energy technology volume 1 issue 4 2013 [doi 10.1002%2fente.201300009] kern, stefan; pfeifer,...

DOI: 10.1002/ente.201300009

Gasification of Low-Grade Coal in a Dual Fluidized-BedSteam GasifierStefan Kern,* Christoph Pfeifer, and Hermann Hofbauer[a]

Introduction

In addition to renewable fuels, such as biomass, and theirthermal utilization,[1] coal will continue to play a major rolein the global primary energy mix to satisfy increasing energydemands; in contrast, the contribution of oil will decreaseafter peak oil is reached. Huge amounts of coal are currentlyavailable; the reserves that can be mined at low cost werecalculated in 2007 to be sufficient to last for at least another148 years.[2] Although coal-based gasification processes forpower generation have not been proven to have an economicadvantage in comparison to combustion power plants thatuse coal or other fossil fuels, the investment costs for inte-grated gasification and combined cycle (IGCC) power plantsare much higher, and power generation by coal gasificationretains certain advantages. For instance, the electrical effi-ciency is greater[3] and if the process is combined with carboncapture and storage (CCS) the costs are lower in terms ofboth initial investment and operation.[2] For the gasificationof coal, commercially available reactor designs include en-trained-flow, moving-bed, and fluidized-bed gasifiers. In con-trast to biomass gasification, for which fluidized-bed reactorsare used in most industrial-sized plants, for coal gasification,the most commonly employed reactor design is currently theentrained-flow reactor.[4] This type of reactor offers the bene-fits of a gas that is free of condensable hydrocarbons (tars)as a result of its operation at very high temperatures (12001600 8C) and it is also able to operate with liquid fuels. In ad-dition, as a result of the high temperatures and fine fuel par-ticles, the carbon conversion to gas is nearly complete ina single stage. The most significant disadvantages of these re-actors are that the fuel has to be ready to be ground toa small particle size (which is not possible for some fueltypes) and that pure O2 is required to drive the process auto-thermally, which means that the heat for gasification is pro-vided by partial oxidation in the gasifier. To maintain a nitro-gen-free product gas, steam or CO2 can be used as a gasifica-tion agent. The advantage of using steam instead of CO2 is

that the reactivity of steam is on average approximately fourtimes higher than that of CO2,

[5] therefore, residence times ofthe char in the gasification section are shorter and the gasifi-cation efficiency is higher. As the use of steam or CO2 asa gasification agent results in an allothermal process, theheat required for endothermic gasification reactions has tobe provided externally. As a solution to the problem of of in-troducing such heat externally on an industrial scale, dualfluidized-bed (DFB) gasification technologies can be utilized,which separate the combustion zone (which provides theenergy for gasification) from the gasification zone, and puresteam is used as the gasification agent. By circulating the bedmaterial between these two reactors, the heat from the com-bustion reactor is carried to the gasification reactor. As withsingle-stage fluidized-bed reactors, this two-stage processneeds to be operated at temperatures below the ash-soften-ing temperature of the fuels. As a result, the highest temper-ature that can be reached by fluidized-bed gasification is ap-proximately 1000 8C. DFB steam gasifiers usually operate atgasification temperatures in the range of 800900 8C. Thelower gasification temperatures employed during fluidized-bed gasification with respect to entrained-flow gasificationlead to lower carbon conversions[6] and residual char is oftenleft behind in a single-stage reactor. As a consequence,highly reactive coal with a comparably high content of vola-tile components, such as lignite, is favored.[4] Through the useof a separate combustion reactor, the residual char can beused to provide the heat for gasification as it is applied inthe DFB gasifier.

The gasification of a bituminous coal was performed ina dual fluidized-bed (DFB) steam gasifier at a fuel power of90 kWth. The main aim of these tests was to prove the suita-bility of the DFB system for cheap low-grade coals that areoften delivered with a wide range of particle sizes and ashloads. The fine fraction in particular can become entrained inthe gas stream and thus be responsible for changes in the

system performance and gas quality. The selected gasifierbed material was olivine, with an applied particle size of375 mm, the steam-to-carbon ratio was 1.3 kgH2OkgC

1, andthe gasification temperature was 850 8C. In addition to stan-dard online analysis of the main components of the productgas, impurities that included NH3, H2S, and tar were also ex-amined.

[a] S. Kern, Dr. C. Pfeifer, Prof. Dr. H. HofbauerVienna University of TechnologyInstitute of Chemical EngineeringGetreidemarkt 9/166, 1060 Vienna (Austria)Fax: (+43)15880116699E-mail: [email protected]

[email protected]

Energy Technol. 2013, 1, 253 264 2013 Wiley-VCH Verlag GmbH&Co. KGaA, Weinheim 253

Another approach to coal gasification is the production ofvaluable products (e.g., FischerTropsch fuels, syngas, metha-nol) from cheap, low-quality coal. Coals with very high ashcontents are usually considered undesirable as an increase infuel ash leads to a decrease in gasification and combustionefficiency because as the heat in the ash cannot be fully re-covered. Furthermore, a high ash content can lead to foulingand increase the amount of waste for disposal.[4] Neverthe-less, if the ash content is not high enough to cause the above-mentioned problems, ash can actually be beneficial for theprocess if it contains catalytically active matter. Although al-kaline metals often cause problems with ash melting, theyare also suitable to catalyze gasification and reforming reac-tions. In this regard, the following order of activity has beendetermined: K2CO3> Na2CO3> trona [Na3(CO3)-(HCO3)2H2O]>borax (Na2B4O710H2O).

[7] Calcium hasbeen identified to play a major role in the promotion ofwatergas-shift reactions if present as CaO. Clemens et al.found that for gasification with steam, the Ca content of coalash both promotes the watergas-shift equilibrium and influ-ences the composition of the product gas.[8] This effect is pro-duced if the gasification products remain in the vicinity ofthe char surface, which indicates that improved gaschar con-tact promotes the catalytic activity. In their comparison ofCaO and MgO for the steam reforming of biomass tars, Del-gado et al. determined an activity series in which dolomitewas the most active material followed by magnesite and cal-cite.[9] CaO is, in addition to MgO, one of the major compo-nents of dolomite, a cheap and disposable catalyst commonlyused as an in-bed, catalytically active additive that containsSiO2, Fe2O3, and Al2O3 as trace minerals.

[10] Low-qualitycoals are also often characterized by high sulfur contents,which is significant given that sulfur compounds must be re-moved from the product gas in the majority of cases. The pri-mary form of sulfur in the product gas is H2S,

[11] althoughminor concentrations of COS are also found. CS2 is generallya product of secondary reactions of COS and H2S, but onlyto a negligible extent as these reactions occur above 850 8C.As a result, CS2 formation is only a matter of interest in en-trained-flow gasification, in which these high temperaturesare observed, and not in fluidized-bed gasification.[12]

To examine the potential for the DFB gasification of high-ash coals, a coal with an ash content of 31 wt%db was usedfor gasification with steam at a fuel power of 90 kWth. In ad-dition to its chemical properties, the fuel particle size hasa significant influence on the process performance, in partic-ular the emission of char and dust in the product gas as wellas the formation of tar and the watergas-shift equilibrium.This is significant as the particle size of low-quality coalsoften varies considerably on delivery, and a high coal-dustcontent can lead to the entrainment of particles in the gasproduced by a fluidized-bed gasifier. To assess the influenceof fine particles, two gasification tests were performed underthe same process conditions but with different particle sizedistributions.

Gasification of solid feedstock with steam

For the gasification of a solid carbonaceous feedstock withpure steam, the influence of the amount of steam present inthe reactor is essential for system performance and productgas quality.[13] Therefore, it is mandatory to maintain thesame amount of steam for the gasification of solid carbon forall operating points. For gasification with pure steam, theoverall gasification reaction can be expressed as:

CxHy xH2O! xCO xy2

H2 1

DHR, 850 > 0

Equation (1) reveals that for each mole of carbon in thefeedstock, one mole of water is required for stoichiometricfuel conversion. However in reality, the main elements ofa feedstock include carbon, hydrogen, and oxygen. Inclusionof this in the overall steam gasification reaction results inEquation (2):

CxHyOz x z H2O! xCO xy2 z

H2 2

for x > z; DHR, 850 > 0

Based on Equation (2), theoretically, the stoichiometricdemand of steam is equal to xz if complete conversion ofthe solid fuel (CxHyOz) is assumed. By using the equationsabove, it can be estimated that a different amount of steampresent in the system can push the reaction towards theproducts or reactants. Therefore, the amount of steam intro-duced into the gasification reactor is referred to as the intro-duced amount of fuel or the introduced amount of carbon bythe solid fuel. These ratios are the steam-to-fuel ratio (fSF)[Eq. (3)] and the steam-to-carbon ratio (fSC) [Eq. (4)].

fSF;wt _msteam vH2O _mfuel

1 vH2O vash

_mfuel3

fSC;wt _msteam vH2O _mfuel

vC _mfuel4

Equation (4) can be employed to determine the minimumamount of steam that has to be present during the processfor a known fuel composition. Equation (5) can then be usedto calculate the stoichiometric steam demand fH2O:

H2O x z 5

In the tests performed here, the feedstock can be ex-pressed as the system CxHyOz, which is free of sulfur and ni-trogen. From the fuel analysis (Table 1), the molarities of C,H, and O were determined as CH0.73O0.12 and fH2O as60.04 molH2Okgdaf,N,S,Cl-free

1 (1.08 kgH2Okgdaf,N,S,Cl-free1).

In most cases of gasifier operation, the actual amount ofsteam in the system differs from that which is stoichiometri-cally required. To quantify the ratio of steam present to that

254 www.entechnol.de 2013 Wiley-VCH Verlag GmbH&Co. KGaA, Weinheim Energy Technol. 2013, 1, 253 264

S. Kern et al.

which is theoretically needed, an equivalent ratio for steamcan be defined as that used for air combustion systems(lH2O).

lH2O _mH2O;actual_mH2O;stoich:

6

As an estimate of the potential of a fuel to form the mainproduct-gas components, the dependence of the calculated[14]

equilibrium compositions on temperature is shown inFigure 1. This graph illustrates a system CxHyOz+fH2OH2O,in which, according to fuel analysis, the molarities of thefeedstock (C, H, and O) are used as reactants together withthe stoichiometric amount of water required for the conver-sion to gas. The thermodynamic equilibrium composition iscalculated by minimization of the Gibbs free energy.[15]

By comparing the stoichiometric equilibrium values(Figure 1) at 850 8C with the measured values presented later(Figure 3), it can be estimated that a thermodynamic equilib-rium will likely not be reached in the gasification reactor asthe residence time is too low at the relatively low tempera-tures used for coal gasification. Therefore, the product-gascomposition in entrained-flow gasifiers is closer to the ther-modynamic equilibrium.

Results and Discussion

Overview

Coal gasification was performed at two operating points(OP1, OP2), the general parameters of which are summar-ized in Table 2. OP1 and OP2 differed only in the fuel parti-cle size distribution, and the other process parameters, such

as gasification temperature, bed material particle size, andfSC, were kept constant. To provide the required heat forgasification at approximately 850 8C, the combustion reactorwas operated at a temperature slightly below 900 8C in bothcases. For OP1 the gasifier was operated with the fuel as de-livered (only pieces larger than 10 mm were sieved out toavoid loop-seal blockage), whereas for OP2 the fine fractionwas reduced by sieving. The two tests were performed sepa-rately, which means that both OPs had to go through thestart-up period of the system until the operating conditionswere reached. These steady-state operating conditions werekept constant for at least 6 h to collect sufficient data formass and energy balances as well as to accomplish the dis-continuous sampling methods. For both OP1 and OP2, thefuel was introduced into the gasifier by using hopper 1, withthe coal introduced directly into the bubbling bed at mediumheight by using a screw conveyor. The particle size distribu-tions of the two coal size fractions are listed in Table 3. Animpression of the behavior of coal gasification by using theDFB system can be achieved by highlighting the tempera-tures in the gasifier at the different operating points as wellas the fluidization conditions in the system. An overview ofthe fluidization regime is presented in Table 4. The velocitiesUmf, Ut, and their ratios to the actual superficial velocity inthe gasification and combustion reactors were used to char-acterize the fluidized-bed system. The process conditionspresented in Table 4 and the fuel particle size distribution

Table 1. Proximate and ultimate analysis of the feedstock.

ContentComponent Dry basis As used

Water content [wt%] 3.6Ash content [wt%] 31.09 29.97C [wt%] 55.13 53.15H [wt%] 3.39 3.27N [wt%] 1.39 1.34O [wt%] 8.18 7.88S [wt%] 0.78 0.75Cl [wt%] 0.04 0.04Volatile matter [wt%] 27.16 26.18Fixed carbon [wt%] 72.84 70.22LHV [MJkg1] 21.00 20.15

Figure 1. Calculated equilibrium product gas composition of the systemCxHyOz+fH2OH2O for the used coal (corresponding to fuel analysis inTable 1).

Table 2. Key parameters of the gasification tests.

Value

Fuel size [mm] (coal) OP1: 0.38OP2: 1.58

Gasification temperature [8C] (bubbling bed) 8482Mean temperature combustion reactor [8C] 8972Fuel power [kW] (coal) 90Fuel mass flow [kgh1] (coal) 16.2fSC [kgH2Okgcarbon

1] 1.30fSF [kgH2Okgfuel,daf

1] 1.05Initial bed material inventory [kg] 100

Table 3. Coal particle size distribution.

Value OP1 OP2

dp10 [mm] 280 1400dp50 [mm] 2600 4990dp90 [mm] 7700 7860

Energy Technol. 2013, 1, 253 264 2013 Wiley-VCH Verlag GmbH&Co. KGaA, Weinheim www.entechnol.de 255

Gasification of Coal in a Dual Fluidized-Bed Steam Gasifier

listed in Table 3 illustrate the problem caused by the pres-ence of fine particles in a fluidized-bed reactor. The gasifica-tion of the fuel particles should take place in the bubblingolivine bed, but to maintain the fluidization regime for a bub-bling bed, a minimum fluidization velocity is required. In thecase presented here, coal gasification was accompanied bya superficial velocity in the gasifier Ug of approximately0.4 ms1.At this fluidization velocity, the terminal velocity Ut of the

olivine particles is far from a value that results in their elu-triation into the product gas, even for smaller particle sizes(dp10). For the fuel particles, however, the situation is differ-ent, as the density of the coal is lower and the particle size iswidely distributed. This is particularly true of the fuel used inOP1, in which some of these particles are entrained in theproduct gas flow. The critical size for coal in the appliedgasification atmosphere is approximately 160 mm, and coalparticles that are smaller than this are elutriated immediatelyafter feeding into the gasifier. For the coal size distributionused in OP1, for which the fine particles were not sieved out,approximately 6.8 wt% of the fuel particles were smallerthan 160 mm. In the case of OP2, for which most of the finefraction was removed, only 2.0 wt% of the initial mass leftthe bubbling bed to enter the gas phase immediately. Howthese particles, which are generally neither completely devo-latilized nor gasified, affect the temperature profile and gasquality will be discussed below.The temperature profile in thegasification reactor is presentedin Figure 2. This diagram illus-trates that the temperaturevaries with the height of the re-actor (relative to the surface ofthe bubbling bed) and betweenthe two OPs as a result of sev-eral factors. The hot bed mate-rial, which transports the heatfrom the combustion reactor tothe gasification reactor, is recy-cled back to the gasification re-actor through the lower free-board section. This producesa local hot spot at which thehighest temperature in the reac-tor is recorded. A cold spot is

produced where the superheated steam for fluidization is in-troduced into the bubbling bed because the steam tempera-ture, which is approximately 300 8C, is significantly lowerthan that of the bed material. However, as a result of the ef-fective intermixing of the steam, bed material, and fuel parti-cles as well as the high heat capacity of the bed material, thetemperature loss is limited. The set point for the gasificationtemperature (850 8C) was chosen at the height of the middleof the bubbling bed, which is where the fuel was also fed. Inthe middle and upper sections of the freeboard, the tempera-ture decreases steadily because of heat losses and energy-consuming reactions as well as the pyrolysis and gasificationof any entrained fine fuel particles in the freeboard. By com-paring the two OPs, we can see that the temperature profileexhibits a greater decrease if finer particles are present inthe freeboard. At the gasifier outlet situated at the end ofthe freeboard, the temperature difference between the twoOPs is approximately 10 8C. This discrepancy is caused bythe higher amount of fine coal particles in the freeboard forOP1 that are heated, release their volatile components, andare partly gasified. A closer analysis of Figure 2 reveals thatthe major difference in temperature between the two OPsoccurs in the splash zone and the lower freeboard section;therefore, we can assume that the fast pyrolysis reactionstake place here, whereas the relatively slow char gasificationtakes place above. Such behavior has already been docu-mented in this field.[16] There is a demand for additional fuelfor the combustion reactor. For OP1, with fine particles inthe fuel, the demand for additional fuel for the combustionreactor is higher than that for OP2. Consequently, the tem-perature of the recycled bed material differs slightly andcauses a deviation in the reference point of the bubbling bedby 4 8C.

Product gas composition

The DFB gasification process yields two separate gasstreams, a product-gas stream (gasification reactor) and

Table 4. Fluidization conditions in the DFB system.

Value OP1 OP2

Umf [ms1] 0.07 0.07

Ut [ms1] 3.36 3.36

Ug [ms1] 0.40 0.41

Ug/Umf 5.51 5.64Ug/Ut 0.12 0.12Uc [ms

1] 9.42 9.23Uc/Umf 130.79 128.19Uc/Ut 2.80 2.74

Figure 2. Temperature profiles during gasification in the gasification reactor.


S. Kern et al.

a conventional flue-gas stream (combustion reactor). Theproduct gas primarily consists of H2, CO, CO2, CH4, C2H4,C2H6, and unconverted H2O. The remaining gas componentsare made up by the following components or impurities:

* A N2 content of

authors argued that the main reason for this difference wasthat approximately 9% of the mixed fuel, mainly sawdust,was transported immediately out of the fluidized bed withthe fluidization steam. As a result, pyrolysis and gasificationtook place in the freeboard section of the gasifier, in whichthe interaction of the pyrolysis products with the hot, catalyt-ically active bed material was absent. This process ultimatelyled to a significant reduction in tar decomposition. Anothercommon interpretation of the higher tar content of fine par-ticles is that they provide less resistance for the devolatilizinggases; in larger particles the pyrolysis gas has to pass a layerof char, and thus secondary catalytic tar cracking reactionscan take place.[23] Table 5 summarizes the relative contribu-tion of each tar compound in the GCMS-detectable tar. Asthe origin of the feedstock used in both OP scenarios wasthe same, the actual composition of the GCMS-detectabletar did not differ much when small fuel particles were en-trained in the freeboard. Nevertheless, a tendency towardsclasses of tar produced by devolatilization can be imaginedfor OP1, as evidenced by the lower contribution of naphtha-lene, a tar component that is mostly a product of secondaryand tertiary tar reactions.The high ash content of the coal makes it necessary to con-

sider its effect on the system. The applied (wet) fuel feedingrate of 16.2 kgh1 corresponds to the introduction of4.86 kgh1 of ash into the gasifier. According to the measure-ments of dust concentration in the product gas (Figure 4),the product gas stream, and the total ash in the product gasand flue gas streams, the balance of the ash is closed. Fig-ures 6 and 7 show the mass flow [kgh1] for inorganic matterthat enters and leaves the gasification reactor, respectively.The most important issue to consider is that the entrainedparticles also include a large amount of inorganic matter. Aslarger coal particles also likely contain larger (sand/gravel)ash particles, it is not surprising that a total of 3.8 kgh1 ofash left the DFB system through the product and flue-gasstreams for OP1, whereas for OP2 only 1.52 kgh1 of ashwas observed. In addition, as the mean ash particle size in-creased after fuel sieving, the terminal velocity Ut of most ofthe ash particles was increased above the superficial gas ve-locity Ug in the gasification reactor. Only the fraction thatwas sufficiently small was able to escape the system. Hence,

for OP1 only 1.25 kgh1 of ash accumulated in the system,whereas for OP2 3.53 kgh1 was unable to leave the fluidizedbed as no ash-removal device was available at the pilot

Figure 5. Amounts of detected tar in the product gas.

Table 5. Relative contribution of tar components to GCMS tar.

Component OP1 OP2

Naphthalene [wt%] 29.62 31.35Acenaphthylene [wt%] 13.07 13.22Indene [wt%] 7.77 7.04Anthracene [wt%] 8.58 8.59Fluoranthene [wt%] 6.46 5.90Fluorene [wt%] 4.30 3.78Phenanthrene [wt%] 3.92 3.38Dibenzofuran [wt%] 3.84 3.91Pyrene [wt%] 3.35 2.874,5-Methylphenanthrene [wt%] 1.94 1.63Carbazole [wt%] 1.68 1.712-Methylnaphthalene [wt%] 1.54 1.49Benzo[a]anthracene [wt%] 1.49 1.391-Benzothiophene [wt%] 1.47 1.36Chrysen [wt%] 1.40 1.37Quinoline [wt%] 1.46 1.30Biphenyl [wt%] 1.00 1.19Styrene [wt%] 0.91 0.82Diethyl phthalate [wt%] 0.89 0.941-Methylnapthalene [wt%] 0.79 0.66Benzo[a]pyrene [wt%] 0.75 0.76Benzo[b]flouranthene [wt%] 0.71 0.85Benzo[k]flouranthene [wt%] 0.55 0.70Isoquinoline [wt%] 0.49 0.51Phenol [wt%] 0.45 0.65Phenylacetylene [wt%] 0.39 0.35Benzofuran [wt%] 0.33 0.32Acenaphthene [wt%] 0.31 0.34Benzo[g,h, i]perylene [wt%] 0.29 0.73Indole [wt%] 0.27 0.28Others [wt%] 0.0 0.6

Figure 6. Ash balance for OP1 [kgh1] .

Figure 7. Ash balance for OP2 [kgh1] .


S. Kern et al.

plant. However, the operation of the DFB pilot plant wasnot jeopardized by accumulating ash as the system was in op-eration for less than 8 h with coal as a feedstock. Neverthe-less, a slight pressure increase in the bubbling bed was re-corded for OP2, which demonstrates the increased amountof inorganic matter in the system. For operations on an in-dustrial scale, an ash-removal system would be required asthe large amount of fuel ash can easily lead to bed-materialagglomeration if species that lower the ash-melting tempera-ture exceed a critical concentration level. Therefore the com-positions of the fuel ash and the ash collected from the gasstreams that exit the reactor were detected by X-ray fluores-cence (XRF) analysis. A benefit of fuel ash is its potentialfor catalytic activity. The main components in the fuel ashwere Si, Al, and Fe with minor amounts of Ca, K, P, Ti, andMg (Table 6). Although fuel ash can play a significant role inincreasing the catalytic activity, no enhancement was ob-served in this study in terms of tar reduction. This suggeststhat the large quantities of Si and Al measured here werenot highly catalytically active. In contrast, the minor compo-nents of K, Mg, and Ca are known to be active catalysts, al-though the extent of this activity depends on their com-pounds. In this study, it is probable that most of these alkalimetals were present as silicates and thus not very active. Themass flow rates of the measured elements into and out of thereactor in the fuel and as particles in gas are summarized inTable 7. The higher output of Mg compared to its inputlikely reflects attrition of the bed material.

Product gas yield and conversion performance

This section focuses on the conversion of the reactants re-quired for the gasification process, carbon and water, as wellas system performance. Table 8summarizes the key data fromthe gasification tests. The prod-uct-gas yield was greater in thetest without fine coal particles(OP2) than that in OP1, whichlikely reflects the longer meanresidence time of the fuel frac-tion in the gasifier and wouldlead to an increased carbonconversion in the gasificationreactor. The operation ofa DFB gasifier is characterizedby the carbon conversion thattakes place only in the gasifica-tion reactor itself and thecarbon conversion that occursin the whole system (gasifica-tion reactor and combustion re-actor together) through to theformation of product gas andflue gas. In the gasification re-actor, the carbon conversioncan be defined as the ratio of

Table 6. XRF analysis of coal and fly ash.

Component Coal Fly ash

SiO2 [wt%] 55.617 51.908Al2O3 [wt%] 21.392 22.066Fe2O3 [wt%] 10.977 13.368K2O [wt%] 3.086 2.930CaO [wt%] 3.402 2.270MgO [wt%] 1.111 3.948TiO2 [wt%] 1.330 1.126V2O5 [wt%] 0.153 0.136Cr2O3 [wt%] 0.159 0.170MnO [wt%] 0.158 0.133Na2O [wt%] 0.375 0.253NiO [wt%] 0.122 0.224ZnO [wt%] 0.147 0.186Others [wt%] 1.971 1.283

Table 7. Mass flow rates of ash components.

Component Input feedstock Output

SiO2 [gh1] 2808.66 789.00

Al2O3 [gh1] 1080.28 335.40

Fe2O3 [gh1] 554.34 203.19

K2O [gh1] 155.83 44.53

CaO [gh1] 171.82 34.50MgO [gh1] 56.11 60.00TiO2 [gh

1] 67.17 17.12V2O5 [gh

1] 7.71 2.07Cr2O3 [gh

1] 8.03 2.58MnO [gh1] 7.99 2.02Na2O [gh

1] 18.96 3.85NiO [gh1] 6.15 3.40ZnO [gh1] 7.40 2.83Others [gh1] 99.55 19.50Sum [kgh1] 5.05 1.52

Table 8. Key data for the tests performed.

Value OP1 OP2

Total product gas yield [Nm3h1] 27.34 28.80H2O content of product gas [vol%] 57.03 51.39Product gas yield [Nm3dbh

1] 12.82 14.47Specific product gas yield [Nm3db kgfuel, daf

1] 1.19 1.34Lower heating value [MnNm3db

1] 11.59 11.78Syngas powerexcl. tar [kW] 41.28 47.34hCPP [%] 44.00 50.72hC;IP [%] 51.80 59.77Water conversion, XH2O, rel [kgH2O kgfuel,daf

1] 0.45 0.50Water conversion, XH2O, rel [kgH2O kgfuel,daf,N,S,Cl-free

1] 0.47 0.51Stoichiometric steam demand [kgH2Okgfuel, daf,N,S,Cl-free

1] 1.10 1.10Stoichiometric steam demand [molH2O kgfuel,daf,N,S,Cl-free

1] 61.16 61.16lH2O [kgH2OkgH2O

1] 0.42 0.47Carbon conversion in the gasification reactor, XC,G [%] 36.41 41.28Overall carbon conversion of the DFB system, XC,DFB [%] 93.13 96.04Specific tar content, GCMS [gkgfuel, daf

1] 9.66 5.99Specific tar content, gravimetric [gkgfuel, daf

1] 7.50 3.18Specific tar content, GCMS [gkgcarbon

1] 12.07 7.49Specific tar content, gravimetric [gkgcarbon

1] 9.38 3.98Tar intensity per kWh of syngas, GCMS [gkWhproductgas

1] 2.51 1.37Tar intensity per kWh of syngas, gravimetric [gkWhproduct gas

1] 1.95 0.73tF [s] 4.1 4.1tC [s] 0.8 0.8



carbon in the product gas stream that exits the gasificationreactor to the amount of carbon introduced in the solid fuel[Eq. (7)]. In contrast, the carbon conversion that occurs inthe whole system includes the ratio of carbon that exits theDFB system in both the product (gasification reactor) andflue-gas streams (combustion reactor) to that introduced inthe solid fuel [Eq. (8)].

XC;G _mC;PGvC _mfuel

7

XC;DFB _mC;PG _mC;FG

vC _mfuel8

The higher water content of the product gas in OP1 thanthat of OP2 is another indicator that the gasification agent,steam, did not react as effectively with carbon particles inOP1. The amount of introduced water consumed or convert-ed during the gasification and steam-reforming reactions canbe considered as an indicator for the whole process. Relativewater conversion is thus defined as the amount of water con-sumed per mass unit of converted fuel [Eq. (9)].

XH2O;rel _mH2O;conv:

1 vH2O vash

_mfuel9

According to this equation, the relative water conversionof OP2 was greater than that of OP1 (Table 8), which sug-gests that the residence time of the fine particles present inthe freeboard at OP1 (tF) was too short for them to reactsufficiently with the steam. However, the values for the rela-tive water conversion are relatively high in both cases com-pared to the gasification of biomass, in which relative waterconversions of approximately 0.15 kgH2Okgfuel,db

1 are usuallyreached[24] and for which the DFB system was originally de-signed. The source of these significantly increased values forcoal can be explained by two main reasons. In general, thestoichiometric water demand of the coal used in this study isapproximately four times higher than that of wood pellets,which makes it clear that gasification reactions that consumesteam dominate over the devolatilization process. Thesecond reason is that the amount of consumed water corre-sponds to the dry and ash-free fuel. As the input mass flowrates for coal are lower than those of wood at the correctfuel power and the ash content here was very high, the totalamount of dry and ash-free fuel was significantly lower. Theeffect on the water conversion rate by the addition of coal toa wood-gasification process in the same system has beendocumented in detail for the cogasification of wood withcoal.[25]

The temperature of the gasification reactor during stan-dard gasification operation is usually controlled by theamount of additional fuel added to the combustion reactor.Compared to the gasification of biomass, which contains lessfixed carbon than the fuel used here, the demand for addi-tional fuel for OP1 and OP2 is relatively low. This is becauseof the high amount of char in the bed that is consequently

transported with the bed material from the gasification reac-tor to the combustion reactor. Nevertheless, there was a sig-nificant difference in the additional fuel demand for the com-bustion reactor observed for the two OPs. As shown inFigure 8, there was more fuel required for OP1 to providethe heat for the gasification reactor. This was already esti-mated by the temperature trend shown in Figure 2, which in-dicated that the temperature of the bed material was lower ifit moved from the gasification reactor to the combustion re-actor during OP1.

An evaluation of the efficiency of the gasification systemwas performed by determining the cold gas efficiency. Thecalculation of this value takes into account that pilot plantsusually do not reach the low ratio of heat losses seen in in-dustrial plants. In the case of the present DFB pilot plant,heat losses amounted to nearly 20% of the fuel input.Stidl[26] calculated the heat losses that occur through radia-tion for the main components of the 10 MWth DFB gasifica-tion plant in Oberwart, Austria.[27] Based on this data, the ra-diation heat loss for a typical industrial plant can be assumedto be 2% of the input fuel power. The cold gas efficiencydata can be effectively used to calculate the parameters foran industrial-scale plant [Eq. (10)]. In contrast to this, thecold gas efficiency for the pilot plant can be expressed asshown in Equation (11), which neglects heat losses.

hC;IP _vPGLHVPG

Pfuel;G Pfuel;C _QPP _QIP 3600 10

hC;PP _vPGLHVPG

Pfuel;G Pfuel;C 3600 11

Conclusion

DFB gasification is a fuel-flexible technology capable of theconversion of carbonaceous feedstock into high-quality prod-uct gas. Even low-grade coal fuel can be successfully used asa feedstock, despite both its low reaction rate at typically ap-plied gasification temperatures (800900 8C) and its difficultyof handling because of its very high ash content and wide

Figure 8. Additional fuel required for the combustion reactor.


S. Kern et al.

particle size distribution. Reasonable product-gas quality andenergetic performance parameters were obtained. In reality,the main obstacle encountered in using this fuel was not itsash content, and the reduced process performance was in-stead caused by the low activity of the coal with the gasifica-tion agent, which led to less-efficient carbon and water con-version. In addition, if a large amount of fine particles arepresent in the fuel, a high level of dust and entrained char isproduced as the fine particles are entrained in the gas phaseimmediately after the fuel is fed into the gasifier. In thisstudy, the inclusion of fine fuel particles reduced the carbonconversion rate further because of the production of ungasi-fied carbon.In its current configuration, the system is limited with

regard to large fractions of fine material in the feedstock.However, a new design has been proposed[28,29] that willenable the conversion of fine material in contact with thehot, catalytically active bed material. The tests presentedhere have generally shown that the DFB-gasification processis suitable even for materials with a low conversion rate andhigh ash content.

Experimental Section

The DFB pilot plant at the Vienna University of Technology

For experiments performed on a pilot scale, the Vienna Universi-ty of Technology operates a 100 kW DFB-gasification reactor.The basic principle of the DFB-gasification process is shown inFigure 9, and a schematic drawing of the pilot rig is shown inFigure 10. The system physically separates gasification and com-bustion, with two fluidized-bed reactors connected by loop seals.The fuel, typically biomass but coal in this study, enters thegasification reactor, a bubbling bed fluidized with steam, inwhich drying, pyrolysis, and heterogeneous char gasification takeplace. The residual char leaves at the base of the gasification re-actor together with the bed material, which circulates betweenthe two reactors through the lower loop seal and into the com-bustion reactor. This reactor is implemented as a fast bed fluid-ized with air to maintain combustion of the residual char and ad-ditional fuel, if required. By burning char and additional fuel (ifrequired) in the combustion reactor, the bed material is heated

and, after particle separation from the flue gas at the exit of thecombustion reactor, it flows back to the gasifier through theupper loop seal.

The pilot plant is equipped with three different hoppers toenable fuel feeding at different positions into the gasification re-actor and to enable the possibility of cogasification of two ormore pure substances at any mixing ratio. During the experi-ments performed in this study, only hopper 1 was used to intro-duce the fuel into the bubbling bed. A more detailed descriptionof the pilot plant, including the geometrical data, can be foundin Refs. [13,15,24]. For safety reasons, the hopper system isflushed with N2. For standard gasification operation with a singlehopper used to feed the fuel, a volumetric flow rate of 0.1 m3h1

of N2 was used.

Analyses

Product gas measurement

The compositions of product and flue gases were measured aftertheir exit from the reactors, of which the main gas componentsCH4, H2, CO, CO2, and O2 were analyzed by using a RosemountNGA 2000 multicomponent gas analyzer, and N2, C2H4, andC2H6 were analyzed by using an online GC (PerkinElmer Clarus500). The detailed arrangement of the measurement system is ex-plained in Ref. [16].

Tar, H2O, and entrained particulate matter measurement

Tar was sampled isokinetically by using impinger bottles and an-alyzed gravimetrically by using GCMS. Tar sampling was ap-plied discontinuously by condensing and dissolving the tar com-ponents. The measurement method was based on the tar protocolaccording to CEN/TS 15439,[30] which concerns tars that originatefrom biomass gasification. The method applied differed in thatFigure 9. Basic principle of the DFB gasification process.

Figure 10. Scheme of the DFB gasification pilot plant at VUT.



toluene was employed as a solvent, whereas CEN/TS 15439 sug-gested the use of isopropanol (IPA). The use of toluene allowedthe simultaneous detection of the product-gas water content,which could be measured as a separate phase in the impingerbottles. However, this meant that tar components with a boilingpoint lower than that of toluene, such as benzene, toluene, andxylene (BTX), could not be detected, although the separationperformance for tar components larger than BTX is higher fortoluene than using IPA. A schematic of the arrangement of thetar-sampling line is shown in Figure 11.

The gas entered the heated sampling line, which consisted of cy-clone-filled and glass-wool-stuffed filter cartridges, in which dustand condensed tar components were deposited. Afterwards, thegas was led through six impinger bottles, five of which werefilled with toluene. The impinger bottles were located in a coolingbath maintained at 8 8C by using a cryostat and the tars andsteam condensed there. The liquid phases in the impinger bottleswere unified, and the aqueous phase was separated from the tol-uene phase. The amount of water in the impinge bottle was de-termined to calculate the water content in the gas stream. Theamount of toluene was also noted, and a sample was withdrawnfor GCMS analysis. The main part of the toluene was evaporat-ed from the sample in a petri dish. To analyze the oil in the filtercartridge, it was necessary to perform a Soxhlet extraction withIPA. Again a sample of the IPA phase was withdrawn for GCMS analysis. The IPA phase was handled in the same way as thetoluene phase. The toluene phase and the IPA phase were com-bined, which gave the amount of gravimetric tar in the productgas. The filter cartridge was reduced to ashes by oxidizing the or-ganic matter in a furnace. By weighing the cartridge before andafter the muffle-furnace treatment, the amount of entrained charand dust was calculated. Finally, the GCMS samples were ana-lyzed to determine the tar composition. This measurementmethod gave the gravimetric tar content, GCMS tar content,GCMS tar composition, water content, char load, and dust load(inorganic matter).

NH3 and H2S measurement

To measure the amounts of NH3 and H2S, the gas was sampled ina similar way to that described above for tar, that is, by using im-pinger bottles, although in this case the solvents employed weredilute H2SO4 for NH3 and an aqueous KOH solution for H2S.The sampling lines (impinger bottles) were operated at 2 8C ina glycol/ethanol mixture cooled by using a cryostatic temperatureregulator.

Analysis of inorganic components

The detection of the bed-material and fuel-inorganic-matter(ash) compositions was performed by XRF analysis by usinga PANalytical Axios Advanced analyzer. This method is basedon the emission of characteristic fluorescent X-rays from a mate-rial excited by bombardment with high-energy X-rays or gammarays. Samples to be analyzed were melted at 1050 8C by usinga Merck Spectromelt and dumped at 400 8C on a stainless-steelplate. Analysis was performed under vacuum by using a Rh

anode, an excitation voltage of50 kV, and a tube current of50 mA. The components werecalculated as oxides.

Bed material

Olivine is a naturally occurringmineral composed of silicate tet-rahedra that also contains Fe andMg in the form (Mg1xFex)SiO2,although the content of these twoelements varies with mining loca-tion. The catalytic-tar-reductioneffect caused by the use of oli-vine as a bed material has been

reported by Koppatz et al.,[13] and precalcination of the olivinecan considerably improve the catalytic activity.[31,32] In this study,the olivine used in the tests was provided by Magnolithe GmbH.The results of the XRF analysis as well as the mechanical proper-ties of the olivine bed material are shown in Table 9. As a resultof its hardness, high heat capacity, and high catalytic activity forsteam reforming, olivine is considered ideal for fluidized-bed ap-plications. A material particle size of dp50=375 mm was selected,which corresponds to particle group B according to Geldart.[33] Agraph of the bed material size distribution is shown in Figure 12.

Feedstock

The coal used in the gasification tests was mined in Puertollano,Spain. As a result of its characteristically high ash content, the

Figure 11. Sampling line for tar, water, and entrained particulate matter (char and dust).

Table 9. Chemical composition and mechanical properties of the used oli-vine.

Composition and properties

Na2O [wt%] 0.43MgO [wt%] 46.76Al2O3 [wt%] 0.40SiO2 [wt%] 39.84P2O5 [wt%] 0.03SO3 [wt%] 0.06K2O [wt%] 0.32CaO [wt%] 0.90Cr2O3 [wt%] 0.28MnO [wt%] 0.15Fe2O3 [wt%] 10.32NiO [wt%] 0.31Cl [wt%] 0.10Others [wt%] 0.11Hardness [Mohs] 67Particle density [kgm3] 2850


S. Kern et al.

calorific value of the coal was low. The results of proximate andultimate analysis of the coal are summarized in Table 1.

Symbols

dp10 [mm] Particle size with mass fraction90%DHR,850 [kJmol

1] Heat of reaction at 850 8CLHVPG [MJNm

3db1] Lower heating value of the product

gas (dry)_mH20;actual [kgh

1] Actual mass flux of steam in thegasification reactor

_mH20;stoich: [kgh1] Actual mass flux of steam in the

gasification reactor_msteam [kgh

1] Mass flux of steam in the gasifica-tion reactor

_mfuel [kgh1] Mass flux of solid fuel into the

gasification reactor_mH20;con: [kgh

1] Amount of water that is convertedto product gas

_mC;PG [kgh1] Carbon flux in product gas stream

_mC;FG [kgh1] Carbon flux in flue gas stream

Pfuel;G [kW] Input fuel power of solid fuel intogasification reactor

Pfuel;C [kW] Input fuel power of fuel for combus-tion reactor

_QPP [kW] Heat loss of the pilot plant_QIP [kW] Heat loss of an industrial size plantT [8C or K] TemperatureUmf [ms

1] Minimum fluidization velocityUt [ms

1] Terminal velocity for a single parti-cle

Ug, Uc [ms1] Superficial gas velocity in gasifica-

tion reactor (g) and combustion re-actor (c)

_VPG [NM3dbh

1] Volumetric flow rate of product gas(dry)

XH2O;rel [kgH2Okgfuel,daf1] Water conversion in the gasifier, re-

lated to the fuel inputXC,G [%] Carbon conversion in the gasifica-

tion reactorXC,DFB [%] Overall carbon conversion of the

DFB systemx [molkgC,H,O

1] Molarity of carbon in the fuel (dry,ash-, N-, Cl-, and S-free basis)

y [molkgC,H,O1] Molarity of hydrogen in the fuel

(dry, ash-, N-, Cl-, and S-free basis)z [molkgC,H,O

1] Molarity of oxygen in the fuel (dry,ash-, N-, Cl-, and S-free basis)

Greek letters

H2O [molH2Okgdaf,N,S,Cl-free1 or kgH2Okgdaf,N,S,Cl-free

1]Stoichiometric H2O demand

hC;IP Cold gas efficiency calculated for an industrial plantwith the same fuel power as the pilot plant

hCPP Cold gas efficiency of the pilot plantfSF;wt [kgH2Okgfuel,daf

1]Steam-to-fuel ratio

fSC;wt [kgH2OkgC1]

Steam-to-carbon ratiolH2O [molmol

1] , [kgkg1]Stoichiometric H2O ratio

tB [s] Product gas residence time in the bubbling bedtF [s] Product gas residence time in the freeboard of the

gasification reactortC [s] Gas residence time in the combustion reactorvi Mass fraction in the fuelvash Ash mass fraction in the fuelvC Carbon mass fraction in the fuelvH2O Water mass fraction in the fuel

Abbreviations and subscripts

BTX Benzene, toluene, xylenec Cold gas (efficiency), combustion reactordaf Dry and ash free basisdb Dry basisDFB Dual fluidized bedg Gasification reactorIP Industrial plantIPA IsopropanolOP Operating pointPAH Polycyclic aromatic hydrocarbonsPP Pilot plantPG Product gasVUT Vienna University of Technology

Figure 12. Size distribution of the olivine bed material.



Acknowledgements

The authors gratefully acknowledge the financial support pro-vided by the European Commission as this study was per-formed within the framework of the Fecundus project, fundedby the Research Fund for Coal and Steel of the EuropeanUnion (Contract No. RFCR-CT-2010-00009).

Keywords: ash industrial chemistry carbon combustion energy conversion

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Received: January 17, 2013Revised: February 18, 2013Published online on April 5, 2013


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energy technology volume 1 issue 4 2013 [doi 10.1002%2fente.201300009] kern, stefan; pfeifer,...

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