experimental investigation and modeling of gasification of sewage sludge in the circulating...

Chemical Engineering and Processing 44 (2005) 717–736

Experimental investigation and modeling of gasification ofsewage sludge in the circulating fluidized bed

I. Petersen, J. Werther∗

Technical University Hamburg-Harburg, Denickestr. 15, D-21073 Hamburg, Germany

Received 31 October 2003; received in revised form 1 September 2004; accepted 1 September 2004Available online 12 October 2004

Abstract

Experiments of sewage sludge gasification were performed in a circulating fluidized bed of pilot plant scale (15 m height, 0.1 m i.d.). Forthe examination of the influence of the air ratio, gasification temperature, feeding height and fluidization velocity several screening tests wereconducted. To understand better the results from the screening experiments, axial profiles of the gas composition were measured. As the mostinfluencing factor for the heating value of the gasification gas the air ratio was found. Additionally, a low feeding height is recommended forg With lowf

te reactionnk for reactors©

K

1

aafihtcdcwt5

handco-olidused

astesonlyat-lrmanwilla-osaly theacil-

en-n the

0d

ood gas quality. While feeding into the lower dense zone of a circulating fluidized bed (CFB), mixing of the fuel particles is better.eeding ports, high velocities are attainable and therefore high fuel throughput can be achieved.

In a second step a model of the CFB gasifier was developed. The fluid dynamics of a CFB were included as well as the compleetwork of gasification. With the measured axial profiles of gas composition during pyrolysis, and gasification with air and a CO2/N2-mixtureinetic rate expressions for sewage sludge gasification under fluidized-bed conditions were determined which may now be usedcale-up calculations.2004 Elsevier B.V. All rights reserved.

eywords:Gasification; Circulating fluidized bed; Reaction kinetics; 1.5-dimensional model

. Introduction

Combustion was the primary method of generating heatnd also power (by a using the heat from the combustion insteam turbine) from renewable fuels. More recently, gasi-

cation of wood, as well as wastes and also sewage sludgeas come into the discussion. The efficiency of the gasifica-

ion process is better, in principle, because the produced gasan be used directly in a power generation process. The onlyrawback to date for this technology is the high tar and dustontent of the synthesis gas produced. Typical tar content forood gasification is in the range of 1–30 g/m3 STP[1]. For

he use in gas-motors or gas-turbines a tar content of only0 mg/m3 STP is permitted[2,3].

∗ Corresponding author. Tel.: +49 40 428783039; fax: +49 40 428782678.E-mail address:[email protected] (J. Werther).

In Germany nowadays the sewage sludge on oneis used energetically in mono-combustion plants; it iscombusted in coal fired power stations and in municipal swaste combustion. On the other hand, sewage sludge isas landfill and in agriculture[4]. From 1 June 2005 on, uselandfill will no longer be permitted. Only pre-treated wascan be deposited. The facilities for pre-treatment havesufficient capacity for the “normal” wastes. Additional trement of sewage sludge is problematic[5]. But agriculturarecycling of sewage sludge is controversial. In some Gestates, a ban is being drawn up. Agricultural recyclingprobably cease by 2020[6]. Therefore operators of wastewter treatment plants are looking for new and reliable disppaths for the sewage sludge having high acceptance bpopulation and also by the operators of co-combustion fities.

In Germany utilization of biomass as renewableergy source is regulated by the German government i

255-2701/$ – see front matter © 2004 Elsevier B.V. All rights reserved.oi:10.1016/j.cep.2004.09.001

718 I. Petersen, J. Werther / Chemical Engineering and Processing 44 (2005) 717–736

Renewable Energy Law (EEG) and the Biomass Ordinance(BiomasseV). The objective of the EEG was to increase thecontribution of renewable energies for the total power con-sumption, and consequently to reach the goal set by the Euro-pean Union to double the contribution of renewable energiesfor the total power consumption until the year 2010. The def-inition of “biomass” according to the EEG is given in theBiomass Ordinance (BiomasseV). The possible processes togenerate power from biomass are also fixed in the Biomas-seV. Sewage sludge itself is not accepted as “biomass” ac-cording to the BiomasseV. But in the processes, which aredefined in the ordinance to be possible for power generationfrom biomass, it is allowed to add synthesis gas from sewagesludge gasification to a portion of up to 10% of the total en-ergy content. In this case a gas with as high as possible energycontent is to be produced in the gasification plant rather thana real tar-free gas.

But besides the use of sewage sludge as renewable en-ergy, gasification of sewage sludge is also interesting for co-combustion in fossil-fired power generation facilities[7]. Theaddition of sewage sludge directly to the combustor is some-times problematic due to the high ash content of the sludgeand the high amount of contaminants (e.g. heavy metals) inits ash. If sewage sludge gasification is performed before co-combusting the produced gas in the utility burner, the coal ashb udgea alueg ertedc d bys theg re-d

n ofs wasc bedsa stionp d itsfl isc in theg ty ofs d, ane tingfl ) atT is,e sent)w ree romt athem py-r inedb iumc thet bus-t

an-t uld

be applicable under fluidized-bed conditions since a majordrawback of kinetic informations about gasification reactionnetworks available in the literature is that they have been ob-tained from single-particle or fixed bed reactor experiments.Since heat and mass transfer in a fluidized bed are quite dif-ferent and since furthermore it is well known from pyrolysisexperiments that the pyrolysis gas yield and the compositionof the pyrolysis gas are strongly dependent on the heating rateit is highly desirable to obtain information about the gasifica-tion kinetics directly from fluidized-bed experiments. In thepresent work a set of kinetic equations is coupled with a fluid-dynamic model of a circulating fluidized-bed riser. Numericalvalues of three kinetic parameters are then determined not bysimply comparing measured and calculated reactor exit con-centrations but by fitting the model to measured axial profilesof the species concentration over the full 15 m of height thegasification reactor. Although, admittedly, inaccuracies of thecomplex fluid-dynamical model will also affect the thus de-termined values it is believed that the fluidized-bed conditionsare essential for the determination of reaction kinetics whichare to be used for the simulation of fluidized-bed gasifiers.

2. Experimental

lysisa e ki-n ctionn ts istl manm asts andd has at theo witho callyb tureo

finalm lone,a gas isc innerd atedb ei fter-b

res-s or in-s g ther .75,2 er. Iti lantb tubei flu-i op

ed is prevented from being polluted by the sewage slsh. In this case, too, the tar amount in the low calorific vas from gasification is unimportant as are some unconvhar particles from the gasifier. Some coal is replaceewage sludge, and additionally the entering gas fromasifier produces a reburning-effect, which results in auction of nitrogen oxides (NOx) [8,9].

The objective of the present paper is the examinatioewage sludge gasification. The circulating fluidized bedhosen as the gasification reactor, because fluidizedre already used mostly in sewage sludge mono-combulants due to the good gas-solid contact, mixing anexibility as far as the fuel composition and conditiononcerned. To understand more deeply the reactionsasification of sewage sludge and to prove the feasibiliewage sludge gasification in the circulating fluidized bexperimental examination was carried out at the circulauidized bed pilot plant (0.1 m riser diameter, 15 m highUHH. To get information about the drying and pyrolysxperiments in an inert atmosphere (only nitrogen preere performed. CO2 gasification and air gasification wexamined thereafter at different operating conditions. Fhese measurements supported by a 1.5-dimensional matical model of a fluidized-bed gasifier, the kinetics of

olysis and the main gasification reactions were determased on kinetics available in the literature. An equilibralculation for the whole gasification reaction network inemperature range commonly used in fluidized-bed comion and gasification technology was also performed.

A major goal of the present work was to derive a quitative description of the gasification kinetics which sho

-

Experimental examination was conducted on pyrond gasification of sewage sludge in order to observe thetics and the mechanism of devolatilization and the reaetwork. The facility used for the gasification experimen

he circulating fluidized bed at TUHH (Fig. 1). The circu-ating fluidized-bed riser is made of stainless steel (Ger

aterial number 1.4841) without refractory lining for ftart-up of the facility. The inner diameters of the riserowncomer are 100 and 80 mm, respectively. The riser

otal height of 15 m. The plant is heated electrically fromutside. The distributor plate is a bubble-cap distributorne bubble cap. The fluidizing gas is preheated electriefore entering the windbox up to a maximum temperaf 800◦C.

The gas and some fly ash leaving the facility pass theeasurement position for the gas quality after the cycnd enter an after-burning chamber where the energy-ompletely burned. The after-burning chamber has aniameter of 300 mm and a total length of 4.3 m and is hey electrical radiation heating to 850◦C. The residence tim

n the after-burning chamber is at minimum 3.3 s. The aurning air can be preheated electrically, too.

The plant is equipped with ports for temperature and pure measurement. In addition there are several ports ferting gas measurement probes. At four positions aloniser height there are ports for fuel feeding (positions: 1.5, 3.5 and 4.6 m). The feeding system is a screw feed

s positioned next to the facility and connected to the py a tube into which the fuel particles are dosed. The

s approximately 1.5 m long, and the particles enter thedized bed due to gravity with their falling velocity. On t

I. Petersen, J. Werther / Chemical Engineering and Processing 44 (2005) 717–736 719

Fig. 1. Flowsheet of the gasification facility.

of the screw feeder the storage is located with enough spacefor fuel to operate the plant for one whole day.

Temperature measurement is carried out with Ni–Cr–Nithermocouples. Only in the after-burning chamber wherehigher temperatures were expected, Pt–Rh–Pt thermocoupleswere used. The riser pressure drop is measured with sensorsfor differential pressures. The mean total riser pressure dropduring the experiments was held at 7000 Pa. The operation ofthe fluidized bed as a gasifier is controlled with gas composi-tion measurements. The complete combustion of the synthe-sis gas in the after-burning chamber is also controlled by gasconcentration measurement. In the after-burning chamber anair ratio ofλ= 1.23 was chosen.

Two basically different sampling locations were used: Themeasurement ports along the riser height enable gas sam-pling from inside the riser to probe the axial profile of the gascomposition. In order to separate the sample gas from thefluidized-bed particles, the gas sampling probe is equippedwith a ceramic filter at the tip of the probe. This ceramic filterprevents the particles from entering the probe. The ceramicfilter is located inside a stainless steel tube to save the fil-ter from erosion. The measured gas concentration is a meanaverage concentration for the cross section. The other mea-surement position is directly after the cyclone at the exit ofthe plant. Because the particles are already separated fromthe gas flow by the cyclone, only some fly ash particles are


Fig. 2. Sampling probes for use inside the riser (left) and for position the measurement after cyclone (right).

expected. Therefore the gas sampling probe is designed dif-ferently. It is just a stainless steel tube with an inner diameterof 4 mm (suction velocity 4–20 m/s). The fly ash that is suckedout is separated from the sample gas by a small measurementcyclone. So at this measurement position not only a gas sam-ple is taken but also a sample of the fly ash is withdrawn.To prevent the particles not separated by this small measure-ment cyclone from plugging the adjacent gas sampling line,an additional filter is positioned behind the measurement cy-clone. The measurement cyclone and filter are heated to 250and 200◦C, respectively. InFig. 2both sampling probes areshown.

Directly after the gas sampling probe and cyclone plusfilter, respectively, is the tar trap. Following the tar trap is thegas sampling line (Fig. 3). The gas sampling duct is made ofTeflon-tubing, which is electrically heated from the outsideto a temperature of 160◦C to prevent the temperature insidefrom dropping below the dew point of water.

The analyzers are fed via a membrane pump. The analyzerfor the water content does not need any more pretreatment of

the sample gas. For the rest of the analyzers, a tar condenserwith glass wash bottles in an ice bath is put in this place. Af-terwards, the gas passes a cooler and the dry gas is distributedto the analyzers. The flow to the analyzers is adjusted to oneliter per minute. The analyzers are supplied in series (Fig. 3).All the main components, which are O2, CO, CO2, H2, CH4,C2H4, and H2O, could be measured online. The oxygen an-alyzer measures with the paramagnetic effect. The analyzersfor CO, CO2, CH4, C2H4, and H2O have a non-dispersive in-frared (NDIR) detector. The instrument for the measurementof the H2 concentration has a thermal conductivity detector(TCD) with cross compensation for CH4, CO, and CO2.

The solid materials used in the experiments were the fueland silica sand as bed material. No additional catalyst orlimestone was used in the riser. The fuel examined was driedpelletized sewage sludge from a municipal waste-water treat-ment plant. InTable 1the composition (proximate and ulti-mate analysis) is given.

The particle size distribution was measured before andafter the feeder. The mean particle diameter before wasx3,50= 2.9 mm. After the feeder the mean particle diameterwas slightly reduced tox3,50= 2.8 mm. The density of thesewage sludge was experimentally determined by helium py-cnometer to 1728 kg/m3.

As bed material silica sand was used. This sand was avail-a 0%v be

TP

P

U

L

Fig. 3. Gas measurement system.

ble with two different particle size distributions. The 5alue of the cumulative mass distribution was found to

able 1roximate and ultimate analysis of the dried sewage sludge

Mean

roximate analysisVolatiles, wt% (waf) 83.4Combustibles,a wt% (raw) 50.6Water, wt% (raw) 7.3Ash, wt% (raw) 42.1

ltimate analysisC, wt% (waf) 50.5H, wt% (waf) 6.6O,a wt% (waf) 34.5S, wt% (waf) 1.2N, wt% (waf) 7.1

ower heating value (LHV), MJ/kg (raw) 10.0a By difference.


x3,50= 180�m andx3,50= 203�m, respectively. In most ofthe experiments finer sand was used. After each experiment,the sand from the facility was completely retained. For thenext experiment a mixture of this retained activated bed ma-terial and new sand was used at the start. The retained bedmaterial to be reused in the next experiment was sieved tohave particle diameters smaller than 70�m only. The densityof the bed material was experimentally determined by heliumpycnometer as 2687 kg/m3. The minimum fluidization veloc-ity umf was determined as 4.4 cm/s for the original sand, and6.5 cm/s for the mixture of fresh and activated bed materialunder ambient conditions with air.

In the experiments the following factors were examined:

• As fluidizing agent air, a CO2/N2-mixture, and pure N2were used, respectively.

• The temperature was varied between 1023 and 1123 K.• To determine the influence of the air ratio, experiments

with λ= 0.3 andλ= 0.6 were carried out.• As superficial gas velocities 3.5 and 5 m/s were adjusted,

respectively.• As feeding height 2.5 and 4.6 m, respectively, above the

distributor plate were examined.

In several screening tests the gas composition was measuredat the measuring position directly after the cyclone at the exito area s ve-l ign[w ed -p m/s,r 2.5 mw

g thes axialp ciallyt oundt restw ttomz highg o thee p tow

moret l day,f ed ato day.T the

TL

F

LV m

operation of the gasifier at the chosen condition. Becausesewage sludge has a high ash content and the ash tends tobake in the course of time and to form agglomerates whichlead to defluidization, the higher temperatures where omittedand the axial profile measurements were carried out at 1023 Konly.

An axial profile of gas concentration was also measuredusing only N2 as fluidizing gas. Therefore only the pyrolysisgas composition was measured, although some homogeneousgas phase reactions will also occur. The pyrolysis experimentis interesting especially with regard to the carbon oxides. Iscarbon monoxide the only product of devolatilization or willcarbon dioxide be released as well to a certain extent?

In order to examine the CO2-reforming reactions in gasi-fication, axial profile measurements were carried out with afluidizing agent mixed from N2 and CO2. Both of them weresupplied from separate gas cylinders containing 100% of car-bon dioxide and nitrogen, respectively. They were mixeddirectly below the distributor plate. The amount of carbondioxide in the mixture was desired to be the same as the oxy-gen in air, so that a switch from air gasification conditions tothe CO2/N2-mixture was just a replacement of the oxygen inthe gasifying agent by carbon dioxide.

At first, it was intended to measure additionally the gascomposition with a H2O/N2-mixture as gasifying agent, buta elec-t t toob tweenflm na op-e amr tent.T om-p -p fors

3

thes in-c ys-t tion,t ddi-t s canb etica sionsi cal-c ex-p odelw uredg eticr ludgeb

f the riser. For each of the four influencing factors, whichir ratio, temperature, feeding height, and superficial ga

ocity, two conditions were chosen (two-level factorial des10,11]), and experiments with all 16 combinations (24 = 16)ere run. For the air ratioλ= 0.3 andλ= 0.6 were chosen, thesired temperature was adjusted to 750 and 850◦C, the suerficial gas velocity at the top of the riser was 3.5 and 5espectively, and as feeding height the ports on 4.6 andere examined. These parameters are also listed inTable 2.To find the reasons for the influences discovered durin

creening tests, for some selected operating conditionsrofiles of the gas composition were measured. Espe

he conditions inside the riser near the distributor and arhe fuel feeding port were examined. It was of great intehether the fuel fed at higher positions reached the boone or was carried away by the gas flow, in particular atas velocities. Answers were also expected with regard txtension of the combustion region in air gasification. Uhich height is oxygen detectable?In contrast to the screening test experiments where

han one condition can be tested during one experimentaor the axial profile measurements the plant was operatne chosen set of operating parameters for the wholehis is also some kind of long-term feasibility test for

able 2ow and high value for screening tests

actor λ T utop hfeed

evel − + − + − + − +alue 0.3 0.6 750◦C 850◦C 3.5 m/s 5 m/s 2.5 m 4.6

s the heat of evaporation must also be supplied by therical heating, this idea has been dropped. This is noad because there was no great difference expected beuidizing with nitrogen and with a CO2/N2- or H2O/N2-ixture, respectively, since Kersten[12] has shown that iconventional non-catalytic atmospheric CFB gasifier,

rated below 1000◦C, carbon dioxide gasification and steeforming reactions do not proceed to a significant exherefore the measurement of the axial profile of gas cosition in gasification with a CO2/N2-mixture is just an exeriment to verify that Kersten’s statement is also trueewage sludge gasification.

. Modeling

It was intended to develop a mathematical model forimulation of the gasifier. As it was intended not only tolude the fluid dynamics of the circulating fluidized-bed sem but also to model the reaction network for gasificahe main gasification reactions have to be considered. Aionally a choice had to be made whether these reactione modeled by equilibrium consideration or with a kinpproach. It was decided to include kinetic rate expres

n the simulation program; nevertheless, an equilibriumulation of gasification was also performed. Kinetic rateressions for the gasification reactions included in the mere chosen from the literature. With the axially measas composition, it should be possible to adjust the kinate constants measured for coal or biomass to sewage sehavior.


The reaction network chosen for the modeling part con-sists of the devolatilization reaction and the homogeneousand heterogeneous oxidation and gasification reactions. Thedrying process was assumed to occur in parallel with thedevolatilization. For modeling purposes it is important toknow the mass fraction of the initial fuel, which is py-rolyzed. It is assumed that the total yield of volatiles equalsthe volatile content of the parent fuel determined by theproximate analysis. Subsequently the composition of thevolatiles becomes interesting. For the sake of simplifica-tion it was assumed that the char consists of pure car-bon. The released volatiles could be assumed to decom-pose according to the following stoichiometrically consistentequation:

CvcHvhOvoSvsNvn

→ vs H2S+ 12vn N2 + ξCOvo CO+ 1

2(1 − ξCO)vo CO2

+[(1 − 2ξC2H4 − 6ξtar)vc − 12(ξCO + 1)vo]CH4

+ [ 12vh − 2(1− ξC2H4 − 4.5ξtar)vc + (ξCO + 1)vo

− vs]H2 + ξC2H4vc C2H4 + ξtarvc C6H6 (1)

In Eq.(1) vi is the mole fraction of componenti (i = C, H,O, S, N) in the volatiles. Eq.(1) assumes that

•ion,tH

• y-s

• xture

• thent

• s

• CO

• tar

• ting

eedr

n

The release flux of the gaseous species H2S, N2, CO, CH4and H2 (and CO2, C2H4, C6H6) are only fractions of the totalmole flux.

The heterogeneous reactions occurring during gasificationare char reacting with oxygen, water vapor or carbon dioxide,respectively. The two reactions with oxygen can be expressedin one equation, as can the two possible reactions with watervapor:

�C + O2 → 2(� − 1)CO + (2− �)CO2, where

1 ≤ � ≤ 2 (1’)

C + �H2O → (2− �)CO + (� − 1)CO2 + �H2, where

1 ≤ � ≤ 2 (2’)

C + CO2 → 2CO (3’)

The splitting factor α in the combustion reactiondetermines the ratio of produced carbon monoxide to car-bon dioxide. There are empirical correlations available for theprediction ofα. The correlation given by Linjewile and Agar-wal [15] was used in the present model giving, e.g.α= 1.3f

wa-ti elw

eouss

C

H

C

I genc n:

C

ion.D e oft en inb ntv

aterv

C

Abi

C

All nitrogen is released as N2. Additionally NH3 could betaken into account as primary product of devolatilizatas van der Drift et al.[13] and Kurkela[14] found thaabout 60% of the fuel-bound nitrogen is converted to N3,but for simplicity reasons the component NH3 is neglectedhere.All sulfur is released as H2S, because in gasification oxgen is scarce and sulfur will be released as H2S and not aSO2.The oxygen might be released as CO only, or as a miof CO and CO2.The rest of the carbon in the volatiles results inproducts CH4, C2H4 and an additional tar compone(C6H6).The hydrogen not consumed by H2S and hydrocarbonwill decompose as H2.

Thus, one has to introduce three splitting factors:

one parameter for the part of the oxygen reacting to(�CO);one parameter for the fraction of carbon reacting to(�tar);finally, one parameter for the fraction of carbon reacto C2H4 (ξC2H4).

The total flux of volatiles in mol/s depends on the fuel fatemfuel and the composition only:

˙vol = mvol

Mvol= mfuelxvol,raw

vcMC + vhMH + voMO + vsMS + vnMN(2)

or T= 1023 K andα= 1.2 atT= 1123 K.In the steam gasification or so-called heterogeneous

er gas shift reaction (reaction(2′)), the splitting factorβ asntroduced by Matsui et al.[16] will be adopted in the modith the valueβ = 1.2.As long as oxygen is present combustion of the gas

pecies CH4, H2 and CO will take place:

H4 + 12O2 → CO+ 2H2 (4’)

2 + 12O2 → H2O (5’)

O+ 12O2 → CO2 (6’)

f all oxygen is consumed, carbon monoxide and hydroould take part in the well-known water-gas shift reactio

O + H2O ↔ CO2 + H2 (7’)

The water-gas shift reaction is an equilibrium reactepending on temperature, the equilibrium is on the sid

he products or the educts. So the reaction can be drivoth directions. As reaction(7′) is a reaction with constaolume, pressure will not influence the equilibrium.

A very pressure sensitive reaction is the reaction of wapor with methane (steam reforming):

H4 + H2O ↔ CO + 3H2 (8’)

lso a CO2-gasification of methane (CO2 reforming) coulde considered. As product gases only CO and H2 are taken

nto account:

H4 + CO2 ↔ 2CO + 2H2 (9’)


In this study of gasification reactions it was not intendedto model the complete reaction scheme of nitrogen and sulfurcompounds especially as biomass fuels do not contain asmuch sulfur as coal does. Therefore the nitrogen and sulfurcontained in the fuel was modeled to be released as volatileN2 and H2S, respectively (Eq.(1)).

3.1. Equilibrium calculation

To solve simultaneously the equilibrium of the nine reac-tion equations, two possible strategies can be pursued. One isthe calculation of the equilibrium composition by minimiz-ing the Gibbs free energy (or Gibbs free enthalpy)[17]. Thesecond possibility is to solve the reaction scheme with theso called relaxation method[18,19]. The relaxation methodwas chosen to calculate the equilibrium composition in thepresent paper. In this method it is assumed that each reactiontakes place in a separate reactor and reacts to equilibrium.This equilibrium composition enters the next reactor, and thenext reaction can reach equilibrium. Leaving the last reactor,the composition enters again the first reactor and the chain ofequilibrium calculation starts a second time. So the equilib-rium composition for all reactions is found iteratively. Theprocedure can be stopped when the molar amount of eachcomponent does not change any more before and after a re-a

. TheG tantw

�

K

achc l ta-b tepso wasc fittedt mc per-a

3

itu inmv , thee

r

ctione con-

sumption of the components, the reaction rate has to be mul-tiplied by the stoichiometric coefficientνij of that componenti given in the reaction equationj. The reaction rates used inthe present paper are simplified overall-reaction rates and ac-count for all mechanisms. This is of course a simplification,but the gasification process is controlled by the chemical re-action [21], and the particle size as well as the initial charmass has no significant effect. As the reactions rates are inmol/(m3 s), for the solids the reaction rates have to be multi-plied with the molar mass of the solid species for the balanceequations.

3.3. Model equations

As the facility used for the experiments is very tall, andhorizontal mixing was complete, the first model presentedin the present paper is a 1.5-dimensional or “pseudo two-dimensional” model. Although concentration changes wereonly considered in the axial direction, the model somewhatdescribes two dimensions, because the riser was divided hor-izontally into different phases. The model is unsteady-stateand semi-empirical. Conservation equations for solids andgas are based on the semi-empirical model by Kruse et al.[22].

To model the circulating fluidized-bed gasifier, the riserw ilutezfl d tob , intw l wasa tankr

ofi wa-t , thed , thec eudos car-b Ninec asifi-cb( -c ast inede r gasr

to bep s andw owsi Thes rsiono only.H nd int nge in

ctor cascade.All reaction equations defined above were considered

ibbs free enthalpy of reaction and the equilibrium consas calculated according to Eqs.(3) and(4), respectively.

g0R =

∑vi �g

0f,i (3)

= exp

(−�g

0R

RT

)(4)

The data for the Gibbs free energy of formation for eomponent was taken from the JANAF thermochemicales[20]. The data is given for several temperatures in sf 100 K. The temperature range from 800 to 1300 Khosen, and the calculated equilibrium constants wereo a function of the Arrhenius type to find the equilibriuonstant for each temperature in-between the given temture steps.

.2. Kinetic rate expressions

For all kinetic rate expressionsr in the model, the unsed is mol/(m3 s). The gas concentrations are all givenol of the model compound and are based on m3 reaction

olume. When in the literature different units were usedxpressions were modified to fit this unit.

= dcidt

(5)

The amount of the reaction rate is related to the reaquations formulated above. In order to determine the

as cut axially into three zones, the bottom zone, upper done and exit zone. As done in Kruse and Werther[23], theuid-dynamic behavior of the bottom zone is consideree similar to that of a bubbling fluidized bed. Therefore

he bottom zone the two-phase theory of fluidization[24]as used. In the upper dilute zone a core-annulus modepplied. The exit zone was modeled as continuous stirredeactor for gas and solids.

The solid material in the fluidized bed is composednert bed material and fuel. The fuel consists of volatiles,er, carbon and ash. To be able to describe the dryingevolatilization, and the carbon gasification separatelyomponents of the fuel were treated as independent “psolids”. So four different solid species were considered:on, ash (or inerts), volatiles, and the solid-bound water.omponents in the gas were chosen to represent the gation gas, namely oxygen (O2), carbon dioxide (CO2), car-on monoxide (CO), methane (CH4), hydrogen (H2), waterH2O), nitrogen (N2), hydrogen sulphide (H2S), and a hydroarbon representing the tar (C6H6). Benzene was chosenhe model compound for the tar because it was determarlier as the major single component in the produceepresenting 60–80% of the total tar[25].

In the bottom zone the bubble phase was believedarticle free. The suspension contains particles and gaas believed to be in minimum fluidization state. Gas fl

n plug flow in both the bubble and suspension phase.olids are evenly distributed and are transported by dispenly. Gas–solid reactions occur in the suspension phaseomogeneous gas reactions take place in the bubble a

he suspension phase. Because of the reaction the cha


volume of gas results in a higher velocity. The suspensionphase gas stays at minimum fluidization, only the bubblephase gas velocity is increased. Thus, there is a net flow ac-cording to Yan et al.[26] that considers the production orreduction of gaseous volume by reaction. This net flow isdirected from the suspension phase to the bubble phase. Inaddition, there is exchange of gas between these two phases.

The conservation equation for the gaseous componenti inthe bubble phase in the bottom zone reads as follows:

εb∂cgb,i

∂t+ ∂(ubcgb,i)

∂z

= −Kbs(cgb,i − cgs,i) + εb∑j(g–g)

vi,jrb,j +∆bzn′′′net flow,i

(6)

In the suspension phase the gas is transported by disper-sion and convection with minimum fluidization velocity. Theconservation equation reads:

(1 − εb)εmf∂cgs,i

∂t− (1 − εb)εmfDg,ax

∂2cgs,i

∂z2

+ (1 − εb)umf∂cgs,i

∂z

=

ans-p r thec

c

l-t et

tionsi

c

f

−

for the solids in the suspension phase of the bottom zone

∂xsbz,i

∂z

∣∣∣∣z=0

= 0 (11)

−cv,bzDs,axρs∂xsbz,i

∂z

∣∣∣∣z=Hbz

={−fdcvdvdρs(xsd,i − xsbz,i)

−(1 − fd)cvlvlρs(xsl,i − xsbz,i)(12)

The feeding of gas and solids (feed and return) was de-scribed by an ousting model instead of absolute mass flowrates. Doing this, it is taken into consideration that for ex-ample during feeding the solid volume concentration can in-crease or decrease in case the solids present are ousted bysolids with lower volume concentration. The distribution isapproximated with the Gaussian error-function (ferf).

m′′′Return,i =

(1 − xsbz,i

xReturn,i

)Gsferf (13)

m′′′Feed,i =

(1 − xsbz,i

xFeed,i

)m′′

Feedferf (14)

In the upper dilute zone a core-annulus model with height-depending core radius is used. The gas flow across the bound-ary between the core and the annulus is described by an ex-change flow, and the change in volume again is modeled bya net flow. The change in solids across the boundary is due toconvective compensation because of the enlarging or short-ening of the core and the annulus region. Additionally, be-cause there will be entrainment of downfalling particles bythe gas stream at the boundary from core to annulus phase,an exchange flow for the solid phase was defined. The frac-tion of the dense phase, i.e. the annulus region, of the totalcross-sectional area is described by the variablefd. The solidsvolume concentration in the two regions is contained incvwith the index l or d for lean and dense phase, respectively.For the velocity of gas the variableu and for the velocityof solidsv is used. As in the dense phase the flow is down-ward, the velocities are defined negative. It follows for thebalance for gas in the dense phase of the upper dilute zonefor componenti:

(1 − cvd)∂[fdcgd,i]

∂t+ (1 − cvd)ud

∂[fdcgd,i]

∂z

= (1 − cv,d)ud∂fd

∂zcgl,i −Kdl,g(cgd,i − cgl,i)

+ fd

(1 − cvd)

∑j(g–g)

vi,jrdense,j

+ cvd∑j(g–s)

vi,jrdense,j

−∆udn

′′′net flow,i (15)

Kbs(cgb,i − cgs,i) + (1 − εb)εmf

∑j(g–g)

vi,jrsusp,j

+ cv,bz

∑j(g–s)

vi,jrsusp,j −∆bzn′′′net flow,i (7)

The solids in the bottom zone were calculated with trort by dispersion only. So the mass balance equation foomponenti of the fuel particle is the following.

v,bzρs∂xsbz,i

∂t− cv,bzDs,axρs

∂2xsbz,i

∂z2

= m′′′Feed,i + m′′′

Return,i + cv,bzMi∑j(g–s)

vi,jrsusp,j (8)

Because the reaction rate is in mol/(m3 s) it has to be muiplied by the molar mass of the solid componenti here, sinche balance for the solids is in kg/m3.

The boundary conditions for the mass balance equan the bottom zone are the following.

For gas in the bubble phase it holds

gb,i|h=0 = cgb,i,IN (9)

or gas in the suspension phase

(1 − εb)εmfDg,ax∂cgs,i

∂z

∣∣∣∣z=0

= +(1 − εb)umf(cgs,i,IN − cgs,i),∂cgs,i

∂z

∣∣∣∣z=Hbz

= 0

(10)


the balance for gas in the lean phase for componenti:

∂[(1 − fd)(1 − cvl )cgl,i]

∂t+ ∂[(1 − fd)(1 − cvl )ulcgl,i]

∂z

= −(1 − cvd)ud∂fd

∂zcgl,i +Kdl,g(cgd,i − cgl,i) + (1 − fd)

×(1 − cvl )

∑j(g–g)

vi,jrlean,j + cvl∑j(g–s)

vi,jrlean,j

+∆udn′′′net flow,i (16)

the balance for solids in the dense phase of the upper dilutezone for componenti:

cvdρs∂⌊fdxsd,i

⌋∂t

+ cvdvdρs∂⌊fdxsd,i

⌋∂z

= cvdvd∂fd

∂zρsxsl,i + fdcvdMi

∑j(g–s)

vi,jrdense,j (17)

the balance for solids in the lean phase for componenti:

ρs∂⌊(1 − fd)cvlxsl,i

⌋∂t

+ ρs∂⌊(1 − fd)cvlvlxsl,i

⌋∂z

∂fd ∑

theg

a

f

a

(

f

a

f

a

(

T nod

tankri

(1 − cv,ez)∂cgez,i

∂tHez

= (−uez + fd(1 − cvd)ud)cgez,i + (1 − fd)(1 − cvl )ulcgl,i

+(1 − cv,ez)

∑j(g–g)

vi,jrez,j + cv,ez

∑j(g–s)

vi,jrez,j

Hez

(23)

the balance for solids in the exit zone for componenti:

cv,ezρs∂xsez,i

∂tHez

= −Gs,i + fdccvvdρsxsez,i + (1 − fd)cvlvlρsxsl,i

+cv,ezMi∑j(g–s)

vi,jrez,jHez (24)

The variables, which are not yet known, are defined ac-cording to the following closure laws. In the mass balancesabove the gas and solids velocities (umf, ub, ud, ul , uez, vd, vl )have yet to be determined, as well as the void fractions (εmf,εb) and solid volume concentrations (cv,bz, cvl, cvd, cv,ez)and the fraction of the dense zone from the cross sec-tional area (fd). The solid density is assumed to be constantρs = 2600 kg/m3. The solids circulation rateGs as well as thesolid volume concentration in the bottom zone and in the exitz d aret

u inputp ensepa elder[ ew tt

c

tioni byt ttomzT

inedu

u

neari

elsi thefl ed tos olidsv s the

= −cvdvd∂zρsxsl,i + (1 − fd)cvlMi

j(g–s)

vi,jrlean,j

(18)

The boundary conditions for these four equations foras in the dense phase:

tz = Hap −Hez : cgd,i = cgez,i (19)

or the gas in the lean phase:

tz = Hbz :

1 − fd)(1 − cvl )ulcgl,i = −fd(1 − cvd)udcgl,i + ubcgb,i

+(1 − εb)umfcgs,i (20)

or the solids in the dense phase:

tz = Hap −Hez : xsd,i = xsez,i (21)

or the solids in the lean phase:

tz = Hbz :

1 − fd)cvlvlxsl,i = −fdcvdvdxsd,i + (fdcvdvd

+(1 − fd)cvlvl )xsbz,i (22)

he exchange coefficientKdl,g is dependent on the fractiof the dense phase of the total cross sectional areafd, asescribed by Schoenfelder et al.[27].

The exit zone was modeled as a continuous stirredeactor. The balance for gaseous componenti in the exit zones obtained as

one, respectively, are input parameters for the model anaken from measurements.

The minimum fluidization void fractionεmf and velocitymf have to be determined for the bed material and arearameters of the model. The velocity of gas in the dhase is assumed to be constant and was set to−0.75 m/sccording to optical probe measurements by Schoenf

28]. The solids volume concentration in the dense phascvdas assumed to be constant over the riser height.cvd was se

o 0.25 according to the findings of Schoenfelder.In the bottom zone it holds

v,bz = (1 − εb)(1 − εmf) (25)

This relationship defines the solids volume concentran the bottom zone, which is the fraction not occupiedhe gas phase. The solids volume concentration in the boone was assumed to be constant and was set tocv,bz = 0.3.hus Eq.(25)allows us to determineεb.

The gas velocity in the dense phase can be determsing the following equation used by Schoenfelder[28]:

d − vd = umf

1 − εmf

(1 − cvd1 − εmf

)nRZ−1

(26)

For the remaining unknowns the same amount of lindependent equations is necessary.

In this model as in the great majority of gasification modn fluidized beds[29], an empirical correlation to describeuid dynamics inside the reactor is used. Thus the neolve the momentum equations is avoided. The mean solume concentration for each height of the riser contain


two parts of solids volume concentration in the dense and inthe lean phase:

cv = (1 − fd)cvl + fdcvd (27)

This mean solids volume concentration can be determinedfrom axial pressure measurements and can be fitted to thefollowing empirical exponential equation according to Kuniiand Levenspiel[30]:

cv = cv,ez + (cv,bz − cv,ez) exp(−αcv(z−Hbz)) (28)

The solids circulation rateGs results from the upwardsflowing solids in the lean phase and the downwards flowingsolids in the dense phase (remembervd to be negative!).

Gs = fdcvdvdρsxsd + (1 − fd)cvlvlρsxsl (29)

The solids velocity in the lean phasevl is contained inan approach given by Martin[31] for the settling velocity ofsingle particles.

Re = 18

[√1 + 1

9

√Ar − 1

]2

(30)

This equation is valid for 0 <Re< 105, with the Reynoldsand the Archimedes number defined by

R

A

v

theh ns. Int n thes duced.A imumfl ocityi tiale theb

At the bottom (z= 0) the incoming gas flowVIN is dis-tributed on the two phases,

z = 0 : VIN = (ub + (1 − εb)umf)At (35)

In the upper dilute zone the homogeneous and heteroge-neous reactions in both core and annulus regions contributeto the velocity in the lean phase, because the velocity of gasin the dense phase is assumed to be constant. The axial profileof the gas velocity in the lean phase can be expressed by

(1 − fd)(1 − cvl )∂ul

∂z

= RT

P

(1 − fd)(1 − cvl )

∑i

∑j(g–g)

vi,jrlean,j

+ (1 − fd)cvl∑i

∑j(g–s)

vi,jrlean,j

+ fd(1 − cvd)∑i

∑j(g–g)

vi,jrdense,j

+ fdcvd∑i

∑j(g–s)

vi,jrdense,j

(36)

With the condition at the boundary to the bottom zone:

ub + (1 − εb)umf|z=Hbz

= fd(1 − cvd)ud + (1 − fd)(1 − cvd)ul (37)

The gas velocity in the exit zone is

uez = fd(1 − cvd)ud + (1 − fd)(1 − cvd)ul

+ RTP

cv,ez

∑i

∑j(g–s)

vi,jrez,j

+ (1 − cv,ez)∑i

∑j(g–g)

vi,jrez,j

Hez (38)

Not only the velocity but also the concentration of eachcomponenti in the gas changes because of reaction. Yan etal. [26] pointed out the significance of a ‘net flow’. So theircalculation of the net flow is adopted in this model. The totalnet flow in the bottom zone is

∆bzn′′′net flow =

∑i

(1 − εb)εmf

∑j(g–g)

vi,jrsusp,j

+ cv,bz

∑j(g–s)

vi,jrsusp,j

(39)

ep = dp(ul − vl )

vG(31)

rp = gd3p

v2G

ρs

ρg(32)

From Eq.(30) it follows

l = ul − 18vG

dp

[√1 + 1

9

√Ar − 1

]2

(33)

The axial velocity profile in the riser is determined byomogeneous and heterogeneous gasification reactio

he bottom zone heterogeneous reactions occur only iuspension phase. Solids are consumed and gas is pros it is assumed that the suspension phase stays in minuidization state, the surplus gas contributes to the veln the bubbles only. Thus, we get the following differenquation for the velocity profile in the bubble phase ofottom zone:

∂ub

∂z= RT

P

εb

∑

i

∑j(g–g)

vi,jrb,j

+ (1 − εb)εmf

∑

i

∑j(g–g)

vi,jrsusp,j

+ cv,bz

∑

i

∑j(g–s)

vi,jrsusp,j

(34)


And the net flow of each componenti in the bottom zone canbe calculated from

∆bzn′′′net flow,i = cgs,i∑

icgs,i

∑i

(1 − εb)εmf

∑j(g–g)

vi,jrsusp,j

+ cv,bz

∑j(g–s)

vi,jrsusp,j

(40)

The corresponding equation for the upper dilute zone forthe total net flow is

∆udn′′′net flow =

∑i

fd(1 − cvd)

∑j(g–g)

vi,jrdense,j

+ fdcvd∑j(g–s)

vi,jrdense,j

(41)

and for the gas speciesi in the upper dilute zone it holds

∆udn′′′net flow,i = cgs,i∑

icgs,i

∑i

fd(1 − cvd)

∑j(g–g)

vi,jrdense,j

4

4

atio,g prod-u ined.T wereg

con-cC iox-i ll thev ssure(t ightc loc-i n ofd port( ingp thea naa

heat-i isa

It is obvious that the parameter with the highest influenceis the air ratioλ. This is not surprising and in agreement withprevious literature[13,32,33]; and the chosen higher value ofλ= 0.6 is the upper limit for an operation condition which canbe called gasification. Vriesmann et al.[34] even stated thatthe lower heating value of gas produced at air ratios higherthan 0.45 is not very useful for combustion purposes.

The parameter temperature shows the trend that withhigher temperature the heating value of the gas increases.This can certainly be understood looking at the tar crackingand the other endothermal gasification reactions, which willperform better at higher temperatures.

For the feeding height, if one looks at the lower heat-ing value of the gas, no clear trend is obvious. For highvelocity and low air ratio the heating value becomes betterthe lower down the fuel is fed into the riser. The parame-ter gas velocity has the least influence on the product gascomposition. Therefore, neither a final statement about thebest feeding height nor about the optimal gas velocity can bemade.

In general, it should be stated here that the feasibilityof sewage sludge gasification in a circulating fluidized bedhas been proven with these experimental findings. The lowerheating value obtained in the experiments, and also the over-all efficiency of the process, fits well with the range knownf rtedi s,

η

o ee 3%f t6

n per-f e of4f luegg

alcu-l andww ion of9g nt ofc Ker-s ha

4

aveb reen-i hich

+ fdcvd∑j(g–s)

vi,jrdense,j (42)

. Results and discussion

.1. Screening experiments

With screening experiments the influence of the air ras velocity, temperature, and the feeding height on thect gas composition and lower heating value was examhe higher and lower levels chosen for these parametersiven inTable 2.

In Fig. 4 all the measured results are shown. Theentration of the combustible components H2, CO, CH4 and2H4 are given as well as the concentration of carbon d

de and the lower heating value of the synthesis gas. Aalues refer to dry gas at standard temperature and preSTP). The two left columns of diagrams inFig. 4 refer tohe experiments at the lower velocity (3.5 m/s), the two rolumns of diagrams show the results for the higher vety (5 m/s). For each of the velocities there is one columiagrams for the measurements with the higher feeding4.6 m, left column of the two), and one for the lower feedort (2.5 m, right column of the two). In all diagrams onbscissa the temperature (750 and 850◦C) is shown, and ill diagrams the results for the two air ratios (λ= 0.3,λ= 0.6)re given.

On the ordinate the gas concentration and the lowerng value in weight percent and kJ/m3 on a water-free basre given, respectively.

or gasification of for instance wood, as these were repon the literature. An efficiency for the gasification proces

eff = VgLHVg (J/m3)

msLHVs (J/kg)(43)

f 58% for λ= 0.3 and 35% forλ= 0.6 was obtained. Thfficiency given in the literature ranges from 46% to 6

or biomass like barley, grass and miscanthus[35] to abou8–87% for wood[12,36–38].

The produced gas from the sewage sludge gasificatioormed in this work had on average a lower heating valu.7 MJ/m3 from the experiments atλ= 0.3 and 1.9 MJ/m3

or λ= 0.6. This compares well with the lower heating vaiven by Kersten[12] who measured 3–5.8 MJ/m3 STP in airasification of wood (λ= 0.4–0.2).

The carbon conversion in the present study was cated from the carbon detected in the gas (without tar)as about 96% for theλ= 0.6 experiments, and atλ= 0.3 itas 85% on average. In the literature, a carbon convers7%[35,38]or even complete carbon conversion[39] is ofteniven for wood and other biomass fuels. But, if the amouarbon in the tar is excluded from the balance given byten[12], a conversion of 87% atλ= 0.3 is obtained, whicgrees with the calculated value in this work.

.2. Axial profile measurements

Axial profile measurements of the gas composition heen performed to better understand the results of the sc

ng tests. Especially the influence of the feeding height, w


Fig. 4. Results of screening tests: influence ofλ, T, utop andhfeed.

Fig. 5. Axial profile of gas composition measured atλ= 0.6,utop = 5 m/s, andhfeed= 2.5 m (H2O calculated from O-mass balance).



showed no uniform trend for different gas velocities, neededmore explanation.

To further examine the influence of the feeding height, inFigs. 5 and 6the results from axial profile measurements atλ= 0.6 with a gas velocity at the top of 5 m/s and a feed-ing height of 2.5 m and 4.6 m, respectively, are shown. Asmeasurements atλ= 0.3 were restricted to the temperature of750◦C because of ash agglomeration problems, the follow-ing diagrams all refer to this latter temperature. The heightof the measurement position above the distributor is given onthe abscissa, and the gas concentration is on the ordinate. Forconvenience an arrow points at the feeding location. In casethere were repeated measurements at one measurement port,as for example at the position after the cyclone inFig. 5, thesingle symbols were shifted to the sides so that the repeata-bility of the measurement is visible.

The concentrations given are all based on the raw gas state.All gases but the water content were measured as dry gas andwere then calculated to the raw gas condition by the watercontent. Because in some cases the water content had not beenmeasured, the H2O-concentration used is calculated from theoxygen mass-balance. This method was used earlier in casea water measurement has not been carried out[40], and acomparison to the measured values showed good agreementin the present work.

measur

Obviously, in the experiments with the high superficialgas velocity, the particles entering at the upper feeding loca-tion did not reach the bottom zone to react there. Pyrolysisproducts such as hydrogen, methane and ethylene were onlymeasured above the feeding point (cf.Fig. 6). Narvaez et al.[33] found already that for a bubbling fluidized bed feedingnear the bottom is recommended for gas of good quality. Be-cause the gas quality from the experiments withλ= 0.6 ispoor, the same axial profiles were intended to be measuredat λ= 0.3. But operation of the plant atλ= 0.3 was not easywith the higher feeding port. With the higher velocity steadystate operation could not be held long enough for axial pro-file measurement, because possibly many unreacted particleswere entrained from the riser and reacted further in the down-comer, which lead to blocking. Therefore, it can be stated thata feeding port in the upper section, even though possible, isnot advisable.

To examine the influence of the velocity, axial profileswere measured with a superficial gas velocity of 3.5 and 5 m/s,respectively, atλ= 0.3 with feeding at the 2.5 m position. Theresults are shown inFigs. 7 and 8.

Comparing the results from the experiments with the samefeeding height but different gas velocities, it is obvious that atthe higher gas velocity the hydrogen concentration is higherwhile the CO concentration remains roughly constant, and

Fig. 7. Axial profile of gas composition
ed atλ= 0.3,utop = 3.5 m/s, andhfeed= 2.5 m.



therefore not only a lower feeding height but also a highervelocity is recommended.

4.3. Pyrolysis and CO2-gasification experiment

The measurements of the gas composition from pyroly-sis and with a CO2/N2-mixture, respectively, have been con-ducted to get information on the direct pyrolysis components,and also to test the kinetic rate expressions taken from litera-ture. Using the 1.5-dimensional model of the CFB, the kineticconstants from different authors were applied and adjustedif necessary to simulate the axial profiles measured duringthe experiments. InTable 3the average exit gas compositionfrom the two experiments are given, respectively.

Looking first at the results from pyrolysis, it is obviousthat carbon dioxide is a primary product of the decompositionreaction. Because no oxygen is added to the reaction chamber,the oxygen in the CO2 can only result from the fuel particles.Therefore the splitting factorξCO is not equal to one.

What else can be seen fromTable 3is that the additionof CO2 to the nitrogen in the fluidizing gas does not leadto increased carbon conversion. The water gas shift reactionis influenced, however, resulting in slightly lower hydrogencontent and increase in the water concentration, whereas achange in the amount of carbon monoxide could not be de-t n thee ctiond ortedb ots tratioi es don roly-sT -free

TE

v

PC

basis, but qualitative agreement is good: The carbon dioxideand carbon monoxide concentration detected in the gas arenearly equal and the hydrogen content is approximately twicethat of the carbon monoxide. Only the methane concentrationobtained by Schuller and Brat is lower.

4.4. Results from the equilibrium calculation

The calculations were carried out for the composition ofthe dried sewage sludge used in the experiments. The de-volatilization products were calculated as presented aboveand were taken as input concentrations. The chosen air ra-tio for the equilibrium calculation wasλ= 0.3. The resultingequilibrium compositions were computed. InFig. 9the con-centration for the main components are shown. They are given

Table 4Pyrolysis gas composition according to Schuller and Brat[41] in volumepercent

Component vol%N-free

H2 21.8–40.9CO 15.55–44.05CO2 18.4–41.0CH4 5.9–14.5

F posi-t

ected. The reason will certainly be the low temperature ixperiments. At these temperatures the Boudouard reaoes not proceed to a significant extent, as already repy Kersten[12]. Also the CO2-reforming of methane does neem to proceed because although the methane concen

s decreasing, the hydrogen and carbon monoxide valuot increase. The results might be compared with the pyis gas composition given by Schuller and Brat[41] (Table 4).he concentrations given by the authors are on a nitrogen

able 3xit gas composition from pyrolysis and CO2-gasification experiments

ol% raw CO2 CO H2 CH4 C2H4 H2O

yrolysis 4.0 4.8 8.1 4.7 1.8 4.7O2-gasification (ψCO2 = 0.2) 16.6 4.8 6.1 3.6 1.5 5

n

ig. 9. Results from equilibrium calculations. Main components, comion in vol% at different temperatures.


Table 5Kinetic rate constants valid for sewage sludge gasification

Reaction Kinetic rate expression, (mol/m3 s) kinetic constants Reference

(1′) αC + O2 → 2(α− 1)CO+ (2 − α)CO2 r(1) = k(1)cO2f(1) [45]

f(1) = 0.05

k(1) = 5.957× 102 (m/s K)Tp exp(− 149,440 (J/mol)

RTp

)6dp

α = 1+2fr1+fr

with fr = 4.72× 10−3 exp(

37,737 (J/mol)RTp

)[15]

(2′) C + βH2O → (2 − β)CO+ (β − 1)CO2 + βH2 r(2) = k(2)cH2O

1+K(2)k H2OcH2O+K(2)

k H2cH2+K(2)

k COcCOf(2) [16]

f(2) = 2

k(2) = 2.39× 102 (m3/s mol) exp(− 129,000 (J/mol)

RT

)ρcharMC

(1 −X)

K(2)k H2O = 3.16× 10−2 (m3/mol) exp

(− 30,100 (J/mol)

RT

)K

(2)k H2O = 5.36× 10−3 (m3/mol) exp

(− 59,800 (J/mol)

RT

)K

(2)k CO = 8.25× 10−5 (m3/mol) exp

(− 96,100 (J/mol)

RT

)β = 1.2

X= 0.5

(3′) C + CO2 → 2CO r(3) = k(3)cCO2

1+K(3)k CO2

cCO2+K(3)k COcCO

f(3) [46,47]

f(3) = 1

k(3) = 4.89× 107 (m3/s mol) exp(− 268,000 (J/mol)

RT

)ρcharMC

(1 −X)

K(3)k CO = 6.60× 10−2 (m3/mol)

K(3)k CO = 1.2 × 10−1 (m3/mol) exp

(− 25,500 (J/mol)

RT

)X= 0.35

(4′) CH4 + 12O2 → CO+ 2H2 r(4) = k(4)c

0.7CH4c0.8O2f(4) [48]

f(4) = 100

k(4) = 1.58× 1010 ((m3)0.75/s mol0.75) exp

(− 202,641 (J/mol)

RT

)(5′) H2 + 1

2O2 → H2O r(5) = k(5)cO2cH2f(5)

f(5) = 0.001 [49]

k(5) = 1.08× 1010 (m3/mol s) exp(− 125,525 (J/mol)

RT

)(6′) CO+ 1

2O2 → CO2 r(6) = k(6)cCOc0.5H2Oc

0.25O2f(6) [48]

f(6) = 1

k(6) = 1.78× 1010 ((m3)0.75/s mol0.75) exp

(− 180,032 (J/mol)

RT

)(7′) CO+ H2O ↔ CO2 + H2 r(7) = k(7)

(cCOcH2O − cCO2cH2

K(7)eq

)f(7) [50]

f(7) = 0.1

k(7) = 2.778 (m3/mol s) exp(− 12,560 (J/mol)

RT

)K(7)eq = 0.022 exp

(3.473×104 (J/mol)

RT

)(8′) CH4 + H2O ↔ CO+ 3H2 r(8) = k(8)

(cH2OcCH4 − cCOc

3H2

K(8)eq

)cSf(8) [51]

f(8) = 0.1

k(8) = 4.916× 10−10T 2 (kg m4/s mol2 K2) exp(− 36,150 (J/mol)

RT

)1

MCρchardp

K(8)eq = 3.106× 1014 (mol/m3)2

exp(− 2.088×105 (J/mol)

RT

)(9′) CH4 + CO2 ↔ 2CO+ 2H2 Neglected (does not proceed, too slow)

(10′) C2H4 + O2 → 2CO+ 2H2 r(10) = k(10)c0.9C2H4

c1.18O2f(10) [52]

f(10) = 1

k(12) = 3.71× 1012 ((m3)1.08/s mol1.08) exp

(− 209,205 (J/mol)

RT

)(11′) C6H6 + 3O2 → 6CO+ 3H2 r(11) = k(11)ctarcO2

k(11) = 1.58× 1012 (m3/s mol) exp(− 202,641 (J/mol)

RT

)


in vol% based on the raw gas for temperatures in the rangefrom 800 to 1300 K.

There is a big change in composition from 800 to 900 K. Attemperatures above 1000 K the composition does not changesignificantly. The main products are H2 and CO in nearlythe same amount; the H2:CO ratio is around one. This ra-tio stays the same for all temperatures, but from 800 to1000 K the H2- and CO-concentration doubles, respectively.The CH4- and CO2-content at higher temperatures is aston-ishingly low (nearly zero). The equilibrium concentration ofwater is nearly unaffected by a change in the temperature(around four volume per cent). Looking at the experimen-tally determined composition of a gasification gas given inthe literature or measured in this work, it can be stated thatthere must be strong kinetic influence that affects the com-position because during gasification with air there are con-siderable amounts of carbon dioxide present, and it has not

Fp

been reported that the hydrogen and carbon monoxide con-centrations calculated above were ever measured in air gasi-fication. The same deviation between measured synthesis gascomposition and the concentrations predicted by equilibriumcalculation has been detected by Li et al.[17] and Kersten[12].

Therefore the kinetic approach was chosen to model the re-action network in the program. The pseudo two-dimensionalmodel was used to check the kinetic rate expressions takenfrom the literature, whether they are also valid for sewagesludge gasification, and whether they are applicable for gasi-fication in circulating fluidized beds, too. To adjust the ki-netics for some reactions of the complete reaction networkseparately from the others, the measurements of the gas com-position along the riser height during pyrolysis and CO2/N2-gasification were taken, respectively. The best fit for thesplitting factor for the distribution of the oxygen-content in

ig. 10. Axial profiles of measured gas composition in (a) pyrolysis, (b) CO2-gaseudo two-dimensional model.

sification, and (c) air-gasification experiments and modeling results from


the fuel to CO and CO2 was obtained forξCO = 0.3. Withnitrogen as the only fluidizing gas in the pyrolysis experi-ments, the heterogeneous gasification reactions can performonly with the steam released during drying and with the car-bon dioxide from the devolatilization. These take part in thewater-gas shift reaction, too; and gasification reactions of thehydrocarbons (reaction(8′) and(9′)) might also play a role.But especially the CO2-gasification experiments showed thata higher amount of carbon dioxide present in the gasifierdid not significantly alter the gas composition in comparisonto the pyrolysis experiment. This has also been detected byGarcıa et al.[42] who achieved only higher conversion andhigher CO-concentration in their experiments with pine saw-dust in a fluidized-bed gasifier at 700◦C with higher amountsof catalyst in the reactor.

As the amount of carbon dioxide strongly influences theequilibrium of the water-gas shift reaction, it can be statedthat the water-gas shift reaction is nearly unaffected andmust therefore be far from equilibrium. Chamberland andLabrecque[40] also found that the final composition of thesynthesis gas is not that expected from the equilibrium con-stants but depends far more on the pyrolysis reactions. Sothe assumption often made in gasification modeling that thewater-gas shift reaction reaches equilibrium is obviouslywrong. Hamel[43] also reported that equilibrium of thew ned.T ateda ranks[ ther allert iss

bove,a hek udgeg

form-i wast om-p r witha ane-o

-b

C

C

dif-f werea hep ione ults.F ure-m ofu ,

where the feed was supplied athfeed= 2.5 m above the dis-tributor plate. Agreement of the measured gas compositionwith the calculation results from the pseudo two-dimensionalmodel is good. The best fit was obtained for the splitting fac-tor for the amount of ethylene withξC2H4 = 0.1, and for theone for the tar amount withξtar = 0.005. With this tar coeffi-cient a benzene concentration of 1460 mg/m3 was calculatedfor theλ= 0.3 case.

5. Conclusions

Combustion, pyrolysis and gasification of sewage sludgein the circulating fluidized bed were examined by both, ex-perimental and modeling studies.

In order to obtain information about the optimal operationparameters, several gasification experiments were performedwith dried sewage sludge (>90% dry matter) in a pilot-scalecirculating fluidized bed. The experimental program was sub-divided into screening tests, axial gas composition measure-ments, and additional pyrolysis as well as CO2-gasificationexperiments. The program for the screening tests was de-veloped based on statistical methods and was conducted todetermine the influence of the temperature, air ratio, feedingheight, and superficial gas velocity. Axial profile measure-ments were performed to better understand the processes in-s usedt ionsa

canb

• e on

e.• e. At

. Butationnd

• ed-theg ofoca-rbon

• ble.

• f the

• nd-alsoion

• cantt ash

ater-gas-shift reaction in fluidized bed is seldom attaihis is why, although the kinetic rate expression is formuls a rate equation for reversible reactions according to F

44] with the equilibrium constant, the kinetic constant foreaction is smaller than one. A kinetic rate constant smhan one signifies that the attainment of the equilibriumlowed down.

The reactions, which have already been presented are listed again inTable 5. They are given together with tinetic rate expressions, which are valid for sewage slasification.

As has been discussed already, the carbon dioxide reng reaction does not proceed to a significant extent andherefore omitted. For the partial combustion of the tar cound (e.g. benzene) a rate expression of second ordekinetic rate constant similar to the one for the meth

xidation was used (see reaction(11′) in Table 5).In Table 5the reactions(10′) and(11′) are the partial com

ustion of ethylene and tar, respectively.

2H4 + O2 → 2CO + 2H2 (10’)

6H6 + 3O2 → 6CO + 3H2 (11’)

For the adjustment of the oxidation reaction kinetics,erent measurements of axial gas composition profilesvailable. InFig. 10 the measured axial profiles from tyrolysis, CO2-gasification and one of the air-gasificatxperiments are shown, together with the modeling resor the air-gasification experiment simulation, the measent atλ= 0.3 with a superficial gas velocity at the top

top = 3.5 m/s at the temperature ofT= 750◦C was chosen

ide the riser; different gasification gases than air wereo learn about the devolatilization and reforming reactnd kinetics.

From the experimental part, the following conclusionse drawn:

The excess air ratio has the most significant influencthe produced gas composition. A value ofλ= 0.3 is a goodchoice for the operation of a gasifier for sewage sludgTemperature has the second most important influenchigher temperatures a more valuable gas is producedthe choice of temperature in sewage sludge gasificis limited due to the risk of melting, agglomeration asintering of the sewage sludge ash.Although no clear trend is obvious for the optimal feing height, for good gas quality a feeding port close tobottom of the riser is recommended, because mixinthe fuel particles is better there. Feeding at higher ltions leads to particle entrainment and incomplete caconversion.With near-bottom feeding, high velocities are attainaand therefore a high fuel throughput can be achievedThe extension of the combustion zone at the bottom oriser is small.In pyrolysis, in spite of the lack of oxygen in the surrouing gas, not only carbon monoxide is produced, butcarbon dioxide is a primary product of the devolatilizatreaction.CO2-gasification reactions do not proceed to a signifiextent at the low temperatures necessary to prevensintering.


• The water-gas shift reaction does not reach equilibrium.

Considering the diameter of 0.1 m and its height of 15 mthe reactor can be described as one-dimensional. A 1.5-dimensional model was developed, which contains the fluid-dynamic characteristics of the circulating fluidized bed withgas and solids upflow in the center and downflow at the wall.A complete reaction network for pyrolysis, combustion andgasification was formulated. Most kinetic rate expressionswere taken from the literature. Three open parameters are re-lated to the devolatilization process, namely the splitting fac-tor ξCO which describes the part of the oxygen in the volatilematter which is released in the form of CO, the splitting fac-tor ξC2H4 for the fraction of carbon in the volatiles which isreleased as C2H4 and the splitting factorξtar for the carbonfraction which is released as tar (and which is balanced asC6H6), respectively. Since the composition of pyrolysis gasis strongly dependent on the heating rate of the fuel particlesthese latter parameters must be determined under fluidized-bed conditions. Numerical values ofξCO, ξC2H4 andξtar weretherefore obtained by fitting the model calculations to threesets of measured axial profiles of the species O2, CO2, CO,H2, CH4, C2H4 and H2O which were taken under conditionsof pyrolysis, CO2-gasification and air-gasification, respec-tively. The model is seen to give a good description of thereactor behavior under any of these conditions. In particular,u ive ag m oft usedf dif-f

A

AAcc res-

c

c

dDfff� po-

GGh�

HHi

k reaction rate constant (m3/(kg s))kcvl parameter for the calculation of thecv-profile

(kcvl = 0.65)Kbs mass transfer coefficient between bubble and sus-

pension phase in the bottom zone (1/s)Kdl mass transfer coefficient between dense and lean

phase in the upper dilute zone (1/s)Kk adsorption constant according to Reed (1981)

(1/atm)K, Kp, Keq equilibrium constant ((various units))LHV lower heating value (J/m3, J/kg)m mass flow (kg/s)m′′ mass flow based on an area (kg/(m2 s))m′′′ mass flow based on a volume (kg/(m3 s))M molar mass (kg/mol)n molar amount (mol)n local cell number in the pseudo two-dimensional

modeln molar flow (mol/s)nRZ Richardson–Zaki exponent (nRZ = 3)�n′′′ molar net flow (convection because of reaction)

(mol/(m3 s))P system pressure (Pa)Pe Peclet numberr reaction rate (mol/(m3 s))RRtTu by

u ilute

uu lute

uv

v

V

w

xx m,

Xz

Gα

α

β

ε

η

η

λ

nder air gasification conditions the model is seen to good description of the combustion zone near the botto

he riser. The thus determined reaction kinetics can beor simulation calculations of fluidized-bed gasifiers witherent geometries[53].

ppendix A. Nomenclature

r Archimedes numbert cross-sectional area of reactor (m2)

concentration (mol/m3, kg/m3)S char or carbon concentration in kinetic rate exp

sion (kg/m3)v solids volume concentrationv average solids volume concentrationp particle diameter (m, mm,�m)

dispersion coefficient (m2/s)d volume fraction of the dense phaseerf Gaussian error function(No.) fitting factor for the reaction specifiedg0

f,i standard Gibbs free energy of formation for comnenti (J/mol)Gibbs free enthalpy (J/mol)

s solids circulation rate (kg/(m2 s))height in one zone (m)

h0R heat of reaction (J/mol)

height of one zone (m)HV higher heating value (J/m3, J/kg)

, j counting variables

ideal gas constant (J/(mol K))e Reynolds number

time (s)temperature (K)

b volumetric gas flow in the bubble phase dividedthe cross-sectional area of the reactor

d interstitial gas velocity in dense phase (upper dzone) (m/s)

ez superficial gas velocity in the exit zone (m/s)l interstitial gas velocity in lean phase (upper di

zone) (m/s)mf minimum fluidizing velocity (superficial) (m/s)

velocity of solids (m/s)i mole fraction of componenti in volatiles) (i = C, H,

O, S, N) (mol/mol)˙ volume flow (m3/s)

water (kg, mol)mass fraction (kg/kg)

3,50 particle diameter of fifty percent mass fraction (m�m)carbon conversioncoordinate in axial direction (m)

reek letterssplitting factor for the reaction of carbon and O2

cv coefficient for the fitting of the ¯cv (m−1)splitting factor for the reaction of carbon and H2Ovoid fractiondynamic viscosity (Pas)

eff efficiencyair ratio


µ chemical potential (J/mol)νg cinematic viscosity of gas (m2/s)νi,j stoichiometric coefficient of gas speciesi in reaction

jψ volume fraction (e.g. of oxygen in air)ρ density (kg/m3)ξ splitting factor

Indicesa, ash ashap apparatus (riser)ax axial directionb bubble phase in the bottom zonebz bottom zonec, C carbond, dense dense phase in the upper dilute zoneeq equilibriumez exit zoneg gasg–g homogeneous gas phase reactiong–s heterogeneous gas-solid reactionh, H hydrogenhor horizontali gas speciesIN inlet of reactorjlmmnopss lance

sstuvw

R

ers:s of, 28

ausemie

2002)

ilanzll und

[5] H. Alwast, J. Hoffmeister, H. Paschlau, 2005 oder “5 vor 12”, Muellund Abfall 1 (2003) 16–29.

[6] K.J. Thome-Kozmiensky, Aufbereitungskonzepte fuer Er-satzbrennstoffe (Processing Concepts for Substitute Fuels),Aufbereitungs Technik 43 (4) (2002) 11–20.

[7] K. Redepenning, A. Hellwig, Klaerschlammverwertung durch Ver-gasung in Kohlekraftwerken, Abfallwirtschafts J. 7 (1995) 442–445.

[8] H. Anderl, A. Mory, T. Zotter, BioCoComb-Vergasung von Biomasseund Mitverbrennung von Gas in einem Kohlenstaubkessel, VGBKraftwerks Technik 3 (2000) 68–75.

[9] L.D. Smoot, S.C. Hill, H. Xu, NOx control through reburning, Prog.Energy Combust. Sci. 24 (1998) 385–408.

[10] G.E.P. Box, W.G. Hunter, J.S. Hunter, Statistics for Experimenters—An Introduction to Design, Data Analysis and Model Building, PartIII: Measuring the Effects of Variables, John Wiley & Sons, NewYork, Chichester, Brisbane, Toronto, Singapore, 1978.

[11] A. Nordin, L. Eriksson, M. Oehman, NO reduction in a fluidizedbed combustor with primary measures and selective non-catalyticreduction—a screening study using statistical experimental designs,Fuel 74 (1) (1995) 128–135.

[12] S.R.A. Kersten, Biomass Gasification in Circulating Fluidized Beds,Dissertation Twente University, Twente University Press, Enschede,Netherlands, 2002.

[13] A. van der Drift, J. van Doorn, J.W. Vermeulen, Ten residual biomassfuels for circulating fluidized-bed gasification, Biomass Bioenergy 20(2001) 45–56.

[14] E. Kurkela, Formation and removal of biomass-derived contaminantsin fluidized-bed gasification processes, PhD thesis, VTT Publications,Espoo, 1996, p. 287.

[15] T.M. Linjewile, P.K. Agarwal, The influence of product CO/CO2

oke12–

[ asi-em.

[ ilib-ach

Fuel

[ eth-Gasevon

[ ags-

[ ork,

[ edce of

[ odelzed

[ rcu-rofile

[ s.),emic

[ m-d.),3, p.

[ ling‘net

reaction number, lean lean phase in the upper dilute zone

ax maximumf minimum fluidization state, N nitrogen, O oxygen

particlesolid phase

, susp suspension phase in the bottom zone in baequation

, S sulfur in fuel compositiont stoichiometric

totald upper dilute zoneol volatileaf water and ash free

eferences

[1] R. Buehler, Internal Combustion (IC) Engines for Biomass GasifiAn Overview, IEA Thermal Gasification Seminar, in: Proceedingthe IC Engines for LCV Gas from Biomass Gasifiers, ZuerichOctober 1997, 1998, pp. 7–12.

[2] S. Bajohr, R. Reimert, Katalytische Reinigung von Rohgasender Pyrolyse und der Vergasung von Biomasse und Muell, ChIngenieur Technik 71 (9) (1999) 1053.

[3] O. Carsjens, Biogas - eine saubere Sache, UmweltMagazin 6 (73.

[4] H. Friedrich, H. Hehrenbach, J. Giegrich, F. Knappe, Oekobder verschiedenen Entsorgungswege des Klaerschlamms, MueAbfall 10 (2002) 558–568.

ratio on the ignition and temperature history of petroleum cparticles in incipiently gas-fluidized beds, Fuel 74 (1) (1995)16.

16] I. Matsui, D. Kunii, T. Fursuawa, Study of fluidized bed steam gfication of char by thermogravimetrically obtained kinetics, J. ChEng. Jpn. 18 (2) (1985) 105–113.

17] X. Li, J.R. Grace, A.P. Watkinson, C.J. Lim, A. Erguedenler, Equrium modeling of gasification: a free energy minimization approand its application to a circulating fluidized bed coal gasifier,80 (2001) 195–207.

18] O. Teichmann, G. Saller, G. Funk, W. Krumm, Die Relaxationsmode zur Bestimmung der Gleichgewichtszusammensetzung derbei der Verbrennung von Kohlenwasserstoffen und VergasungBiomasse, Muell und Abfall 30 (3) (1998) 153–158.

19] J. Gmehling, B. Kolbe, Thermodynamik, second ed., VCH Verlgesellschaft mbH, Weinheim, 1992.

20] JANAF Thermochemical Tables, third ed., Part 1 + 2, New Y1986.

21] J.M. Encinar, J.F. Gonzalez, J.J. Rodrıguez, M.J. Ramiro, Catalysand uncatalysed steam gasification of eucalyptus char: influenvariables and kinetic study, Fuel (2001) 80.

22] M. Kruse, H. Schoenfelder, J. Werther, A two-dimensional mfor gas mixing in the upper dilute zone of a circulating fluidibed, Can. J. Chem. Eng. 73 (1995) 620–634.

23] M. Kruse, J. Werther, 2D gas and solids flow prediction in cilating fluidized beds based on suction probe and pressure pmeasurements, Chem. Eng. Proc. 34 (1995) 185–203.

24] R. Clift, J.R. Grace, Fluidization, in: Davidson, Clift, Harrison (EdContinuous Bubbling and Slugging, second ed., London, AcadPress, 1985 (Chapter 3).

25] P. Simell, E. Kurkela, P. Stahlberg, Formation and catalytic decoposition of tars from fluidized-bed gasification, in: Bridgwater (EAdvances in Thermochemical Biomass Conversion, vol. 1, 199265.

26] H.-M. Yan, C. Heidenreich, D.-K. Zhang, Mathematical modelof a bubbling fluidised-bed coal gasifier and the significance offlow’, Fuel 77 (9–10) (1998) 1067–1079.


[27] H. Schoenfelder, M. Kruse, J. Werther, Two-dimensional modelfor circulating fluidized bed reactors, AIChE J. 42 (7) (1996)1875–1888.

[28] H. Schoenfelder, Modeling of Circulating Fluidized Bed Reactors,Dissertation TUHH, Shaker Verlag, Aachen, 1996.

[29] F. Chejne, J.P. Hernandez, Modelling and simulation of coal gasifi-cation process in fluidised bed, Fuel 81 (2002) 1687–1702.

[30] D. Kunii, O. Levenspiel, Fluidization Engineering, second ed.,Butterworth-Heinemann, Boston, 1991.

[31] H. Martin, Waerme- und Stoffuebertragung in der Wirbelschicht,Chem. Ing. Tech. 52 (3) (1980) 199–209.

[32] C. Chen, M. Horio, T. Kojima, Numerical simulation of entrainedflow coal gasifiers. Part II: Effects of operating conditions on gasifierperformance, Chem. Eng. Sci. 55 (2000) 3875–3883.

[33] I. Narvaez, A. Orıo, M.P. Aznar, J. Corella, Biomass gasification withair in an atmospheric bubbling fluidized bed. effect of six operationalvariables on the quality of the produced raw gas, Ind. Eng. Chem.Res. 35 (1996) 2110–2120.

[34] P. Vriesmann, E. Heginuz, K. Sjoestroem, Biomass gasification ina laboratory-scale AFBG: influence of the location of the feedingpoint on the fuel-N conversion, Fuel 79 (2000) 1371–1378.

[35] H.W. Gudenau, H. Meyrahn, A. Bellin, W. Hahn, Vergasungnachwachsender Rohstoffe fuer die CO2-neutrale Energiegewinnungmit Hilfe der Rheinbraun HTW-Vergasung, Final Report, Bun-desministerium fuer Ernaehrung, Landwirtschaft und Forsten, Bonn,1994.

[36] F.G. van den Aarsen, Fluidised bed wood gasifier—performance andmodelling, Dissertation TH Twente, University Publication, 1985.

[37] M. Ising, U. Balke, C.A. Unger, Energetische Nutzung von Holzund Biomasse durch Vergasung in der zirkulierenden Wirbelschicht,

99.[ Ver-

ndenH,

[ aij,r for

[ gen-ter,rsion,

[41] D. Schuller, B. Brat, Klaerschlammpyrolyse mit Rueckstandsverga-sung, Chem. Ing. Tech. 65 (4) (1993) 401–449.

[42] L. Garcıa, M.L. Salvador, J. Arauzo, R. Bilbao, CatalyticCO2 gasification of biomass in a fluidized bed reactor, In-fluence of catalyst weight/biomass flow rate and CO2/biomassratio, Fluidisation (II), Toulouse, France, 2000, pp. 657–664.

[43] S. Hamel, Mathematische Modellierung und experimentelle Unter-suchung der Vergasung verschiedener fester Brennstoff in atmo-sphaerischen und druckaufgeladenen stationaeren Wirbelschichten,Dissertation Uni Siegen, Fortschritt-Berichte VDI Reihe, VDI Ver-lag, Duesseldorf, 6 (2001) 469.

[44] R.G.E. Franks, Mathematical Modeling in Chemical Engineering—Chapter VI: Reaction Kinetics, John Wiley & Sons Inc., New York,1967.

[45] M.A. Field, D.W. Gill, B.B. Morgan, P.G.W. Hawksley, ReactionRate of Carbon Particles, Combustion of Pulverised Coal, Cheney& Sons Ltd., Banbury, England, 1967 (Chapter 6).

[46] I. Matsui, D. Kunii, T. Furusawa, Study of char gasification by car-bon dioxide. 1. Kinetic study by thermogravimetric analysis, Ind.Eng. Chem. Res. 26 (1987) 91–95.

[47] I. Matsui, D. Kunii, T. Furusawa, Study of char gasification by car-bon dioxide. 2. Continuous gasification in fluidized bed, Ind. Eng.Chem. Res. 26 (1987) 95–100.

[48] F.L. Dryer, I. Glassman, High-temperature oxidation of CO and CH4,in: Proceedings of the 14th Symposium (Int.) on Combustion, Pitts-burgh, Pansylvania, 1973, pp. 987–1003.

[49] E.V. Kerinin, E.I. Shifrin, Mathematical Model of Coal Combustionand Gasification in a Passage of an Underground Gas Generator,Combustion, explosion and Shock Waves, Academy of Sciences,

[ forocess

[ ion,

[ the-ses,

[ in aSci.,

Final Report, UMSICHT, 94 NR 140-F, Institute Publication, 1938] P. Mehrling, R. Reimert, Entwicklung eines Verfahrens zur

gasung von Biomassen nach dem Prinzip der ZirkuliereWirbelschicht, Final Report BMFT-FB-T 86-009, Lurgi GmbFrankfurt, 1986.

39] S.R.A. Kersten, W. Prins, B. van der Drift, W.P.M. van SwaPrinciples of a novel multistage circulating fluidized bed reactobiomass gasification, Chem. Eng. Sci. 58 (2003) 725–731.

40] A. Chamberland, R. Labrecque, Parametric study of an oxyblown fluidised-bed reactor for wood gasification, in: BridgwaKuester (Eds.), Research in Thermochemical Biomass ConveElsevier Applied Science, London, NY, 1988, pp. 1125–1140.

Russia, 1993, pp. 148–154.50] V. Bıba, J. Macak, E. Klose, J. Malecha, Mathematical Model

the Gasification of Coal under Pressure, Ind. Eng. Chem. PrDes. Dev. 17 (1) (1978) 92–98.

51] Y. Wang, C.M. Kinoshita, Kinetic model of biomass gasificatSolar Energy 51 (1) (1993) 19–25.

52] R.A. Srinivasan, S. Siramulu, S. Kulasekaran, P.K. Agarwal, Mamatical modeling of fluidized bed combustion-2 combustion of gaFuel 77 (9–10) (1998) 1033–1049.

53] J. Petersen, J. Werther, Modeling three-dimensional effectscirculating-fluidized-bed gasifier for sewage sludge, Chem. Eng.submitted for publication.

experimental investigation and modeling of gasification of sewage sludge in the circulating...

Documents