A paper in the proceedings of a conference on fluidization 103 16–17 November 2011, Johannesburg, South Africa

IFSA 2011, Industrial Fluidization South Africa:103–123 Edited by A. Luckos & P. den Hoed Johannesburg: Southern African Institute of Mining and Metallurgy, 2011

Particle fluidization and reaction engineering activities at the Institute of Chemical Engineering, TU Wien

F. Winter, C. Pfeifer, T. Pröll, R. Rauch, A. Reichhold and H. Hofbauer Vienna University of Technology, Institute of Chemical Engineering

Vienna, Austria

Keywords: gasification, CO2 capture, chemical looping, FCC, synthetic biofuels, combustion, agglomeration, iron ore reduction, fluidized bed

Abstract—This paper gives an overview of the research activities and highlights on particle fluidization and reaction engineering at the Institute of Chemical Engineering at Vienna University of Technology. It is structured into steam gasification in dual fluidized bed, CO2 capture including oxy-fuel combustion, chemical looping combustion and reforming, synthetic biofuels including synthetic natural gas, Fischer-Tropsch synthesis, mixed alcohols and the production of hydrogen. The chapter is followed by the studies of agglomeration under fluidized bed combustion conditions and catalytic cracking of feeds from biological sources. The paper concludes with iron ore reduction kinetics.

GASIFICATION OF CARBONACEOUS FEEDSTOCK IN A DUAL FLUIDIZED BED STEAM GASIFIER

The research group for Gasification and Gas Cleaning1 focuses on new approaches thermo-chemical conversion of carbonaceous feedstock, primarily different types of biomass. In the “classical” dual fluidized bed gasifier, heat is provided in direct contact with hot bed material particles externally heated in a combustion reactor. This type of gasification system is demonstrated in Güssing and Oberwart (Austria) and yields a high quality product gas. Thus, the gas is well suited for synthesis processes.

Comprehensive reviews about classical dual fluidized bed gasifiers have been published by Corella et al.2 and Göransson et al.3 whereas the dual fluidized bed steam gasification at Vienna University of Technology is described in detail in the following.

Dual fluidized bed steam gasification Biomass steam gasification allows for the conversion of solid biomass to a medium calorific gas (12–14 MJ/Nm³) consisting mainly of H2, CO, CO2, CH4 and H2O. At Vienna University of Technology, the dual fluidized bed steam gasification technology has been developed to provide the heat for the gasification reactor by circulating the bed material. This system is a further development of the so-called FICFB technology (fast internally circulating fluidized bed) which can be found described in references 4 and 5.

Figure 1 shows the principle of the dual fluidized bed steam gasification process and Figure 2 shows how this principle is implemented. The biomass enters a bubbling fluidized bed gasifier where drying, thermal degasification and partially heterogeneous char gasification take place at bed temperatures of about 850–900°C. Residual biomass char leaves the gasifier together with the bed material through an inclined, steam-fluidized chute towards the combustion reactor. The combustion reactor is used for heating up the bed material and is designed as a fast fluidized bed (riser). Air is used as the fluidization agent in the riser. After particle separation from the flue gas in a cyclone, the hot bed material flows back to the gasifier via a loop seal. Both connections, the loop seal and the chute are fluidized with steam, which effectively prevents gas leakage between the gasification and combustion zones and also allows

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for high solid throughput. The temperature difference between the combustion and the gasification reactor is determined by the energy needed for gasification as well as the bed material circulation rate. The system is inherently auto-stabilising, since a decrease of the gasification temperature leads to higher amounts of residual char which result in more fuel for the combustion reactor. This, in turn, transports more energy into the gasification zone and thereby stabilises the temperature. In practical operation, the gasification temperature can be influenced by the addition of fuel (e.g. recycled product gas, saw dust, etc.) into the combustion reactor (shown in Figure 1 as Add. fuel). The pressure in both the gasification and combustion reactors is close to atmospheric conditions. The process yields two separate gas streams, a high quality product gas and a conventional flue gas, at high temperatures. The product gas is generally characterised by a relatively low content of condensable higher hydrocarbons (4–8 g/m³ of so-called tars, heavier than toluene), low N2 (< 1 vol.-%db) and a high H2 content of 36–42 vol.-%db. For practical use, olivine, a natural mineral, has proven to be a suitable bed material with enough resistance to attrition and moderate tar cracking activity.

Figure 1. Principle of steam gasification without selective transport of CO2

Figure 2. Classical dual fluidized bed steam gasifier

Steam Air

Biomass

Gasification (~ 850 °C)

Gasification (~ 850 °C)

Combustion (~ 920 °C)

Combustion (~ 920 °C)

Producer Gas(CH4, CO, H2, CO2, H2O) Flue gas

Circulation (bed material,

char coal)

Heat

Add. fuel

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Results Different kinds of feedstock have been investigated successfully so far at the 100 kW pilot plant (see Table 1). Additionally to the arranged biomass fuels, different coals and lignites have also been investigated. For all the presented investigations, the bed inventory of 100 kg olivine was kept constant as well as the gasification temperature of 850°C. It can be seen that the gas composition for the different biofuels is in the same range. Coal and lignite show generally higher values for hydrogen and lower methane as well as other gaseous hydrocarbon levels. Coal has been tested in mixtures with wood pellets in ratios of 0 to 100%, and generally the tar content in the product gas of coal gasification is about half the value as for wood gasification. This experimental campaign is detailed by Aigner and co-workers.6 The main components of the GC-MS tar are naphthalene, indene and acenaphthalene, which are the same (components as well as composition) for coal as for wood.

An overview about variation of process parameters, bed materials and feestock is given by Pfeifer et al.7,8

Table 1. Product gas composition for the experiments with different fuels at 850°C gasification temperature

Product gas composition, vol.-% Fuel CO CO2 CH4 H2 Wood pellets 26.1 21.3 9.9 40.3 Wood chips 24.4 21.7 11.32 37.4 Bark 23.3 18.3 8.0 44.3 Wood chips (willow) 21.9 24.8 10.7 39.2 Wood/straw mixt. 80/20wt.-% 20.3 24.3 9.9 40.4 Wood/straw mixt. 60/40wt.-% 22.4 21.5 10.0 41.8 Sewage sludge 16.8 26.7 8.0 41.5 Lignite 23.0 17.2 5.3 54.5 Wood/coal mixt. 50/50energy% 7.7 23.85 17.7 45.8 Coal 11.5 18.0 7.9 57.9

G-volution gasifier The novel dual circulating fluidized bed (DCFB) concept with countercurrent effect in the gasification section and with multistage (additional) solids separation systems, for coarse and fine particles on each side, is displayed in Figure 3. The outstanding improvement of the G-volution design is the special construction of the gasification section with the fluid dynamics in this reactor. Thus, the gas-solids interaction is significantly improved. Gas as well as solids residence time is increased with regard to contact of bed material and the gaseous phase. Fluid dynamics in the reactor can be expected to resemble a multi-stage cascade of stirred vessels. The gasification reactor can also be described as a plug flow reactor for gas and a column of stirred vessels for solids, with the special characteristic that the gaseous phase and solids move in countercurrent directions. This includes that fresh, regenerated and hot bed material from the combustion reactor (CR) is present in the upper part of the gasification reactor (GR) before the product gas is leaving the gasifier. Thus maximize chemical and physical driving forces, ensure high conversion rates and low tar contents in the product gas. With increasing hold up of bed material the pressure drop is increasing in the gasification reactor. The pressure difference between the lower parts of the combustion and gasification reactor can be used to replace the externally auxiliary fuel input. Therefore a bypass directs hot product gas in a defined quantity from the gasification to the combustion reactor (see Figure 3). Detailed results about the conversion of methane, hydrocarbons and tars, as well as cold flow modeling of the key modifications of the G-volution concept have been presented by Schmid et al.9,10, Pfeifer et al.11 and Guìo-Pèrez et al.12 Based on those findings the following main advantages of the new gasification system can be expected: • Smaller particle sizes of bed material reduces the necessary gas velocity and in the

following attrition and abrasion effects • Increases of the residence times for fuel particles as well as gases with regard to gas-solids

interaction • Feedstock can be fed at several positions depending on the fuel parameters (e.g.

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gaseous/liquid/solid, amount of volatiles) • Global circulation rate predominantly controlled with fluidization rate of the combustion

reactor • Fast fluidized combustion reactor combined with a moderately fluidized gasification

reactor with zones • Turbulent fluidization in each stage (excellent gas-solids contact) • Solids residence time distribution resembles a cascade of stirred vessels (dispersed

downward movement of solids) • Increased bed material hold up (and residence time) in the gasification reactor • The countercurrent effect of solids and gas in the gasification reactor maximizes chemical

and physical driving forces over height • Simple geometrical changes (good applicability to refractory-lined units) • Entrained fines (like fine char and CaCO3) out of the gasification reactor are fed back to the

reactor system through a solids separator system • A combination of hard coarse particles and softer fines is possible • Classifying effect and countercurrent movement of coarse particles (downward) and fines

(upward) in the gasification reactor • Internal utilization of hot product gas to control process parameters (no need of external

auxiliary fuel input to combustion reactor)

Figure 3. G-volution gasification system

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CO2 CAPTURE IN FLUIDIZED BED FUEL CONVERSION

Carbon capture and storage general aspects Carbon capture and storage (CCS) is a key part of the IEA’s mid-term strategy to limit the CO2 concentration in the atmosphere and could contribute about 10% of total CO2 emission reduction already by 2030 with further increase in share after 2030.13 Since the main efficiency loss is contributed by the process of CO2 capture, extensive research is going on to find effective ways of separation and concentration of CO2 from fuel conversion processes.

Fluidized bed technology offers a great potential with respect to CO2 capture. On the one hand, oxy-fuel combustion is discussed as an early technology for CO2 capture. Here, the use of circulating fluidized bed (CFB) technology can allow for highly oxygen-enriched operating conditions compared to pulverized coal boilers where the risk for creation of hot spots in the furnace limits the achievable oxygen concentration. Lower exhaust gas recycle rates make the overall units more compact and efficient. Another typical fluidized bed technology with a tremendous potential for CO2 capture is chemical looping combustion (CLC). Here, inefficient gas-gas separation is inherently avoided. A variant of CLC at global lack of oxygen conditions is chemical looping reforming (CLR) for CO2 syngas or hydrogen generation.

Both CFB oxy-fuel combustion and CLC/CLR are investigated in the research group on Zero Emission Technologies at the Institute of Chemical Engineering. A 100 kWth CFB pilot plant is available for oxy-fuel combustion and a 120 kWth pilot plant for CLC/CLR of gaseous fuels, the latter being the largest CLC unit successfully operated so far.

Circulating fluidized bed oxy-fuel combustion Oxy-fuel is a combustion technology where the combustion air is an artificial mixture of recycled flue gas and oxygen. The output of the combustion is a CO2-rich flue gas which is best suited for capture and storage. Due to the mixing of flue gas and oxygen, no nitrogen is introduced and capture of carbon dioxide is facilitated. The concept of an oxy-fuel process is shown in Figure 4.

combustionchamber (CFB)

boiler exhaust gas cleaning

fuel

pure-O2

reci-gas

water

ASUair N2

Figure 4. Oxy-fuel concept.

Application of pure oxygen in industrial furnaces has the potential to provide concentrated CO2 exhaust streams for carbon capture and, additionally, adds a relevant degree of freedom to the combustion process because of the variable oxygen content of the synthetic air mixture (combustion gas). The main advantage of fluidized bed boilers is the relatively isothermal combustion chamber due to the heat capacity of the bed material. It is possible to extract heat directly out of the fluidized bed material. This means that the oxygen content in the combustion

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gas may be increased without increasing the combustion temperature. Less exhaust gas needs to be recycled. The oxy-fuel process offers a supplementary degree of freedom, in the terms of O2 concentration in the feed which allows improved operation especially for low calorific fuels.

Oxygen enrichment of combustion air has been applied for low heating value fuels like sewage sludge in bubbling fluidized bed combustors.

The present project is to some extent a combination of the above approaches and paves the way for capture ready waste incineration with immediate economic benefit. Such benefit comes e.g. from the avoidance of fossil fuel co-firing when converting low calorific fuels.

In a final stage of development with infrastructure for CO2 compression, transport and storage, the technology offers the potential of below zero emission spots if the fuel mix has a non-fossil carbon content.

The pilot rig (Figure 5) is located at the site of an air separation facility where technical gases are readily available. The plant is designed for hard coal operation but with additional feeding devices for alternative solid fuels such as sewage sludge and wood chips.14 The nominal operating point is 100 kWth for hard coal at 21 vol-% O2 in the combustion gas. Excess heat can be withdrawn from the bed material return loop in a controlled way. For higher O2 content in the combustion gas the necessary fuel power may be higher than 100 kWth to keep the fluidization conditions in the CFB riser. The variable O2 content in the combustion gas allows for operation of low calorific fuels such as mechanically dewatered sewage sludge.

fuelO2

CO2

post

com

bust

ion

heat

exc

hang

er

reac

tor

bed material cooler

water

water

bag filter

water

water

cyclone

hot g

as fa

n

Figure 5. OxyFuel Pilot plant 100 kWth

Experimental campaigns for the oxy-fuel combustion of coal15 and sewage sludge with 40% water content16 was performed. Further work will focus on investigation of the oxy-fuel process for other alternative fuels such as biomass and domestic waste.

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Chemical looping combustion and reforming Chemical looping combustion (CLC) refers to a process where a solid oxygen carrier material is used to selectively transport oxygen from an air reactor (AR) to a fuel reactor (FR) with the purpose to fully oxidize a hydrocarbon fuel in the fuel reactor to CO2 and H2O. The oxygen carrier leaves the fuel reactor in a reduced state and is re-oxidized again in the air reactor. The principle of CLC is illustrated in Figure 6(a).

Chemical looping reforming (CLR) refers to a process where less air is supplied to a CLC system than needed for full oxidation of the fuel. This requires the fuel reactor exhaust stream to contain CO and H2 besides CO2 and H2O. The principle of CLR is illustrated in Figure 6(b).

a)

Air reactor(AR)

Fuel reactor(FR)

MeOx

MeOx-1

Air

N2, Ar, (O2)CO2, (H2O)

Fuel

Cooling/ condensation

CO2

H2O

b)

Air reactor(AR)

(Regenerator)

Fuel reactor(FR)

(Reformer)

MeOx

MeOx-1

Air

N2, Ar, (O2) CO, H2, CO2, H2O

Fuel

Figure 6. Chemical looping combustion (a) and chemical looping reforming (b). At Vienna University of Technology, the dual circulating fluidized bed (DCFB)

configuration has been proposed. 17,18 The DCFB principle is sketched in Figure 7(a). The global solids loop starts in one of the two reactors (air reactor) where solids are entrained, separated from the gas in a cyclone and sent to the other reactor (fuel reactor) through a fluidized loop seal. From the fuel reactor, the solids are flowing back into the air reactor through a second loop seal connecting the bottom regions of the two reactors. The fuel reactor features a circulation loop in itself (fuel reactor cyclone and loop seal) and may be optimized with respect to good gas-solid contact and low particle attrition.

The global circulation rate can be effectively controlled by staged fluidization of the air reactor. The direct hydraulic communication of the two circulating fluidized bed reactors in the lower loop seal allows stable solids distribution in the system. The implementation in the pilot plant is sketched in Figure 7(b) including the controllable air reactor cooling device.

More than 500 hours of chemical looping operation have been achieved at the pilot plant between 2008 and 2011. Different oxygen carrier materials ranging from natural minerals such as ilmenite19 to designed nickel-based materials.20,21 The most relevant operating parameters and results obtained with nickel-based oxygen carriers are summarized in Table 2. In chemical looping reforming of dry natural gas full conversion to equilibrium was achieved in the fuel reactor without any carbon loss to the air reactor where an oxygen-free exhaust gas stream consisting of N2 and Ar is obtained.22

In conclusion, CLC of gaseous fuels at atmospheric pressure is feasible today. The DCFB system configuration is simple and robust proposing good scalability to larger size. CLR has been most successfully demonstrated at atmospheric pressure for CO2 ready production of hydrogen. Ongoing research at the institute focuses on scale-up of CLC for CO2ready industrial steam generation23, testing of innovative oxygen carriers for CLC24, direct CLC of solid fuels25, and chemical looping reforming.26

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a)

airr

eact

or(A

R)

exhaustAR

fuel

reac

tor(

FR)

exhaustFR

air fuel

LS

LS

LS

b)

air

steam

steam

steam

fuelgas

FR exhaustAR exhaust

steamsteam

air/ steam

air/ steam

Figure 7. Dual circulating fluidized bed(DCFB) principle(a) and sketch of the 120 kWth pilot plant(b).

Table 2. Measured performance data for CLC of natural gas using an oxygen carrier containing 40 wt.% active NiO on a NiAl2O4/MgAl2O4 support (values in brackets refer to standard operation).

Parameter Unit Value, Range Fuel power kW 60-150 (140)a Global air/fuel ratio - 0.4 – 1.3 (1.1)b Specific total oxygen carrier inventory kg/MW 500a,b Difference in degree of oxidation ΔXS - 0.07 – 0.13 (0.07)a CH4 conversion % 99.3b CO2 yield % 92.5b Combustion efficiency % 93.6b Carbon loss to AR % 0a,b Dilution of FR exhaust with AR exhaust % < 1.2a Cumulative operating time hours > 200b Loss of solids (fines) %/hour not determined References a) 19, b) 21

SYNTHETIC BIOFUELS In Austria there was much R&D done in the past to develop different technologies for thermochemical conversion based on gasification of biomass. With the development of the biomass gasification technology at the Institute of Chemical Engineering and the successful implementation at the biomass CHP Güssing also other R&D topics than usage of the valuable product gas in a gas engine for CHP were considered. On the following technologies, R&D is done by Vienna University of Technology in cooperation with Bioenergy 2020+:27

• BioSNG to produce synthetic natural gas from wood • Fischer-Tropsch synthesis for production of diesel

• Synthesis of mixed alcohols • Production of hydrogen from biomass

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Biomass CHP Güssing A demonstration plant was built in Güssing, region Burgenland, with an installed fuel input of 8 MW and a dual fluidized bed reactor28 to gasify biomass and to produce syngas, used currently for heat and electricity generation. The start-up of the biomass CHP was end of 2001 and currently this biomass CHP reaches about 7000 hours of operation per year. The biomass CHP includes a gas cleaning, which consists of a particle filter and a RME scrubber, to clean the gas to a level, that it can be used without any problems in a gas engine. For the synthesis gas applications slip streams of this product gas, after the gas cleaning, before the gas engine are taken (see Figure 8).

Figure 8: Flow chart of biomass CHP Güssing

From the history, the R&D on BioSNG was the first synthesis, which was investigated at the biomass CHP Güssing. Here the work started already in 2002. The second synthesis was the Fischer-Tropsch synthesis (FT) where the work started within the EC-project RENEW in 2004. In 2007 several Austrian R&D organisations in the area of thermochemical biomass conversion formed the competence centre “Bioenergy 2020+”, which is funded by the Austrian COMET funding program. Bioenergy 2020+ has a very close cooperation with the Institute of Chemical Engineering at Vienna University of Technology and the R&D on FT synthesis, mixed alcohols or hydrogen production is done within this cooperation. Within Bioenergy 2020+ also a new Technikum was build beside the biomass CHP Güssing, which gives excellent frame conditions for the R&D work done. In the following section the main synthesis gas applications are described in more detail.

BioSNG (synthetic natural gas) Within the EC demonstration project BioSNG the complete chain from wood chips to final BioSNG including usage of the BioSNG in CNG-cars was demonstrated. This project was finished by end of 2009 and actually the industrial partners try to implement this technology into the market. After successful demonstration of the whole chain in 2009, no ongoing project is done at the moment.

To synthesis gas applications

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Fischer-Tropsch synthesis Fischer-Tropsch synthesis converts hydrogen and carbon monoxide to n-paraffin’s, which can be used as high grade diesel. This synthesis is investigated in a scale of 5kg/day to develop the optimal chain from wood chips to diesel. Here in the first years the optimal gas treatment technology was developed and actually the optimization of the FT synthesis itself and the upgrading of the raw FT product are investigated.

The FT synthesis consists of the additional gas treatment, a slurry FT reactor and the product separation. In the gas treatment the main steps are sulphur removal, reforming to convert all hydrocarbons to hydrogen and carbon monoxide and to adjust the H2:CO ratio, and the compression of the gas. As FT reactor a slurry reactor is used, as this type of reactor has an excellent heat transfer, which is important for strong exothermic reactions like FT. The product separation is done first by condensation and then by removal of aerosols by scrubbing. A flow chart is given in Figure 9.

Figure 9. Flow sheet of the Fischer-Tropsch installation including the gas sampling points

The FT synthesis is operated fully automatically and several thousand hours of operation

were done in the past. In the first years the main work was optimisation of the gas treatment, which is now a level, that a FT catalyst can be operated for more than 1500 hours without any poisoning (no sulphur detected at the surface of the catalyst after operation).

The last work was a parameter variation with a cobalt based FT catalyst, where some results are given in Table 3.29

Parallel the hydroprocessing was investigated in cooperation with the refinery OMV. There the waxes were converted to gasoline and diesel with an excellent quality (CFPP below -60°C).30

At the moment the integration of the hydroprocessing into the FT synthesis is going on, parallel to the up-scaling of the technology.

113 Table 3. Results

16 bar 20 bar 24 bar H2/CO 1.5 1.6 2.0

CO conv. % 44.1 52.5 63.7 α value 0.89 0.90 0.89 C5+ % 90.6 91.7 90.2

Par/Ole 11.4 14.1 13.8

Mixed alcohols The above discussed synthesis are very sensitive to poisoning e.g. by sulfur. One synthesis which is very robust, as it uses a molybden sulfide catalyst, is the production of mixed alcohols. To investigate this synthesis with real synthesis gas from wood, a slip stream of 5 Nm³/h from the biomass CHP Güssing is used.

The main components of the MAS pilot plant are – the steam reforming unit, – the gas drying unit, – the gas compression unit, – the MAS fixed-bed reactor and – the product separation unit.

The technical equipment of the mixed alcohols plant is shown in Figure 10. Here the first experiments were done in spring 2011 and the first mixed alcohols were produced.31 As the results do not fit to literature data, here no detailed results were published till now.

Figure 10. Flow chart of mixed alcohol synthesis

Hydrogen from biomass Refineries are searching for alternative input material to reduce CO2 emissions. As actual refineries are using a high amount of hydrogen, this is the most economic option to bring in renewables into a classical refinery.

In a study leaded by OMV the technology to produce hydrogen for a refinery was investigated. Based on the indirect gasification process developed by the Institute of Chemical Engineering a flow chart of the hydrogen production process was developed and the mass- and energy balances were calculated.32

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Figure 11 shows the process design for the generation of high purity hydrogen for the integration in a refinery. The displayed process design is based on a dual fluidized bed steam gasification system which converts solid biomass into syngas with high hydrogen content of about 40 vol.-%.

Downstream of the gasifier the produced syngas is cooled in a heat exchanger and cleaned in a bag filter where the dust is separated from the syngas stream. After that, the syngas is mixed with the recycled PSA tail gas. Subsequently, the gas stream is heated up to a temperature of about 300°C before entering a CO-shift unit. Herein, the CO content of the syngas is converted with steam to hydrogen and CO2 by a sour gas CO-shift-catalyst.

When the syngas has passed the shift stage the gas stream enters a scrubber operated with rape seed oil methyl ester (RME) followed by a scrubber operated with water. In these steps, condensable hydrocarbons (tars) are removed and water is condensed. The scrubber fluids are to be cooled to low temperatures to ensure the necessary separation efficiencies and to protect the subsequent PSA unit against condensates. After that, the syngas is compressed to a pressure of 23 bar and 85% of the CO2 is removed in a pressurized water scrubber unit. After all these steps, hydrogen is recovered from the syngas using pressure swing adsorption (PSA). PSA units reach high separation efficiencies and deliver hydrogen with high purity at high pressure.

An important aspect for high overall plant efficiency is the utilization of the left over tail gas from the PSA unit. The remaining tail gas has a high content of CH4, C2H6 and C3H8. As can be seen in the displayed process design, a part of the tail gas (~35%) is used as additional fuel for the combustion chamber of the gasification system.

The main part of the tail gas is fed into a steam reformer and recycled into before the CO-shift unit to raise the overall hydrogen output of the plant.

Figure 11. Flow chart of hydrogen production The economic evaluation of the process shows that it is feasible and therefore in a next step

the whole chain will be piloted in a slip stream of the biomass CHP Güssing.

Conclusions By all these R&D projects the biomass CHP Güssing developed to an international known platform for synthesis gas applications. As the biomass CHP is for 7000 hours per year in operation, synthesis gas in different qualities (after gasifier, after filter, after scrubber) can be easily used for R&D. The Institute of Chemical Engineering and Bioenergy 2020+ have there a platform and also the necessary environment to do R&D on the different possibilities for the further usage of the product gas from biomass and the whole chain from wood chips till the final usage of the product can be investigated. So in the R&D work of BioSNG and FT the whole chain from wood chips till usage of the product in engines was investigated and this will be also done in future with other syntheses.

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INVESTIGATIONS ON AGGLOMERATION BEHAVIOR OF BED MATERIALS IN FLUIDIZED-BED COMBUSTION

For fluidized bed combustion systems stabile fluidizing conditions are of importance. Agglomerations generated during the process may influence the fluidization and can even cause problems in the operating and damages on the combustion facility. The knowledge of the agglomeration potential of different bed materials is necessary. The generation of agglomerates is influenced by different parameters: the operating condition, type of fuel and of possible additives. At standard operating conditions the agglomeration formation is avoid by the movement of the particles, but in case of an emergency shut-down the risk of agglomeration is much higher, this may preclude a restart of the bed. Elements (like Ca, Mg, Al, Fe, especially K) added by the fuel and the additives or are even in the bed material, are able to build compounds which deposit on the bed material particles and may cause sintering and agglomeration as well. Therefore the fuel quality, the choice of bed material and the additives are of importance. Investigations on agglomeration using different combustors, materials additives and fuels have been carried out by several research groups.33–38

This research deals with the influence of two different fuels on the bed agglomeration during emergency shut-downs, whereby two different silica sands were investigated. The first sand (called sand A) shows a grain size distribution of 560 to 2000 µm, the second one (sand B) has a distribution of 800 to 1600 µm and shows a lower potassium content than sand A. The used fuels are: mineral coal (low K content) with a grain size from about 5 to 10 mm and paper-mill sludge (dried, higher K content). The sludge particles showed an irregular shape and are deformable, so grain size was estimated to lower than 10 mm.

The experiments were carried out in a laboratory-scale fluidized bed reactor (refer to Figure 12), which is electrically heated up to 900° C (tube diameter: 70 mm, maximum volumetric flow 950 Nl/min). To simulate an emergency shut down during a typical combustion process, coal or sludge was inserted to the fluidized bed at about 800°C. During the combustion the temperature rises up between 900 and 1000°C. After 5 minutes of combustion the fluidization air supply was shut-down, whereby the heating shells still delivered heat to the system to simulate the heat capacity of the bed and the combustor walls.

Figure 12. The laboratory-scale fluidized bed reactor

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Figure 13 shows different types of agglomerations after different shut-down residence times.

For the experiments a to d sand A was used as bed material for e and f sand B. The black particles in a, b, e, f are coal-ash particles, which are formed during the shut-down combustion process, whereby in the cases a and b the temperatures raised up to 1200° C in case of e and f the temperatures remained around 950°C. This temperature rise during the emergency shut-down experiments is just occurring if sand A was used. The temperature has strong influence on the agglomeration process, at higher temperatures the ash may soften or even melt and can so generate agglomerations. In experiment c the white particles are sludge-ash-sand agglomerates. These occurred during the shut-down combustion experiments and are destroyed if the fluidization is restarted, because of the quite week connection. During the experiment d more stable sand agglomerates were formed.

Figure 13. Different agglomerations generated during the shut-down experiments; a: sand A, coal , 30 min; b: sand A, coal, 60 min; c: sand A, paper mill-sludge, 10 min; d: sand A, coal, 60 min; e: sand B, coal, 10

min; e: sand B, coal, 60 min

Table 4. Agglomeration types generated during different experiments

bed material agglomeration type time 10 min 30 min 60 min

sand A none, coal-particles ash-agglomeratesash-agglomerates

sand-agglomerates

sand B none, coal-particles ash-agglomerates ash-agglomerates

Table 4 shows a summary of the agglomeration behavior at different experiments. Generally

can be said Sand A shows higher potential of building agglomerations than sand B, whereby the regime has to be quite raw (highest temperatures and longest shut-down times). This behavior may change if the sand is used for longer terms and ash is cumulated in the bed, when higher amounts of the elements mentioned before, especially potassium, are existing. But as general conclusion the agglomeration potential of sand A is higher than of sand B.

CATALYTIC CRACKING OF FEEDS FROM BIOLOGICAL SOURCES IN A CFB Fluid catalytic cracking (FCC) is one of the most important refinery technologies. Long chain hydrocarbons (usually hydrated vacuum gasoil VGO) are converted to fuels in the diesel and gasoline boiling range and light hydrocarbons for the petrochemical industry. Feedstocks from renewable biological sources are gaining interest because of decreasing crude oil reserves and

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subsequently increasing oil prices as well as the need to reduce fossil CO2 emissions. Therefore, the suitability of different vegetable oils and other biogenous liquids as feedstock for the FCC process were tested at TU Vienna (refer to 39–44).

The experiments are conducted in a fully continuous small scale FCC pilot plant (see Figure 14) with internal CFB design. In the feed inlet zone the preheated feed comes in contact with the hot catalyst and evaporates instantaneously. The cracking reactions take place in the riser with short contact times of about 1s. The products leave the reactor at the top while the catalyst (deposited with formed coke) flows to the regenerator where the coke is burned with air. Liquid products are condensed and analyzed via a simulated distillation in a GC. The gasoline composition as well as RON and MON are determined by an FTIR fuel analyzer. The cracking gas composition is analyzed by a GC.

Heating system

Inert gas N2

Dry pressured air

Regenerator zone

Siphon

Feed inlet zone

Return flow tube

Particle separator

Riser

Inert gas N2

Flue gas

Product gas

Oil- feed

Preheating oven

Figure 14. Scheme of the small FCC pilot plant

Extensive investigations were carried out with rapeseed oil, soy oil, palm oil, sunflower oil,

and jatropha oil. The vegetable oils were added to VGO in steps of 20 wt.-% up to 100 wt.-%. Thorough analysis of the basic applicability of these vegetable oils in an FCC-unit in terms of cracking reactions and subsequent effects on catalyst circulation and obtained product distribution were performed. Special consideration was given to the effect of the vegetable oils on the octane numbers of the obtained products in the gasoline boiling range. Pure vegetable oils yielded more than 40 wt.-% gasoline and 20–30 wt.-% cracking gas at 550°C. Oxygen contained in the feed reacts mainly to form water and smaller amounts of CO and CO2. The gasoline is oxygen free at high octane numbers. The cracking gas contains high amounts of valuable gaseous olefins for the polymer industry.

Furthermore, new possible feedstocks which are generally available for lower prices like fatty acids, waste vegetable oil, tall oil and pyrolysis oils were tested. Some of these feeds form higher amounts of coke during the FCC process which leads to limitations in the regenerator capacity of the FCC pilot plant. Therefore, a new FCC pilot plant with higher capacity ratio between regenerator and riser reactor was designed and constructed. To increase the flexibility of the system regarding low temperature cracking to higher boiling products a special heat exchanger system was implemented.

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IRON ORE REDUCTION KINETICS UNDER FLUIDIZED-BED CONDITIONS Iron ore reduction kinetics were studied at conditions relevant to the FINEX® and FINMET® process (refer to 45).

The FINEX® process is based on a series of four fluidized bed reactors followed by a melter-gasifier. This process allows the direct utilization of fine iron ores and coal. The operating pressure is about 3.5 bar and the temperatures are up to 850°C. The intermediate product which is direct reduced iron is briquetted to hot compacted iron (HCI). The HCI is fed into to melter-gasifier. The final product of the melter-gasifier is liquid iron. The gas which is used for the prereduction of the iron ore fines is from the melter-gasifier and rich in CO, besides H2, CO2, CH4 and others.

In the FINMET® process, iron ore fines are reduced by H2-rich reduction gas in four serial fluidized bed reactors. The reduction gas is provided by natural gas steam reforming and led in counter current mode with the iron ore. After the ore is dried, it is conveyed to the hopper system, where the pressure is increased to operating conditions, and it is then moved into the reactor system, passing through the reactor cascade. The final product is hot briquetted iron (HBI) by briquetting hot reduced iron ore fines. The operating pressure is about 13 bars and the temperature ranges from 450 up to 800°C depending on the specific reactor.

For optimized plant performance, detailed knowledge about reaction kinetics and thermodynamic stability of iron ores is necessary.

Experimental To be able to perform experiments under industrial plant operating conditions, a laboratory-scale pressurized fluidized bed reactor was built after the “reactor in pressure vessel” principle and according to the chemical similarity rules. The set up of the reactor system shown in Figure 16 facilitates pressures up to 10 bars in a temperature range from 400 to 900°C with reducing gas mixtures containing H2, CO, H2O, CO2, CH4 and N2.

For kinetic measurements, the reactor system is equipped with a sampling system, which allows sampling of the bed material at operating conditions and certain residence times. The obtained samples were analyzed by the iron-chloride method on the degree of reduction and by mercury porosimetry on specific surface area, mean pore diameters and porosity. Furthermore polished sections of selected samples were investigated by light optical microscopy. For the

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distinction between magnetite and wuestite the samples were etched with diluted hydrochloric acid.

Figure 16. Flow scheme of the laboratory-scale pressurized fluidized bed reactor

Results Typical iron ore reduction results can be seen in Figure 17. The degree of conversion, the porosity, the specific surface area as well as the mean pore diameter are shown. For further information refer to 46–53.

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Figure 17. Experimental results: Reduction of magnetite to wuestite (refer to 46)

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