14 1.13
Carbon conversion [%] 98.7 101.1
Results
Catalytic Effect of the Reactor Wall
The material of construction of the SCWG reactor is suspected to influence the conversion results in
case reactors with a high specific wall area are compared. By adding Inconel powder to the quartz
capillaries, the effect of an Inconel reactor wall was mimicked. In these experiments the catalytic
activity of Inconel was clearly observed (see Chapter 3). In well-defined experiments in the micro
continuous flow reactor the influence of the reactor wall material on the SCWG process has been
confirmed and quantified. Using Inconel instead of stainless steel 316 as construction material resulted
in an increase of the carbon conversion from 62 to 82% when processing a 10 wt % glycerol solution
at 600oC. Also the product gas composition is affected significantly; SS316 seems to promote the
water-gas shift reaction (low CO yield), Inconel not (see Figure 5.2). These results on the catalytic
influence of the reactor wall are inline with earlier observations of Yu et al. (Yu et al., 1993), who
reported results of gasification of glucose in the same experimental set-up under identical conditions,
but using other metals for the reactor. They found large differences between experiments carried out in
Inconel, Hastelloy and corroded Hastelloy reactors.
0
20
40
60
80
100
H2 CO2 CO CH4 C2H6 Inconel SS --0.0
0.4
0.8
1.2
1.6
gas
yiel
d [m
ole
gas/
mol
e gl
ycer
ol] Inconel 625
SS 316
car
bon
conv
ersi
on [%
]
Figure 5.2 Influence of the reactor wall material on the gas production and carbon conversion for 10
wt % glycerol, T = 600 C, <p> = 300 bar and 60 s reaction time.
103
Effects of the Process Conditions
Concentration organics in the feed solution
From the capillary tests it has become clear that for non-catalytic SCWG the concentration of organics
in the feedstock affects the carbon conversion to a large extent (see Chapter 3). At 650oC, a high
carbon conversion (> 98%) turned out to be possible only with low concentrated feedstocks (< 2 wt
%). Higher concentrations resulted in an incomplete carbon conversion. In the Inconel micro
continuous flow reactor, the observed carbon conversions are higher than in the quartz capillaries for
identical conditions, probably due to the catalytic activity of the wall. However, the trend with respect
to the influence of the concentration on the feedstock is also observed in the continuous Inconel
reactor (see Figure 5.3). At 600°C glycerol solutions of up to 5 wt % can be gasified nearly
completely, while for a 10 wt % solution the carbon conversion is limited to 82%.
H2 CO2 CO CH4 C2H6 Cefficiency --0
20
40
60
80
100
gas
com
posi
tion
[%] 1 wt%
5 wt% 10 wt%
glycerol 600 oC, 300 bar, 60s
car
bon
conv
ersi
on [%
]
Figure 5.3. Influence of the glycerol concentration on the gas distribution and carbon conversion.
Inconel reactor, T = 600 oC, P = 300 bar and 60 s reaction time.
Process pressure
Although it has been suggested that supercritical water properties would be beneficial for the
gasification reaction, experimental results from the capillary reactors (see Chapter 3) showed that the
pressure has no influence on the carbon efficiency and product distribution. In the SS 316 micro
continuous flow reactor these tests have been repeated. Again, it can be concluded that the pressure
has no significant effect on the carbon conversion and product distribution. Figure 5.4 shows the two
extremes of the investigated pressure range (100 to 300 bar) for 10 wt % glycerol at 700oC and 60 s
reaction time.
104
H2 CO2 CO CH4 C2H6 Cefficiency0
20
40
60
80
100
gas
com
posi
tion
[%]
carb
on c
onve
rsio
n [%
]
100 bar 300 bar
glycerol 10 wt%, 700oC, 60s
Figure 5.4 Influence of the pressure on the product gas composition and carbon conversion. 10 wt %
glycerol, stainless steel 316 reactor, T = 700oC and 60 s reaction time.
Catalytic fixed bed reactor
These tests were performed with glycerol as feedstock. When glycerol is not completely gasified, only
liquids are formed as by-products never carbonaceous deposits (Buhler et al., 2002). This makes
interpretation of the results easier.
The potential and necessity of heterogeneous catalysis for the SCWG process has been demonstrated
already with the batch capillaries (Chapter 3). However, the undefined contacting efficiency of the
catalyst particles in the capillary made the experiments difficult to interpret and to compare with the
limited available results form literature. In this chapter well-defined catalytic experiments have been
done in a fixed bed reactor with a length of 30 cm and an internal diameter of 1 mm. Charcoal of
beech wood (produced at 600°C) and 3 wt % Ru on TiO2 were tested as catalysts. A 10 wt % glycerol
solution was used as feedstock. In these tests no carbonaceous deposits products (not soluble in
acetone) were observed in the collecting vessel and the reactor after the experiments. When passing
the (intermediate) reaction products over a charcoal bed with a residence time of ca. 10 s, the carbon
conversion increases slightly viz. from 65 to 74% (see Figure 5.5). Antal and coworkers (Xu et al.,
1996) observed complete carbon conversion by leading the effluent of their tubular reactor over a bed
of coconut shell activated carbon at 600°C and 15 s residence time. The results are difficult to compare
because different types of charcoal were used under different process conditions (residence time in
catalyst bed, location of catalyst bed, temperature profile in the reactor).
Using the Ru catalyst, the carbon conversion increased to nearly 100%. Another interesting
observation is that with this cat C2-compounds are not found in the product gas, while the yield of CO,
CO2, CH4 and H2 increases. These results indicate that the Ru catalyst cracks/reforms oxygenated
105
hydrocarbons, including those being liquids at ambient conditions, to smaller molecules such as CO,
CO2, CH4 and H2.
0
20
40
60
80
100
H2 CO2 CO CH4 C2H6 10% 10%+C10%+Ru0.0
0.5
1.0
1.5
2.0
2.5
carb
on c
onve
rsio
n [%
]
yiel
d [m
ole
gas/
mol
e gl
ucos
e]
glycerol 10 wt.% glycerol 10 wt.%+ char glycerol 10 wt.% + Ru/TiO2
Figure 5.5 Influence of different catalysts on the gas production and the carbon conversion for 10 wt
% glycerol, SS 316 reactor, T = 600oC, P = 300 bar and 60 s reaction time.
Burn-off of carbonaceous deposits
In a real system, using waste streams as feedstock, the formation of fouling in the form of
carbonaceous deposits is inevitable. To study the SCWG process under such conditions glucose was
used as feedstock, because it is know to produce carbonaceous deposits. Combustion of these deposits
is proposed as a feasible remedy in an industrial reactor (see Chapter 7). In this way, a by-product is
removed from the system and heat is generated to drive the endothermic process, provided a suitable
reactor/process design is made.
At 600°C a 10 wt % glucose solution was processed without catalyst. The carbon conversion during
this test was ca. 63% (see Table 5.2, first column). After passing air through the system an additional
15% of carbon was recovered in the form of CO2 and tracer of CO (see Table 5.2, second column).
However, in an identical experiment but with the Ru catalyst bed included, 100% of the carbon in the
feedstock is recovered. 85% is directly gasified and 15% is recovered after burn-off (see Table 5.2,
third and fourth column). This result indicates that carbonaceous deposits do not slip trough the
reactor, but apparently remain on the reactor wall and the catalyst particles. The increase in carbon
conversion when using the Ru catalysts again points in the direction that this catalyst converts
oxygenated hydrocarbons to permanent gases. Without Ru catalyst these oxygenated hydrocarbons
would have slipped through the system and would be recovered as liquid products after the
experiment.
106
Table 5.2 Results of gasification and gasification and combustion process as for 10 wt % glucose at
600oC, ± 280 bar and ± 60 s reaction time
Glucose 10 wt % Glucose 10 wt % + Ru/TiO3
gasification gasification +
combustion
gasification gasification +
combustion
C conv.[%],
C-recovery
63.3
77.4
85.2
103.1
Gas composition
H2 26.6 41.2
CO2 39.9 42.5
CO 14.0 1.1
CH4 12.8 15.1
C2H4 0.8 0.0
C2H6 3.7 0.0
C3H6 1.0 0.0
C3H8 1.1 0.0
In the capillary tests, complete gasification of 10 wt % glucose was achived at 600oC when adding the
Ruthenium catalyst. The 15% difference in conversion between the batch and continuous tests may be
explained by:
● Deactivation of the catalyst in the continuous reactor. In the batch tests always fresh catalyst
was used and the test only lasted for 60 s, whereas the continuous reactor was operated for two
hours without regeneration of the catalyst. The life time of the catalyst was not investigated by
the developers (Elliott et al., 1997) under SCWG condition. However, it was reported that at
the designed conditions (24 MPa, 360oC) the catalyst showed a tendency towards de-activation
(Elliott et al., 2004).
● The presence of carbonaceous deposits in the inert section before the catalytic section. When
these deposites are not entrained with the main flow through the catalytic bed, their catalytic
conversion is obviously not possible.
107
108
Conclusion
A simple continuous micro-scale SCWG reactor system has been designed and demonstrated for fast
and cheap experimentation. It avoids the use of a cumbersome pressure relief valve and is suitable for
empty tube, packed bed and potentially fluid bed reactors. The conclusions from this work can be
summarized as follows:
● The catalytic activity of metal reactor construction materials has been confirmed in the empty
tube continuous flow reactor. To different materials, viz. Inconel and stainless steal showed
considerable differences in carbon conversion and product distribution under identical process
conditions.
● Despite the catalytic influence of the wall, all trends found in the inert quartz batch capillaries
were observed also in the metal continuous flow reactors. The most important ones being that
the pressure does not affect the product distribution and carbon conversion and that increasing
the concentration of the feedstock results in a lower carbon conversion and less hydrogen
production.
● The tested Ru catalyst increases the carbon conversion considerable. There is evidence that this
catalyst cracks/reforms organic components that without catalyst would slip through the reactor
and end-up as liquid product.
● Glucose produces carbonaceous by-products in the SCWG process. These deposits can be
combusted with air under realistic process conditions, which may be an important aspect of an
industrial SCWG reactor system.
Notation
N c, product-gas,i Number of moles C in product gas i produced, mol
N c, feed Number of moles C in feedstock fed, mol
N product-gas,i Number of moles product gas i component produced, mol
N feed Number of moles of feedstock fed, mol
Xc Carbon conversion, %
Yproduct-gas,i Yield of product gas component i, mol produced gas/mol feed
109
References
Buhler, W.; Dinjus, E.; Ederer, H.J.; Kruse, A.; Mas, C., Ionic Reactions and Pyrolysis of Glycerol as Competing Reaction Pathways in Near- and Supercritical Water. J. Supercrit. Fluids 2002, 22, 37. Elliott, D.C.; Sealock, L.J. Jr.; Baker, E.G., Method for the Catalytic Conversion of Organic Materials into a Product Gas. U.S. Patent 5616154, 1997. Elliott, D.C.; Neuenschwander, G.G.; Hart, T.R.; Butner, R.S., Low-Temperature Catalytic Gasification of Wet Biomass Residues. In Proceedings of the 2nd World Biomass Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 746. Potic, B.; Kersten, S.R.A.; Prins, W.; Assink, D.; van de Beld, L.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water: Results of Micro and Pilot Scale Experiments, In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection, van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 742. Potic, B.; Kersten, S. R.A.; Prins, W.; van Swaaij, W.P.M. A High-throughput Screening Technique for Conversion in Hot Compressed Water. Ind. Eng. Chem. Res. 2004, 43, 4580. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Gasification of Model Compounds and Wood in Hot Compressed Water. Submitted to Ind. Eng. Chem. Res., 2005. Xu, X.; Matsumura, Y.; Stenberg, J.; Antal, M.J., Carbon-Catalyzed Gasification of Organic Feedstocks in Supercritical Water. Ind. Eng. Chem. Res. 1996, 35, 2522. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels 1993, 7, 574. Yu, J.; Savage, P.E., Decomposition of Formic Acid under Hydrothermal Conditions. Ind. Eng. Chem. Res. 1998, 37, 2.
110
111
at h
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Abstract
In this chapter the concept of micro-fluidized beds is introduced. A cylindrical quartz reactor with an
internal diameter of only 1 mm is used for process conditions up to 500oC and 244 bar. In this way,
fast, safe, and inherently cheap experimentation is provided. The process that prompted the present
work on miniaturization is gasification of biomass and waste streams in hot compressed water
(SCWG). Therefore, water is used as fluidizing agent. Properties of the micro-fluid bed such as the
minimum fluidization velocity (Umf), the minimum bubbling velocity (Umb), bed expansion, and
identification of the fluidization regime are investigated by visual inspection. It is shown that the
micro-fluid bed requires a minimum of twelve particles per reactor diameter in order to mimic
homogeneous fluidization at large scale. It is not possible to create bubbling fluidization in the
cylindrical micro-fluid beds used. Instead, slugging fluidization is observed for aggregative conditions.
Conical shaped micro-reactors are proposed for improved simulation of the bubbling regime.
Measured values of Umf and Umb are compared with predictions of dedicated 2D and 3D discrete
particle models (DPM) and (semi)-empirical relations. The agreement between the measurements and
the model predictions is good and the model supports the concept and development of micro-fluid
beds.
This chapter has been published in Chemical Engineering Science, 60 (2005), 5982-5990.
112
Introduction Miniaturization of chemical reactors is receiving increased interest (Kolb and Hessel, 2004; Maharrey
and Miller, 2001). Micro-reactors are specifically suitable for high throughput screening, on-demand
chemical synthesis, and carrying out reactions under conditions that would normally lead to unsafe
operation. In general, miniaturization leads to lower capital and operational costs, less energy
consumption, improved safety, and less pollution. In this chapter, the miniaturization principle is
applied to fluid bed research. At present, laboratory (10 - 100 mL reactor volume) and pilot-scale (0.1
- 1000 l reactor volume) research and development concerning fluidised bed systems is expensive and
labour intensive. In fact, it is anticipated here that this type of research will become too expensive in
the future, especially if it involves high temperature and pressure. The objective of this work is,
therefore, to develop a new experimental method based on micro-fluid bed reactors with an internal
diameter of only a few millimetres. Possible applications of micro-fluid beds are high-throughput
screening of catalysts, establishing the operating window of a certain process, or study of specific
hydrodynamic phenomena. Investigating chemical fluid bed processes in micro-fluid beds, for
instance, is preferable over testing in high-throughput fixed beds. Translating fixed bed data to fluid
bed operation is not straightforward, particularly when adsorption/desorption of reactants and
products, decompositions on the catalyst, and fast-aging catalyst are involved. Under these conditions,
fixed beds suffer from the fact that strong spatial gradients in the composition of the fluid phase and
the catalyst properties will make interpretation of the data troublesome.
The typical characteristics and conditions of a large-scale fluid bed, being 5 to 15 s gas residence time,
2 - 3 mixing units, and 1 - 5 mass transfer units (De Vries et al., 1972; Van Swaaij and Zuiderweg,
1973), should be mimicked in micro-systems. It is, however, not necessary that exact mapping is
achieved at every level. The similarity should be based on the overall (chemical) performance of the
process under consideration. The micro-fluid bed can have every desired shape, ranging from a
uniform cylindrical tube to a conical bed and a staged reactor configuration (see Figure 6.1). In
cylindrical micro beds of only several millimeters, it will be difficult to create stable bubbling
fluidization, because slug formation will occur immediately after entering the aggregative regime.
Conical micro-fluid beds are considered to suppress slugging and to simulate a bubbling bed.
Somewhat larger conical fluid beds (1 - 4 cm), showing intense back mixing of both the solid and the
fluid phase, have been discussed by Kersten et al. (Kersten et al., 2003). Conical fluid beds can also be
applied if more fluid phases are to be processed together.
The process that prompted the present work on miniaturization is gasification of biomass and waste
streams in hot compressed water (SCWG). SCWG is foreseen to require fluid bed technology operated
in severe conditions (Kersten et al., 2004). Gasification in hot compressed water is a novel process for
the conversion of wet biomass and waste streams (> 80 wt % moisture) into hydrogen or methane rich
gas. The operating pressure is typically above 200 bar and temperatures in the range of 300 - 650°C
are considered (Antal et al., 2000; Elliott et al., 1994; Hao et al., 2003; Kruse, 1999; Potic et al., 2004;
Potic et al., 2004; Schmieder et al., 2000; Yu et al., 1993). The SCWG process is at an early stage of
development. In the laboratory and pilot plant work carried out so far, next to batch autoclave reactors,
tubular flow reactors have been used in connection with a shell and tube heat-exchanger. Not much
information has been published on the practical design of a reactor for a commercial plant. Simple
non-fouling model compounds have been used frequently as feedstock. However, thermal composition
of real biomass (e.g. verge grass paste) results in carbon formation and ash deposition (Yoshida et al.,
2004). This implies that catalysis is required to achieve complete conversion to gases and to steer the
product distribution (Antal et al., 2000; Potic et al., 2005). To cope with such fouling conditions in
combination with catalysis an integrated reactor/heat-exchanger concept, based on fluid beds and
circulating solids, has been proposed (Kersten et al., 2004). For quick, safe and cheap experimentation
with such a system the micro-fluid bed technique has been developed. Sufficient mass balance closure
of experiments with a chemical reaction using such a small system has been established already with a
comparable batch method (Potic et al., 2004).
Available data on fluidization with high-pressure CO2 cover a broad range of fluid densities and
viscosities (Liu et al., 1996; Marzocchella and Salatino, 2000) and can be used, as a starting point, for
the investigation of fluidization with hot compressed water. In fact, using CO2 of much less severe
conditions can approximate the density and viscosity of water used for the gasification process. To
study the chemical process it is, however, essential that the actual reaction medium, viz., hot
compressed water, can be used in the experimental method.
Cylindrical micro two phase reactor
Conical micro two phase reactor
Conical three phase reactor
Figure.6.1 Different microreactors in the fluid bed program.
In the present contribution, cylindrical micro-fluid beds are discussed. It is investigated whether or not
the classical design equations for Umf, Umb, Ut and bed expansion can be used for the micro-system.
Fluidization of A and B powders with water in the range of 1 - 244 bar and 20 - 500°C is reported.
Also the differences and similarities between gas, liquid, and supercritical fluids in fluidization are
shown. The experimental results are compared with both literature correlations and the predictions of a
state of the art 3D Euler-Lagrange (DPM) model (Ye et al., 2004, 2005). 113
114
Experimental
The central part of the micro reactor system presented here is a fluid bed with an internal diameter of
only a few millimetres (say one to three millimetres). For testing under high pressure and temperature,
all kinds of special alloys (e.g. Inconel) and quartz can be used as construction material of the fluid
bed. In case of quartz, visual inspection of the fluid bed reactor under process conditions is possible
even under extreme circumstances. Due to their small size, the quartz capillaries used are strong
enough to withstand extremely high pressures (300 bar) and high temperatures (900oC). Other
advantages of the small dimensions are intensive heat transfer from the surroundings (oven) to the
reactor, the very small amount of catalyst needed, low costs, and the fact that the system is inherently
save.
The experimental system used for this investigation on fluidisation consisted basically out of a HPLC
piston pump, a quartz capillary used as fluid bed, an electrical oven, and a pressure vessel. Figure 6.2
is a schematic representation of the experimental set-up. The flow rate of water was controlled
accurately with a HPLC pump (LabAlliance). This pump could deliver flow rates of water in the range
of 0.01 - 24 mL/min at a maximum pressure of 500 bar. A quartz capillary of 1 mm ID, 6 mm OD and
a length of 1 m was used as fluid bed. The actual settled height of the bed was 20 mm. These
dimensions lead to an aspect ratio of the bed of 20, which is high compared to industrial practice.
Twenty millimetres fixed bed height was selected in order to keep the error on the measurement of the
fluid bed level below 10%. This was necessary to obtain accurate estimates of Umf and Umb from the
expansion curve. In case of homogeneous fluidization, it is known that the aspect ratio does not
influence the fluidization behaviour to a large extent. For bubbling fluidization, the aspect ratio is an
important parameter. However, the overall heat and mass transfer rates of beds with high and low
aspect ratio may approach each other by tuning of the operating conditions.
Some 30 cm of length before and after the bed were required to account for heating the incoming
water and to prevent back-mixing effects of cold water from the down stream side.
Water tank
PI
TI
TI
TI
Section“A”
HPLC pump
Electrical oven
Glass window
PressurizedvesselReactor
Pressure sensor
Thermocouple
Section“A” Window
Distributor
Thermocouple
Reactor
Figure 6.2 Scheme of the experimental set-up.
Application of a conventional gas distributor turned out not to be possible in such a small system. To
solve this problem, a quartz ball (~ 0.9 mm) that just fitted inside the 1mm i.d. capillary placed on a
thermocouple was used as an alternative to a regular distributor plate. Raining of bed particles through
the gap between the inner wall and the ball distributor was never observed. The quartz capillary set up
was heated in an electrical oven. A glass window was placed in this oven at the location where the
fluid bed was positioned. This allowed for visual inspection of the fluid bed and photography. A ruler
was placed behind the fluid bed for measurement of the bed height. Connections of the quartz
capillary to the steal tubing, connecting it to the pump and the pressure vessel, were placed outside the
oven. The capillary was connected to the pressure vessel via a one-way valve in order to prevent that
the content of the pressure vessel could flow back into the reactor. Nitrogen was used to pressurize this
vessel. Upstream the pressure vessel, the system was pressurized by pumping water, with the valve
closed, into the quartz capillary and tubing until the pressure prevailing inside the pressure vessel was
reached. Then the valve was opened and the nitrogen in the vessel maintained the pressure in the
system. Due to the much higher volume of the vessel compared to the capillary (500:1), the pressure 115
remained almost constant during operation. By employing this semi-batch method of operation, the
need for a complex two-phase pressure-release valve was omitted and large samples could be collected
at nearly constant conditions. Monitoring of the pressure was done with the built-in sensor of the
HPLC pump and a pressure sensor placed at the entrance of the pressure vessel. During operation, the
temperature was measured with three thermocouples (tip = 0.5 mm) placed on the outer wall of the
capillary and one thermocouple (tip = 0.5 mm) inside the reactor. The latter was the one used as
support for the quartz ball serving as a distributor. One of the three thermocouples on the outer wall
was placed at the location of the fluid bed. The temperature difference between this thermocouple and
the one inside the bed never exceeded more than 3°C during fluidization and their average value was
defined as the bed temperature. Additionally, the axial temperature profile inside the quartz tube was
measured, in the absence of the fluid bed, by a movable thermocouple inserted via the top. By these
measurements it was shown that up to 20 cm downstream the ball distributor the temperature inside
the capillary was constant for all investigated flow rates.
Sand with a density of 2450 kg/m3 was as bed material. Four different particle sizes were used: 60 - 70
µm, 80 - 90 µm, 100 - 150 µm, and 150 - 250 µm.
To investigate the effect of the reactor diameter the 1 mm diameter quartz tube was replaced by tubes
of 12 and 26 mm.
Measurement Procedure
A digital camera was used to take snapshots of the fluidization behaviour. All derived numbers in this
work were based on visual observations.
In case of homogeneous fluidization (investigated range ρf = 500 - 1200 kg/m3), Umf, Ut, and n were
determined from the bed expansion curve. The superficial velocity was fixed by the flow rate of the
HPLC pump and the prevailing temperature and pressure in the fluid bed. The minimum velocity
increment corresponded to an increase of the flow rate of liquid water by 0.01 mL/min. The bed
expansion curve was constructed by reading the height of the bed at different velocities. The error on
reading the bed level was 1 - 2 mm. By measurement of the amount of bed material and the fixed bed
height, the initial (fixed bed) porosity was determined. The voidage and the superficial velocity were
correlated according to the Richardson and Zaki equation:
n
tUU ε= (1)
Figure 6.3 shows a typical measured expansion graph and the derived Richardson and Zaki curve. Umf
was determined by extrapolating the expansion curve to the fixed bed height. Extrapolation was
necessary, because in all cases, Umf corresponded to flow rates below the lower limit of the HPLC
pump. Due to the need for extrapolation and the error on reading the bed height, the statistical error on
116
Umf was around 50%. Extrapolation was not required for determining Ut and n, and consequently the
related error was much lower, namely ca. 10%.
Bed expansion was measured visually also in case of fluidization with water vapor (investigated range
ρf = 16 – 230 kg/m3). For superficial velocities higher than 12 mm/s aggregative fluidization, in the
form of slugs, was observed in this regime. In case of slugging, the height of the expanded bed was
more difficult to determine. A time-averaged bed height taken over a time period of 1 min was then
used.
0 5 10 15 20 250.0
0.5
1.0
1.5
(H-H
0)/H0
[-]
U [mm/s] -0.6 -0.5 -0.4 -0.3 -0.2
1.0
1.5
2.0
2.5
3.0
ln (U
)
ln(ε) (a) (b)
Figure 6.3 Example of the liquid fluidization: T = 320oC, P = 160 bar, ρ = 682 kg/m3: a) Expansion
curve. b) Plot of the Richardson-Zaki equation constructed expansion curve.
Due to the limitation of the pump with respect to minimum achievable flow rate (0.01 mL/min),
determination of Umf was principally not possible for densities below 100 kg/m3. The minimal flow
rate of 0.01 mL/min corresponded to a linear velocity in the range of 2 - 14 mm/s for densities below
100 kg/m3, whereas according the Abrahamsen and Geldart relation, Umf should range between 2 and 3
mm/s for these conditions. For the density range of 100 - 230 kg/m3 it turned out that Umf could also
not be determined. The reason for this was that the observed bed expansion between the fixed bed
state and the slugging velocity was only 2 - 3 mm, which was too little to serve as basis for an accurate
estimate of Umf. The minimum bubbling velocity was defined as the average value of the velocity at
which the first slugs were observed and the velocity before that one.
Discrete Particle Model
The micro-fluid bed was modeled by the volume-averaged Navier-Stokes equations (Kuipers et al.,
1992) combined with the Newtonian equations of motion for a single particle taking into account
possible collisions with other particles and the confining walls. In this study, the contact force between
two particles was calculated by use of the soft-sphere model (Cundall and Strack, 1979). The details
117
can be found elsewhere (Ye et al., 2004, 2005). Compared to the code recently described (Ye et al.,
2004, 2005), one additional force, namely the added mass force, was added to allow simulation of
systems involving high-density fluidizing agents. The added mass force is essential if the fluid is a
dense gas or a liquid. There are different expressions for the added mass forces in the literature (Auton
et al., 1988; Chang and Maxey, 1995; Clift et al., 1978; Maxey and Riley, 1983). In this study, we
followed the derivation of Maxey and Riley, where the added mass force was given by:
DtDVCF pfmm
)uv( −−= ρνν (2)
with:
∇⋅+∂∂
= utDt
D (3)
In most previous studies the added mass coefficient was simply assumed to equal 0.5, being the
theoretical value for a sphere moving in an unbounded fluidum. This value, however, is in addition
only valid for very dilute systems where the influences of neighbouring are not important. In this
research, the added mass coefficient was taken from the Zuber equation for the added mass force of a
particle in a homogeneous dispersion in an non-rotational flow (Batchelor, 1988), given by:
vmC
3 22vmC εε
−= (4)
Together with the added mass force, there also exists a history force, which is also known as memory
force or Basset force. It takes into account the vorticity diffusion in the surrounding fluid, and the
disturbance effect caused by the acceleration of the sphere. The calculation of this force is complicated
since it involves the integral of the derivative of fluid velocity and particle velocity. In all simulations
reported in this paper the Basset history force was neglected.
In the presented simulations, the geometries of the system were taken as either a box with a size of 1 ×
1 × 40 mm3 in 3D or a rectangle with a size of 1.5 × 50 mm2 in 2D. The choice of such simplified
geometries was made to enable the DPM simulations. The parameters used in the simulations are
summarized in Table 1. The water properties were taken from the database of the National Institute of
Standards and Technology (NIST) in the form of polynomials.
118
Table 6.1 Parameters used in the DPM simulations.
Parameters Value
Particle number 36000 (3D) and 3440 (2D)
Particle diameter 85 µm (3D) and 80 µm (2D)
Particle density 2495 kg/m3 (3D) and 2600 kg/m3 (2D)
Initial bed height 2.00 cm (3D) and 1.36 cm (2D)
CFD time step 2.0×10-6 s
Number of CFD cells 4×4×100 (3D) and 4×100 (2D)
Coefficient of restitution 0.9 (3D) and 0.95 (2D)
Coefficient of friction 0.2 (3D) and 0.1 (2D)
Fluid temperature 873.16 K; 773.16 K; and 673.16 K
Fluid pressure Varying from 0.6 to 28.4 MPa
119
Results
Visual Presentation
In Figure 4 snapshots of the micro-fluid bed in operation are presented. Figure 6.4a shows a typical
series of homogeneous fluidization experiments with liquid water in which the velocity is increased.
The length of the expanded bed for different velocities is used for the derivation of the expansion
curve. In the case of fluidization with water vapor slugging is observed. The slugs can be clearly seen
in the capillaries (see Figure 4b). The bed level is fluctuating under slugging conditions (see Figure
6.4b).
(a) (b)
Figure 6.4. Photographs of (a) homogeneous fluidization at different flow rates and, (b) slugging
fluidization, same bed at different time.
Minimum Number of Particles per Reactor Diameter
It is expected that there should be minimum number of particles present per reactor diameter for a
micro-fluid bed to mimic a large-scale bed under otherwise identical conditions. To investigate this,
glass balls of 60 - 200 µm diameter have been fluidized with water at ambient conditions in three glass
tubes with different internal diameters, viz. 26 mm, 12 mm and, 1mm. The homogeneous expansion
curves of these fluid beds have been recorded while using sand particles of different size.
Unfortunately, the data cannot be compared over the full velocity range. Due to limitations of the
pump with respect to its minimum achievable flow rate, the lowest velocities in the 12 mm and 26 mm
bed could not be reached in the 1mm capillary.
120
Figure 6.5a shows the bed expansion and the voidage versus the velocity for the 1 mm and the 12 mm
ID bed while using particles of 60 - 70 µm diameter. It can be seen that the curves obtained in the 1
mm and the 12 mm bed are nearly identical. Also for the particles of 80 - 90 µm, there is a good
agreement between the expansion curves obtained in the three beds of different sizes, although the
agreement seems less satisfactory than for the 60 – 70 µm particles. In case of particles of 100 - 150
µm, expansion in the 1 mm bed differs significantly from expansion in the 12 mm bed. This deviation
is ascribed to wall effects. Finally, when using particles larger than 150 µm, it has been observed that
the particles appear to move irregularly in clusters in stick-slip flow and a fluidized state cannot be
clearly recognized anymore. These results lead us to the conclusion that micro-fluid beds need 12 or
more particles per diameter in order to resemble a large-scale fluid bed in case of homogeneous
fluidization.
-10 -9 -8 -7 -6 -5
-2
-1
0
1
-10 -9 -8 -7 -6 -5
-2
-1
0
1particle size 60-70micron
ln(e
), ln
(H-H
0)/H
0
ln(U)-5-10 -9 -8 -7 -6
-2
-1
0
1
-10 -9 -8 -7 -6 -5
-2
-1
0
1
-10 -9 -8 -7 -6 -5
-2
-1
0
1
-10 -9 -8 -7 -6 -5
-2
-1
0
1particle size 80-90micron
ln(e
), ln
(H-H
0)/H
0
ln(U)
(a) (b)
-10 -9 -8 -7 -6 -5
-2
-1
0
1
-10 -9 -8 -7 -6 -5
-2
-1
0
1 1mm voidage 1mm expansion
particle size 100-150micron
ln(e
), ln
(H-H
0)/H
0
ln(U)
o 26mm expansion • 26mm voidage
12mm voidage 12mm expansion
(c)
Figure 6.5 Expansion and voidage curve for the different i.d. reactors with particles of: (a) 60 - 70
micron, (b) 80 - 90 micron and (c) 100 - 150 micron. Homogeneous systems.
121
Fluidization with Water: ρf = 500 to 1020 kg/m3
Measurements have been done in the range of 20 - 360°C and 1 - 221 bar. In this density range, the
observed fluidization is always particulate. With increasing velocity, homogeneous expansion of the
bed is observed and bubbles (slugs) are never present. These observations are in agreement with the
discrimination number (Dn) criterion of Liu et al. (Liu et al., 1996) which states that for Dn < 104,
fluidization is homogeneous. For the conditions presented here, Dn is always smaller than 3500.
In the considered range, the viscosity of water at a certain density hardly depends on the prevailing
temperature and pressure. This allows constructing a characteristic velocity map with the fluid density
on the x-axis. Such a plot for particles with dp = 85 µm and ρp = 2450 kg/m3 is presented in Figure 6.
As mentioned before, the measurement error for Umf is rather large. Nevertheless, it can be concluded
that the measured Umf values are in reasonable good agreement with predictions, obtained on basis of
the Ergun equation (Ergun, 1952) provided that the measured fixed bed porosity of 0.5 is used in the
calculations. The sharp increase of one decade in Umf when going down in density from 1020 to 900
kg/m3 as predicted by theory is also observed experimentally. Values of Ut derived from the expansion
graphs are close, although systematically higher, to theoretical values calculated according to Haider
and Levenspiel (Haider and Levenspiel, 1989).
Correlations for Umf and Ut are based on measurements in larger fluid beds ranging from several
centimetres to metres. Because the measurements obtained in the 1 mm fluid bed are reasonably well
described by these correlations, it can be concluded that large-scale fluidized operated in the
homogeneous regime can be mimicked in micro-fluid beds. However, for this to hold the micro-fluid
bed should to satisfy the condition of D/dp > 12 (see above).
500 600 700 800 900 1000
1E-4
1E-3
0.01
0.1
Ergun (ε0=0.50)
U [m
/s]
fluid density [kg/m3]
Umf
Ut
Haider & Levenspiel
Figure 6.6. Umf and Ut measured for water in the 1 mm i.d. FB versus ρf. ρf = 500 - 1050 kg/m3, ρp =
2450 kg/m3, dp = 85 µm. Measurements as well as correlations are plotted. 122
123
Fluidization with Water Vapor: ρf = 16 to 230 kg/m3
Measurements have been done as well in the range of 345 - 485°C and 50 - 245 bar. In the density
region below 90 kg/m3, only slugging is observed. This region corresponds to discrimination numbers
exceeding 4 x 104, which is indicative for the aggregative fluidization regime. For higher densities,
two regimes are observed. From zero velocity up to approximately 2xUmb, fast slugging is observed.
Upon increasing the velocity, the slugs become smaller and their frequency increases. Eventually, this
leads to fluidization in which the slugs cannot be observed anymore and the bed level becomes
constant. For instance, for water vapor of 113 kg/m3 such a regime is observed for velocities higher
than 26 mm/s. This phenomenon has been described also by Li on the basis of simulations with a
Euler-Lagrange CFD model (Li, 2003). The corresponding Dn values for ρ ≥ 100 kg/m3 and U ≥
2xUmf range from 0.5 x 104 up to 1.8 x 104 describing homogeneous and transitional regimes.
It is not possible to create bubbling fluidization in the cylindrical micro-fluid beds used. Conical
shaped micro-reactors, geometrical similar to those of Kersten et al. (Kersten et al. 2003), will be
developed for improved simulation of bubbling and three-phase fluidization as may prevail in
catalytical gasification in hot compressed water.
As mentioned before, Umf cannot be determined for these conditions and Umb is estimated from the
velocity at which the first slugs are observed (see section 2.1). Measured values of Umb are compared
with the Abrahamsen and Geldart relation and predictions of a 3D DPM model (see Figure 6.7). It
should be noticed that the A and G correlation must be extrapolated outside the validated regime for
the conditions investigated here. In the DPM model, the fluid velocity is increased linearly with time.
At the point where the overall pressure drop equals the bed weight per unit area, the minimum
fluidization velocity is reached. Via visualization of the computed results using animation techniques,
the minimum bubbling velocity can be established. At Umb larger voids appeared in the generated
pictures of the fluid bed. The details of the DPM model are presented elsewhere (Ye et al., 2004,
2005). Both the A and G relation and the DPM model predict that, in the investigated density regime,
Umf and Umb hardly (10%) depend on the temperature (viscosity) for a certain water vapor density
(pressure). Within the experimental error, the experimental estimates of Umb agree with the DPM
results, which are also somewhat higher than predictions from the empirical relations of A and G (see
Figure 6.7).
0 50 100 150 200 2500
5
10
15
20
25 experimental data for Umb
density [kg/m3]
Umb (A&G correlation)
Umf (A&G correlation)
DPM model Umb 500 oC
supe
rfici
al fl
uid
velo
city
[mm
/s]
DPM model Umb 600 oC
DPM model Umf 400 oC DPM model Umf 500 oC DPM model Umf 600oC
DPM model Umb 400 oC
Figure 6.7 Comparisons of the Umf and Umb calculated from empirical correlations with DPM model
results and experimental values.
2D Simulation Results
In addition, 2D simulations have been carried out to visualize the flow patterns and the transition of
the flow regimes. Snapshots of the simulated fluidization behavior at a set pressure of 153.0 bar are
shown in Figure 6.8. At first, homogeneous bed expansion is observed. Then void structures in the top
and bottom of the bed can be clearly observed, but no obvious bubbles appear. At a velocity of 36
mm/s, more or less clear bubbles can be found. For a still higher fluid velocity, the bubbles develop
quickly and form clear slugs. The snapshots of the fluidization at the same fluid velocity but at
different fluid pressures are shown in Figure 6.9. It can be seen that with an increasing pressure, the
fluidization behavior becomes more ”smooth” and large bubbles or slugs are transformed into micro-
voids. Visually this may be interpreted as homogeneous fluidization.
124
Figure 6.8 Snapshots of the simulation results in a 2D supercritical water fluidized bed. Particle
diameter d = 80 µm, particle density = 2600 kg/m3, fluid temperature T = 783 K, fluid pressure p =
153.0 bar, and fluid density 50 kg/m3. Superficial fluid velocities U are 0.0, 0.012, 0.024, 0.036, and
0.048 m/s respectively.
Figure 6.9 Snapshots of the simulation results in a 2D supercritical water fluidized bed. Particle
diameter d = 80 µm, particle density = 2600 kg/m3. Fluid temperature T = 783 K. Fluid pressures are,
from left to right, 34.5, 67.1, 97.5, 126.3, 153.0, 178.1, 201.4, 223.2, and 262.4 bar, respectively,
which corresponds with fluid densities of, from left to right, 10, 20, 30, 40, 50, 60, 70, 80, and 100
kg/m3. Superficial fluid velocity U is 0.036 m/s.
125
126
Conclusions
The conclusions of this work can be summarized as follows:
● A quartz micro-fluid bed set-up has been developed and operated up to 5000C and 244 bar.
With these reactors, biomass gasification in hot compressed water can be studied. The method
is safe and cheap. Visual observation of the bed position and fluidization regime is possible as
well as the occurrence of different phases, carbon deposition in case of gasification, etc. Full
control of pressure, temperature, and flow rates is possible and relatively easy.
● Homogeneous fluidization, slug flow and more or less homogeneous turbulent beds have been
observed roughly in line with the Dn criterion of Kwauk and Li and empirical correlations.
● Despite the small diameter of the fluid bed, the expansion of the homogeneous fluid bed is
similar to those of larger beds provided Dt/dparticles >12.
● 2D and 3D soft-sphere simulations, including added mass effects for high density fluids, were
able to simulate the micro-fluid beds well. In the future, this model will be used to assist the
further development of micro-fluid bed reactors.
Notation
Ar
Cνm
D
Dn
dp
Fνm
g
k
mf
n
Remf
t
u
U
Umb
Umf
Ut
Archimedes number, Ar = gdp3 (ρp - ρf)ρf/µ2, dimensionless
added mass coefficient, dimensionless
reactor inside diameter, m
discrimination number , Dn= Ar/ Remf (ρp - ρf)/ ρf, dimensionless
particle diameter, m
added mass force, N
acceleration of gravity =9.81 m/s2
slope,m/s2
mass, kg
numerical exponent in Richardson-Zaki equation, dimensionless
Reynolds number at minimum fluidization velocity, Remf= Umf dp ρf /µ,
dimensionless
time, s
local gas velocity, m/s
fluid velocity, mm/s
minimum bubbling velocity, mm/s
minimum fluidization velocity, mm/s
terminal velocity, mm/s
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v
Greek letters
ε
Ρf
Ρp
µ
particle velocity, m/s
void fraction, dimensionless
fluid density, kg/m3
particle density, kg/m3
fluid viscosity, Pa.s
References
Antal, M.J.; Allen, S.G.; Schulman, D.; Xu, X.; Divilio, R.J., Biomass Gasification in Supercritical Water. Ind. Eng. Chem. Res. 2000, 39, 4040. Auton, T.R.; Hunt, J.; Prud'Homme, M., The Force Excerted on a Body in Inviscid Unsteady Non-Uniform Rotating Flow. J. Fluid Mech. 1988, 197, 241. Batchelor, G. K., A New Theory of the Instability of a Uniform Fluidized Bed. J. Fluid Mech. 1988, 193, 75. Chang, E.J.; Maxey, M.R., Unsteady Flow about a Sphere at Low to Moderate Reynolds umber. Part 2: Accelerated Motion. J. Fluid Mech. 1995, 303, 133. Clift, R.; Grace, J.R.; Weber, M.E., Bubbles, Drops and Particles. Academic Press, New York, 1978. Cundall, P.A.; Strack, O.D., A Discrete Numerical Model for Granular Assemblies. Geotechnique 1979, 29, 47. De Vries, R.J.; van Swaaij, W.P.M.; Mantovani, C.; Heijkoop, A., Design Criteria and performance of the Commercial Reactor for the Shell Chlorine Process. In Proceedings of the 5th Eur. Symposium on Chemical Reaction Engineering; Elsevier Publ. Corp.: Amsterdam, The Netherlands, 1972. Elliott, D.C.; Sealock, L.J., Jr.; Baker, E.G., Chemical Processing in High-Pressure Aqueous Environment. 3. Batch Reactor Process Development Experiments for Organics Destruction. Ind. Eng. Chem. Res. 1994, 33, 558. Ergun, S., Fluid Flow Through Packed Columns. Chem. Engng. Prog. 1952, 48, 89. Haider, A.; Levenspiel, O., Grag Coefficient and Terminal Velocity of Spherical and Non-Spherical Particles. Powder Technology 1989, 58, 63. Hao, X.H.; Guo, L.J.; Mao, X.; Zhang, X.M.; Chen, X.J., Hydrogen Production from Glucose used as a Model Compound of Biomass Gasified in Supercritical Water. Hydrogen Energy 2003, 28, 55. Kersten, S.R.A.; Prins, W.; van der Drift, B.; van Swaaij, W.P.M., Principles of a Novel Multistage Circulating Fluidized Bed Reactor for Biomass Gasification. Chem. Eng. Sci. 2003, 58, 725.
128
Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Reactor Design Considerations for Biomass Gasification in Hot Compressed Water. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection, van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 777. Kolb, G.; Hessel, V., Micro-stuctured Reactors for Gas Phase Reactions. Chem. Eng. J. 2004, 98, 1. Kruse, A. Abeln, J.; Dinjus, E.; Kluth, M.; Petrich, M.; Schacht, E. Sadri, E.; Schmieder, H., Gasification of Biomass and Model Compounds in Hot Compressed Water. Presented at the International Meeting of the GVC-Fachausschuβ "Hochdruckverfahrenstechnik", Karlsruhe, Germany, 1999; Paper No. 107. Kuipers, J.A.M.; Duin, K.J.; van Beckum, F.P.H.; van Swaaij, W.P.M., A Numerical Model of Gas-fluidized Beds. Chem. Eng. Sci. 1992, 47, 1913. Li, J., Euler-Lagrange Simulation of Flow Structure Formation and Evolution in Dense Gas-Solid Flows. University of Twente, The Netherlands, 2003. Liu, D.; Kwauk, M.; Li, H., Aggregative and Particulate Fluidization - The Two Extremes of a Continuous Spectrum. Chem. Eng. Sci. 1996, 51, 4045. Maharrey, S.P.; Miller, D.R., Quartz Capillary Microreactor for Studies of Oxidation in Supercritical Water. AIChE J. 2001, 47, 1203. Marzocchella, A.; Salatino, P., Fluidization of Solids with CO2 at Pressures from Ambient to Supercritical. AIChE J. 2000, 46, 901. Maxey, M.R.; Riley, J.J., Equation of Motion for a Small Rigid Sphere in a Nonuniform Flow. Phys. Fluids 1983, 26, 883. National Institute of Standards and Technology (NIST), Gaithersburg, MD. Potic, B.; Kersten, S.R.A.; Prins, W.; Assink, D.; van de Beld, L.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water: Results of Micro and Pilot Scale Experiments. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 742. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., A High-throughput Screening Technique for Conversion in Hot Compressed Water. Ind. Eng. Chem. Res. 2004, 43, 4580. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Gasification of Model Compounds and Wood in Hot Compressed Water. Submitted to Ind. Eng. Chem. Res., 2005. Schmieder, H.; Abeln, J.; Boukis, N.; Kruse, A.; Kluth, M.; Petrich, G.; Sadri, E.; Schacht, E., Hydrothermal Gasification of Biomass and Organic Wastes. J. Supercrit. Fluids 2000, 17, 145. Van Swaaij, W.P.M.; Zuiderweg, F.J., Investigation of Zone Decomposition in Fluidized Beds on the Basis of Two Phase Model. In Int. Symposium on Fluidization and its Applications, Toulouse, France, 1973. Ye, M.; van der Hoef, M.A.; Kuipers, J.A.M., A Numerical Study of Fluidization Behaviour of Geldart A Particles Using a Discrete Particle Model. Powder Technology 2004, 139, 129.
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Ye, M.; van der Hoef, M.A.; Kuipers, J.A.M., The Effects of Particles and Gas Properties on the Fluidization of Geldart A Particles. Submitted to Chem. Eng. Sci., 2005. Yoshida, T.; Oshima, Y.; Matsumura, Y., Gasification of Biomass Model Compounds and Real Biomass in Supercritical Water. Biomass Bioenergy 2004, 26, 71. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels 1993, 7, 574.
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131
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Abstract
This chapter deals with the design of a reactor for biomass gasification in hot compressed water.
Starting from the experimental results obtained in this thesis and those available in the open
literature, the requirements for an optimal design are identified, listed and discussed. An integrated
reactor/heat-exchanger concept, based on fluid beds and circulating solids, is proposed that can cope
with catalysis under fouling conditions, combustion of carbonaceous deposits, and ash removal.
132
Introduction
Biomass gasification in supercritical water (SCWG) is a novel process, under development after the
pioneering work of Modell (Modell, 1985) and Antal and co-workers (Yu et al., 1993) since the mid
eighties of the last century. In the laboratory (Antal et al., 2000; Hao et al., 2003; Kruse et al., 2003;
Lee et al.; 2002, Yu et al., 1993) and pilot-plant work (Boukis et al., 2002; Van de Beld et al., 2003)
that was reported so far, metal tubular reactors were used frequently in connection with a shell and
tube heat-exchanger. This thesis adds the following items to the experimental knowledge on SCWG:
The work described in Chapters 2 and 3 deals with mapping of the SCWG operating window under
catalytically inert conditions in batch microreactors. Continuous test including catalyst testing and
burn off of carbonaceous deposits is investigated at micro (Chapter 5) and pilot scale (Chapter 4).
Fluidization under SCWG conditions is investigated in Chapter 6. Experiments have been performed
with glycerol and glucose solutions. Glycerol does not produce carbonaceous deposits under SCWG
conditions and is consequently a model compound that represents feedstock materials that do not
cause fouling problems (e.g. very diluted streams). Glucose does produce carbonaceous deposits and
thus embodies more complex feed streams. Tests with glucose solutions have provided information on
the problems (fouling and blocking) that such feedstock materials can cause in the heat-exchanger and
reactor (Chapters 3 and 5).
Not too much attention has been paid to the practical design of a SCWG reactor (including heat-
exchanger) for continuous operation in a future commercial plant. While looking for the best reactor, a
number of requirements, characteristic for the SCWG process, should be taken into account. They are
listed and discussed in this chapter. Based on the list of these design aspects, a reactor concept for
SCWG is proposed. Such an analysis has been presented earlier by Kersten et al. (Kersten et al., 2004)
in the initial stage of the SCWG project at the University of Twente, based primarily on data available
in the open literature. In this chapter, the results of this thesis are added to the discussion, in particular
those results that provide new insights. Only the reactor/heat-exchanger is discussed here, other
challenging units of SCWG plant like the feedstock pump are not considered in this analysis.
133
Design aspects
1. Temperatures of above 800oC are needed in the reactor to allow non-catalytic thermal degradation
of the biomass molecules, and subsequent cracking or gasification of the intermediate sized fragments
to only the desired small molecules H2, CO, CH4 and CO2 (Antal et al., 2000 and Chapter 3 and 5).
Complete conversion of the model solutions to the gas-phase is only possible under inert conditions
for solutions with a concentration of organics of less than 5 wt %; higher concentrations produce
liquid and solid by-products (Chapters 3 and 5). By applying catalysis, the required operating
temperature can be reduced to 600oC and more concentrated solutions (up to 17 wt % glucose has been
tested) can be gasified completely (Chapter 3). Recently, Cortright et al. (Cortright et al., 2002) and
Elliot et al. (Elliott et al., 1993) demonstrated the potential of catalysis. Cortright et al., for instance,
obtained a high yield to hydrogen from diluted solutions (1 wt % model compound in compressed
water) at temperatures as low as 265oC by using Pt/Al2O3. It is, however, not evident that at such low
temperatures also high concentrations of organic matter can be treated.
2. High pressures. Although it has been suggested sometimes that supercritical water properties would
be beneficial for the gasification reactions, experiments at micro scale have shown little influence of
the pressure on either the carbon conversion degree, or the product yield and distribution (Chapter 3).
It is crucial for the process that the heat content of the reactor effluent is utilized as far as possible to
pre-heat the feedstock stream (mainly water) to reaction conditions. However, heat exchange between
these streams is not practical at low pressure, because of the high heat of evaporation at nearly
isothermal and isobaric conditions. Heating of the feedstock stream to the desired gasification
temperatures in a heat-exchanger without evaporation, requires operation at high pressures. This is the
true incentive of the high pressures involved in wet gasification (Chapter 1).
3. Additional heat. In practice, a hundred percent transfer of the available heat in the reactor effluent to
the feedstock stream is impossible. In fact, efficiencies of ca. 75% are typical for liquid-liquid shell
and tube heat-exchangers (Woods, 1995) resulting in a heat-exchanger outlet temperature of the feed
stream of ca. 450oC. The required temperature rise to the reactor conditions may be generated by the
exothermicity of the reactions. This is however only possible for highly concentrated feedstocks in
combination with a considerable methane production being a combination that is only possible by
applying catalysis. In all other cases additional heat input is required for the reactor. This heat can be
supplied externally and/or by in-situ generation by the heat of oxidation reactions. For the latter,
oxygen has to be added to the process. In-situ oxidation is preferable over external heating because of
efficiency and construction reasons. But it is then crucial to do it selectively, that is without consuming
any desired products. This is discussed further in the next part.
134
4. Incomplete conversion of the biomass is something that has to be taken into account in reactor
design considerations. Thermal decomposition of most wet feedstocks will result in carbon formation,
either simultaneously with the production of tars (condensable vapors) and gaseous compounds, or as
a result of secondary reactions of tar. It was found that, even under the high water pressure prevailing
in the SCWG process, steam gasification of char is still extremely slow (Van Swaaij, 2003). This
result excludes recycling of the carbon from the water outlet stream, or applying very long residence
times, as a reactor design option. In-situ oxidation of the carbon looks indeed interesting, because of
the corresponding heat production inside the reactor. Besides, in case of catalyst addition (discussed in
the next part), regeneration will be required to burn off any carbon deposited on the catalyst surface.
While considering in-situ oxidation of the carbon, it should be realized that oxygen introduced into the
SCWG reactor will react preferentially with gaseous products and dissolved organic molecules, while
leaving the carbon largely unconverted. Reactor staging, or creating separate combustion and
gasification zones inside a single reactor, may be a solution (Kersten et al., 2003). Alternatively, a
secondary wet-oxidation reactor can be used to clean the effluent water from the SCWG reactor. Burn-
off of carbonaceous deposits on the reactor wall and /or catalyst particles has been tested successfully
in a continuous micro reactor (Chapter 5) under realistic process conditions.
5. Catalysis is needed to lower the required gasification temperatures for complete conversion and to
improve the selectivity to either H2, or CH4 in case substitute natural gas is aimed at (Elliott et al.,
1993). From the work done in Japan, by researchers of NIRE (Minowa et al., 1998), it appears that
alkali metals as well as nickel-based catalysts promote the production of hydrogen at the cost of the
carbon monoxide yield. Elliott and co-workers from the Pacific Northwest National Laboratory
(Sealock et al., 1996) used a ruthenium catalyst for methane production from organic waste streams. In
both cases the conditions were subcritical, that is within the temperature range from 200 to 350oC and
at pressures up to 350 bar. Catalysis for supercritical conditions has been applied by Antal (Antal et
al., 2000) who found in his laboratory reactor increased carbon conversions when activated carbon
derived from coconut shells was used as a catalyst in a fixed bed. By adding 3 wt % Ru on TiO2, a
catalyst developed for conversion of waste water (Elliott et al., 1999), to the quartz capillary reactors
complete conversion of 1 to 17 wt % glucose solutions has been achieved at 600oC (Chapter 3). These
experiments clearly show the potential of heterogeneous catalysis. However, this catalyst may not
suitable for the actual process because of degeneration problems as discussed in Chapter 3 of this
thesis and by Elliot et al. (Elliott and Chapter 3). It is important to realize that, at the time of writing,
the potential of noble metal catalysts has been shown, but that stable systems with a long life time are
not available yet. This chapter focuses on catalytic gasification of feedstock stream containing more
than 5 wt % organics at 600oC, being a realistic temperature in view of the current status of catalyst
development. For reactor design the application of a heterogeneous catalyst has complicating
135
consequences, e.g. regeneration is required (see above). On the other hand, it also creates the
opportunity for additional removal of minerals that are deposited together with the coke on the catalyst
surface. Moreover, particle circulation can be used further to prevent the heat-exchanger from coking
and plugging. The formation of ash and coke occurs in a temperature range of roughly 200 - 400°C,
and perhaps mainly in a confined region of the tubular heat-exchanger.
6. Mixing pattern inside the reactor. At present, most of the continuous reactors in operation are
tubular with little or no backmixing. Available literature results (Antal et al., 2000; Kruse et al., 2003;
Lee et al., 2002), and our own experimental data (Potic et al., 2004 and Chapter 3), demonstrate
clearly that when the weight percentage biomass is increased, less permanent gases and more liquids
and chars are produced. As this situation is undesired, low feed concentrations are preferred. However,
it is not the concentration in the feed that is important, but the (local) concentration in the reactor.
When using an ideally stirred reactor it is possible to have a low reactant concentration in the reactor
while feeding a high concentration. It has been observed that, for non-catalytic gasification, the time
needed to reach the maximum conversion at temperatures above 650oC is about 30 s (Chapter 3).
Following the rules of David and Villermaux (David and Villermaux, 1978), this means that the
mixing time in the stirred cell should be around 3 s when ideal micro-mixing is aimed at. The question
is however, whether or not ideal micro-mixing is really required. Probably, a less-ideally mixed
reactor (partial segregation) could also result in higher conversions than can be achieved in reactors
with plug-flow. This will have to be established experimentally. A fluidized bed with turbulent solids
mixing (e.g. a spouted bed) is known to have reasonably good mixing of the fluid phase.
7. Ash treatment is an important issue in SCWG as most ashes are completely unsolvable in
supercritical water. Ash removal from the reactor as deposits on bed particles is a better option than
entraining it with the reactor effluent. Removal of these deposits together with the bed particles may
prevent blockage and corrosion/erosion problems.
8. Fouling and Corrosion. Corrosion is a serious problem at high temperatures, especially if oxygen
and sulfur are present. This is one of the reasons to develop catalysts that can reduce the required
gasification temperatures. With respect to the selection of the construction material for a commercial
plant, a two-barrier solution could be adapted from the technology of SCW oxidation that has been
developed earlier (Fauvel et al., 2003). Up to now it is recommended to construct the essential parts
from a custom-made alloy like for instance Inconel 625. A problem of general nature in SCWG is the
required heat exchange between the reactor outlet and inlet streams. Heating of the biomass slurry in a
heat-exchanger is likely to cause fouling/plugging problems because the thermal decomposition starts
already at ca 200oC producing oily products (tars) and, more seriously, polymers (char).
Reactor design
The following integrated reactor/heat-exchanger concept is proposed here for the SCWG process (see
Figure 7.1). It is assumed that heat is required in the main reactor.
This concept consists of:
A fluidized bed (FB) reactor is used as main reactor to carry out the gasification/cracking of biomass
to permanent gases and by-products. Fluid bed operation has been tested and demonstrated in this
thesis (Chapter 6). It has become clear that a supercritical fluid bed can be designed with available
design relations derived for much less severe conditions. The particles in the FB are catalytically
active in order to lower the temperature at which maximal conversion to the desired product is reached
and to steer the product distribution. The reactor is operated at ca. 600oC; the total system runs at
approximately 230 bar. This fluidized bed is designed and operated with a conical bottom and/or draft-
tube to ensure good mixing of the phases.
Products
Fluidized BedReactor
SolidsPurge
Product
O2 / air
Feed Heat exchanger
Off gas
Off gas
Figure 7.1 Proposed reactor/heat-exchanger for SCWG.
In the upper section of the main reactor, the solids are transported via a drainpipe to a separate
combustor. The fluid phase is stripped from particles by conventional means (e.g. cyclones) and is
available for heat exchange.
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137
The carbon deposits are removed from the catalyst particles in the combustor by reaction with oxygen.
Also for this reaction a fluidized bed is a suitable reactor. In the combustor, heat is produced and the
catalyst particles are regenerated. Separate reactors for the main reaction and heat generation/catalyst
regeneration, ensures that combustion of the desired product does not take place. This results in a
higher thermal efficiency and effective conversion of the by-product. For an operating temperature of
ca. 600oC in the main reactor, a suitable temperature of the combustor would be around 650oC. Just
like in the main reactor, the solids are transported out of the reactor via drainpipe to the heat-
exchanger and the main reactor. The solids transported to main reactor provide the heat that is
necessary to raise the temperature of the feed stream after leaving the heat-exchanger to reactor
conditions. The off gas is available for heat exchange after separation of the solids.
At the tube side of the heat-exchanger, the solids phase from the combustor is mixed with the fresh
feed. It is wise to mix the particle phase with the fresh feed when the temperature level of this feed is
around 200oC. Below this temperature, biomass does not decompose. Above 250oC, normally,
produced char particles and tar deposits would cause blocking and fouling problems. However, in the
proposed design this is counteracted by the scouring action of the particles, which keeps the heat-
exchanger clean.
At the shell side of the heat-exchanger, the effluent from the main reactor and the off gas from the
combustor are available. A special layout must be considered to keep the product and the off gas from
the combustor separated.
Ashes deposited on the bed particle’s surface can be withdrawn from the system at several locations in
the solids-loop.
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Conclusions
From the results obtained in this thesis on the SCWG of biomass and from previous literature results, a
preliminary plant design concept is described which can cope with heat exchange and reaction under
fouling conditions dealing with catalysis, coke combustion, and ash removal. The concept is based on
fluid beds and circulating solids. The system will be further developed in micro systems.
References
Antal, M.J.; Allen, S.G.; Schulman, D.; Xu, X.; Divilio, R.J., Biomass Gasification in Supercritical Water. Ind. Eng. Chem. Res. 2000, 39, 4040. Boukis, N.; Diem, V.; Dinjus E.; Galla U.; Kruse A., Advances with the Process of Biomass Gasification in Supercritical Water. Presented at the 4th International Symposium on High Pressure Technology and Chemical Engineering, Venice: Italy, 2002. Cortright, R.D.; Davda, R.R.; Dumesic, J.A., Hydrogen from Catalytic Reforming of Biomass-derived Hydrocarbons in Liquid Water. Nature 2002, 418, 964. David, R.; Villermaux, J.J., De Chimie Physique 1978, 75, 656. Elliott, D.C.; Sealock, L.J.; Baker, E.G., Chemical Processing in High-Pressure Aqueous Environment: 2. Development of Catalyst for Gasification. Ind. Eng. Chem. Res. 1993, 32, 1542. Elliott, D.C.; Neuenschwander, G.G.; Phelps, M.R.; Hart, T.R.; Zacher, A.H.; Silva, L.J., Chemical Processing in High-Pressure Aqueous Environment. 6. Demonstration of Catalytic Gasification for Chemical Manufacturing Waste Water Cleanup in Industrial Plants. Ind. Eng. Chem. Res. 1999, 38, 879. Elliott, D.C.; Neuenschwander, G.G.; Hart, T.R.; Butner, R.S., Low-Temperature Catalytic Gasification of Wet Biomass Residues. In Proceedings of the 2nd World Biomass Conference, Biomass for Energy, Industry, and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 746. Fauvel, E.; Joussot-Dubien, C.; Pomier, E.; Guichardon, P.; Charbit, G.; Charbit, F.; Sarrade, S., Modeling of a Porous Reactor for Supercritical Water Oxidation by a Residence Time Distribution Study. Ind. Eng. Chem. Res. 2003, 42, 2122. Hao, X.H.; Guo, L.J.; Mao, X.; Zhang, X.M.; Chen, X.J., Hydrogen Production from Glucose Used as a Model Compound of Biomass Gasified in Supercritical Water. Hydrogen Energy 2003, 28, 55. Kersten, S.R.A.; Prins, W.; van der Drift, B.; van Swaaij, W.P.M., Principles of a Novel Multistage Circulating Fluidized Bed Reactor for Biomass Gasification. Chem. Eng. Sci. 2003, 58, 725. Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Reactor Design Considerations for Biomass Gasification in Hot Compressed Water. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection;, van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 777.
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Kruse, A.; Henningsen, T.; Sinag, A.; Pfeiffer, J., Biomass Gasification in Supercritical Water: Influence of the Dry Matter Content and the Formation of Phenols. Ind. Eng. Chem. Res. 2003, 42, 3711. Lee, I.; Kim, M.-S.; Ihm, S.-K., Gasification of Glucose in Supercritical Water. Ind. Eng. Chem. Res. 2002, 41, 1182. Minowa, T.; Zhen, F.; Ogi, T., Cellulose Decomposition in Hot-Compressed Water with Alkali or Nickel Catalyst. J. Supercrit. Fluids 1998, 13, 253. Modell, M., Gasification and Liquefaction of Forest Products in Supercritical Water. In Fundamentals of Thermochemical Biomass, Overend, R.P.; Milne, T.A.; Mudge, L.K., Eds.; Elsevier Applied Science Publishers Ltd., London, 1985; p. 95. Potic, B.; Kersten, S.R.A.; Prins, W.; Assink, D.; van de Beld, L.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water: Results of Micro and Pilot Scale Experiments. In Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 742. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., Gasification of Model Compounds and Wood in Hot Compressed Water. Submitted to Ind. Eng. Chem. Res., 2005. Potic, B.; Kersten, S.R.A.; Ye, M.; van der Hoef, M.A.; Kuipers, J.A.M.; van Swaaij, W.P.M., Fluidization of Supercritical Water in Micro Reactor. Chem. Eng. Sci 2005, 60, 5982. Sealock, L.J. Jr.; Elliott, D.C.; Baker, E.G.; Fassbender, A.G.; Silva, L.J., Chemical Processing in High-Pressure Aqueous Environments. 5. New Processing Concepts. Ind. Eng. Chem. Res. 1996, 35, 4111. Van de Beld, L.; Wagenaar, B.M.; Assink, D.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; Penninger, J.M.L., Biomass and Waste Conversion in Supercritical Water for the Production of Renewable Hydrogen. Presented at the 1st European Hydrogen Energy Conference, Grenoble, France, 2003. Van Swaaij, W.P.M., Technical Feasibility of Biomass Gasification in a Fluidized Bed with Supercritical Water. University of Twente: The Netherlands 2003, 160. Woods, D.R., Process Design and Engineering Practice. PTR Prentice Hall, New Jersey, 1995. Yu, D.; Aihara, M.; Antal, M.J., Hydrogen Production by Steam Reforming Glucose in Supercritical Water. Energy Fuels 1993, 7, 574.
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141
ACKNOWLEDGEMENTS
And now, after five years of learning how to fit into a new scientific environment, but also how to fit
in every-day Dutch life, I have the opportunity to express my gratitude to the many people that helped
me on my way to the “dr” title.
To start with, I would like to thank my promoter prof.dr.ir. Wim van Swaaij, and the assistant-
promoters dr.ir. Wolter Prins and dr. Sascha Kersten for giving me an opportunity to learn from them
and to develop my expertise in reactor engineering and particularly biomass gasification processes.
I am also grateful to prof.dr.ir. J.A.M. Kuipers, dr.ir. M.A van der Hoef, and dr. M. Ye for
collaboration and support in CFD modeling of supercritical water fluidization, which resulted in a
joint paper.
Dr. Sascha Kersten, who also became a good friend of mine, should be mentioned especially. I would
like to thank him for his generous moral support, patience and understanding. Sascha, I really learned
a lot from you. Not just how things should be done in a PhD project, but also how an honest friendship
can enrich one’s life.
Regarding the technical support, I was really lucky to have in the High Pressure Laboratory the kind
assistance of Johan Agterhorst, who was always ready to help me at once. Daan Asink, an expert from
BTG, enabled the pilot plant investigation and I sure need to thank him for his valuable support.
Mrs. Yvonne Bruggert-ter Huurne and Wies Elfers, both secretaries in the TCCB group, helped me a
lot with everyday problems at work. During the last few months Yvonne has put much effort in editing
the style and text of my thesis for which I am very grateful. It was a great pleasure to work with her.
My first contact at the UT was Rik Akse. He always took care of the financial and administrative
aspects of my PhD project, which made my PhD-life a lot easier.
I had a great time with all the other colleagues in the TCCB group, Xiaoquan Wang, Mariken Bleeker,
Guus van Rossum, Dragan Knezevic, Marco Rep. And I should of course also acknowledge the staff
of the High Pressure Laboratory, Gert Banis, Fred ter Borg, Karst van Bree and Geert Monnink for
their pleasant support. Last but not least I would like to mention my students Mariken, Johny, Stefan
and Sander who made a significant contribution to my research results.
142
Apart from colleagues from the University, I received a lot of support and help from my personal
friends. My best Dutch girlfriend and paranimf Diana Wulms enriched my life in many aspects, in
particular regarding the Dutch culture. She accepted my family and me and helped us to adapt to a
Dutch lifestyle. Thank you Diana!
During the last year of my PhD work, Dennis Schipper fortunately opened to me the door of his
company Demcon, and with that to the real professional life. I have now the opportunity to work with
highly skilled people who are also very kind and helpful. Dennis, Bianca, Helena, HenkJan, Twan,
Rini, Peter, Sjoerd, Mathijn, Philia, Reinier, Gert (Nanomi bv), Jeroen (Medspray bv) and all the
others, I would like to thank you here as well.
Finally I would like to thank Harry Slaghuis who played a special role in my family’s life.
Harry, only you know how many problems my family and I had since we arrived in Enschede. You
brought to us a life that we never could dream of, peaceful and organized with lots of joyful moments.
Having you as a friend makes me a rich person.
To my Serbian friends (in Serbian)
Hvala svima koji ste se potrudili da ulepšate vreme koje smo proveli zajedno. Naročito bih želela da se
zahvalim Vesni Smiljanić na dizajnu koji je osmislila za ovu knjigu.
To my parents Bora i Stojanka Višnjički (in Serbian)
Stigli smo do kraja i ove uzbudljive priče. Posle toliko godina vaše bezuslovne podrške a pre svega
vaše beskrajne ljubavi, dobila sam šansu da vam se zahvalim.
Mama, ta kap vode na dlanu koju si toliko čuvala postala je vir oslikan na koricama ove knjige. Hvala
ti što si uvek bila uz mene.
Tata, postoji samo jedan način da ti kažem koliko sam ponosna na sve to što smo zajedno prošli:
“Naša škola ide napred!”
This book I dedicate to my daughters Luna and Mila and to my husband Goran. They give me all the
strength and joy in my life. Their support and understanding also helped me to arrive at this very last
part of my book. Luna, Mila and Goran, you are my life and my orientation.
I love you!
Biljana Potic
Enschede, May 2006.
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LIST OF PUBLICATIONS
Potic, B., van de Beld, L., Assink, D., Prins, W., van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water. In Proceedings of the 12th European Conference and Exhibition on Biomass for Energy, Industry and Climate Protection; Palz, W.; Spitzer, J.; Maniatis, K.; Kwant, K.; Helm, P.; Grassi, A.; Eds.; ETA Florence, WIP Munich., Amsterdam, 2002, p.777. Van de Beld, L., Wagenaar, B.M., Assink, D., Potic, B., Kersten, S.R.A., Prins, W., van Swaaij, W.P.M., Penninger, J.M.L., Biomass and Waste Conversion in Supercritical Water for the Production of Renewable Hydrogen. Presented at the 1st European Hydrogen Energy Conference, Grenoble, France, 2003. Potic, B.; Kersten, S.R.A.; Prins, W.; Assink, D.; van de Beld, L.; van Swaaij, W.P.M., Gasification of Biomass in Supercritical Water: Results of Micro and Pilot Scale Experiments. Proceedings of the 2nd World Conference and Technology Exhibition on Biomass for Energy, Industry and Climate Protection; van Swaaij, W.P.M.; Fjallstrom, T.; Helm, P.; Grassi, A., Eds.; ETA-Florence and WIP-Munich: Rome, Italy, 2004; p. 742. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M., A High-throughput Screening Technique for Conversion in Hot Compressed Water. Ind. Eng. Chem. Res. 2004, 43, p.4580. Matsumura, Y.; Minowa, T.; Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; van de Beld, L.; Elliott, D.C.; Neuenschwander, G.G.; Kruse, A.; Antal, M.J.; Biomass Gasification in Near- and Supercritical Water: Status and Prospects. Biomass and Bioenergy 2005, 29(4), 269. Potic, B.; Kersten, S.R.A.; Ye, M.; van der Hoef, M.A.; Kuipers, J.A.M.; van Swaaij, W.P.M., Fluidization of Supercritical Water in Micro Reactor. Chem. Eng. Sci. 2005, 60, 5982. Potic, B.; Kersten, S.R.A.; Prins, W.; van Swaaij, W.P.M.; Gasification of Biomass in Supercritical Water: Catalytic and Non-catalytic Batch Results. Accepted for publication in Ind. Eng. Chem. Res.
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CURRICULUM VITAE
Biljana Potic was born on September 20th 1970, in Novi Sad, Serbia and Montenegro. She studied
Chemical Engineering at the University of Belgrade, Serbia and Montenegro and graduated in
September 1994. Her master studies she did at the same University. During that period, she worked as
research assistant in “Process Thermodynamics” group headed by Prof. Serbanovic. On 8th March
1998 she obtained her master degree on the subject of “Heat transfer phenomena in the process
furnaces”.
From 1996 till 2001 she worked as process engineer for “Stark”, Belgrade, Serbia and Montenegro. In
2001 she started her PhD work at the University of Twente with Prof. Van Swaaij and Dr. Prins on the
subject: ”Biomass Gasification in Supercritical Water”. Since February 2005 she works as a Project
Manager in Demcon Twente B.V.