Fluidised bed combustion differs from conventional pulverised coal combustion. It uses a continuous stream of air to createturbulence in a mixed bed of fuel, inert material and coarse fuel ash particles. Combustion occurs at lower temperatures typicallybetween 800°C and 900°C. This report reviews the fundamental aspects of the technology. It provides an up-to-dateunderstanding of the mechanisms and chemical reactions involved in fluidised bed combustion and how process parametersand/or fuel quality may affect fluidisation, combustion and pollutant formation.
Acromyms and abbreviations
BFB bubbling fluidised bedBFBC bubbling fluidised bed combustionCFB circulating fluidised bedCFBC circulating fluidised bed combustionDTI Department of Trade and IndustryEU European UnionFBC fluidised bed combustionFET flame extinction timeNOx nitrogen oxide and nitrogen dioxide (NO and NO2)PBFBC pressurised bubbling fluidised bed combustionPFBC pressurised fluidised bed combustionPCFBC pressurised circulating fluidised bed combustion
2 IEA CLEAN COAL CENTRE
Nomenclature
A correlation constant, s/mmnAb bed cross-sectional area, m2
Ao distributor area per orifice, m2
dc initial diameter of char particle, mdp initial coal particle diameter, mmD bed diameter, mDb bubble diameter, mDb0 initial bubble diameter, mDbm maximum bubble diameter, mDo molecular diffusivity of oxygen, m2/sDp particle diameter, mg acceleration due to gravity, m/s2
h overall heat transfer coefficient, W/m2Khg heat transfer coefficient due to the gas phase, W/m2Khp heat transfer coefficient due to the particulate phase,
W/m2Kka attrition rate constantkc reaction rate on the external surface of char, kg of
carbon/m2 skm mass transfer coefficient of oxygen, kg/m2 .Pan power-law correlation constantq burning rate on the external surface, kg of carbon/m2 s Q heat flux, W/m2
R universal gas constant, 8.315 kPa.m3/kmol.K Re particle Reynolds numberSh Sherwood numbertb burnout time of a char particle, stv devolatilisation time, s Tm mean temperature of the diffusion layer around a
carbon particle, K u fluidising air velocity, m/sub bubble velocity, m/sub` velocity of an isolated bubble, m/sumf minimum fluidising velocity, m/sut particle terminal velocity, m/s
W rate of weight loss of a burning char due tocombustion, kg/s
Wc overall weight loss of a burning char, kgz height above air distributor, m
Fluidised bed combustion (FBC) technology was first used toburn petroleum coke and other refinery wastes. It was only inthe 1960s that coal was considered as a fuel for FBC. FBCtechnology has good fuel flexibility and an ability to reduceemissions. During the past three decades, FBC in its variousforms has been used to burn all types of coals, coal wastesand a wide variety of other low-quality fuels (such asbiomass and leather wastes), either singly or cofired withcoal. In particular, it is suitable for those coals which aredifficult to mill and fire in pulverised coal combustionboilers. FBC technology is commercially available formodules up to 300 MWe, with a number of projects beingplanned or presently implemented in the 250–350 MWe sizerange (World Bank, 2002). At present, FBC represents onlyabout 2 per cent of the total coal-fired capacity worldwide.However, it is expected to continue to be adopted for avariety of applications ranging from small industrial boilersand furnaces to large power generation units.
In the last decades, research has been undertaken to investigatevarious issues associated with fluidised bed combustion. Thisreport reviews the fundamental aspects of the technology. Itcomplements earlier reports from the Clean Coal Centre onUnderstanding pulverised coal combustion (Morrison, 1986)(currently being updated) and Understanding coal gasification(Kristiansen, 1996). The report is centred on coal combustionbut, where appropriate, also includes firing or cofiring withalternative solid fuels. It provides an up-to-date understandingof the mechanisms and chemical reactions involved influidised bed combustion and how process parameters and/orcoal quality may affect fluidisation, combustion and pollutantformation. Technological developments, commercialapplications and economics of FBC have been detailed inearlier Clean Coal Centre reports by Scott (2001), Scott andNilsson (1999) and Takeshita (1994), and are thereforeexcluded from this review.
The report features four main chapters. Chapter 2 discussesthe main characteristics of FBC, and briefly describes thefour main derivatives of FBC, namely bubbling fluidised bedcombustion, circulating fluidised bed combustion,pressurised bubbling fluidised bed combustion andpressurised circulating fluidised bed combustion, to whichdiscussions in subsequent chapters are referred. In Chapter 3,the basic phenomena and processes of gas-solids fluidisationare reviewed. This includes discussions on the mechanismsof fluidisation, gas flow characteristics, bed particledynamics, bubble dynamics, and gas-solids mixingcharacteristics. Chapter 4 examines the processes ofcombustion, including heating and drying, devolatilisation,combustion of volatile matter, and combustion of the residualchar particles. It also discusses how process parameters andcoal quality may affect combustion efficiency. This isfollowed by Chapter 5 which looks at the formation andcontrol of pollutants in fluidised bed combustors such asnitrogen oxides, nitrous oxide, sulphur oxides, particulates,carbon monoxide, and trace elements. Factors whichinfluence these emissions are also discussed.
5Understanding fluidised bed combustion
2 Fluidised bed combustion systems
Fluidised bed coal combustion differs from bothconventional stokers and pulverised coal combustion. It usesa continuous stream of air to create turbulence in a mixedbed of inert material and coarse fuel ash particles. Thevelocity of the gas stream ensures that the particles remainsuspended and move about freely. In this state they behavelike a fluid, in other words, the bed becomes fluidised. Whenfuel is added to a hot fluidised bed, the constant mixing ofparticles encourages rapid heat transfer and completecombustion. It also allows a uniform temperature to bemaintained within the combustion zone. The ash producedaccumulates in the bed, eventually forming the bulk of theparticles. Surplus ash is drawn off at intervals to maintain thebed at the correct depth. Most of the heat generated isrecovered by a water/steam system, usually via water tubesimmersed in the bed.
FBC offers a number of significant advantages overconventional combustion systems, including (DTI, 1999)● the bed can be operated at temperatures typically
between 800°C and 900°C above which ash may meltand form slag;
● the scrubbing action of the moving particles on theimmersed water tubes increases the rate of heat transfer;
● the bed has a substantial thermal capacity which allowsa variety of fuels to be burned, including low- qualityfuels with a high level of unwanted mineral matter ormoisture content as well as mixed fuels;
● NOx formation is reduced because the combustion takesplace at relatively low temperatures and the systemoffers opportunities for air staging;
● up to 90% of the SO2 released during combustion can becaptured in the bed by adding a suitable sorbent such aslimestone or dolomite.
However, there are also some disadvantages with thetechnology (DTI, 1999; Anthony, 2003):● commercially proven only at relatively small scale
compared with pulverised coal combustion;● relatively large amounts of solid residues generated
(with sorbent addition), some of which require specialmeasures for disposal;
● higher carbon-in-ash levels than those from pulverisedcoal combustion;
● increased N2O formation (with coal/coke or similarfuels) due to the lower combustion temperatures.
There are currently four different forms of FBC. In terms ofthe fluidising gas velocity, FBC can be divided into twogroups: bubbling fluidised bed combustion (BFBC) whichtakes place at low gas velocities; and circulating fluidisedbed combustion (CFBC) which occurs at high gas velocities.Both technologies are commercially available for modules upto 300 MWe (World Bank, 2002). In terms of operatingpressure, there are also pressurised bubbling fluidised bedcombustion (PBFBC) and pressurised circulating fluidisedbed combustion (PCFBC). PBFBC is being demonstrated atseveral plants worldwide (EU, 2002; Scott, 2001). There is
6 IEA CLEAN COAL CENTRE
significant scope for both technological and economicimprovements with the technology. PCFBC is some waybehind PBFBC and remains at an earlier stage ofdevelopment (Minchener and others, 2000).
Each of the combustion systems is briefly described below.The differences in operating conditions and performancebetween the four systems are discussed.
2.1 Bubbling fluidised bedcombustion
A schematic diagram of a BFBC is shown in Figure 1. Fuelis fed onto a packed bed of inert particles (most commonlygraded sand). An upward air flow is introduced into the bedvia a distributor plate (sometimes a series of closely packedtubes) so that the particles are fluidised. The distributorprovides a uniform flow of air across the whole base area ofthe fluidised bed. It also supports the bed particles, withoutallowing them back into the plenum chamber, if the upwardair flow is turned off and the particles become defluidised orslumped. Fuel is burned in the bed and at times in thefreeboard above the bed. The combustion takes place attemperatures of around 800–900°C which are considerablylower than in pulverised coal combustion. The temperaturesare achieved and stabilised by the opposing effects of theheat input from the burning fuel and outgoing heat in theflue gases and heat transferred to immersed water tubes. Inthe case of a boiler, such tubes form a part of the boilerconstruction. When BFBC is used for applications other thana boiler, such as a hot gas furnace or incinerator, there are nosuch water cooled surfaces. In this case the bed temperatureis stabilised by passing excess air through the bed(Minchener and others, 2000).
Flue gases leave the bed and pass through a section knownas the freeboard. The freeboard can be designed with a much
fluidisedbed
airinlet
coalfeed
cyclone
gas exit
hotwater
orsteam
convectivesectionboiler
freeboard
fine ashstandpipe
distributor plate
bed offtake
plenumchamber
Figure 1 Bubbling fluidised bed combustionsystem (Merrick, 1984)
larger diameter than the bed. In this case, a reduced gasvelocity in the section ensures that coarse solid particlesentrained in the gas flow fall back to the bed by gravity.However, fine particles, which are contained in the feedmaterial and produced by the processes of combustion andabrasion in the bed, will be carried out in the flue gas streamfrom the combustor. Additional air, known as secondary air,is often introduced into the freeboard region to ensurecomplete combustion of the fuel.
After exiting the combustor, the flue gases pass into aconvective section where heat is further recovered and thegases are cooled to below 200°C. They then further passthrough a particulate control unit that can be a cyclone or amore efficient device such as a bag filter or an electrostaticprecipitator. The cleaned gases are finally discharged into theatmosphere through a stack.
Combustion leaves behind the mineral matter in the coal andthe spent sorbent (if added when sulphur removal isrequired) in the bed. As discussed above, fine materials areentrained with the flue gases leaving the combustor.However, the remaining coarse particles can accumulate inthe bed. Any excessive material is removed by an off-takefrom the bottom of the bed or via an overflow weir. Thisensures the bed is maintained at the designed depth.
2.2 Circulating fluidised bedcombustion
CFBC (Figure 2) is a development of bubbling bedcombustion technology. As with the latter, combustion air iscontinuously blown through the bed from the bottom of thecombustor. Combustion takes place at temperatures ofaround 800–900°C, similar to those in a bubbling bed.However, CFBC employs a higher fluidising velocity and afiner bed material. As a result, the particles rapidly becomeentrained in the gas flow and pass through the whole of themain combustion chamber. Large particles remain near the
7Understanding fluidised bed combustion
Fluidised bed combustion systems
base of the bed, and may gradually be worn down by theturbulence. Finer particles occupy the upper part of the bedand are carried out of the combustion chamber in the gasstream.
A heavy-duty cyclone is used to separate the entrainedparticles from the flue gas and return them to the bed. Thisallows the particles to remain within the system for longenough to ensure both effective combustion of the carbonand maximum sulphur capture. The finest particles, however,are not recirculated and escape from the cyclone.
The method of extracting heat from a CFBC varies withdesign. Heat transfer surfaces may be incorporated into thewalls of the main combustion chamber and cyclones.Alternatively, a secondary fluidised bed heat exchanger, withimmersed tubes, may be used. In the latter case, theexchanger is fluidised by air at a low velocity, this air thenbeing used as pre-heated secondary combustion air for themain combustor. On exiting the exchanger, the particles arereturned to the main combustor.
After leaving the cyclone, the hot flue gases enter aconvection pass which includes a superheater, an air heaterand, in some of the more recent installations, a reheater. Inthis section much of the remaining heat is extracted. Thecooled gases pass through a baghouse for removal of fineparticles and finally are discharged to a stack.
CFBC has a number of advantages over BFBC including(Merrick, 1984):● combustion and sulphur retention efficiencies are
improved, which results from the use of finer particles,turbulent gas-particle mixing and a high recycle rate;
● the use of high fluidising velocities makes a CFBCmuch smaller in bed area than a BFBC of the sameoutput. However, this advantage is offset by itsadditional height and the size of the fluidised bed coolerwhich operates at a lower fluidising velocity;
● the number of fuel feed points is reduced substantiallydue to the small combustor size and turbulent mixingconditions;
● erosion and corrosion of heat transfer tubes are reduceddue to: the tubes immersed in the fluidised bed coolerare subjected to significantly lower gas and particlevelocities than in a BFBC; oxidising conditions prevailthroughout the cooler whereas reducing conditions occurnear the fuel feed points in a BFBC;
● convective heat transfer coefficients are increasedbecause of the use of smaller particles. As a result, lessheat transfer tubing would be required in the fluidisingbed cooler than in the bed of a BFBC.
In spite of the above attractions, there are also some areas ofconcern with CFBC technology (Merrick, 1984). Firstly, thepressure drop across a CFBC is generally greater than for aBFBC. This results in increased fan power requirements.Secondly, the large recycle rates require high efficiencies ofcyclone for the recovery of the bed solids from the gasstream. In addition, the high gas velocities combined withthe high particulate loadings may lead to erosion in thecombustor, cyclone and associated ducting.
waterwall heatrecovery
superheaterheat recovery
coal andlimestone
combustor
economiserheat recovery
to baghouse
convective passcyclone
secondary airprimary air
ash
fluidised bed heat recovery
Figure 2 Circulating fluidised bed combustionsystem (DTI, 1999)
2.3 Pressurised bubbling fluidisedbed combustion
With a PBFBC system, the combustion takes place within alarge pressure vessel at pressures of 1–1.5 MPa andtemperatures of around 800–900°C. As the system operatesunder pressure, it can be physically substantially smallerthan a BFBC unit. In a PBFBC unit, combustion air for theprocess is supplied to the pressure vessel and combustor. It isthen forced through a distributor plate into the bottomsection of the fluidised bed. An increased pressure results insmaller and more frequent bubbles which, in turn, providesfor smoother fluidisation. This leads to better gas and solidcontact (mixing). Therefore, the higher pressure willeventually produce a higher combustion efficiency.
In a PBFBC, both coal and sorbent have to be fed across thepressure boundary, and similar provision for ash removal isnecessary. As a result, various arrangements have beendeveloped to meet the need. Crushed coal can be transferredby lock hoppers from ambient pressure to the operatingpressure. It is fed to the bottom of the bed by a pneumatictransport line. Lock hoppers are also used to remove thecyclone fines and excess bed material after cooling.
In combined cycle power applications, about 80% of theelectricity is generated in a conventional steamturbine-generator set. Steam for the steam cycle is producedin tube bundles immersed in the fluidised bed combustor. Atthe same time, the hot flue gases leaving the combustor areunder pressure. They are fed to high-efficiency cyclones orhigh-temperature filters that remove about 99.9% of thesolids. The cleaned gases are then expanded through a gasturbine set to generate the balance (20%) of the electricity.As the gases still contain some particulates, the gas turbinemust be erosion-resistant. It uses blades of a more robustdesign than those found in a conventional gas turbine. Inaddition, the blades are also coated with erosion-resistantmaterial. After leaving the gas turbine, the exhaust gasesfurther pass through the economiser where more heat isrecovered. Finally, they are fed to an electrostatic precipitatoror a bag filter for removal of residual particulates prior torelease to the stack.
Compared to BFBC, PBFBC achieves considerably higherefficiencies of combustion and sulphur retention. This is duemainly to the following factors (Merrick, 1984):● at elevated operating pressure fluidisation tends to
become smother, resulting in better gas-solidscontacting;
● with elevated pressure operation, deeper beds of 2–4 mcan be used. This is because the additional pressure dropassociated with a deep bed has less effect on thecompression energy requirements than in the case ofatmospheric pressure operation. In the latter case, deepbeds are impractical due to the high fan powerrequirements;
● as the throughput per unit of bed area increases with thepressure, adequately high throughputs can be obtained atmoderate fluidising velocities of 1–2 m/s.
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Fluidised bed combustion systems
2.4 Pressurised circulatingfluidised bed combustion
As with PBFBC, in a PCFBC system, the combustor and hotgas cyclones are held inside a large pressure vessel. Thesystem also operates at higher pressures of 1–1.5 MPa andcombustion temperatures of around 800–900°C. Fuel is fedin a slurry/paste form utilising conventional solids handlingtechnologies. Sorbent is supplied pneumatically usingconventional lock hoppers or fed in paste form along withthe fuel.
In combined cycle power applications, steam produced in thefluidised bed combustor and the convective heat recoverysection contributes to about 80% of the electricity generated.The gas turbine generates the remaining 20% of theelectricity. PCFBC systems employ a ceramic filter forremoval of particulates in the flue gases before entering thegas turbine. As the ceramic filter will remove virtually all theparticulates, a fairly standard gas turbine can be used (EU,2002).
3 Fluidisation
This chapter reviews the basic phenomena and processes offluidisation. The mechanisms of fluidisation are firstexamined, highlighting the differences and common featuresbetween a bubbling bed and a circulating bed. This isfollowed by detailed discussions on gas flow characteristics,bed particle dynamics, bubble dynamics, and solids mixingcharacteristics. Finally, heat transfer in fluidised beds isaddressed.
3.1 Mechanisms of fluidisation
A fluidised bed consists of a packed bed of particles above agrid so that a fluid (air in the case of fluidised bedcombustion) can be passed upwards through the bed. Thebed material commonly includes the following components:● coal (present mainly as low volatile char particles,
accounting for less than 1% of the bed material ifcrushed coal is used and less than 5% if lump coal isused);
● limestone or dolomite (if added) which is present mainlyas calcined and partially sulphated stone;
● coal ash which is derived from the mineral matteroriginally present in the coal;
● inert additive (most commonly, graded sand) which isused when insufficient ash, limestone or dolomite arebeing fed to maintain the design mass of bed material(for example, when a washed coal containing only fineinherent ash is burned without limestone or dolomiteaddition).
As the gas velocity increases, the bed undergoes threedistinct stages of fluidisation namely: fixed bed, bubblingfluidised bed, and circulating fluidised bed (illustrated inFigure 3). The characteristics of each stage of fluidisation arediscussed below.
9Understanding fluidised bed combustion
Fixed bedAt low gas velocities, the gas simply flows through theparticle interstitial space without disturbing the particlessignificantly. The particles remain stationary, the bed acts asa porous medium and is called a fixed bed. As the gasvelocity increases, there will be a point at which thegas-particle drag force compensates for the bed weight. Atthis point, inter-particle distances increase, the bed expands,and the particles appear to be suspended in the gas. Thismarks the onset of fluidisation. The gas velocity at this pointis referred to as the minimum (or incipient) fluidisingvelocity and depends on the size of the bed particles.
Bubbling fluidised bedWhen the gas velocity exceeds the minimum fluidisingvelocity, the excess gas passes through the bed as bubblesand the remainder of the gas leaks through the bed material.The bed is then considered to be heterogeneous or bubbling.In practical applications, bubbling fluidised beds areoperated at gas velocities that are several times higher thanthe minimum fluidising velocity. Under these conditions, thebubbles passing through the bed typically occupy 20–50% ofthe bed volume. The passage of the bubbles, in upwards andsideways coalescing movements, gives intensive agitationand mixing of the bed particles. In this stage, however, theparticles still remain in relatively close contact and are notcarried upwards to any significant degree. As a result, thebed maintains a well-defined upper surface, with air bubblespassing through the bed bursting at the surface.
Circulating fluidised bedWhen the gas velocity is further increased to a pointexceeding the particle free-fall velocity, the bed particlesbecome entrained in the gas stream and are carried out of thecontaining vessel. At this point, the gas velocity is referred to
circulating fluidised bed
high gas velocity
gas and solids
recirculated solids
bubbling fluidised bed
medium gas velocity
gas
fixed bed
low gas velocity
gas
Figure 3 Stages of fluidisation (Merrick, 1984)
as the particle terminal velocity. A steady-state operation ofthe fluidised bed, at velocities greater than the particleterminal velocity, is possible if the bed particles arecontinuously supplied. This is generally achieved byseparating the entrained solids from the gas stream using acyclone and recycling them to the bed. Fluidised bed systemsthat operate in this mode are termed fast fluidised beds orhereafter circulating fluidised beds.
A circulating fluidised bed differs from a bubbling fluidisedbed in a number of aspects (Minchener and others, 2000;Merrick, 1984):● higher gas velocities are employed;● finer bed particles are used;● the containing vessel is filled with a turbulent cloud of
particles that no longer remain in close contact witheach other;
● no bubbles are formed; instead clusters and strands ofparticles are turbulently mixed with the gas stream andare continuously breaking up and reforming;
● the density of the bed depends on the flow rate ofrecycled particles. If this is reduced, the system changesfrom a circulating fluidised bed to a dilute phase system.
3.2 Gas flow characteristics
Fluidising velocityFluidising velocity is one of the key parameters for fluidisedbed design. Its choice affects most of the other designparameters. For bubbling fluidised beds, the fluidisingvelocity usually varies from 1 to 3 m/s (Minchener andothers, 2000). In circulating fluidised beds, however, higherfluidising velocities of up to 8 m/s are used (DTI, 1999).
The actual value of fluidising velocity depends on acompromise between the capital cost, the bed pressure dropand the efficiencies of combustion and sulphur retention(Merrick, 1984). A high fluidising velocity reduces the bedarea required for a given output and thereby decreases thesize of the bed containment and distributor plate and thenumber of fuel feed points. Hence, capital costconsiderations favour a high fluidising velocity. However, ahigh fluidising velocity gives a high heat release per unitplan area of the bed. As a result, a deeper bed would berequired in order to accommodate the additional heat transfersurface. This increases the bed pressure drop and thereforefan power requirements and operating costs. In addition, ahigh fluidising velocity leads to more combustion in thefreeboard and increases the carry-over of unburned char andunconverted sulphur acceptor. This subsequently reduces thecombustion and sulphur retention efficiencies.
Air distributionIn spite of the use of an air distributor, the air flow in thecross-section of a fluidised bed is not ideally uniform. It isgenerally seen that the air velocity has the maximum at thebed centre and decreases gradually with increasing thedistance from the centre (Moran and Glicksman, 2001).
Solids effectsMoran and Glicksman (2000, 2001) studied the effects of
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Fluidisation
solids on air velocity in a CFB. The air velocities at the bedcentre, with and without solids, were recorded over a periodof about 7 seconds. It appeared that the presence of solidsproduced a much greater fluctuation in air velocity with afrequency of 0–200 Hz. Without solids flow the velocityfluctuation was about 0.5 m/s while it increased to 3 m/s at asolids concentration of 0.924%. This was most probablyattributed to particle clustering inside the CFB riser.
In the above measurements, the mean air velocity at the bedcentre was also found to increase from 2.06 m/s to 4.14 m/sdue to the introduction of solids. This phenomena wasfurther investigated by measuring the mean air velocitiesacross the bed diameter for solids concentrations of 0.187%,0.384%, and 0.77% (Moran and Glicksman, 2001). It wasshown that the mean air velocities increased with increasingthe solids concentration at the same air flow rate of0.059 m3/s. This can be understood by looking at the solidsflow pattern in the riser (further discussed in Section 3.5). Ina CFB riser, a more dilute layer of solids moves upwards inthe core zone whereas a denser layer of solids travelsdownwards in the wall zone. It can be assumed that thedownward solids flow drags and entrains a certain amount ofair, forming a downward moving air layer near the wall.Also, this boundary layer is expected to increase withincreasing local solids concentration. It accordingly reducesthe effective cross-sectional area and therefore increases theair velocities through a narrower core region (Moran andGlicksman, 2001).
3.3 Bed particle dynamics
In practical applications, the bed particles are present in arange of sizes and have irregular shapes. Both the particlesize and shape affect the particle drag and terminal velocityin the air and, eventually the particle fluidising behaviour.
Particle sizeFor bubbling fluidised beds, the mean diameter of the bedparticles typically ranges from 0.5 mm to 1.5 mm (Merrick,1984). The actual value depends mainly on the choice offluidising velocity, following a linear relationship. A higherfluidising velocity generally requires a coarser bed in orderto obtain good fluidisation and to avoid excessive carry-overof bed material. For circulating fluidised beds, much finerbed particles with a mean diameter about 150 µm are used.This, in combination with a high fluidising velocity, ensuresthat the solids are entrained in the gas flow and circulated inthe system.
The bed particle size is determined by the particle size offeed materials such as crushed coal, sulphur acceptor or inertadditive. In practice these feed materials are in a wide sizerange. For bubbling fluidised beds, it is usual to limit themaximum size of feed materials to 2–6 mm in order toobtain the required mean bed particle size of 0.5–1.5 mm(Merrick, 1984).
Particle terminal velocityThe particle terminal velocity is defined as the constantvelocity reached by a free-falling particle in stagnant air. It
has a physical significance that for gas velocities greater thanthe terminal one, particles will be carried out of the bed.Hence, it gives the maximum operating gas velocity for aBFB and the minimum one for a CFB.
The particle terminal velocity, ut, depends on the particle sizeand density. It is obtained by balancing the gravitational,buoyancy and drag forces. For a spherical particle, ut is givenby the following relationships (Fueyo and Dopazo, 1995):
ut = g(rp - rg)Dp2/(18µ) for Re<0.4 (1)
ut = Dp [4g(rp - rg)2 g2 /(225µrg)]1/3 for 0.4<Re<500 (2)
ut = Dp [3.1g(rp - rg)/rg)]1/2 for 500<Re<2×105 (3)
where g is the acceleration due to gravity, rp is the density ofthe particle, rg is the density of the gas, Dp is the particlediameter, µ is the gas dynamic viscosity, and Re is theparticle Reynolds number.
An irregularly shaped particle will generally experience agreater drag force than a spherical one of the same volumeand hence settle slower (Wu and Colbeck, 1996; Wu, 2000).It is therefore necessary to take into account the effect ofparticle shape. A number of correlations have been publishedin the literature and this is detailed in a review by Fueyo andDopazo (1995).
Particle fluidisation behaviourThe behaviour of fluidisation depends on the characteristicsof the particles. Geldart (1973) suggested that the particlesize and particle-gas density difference are the keyparameters affecting fluidisation behaviour. He classifiedparticles into four different groups (see Figure 4) asdescribed below.● Group C includes particles with high cohesivity, very
small size and strong electrostatic charges. Theinterparticle forces are greater than those which the fluidexerts on the particles. As a result, the gas passes upvoids extending from the distributor to bed surface.
11Understanding fluidised bed combustion
Fluidisation
Fluidisation with such particles is therefore extremelydifficult and particle mixing and consequently heattransfer between a surface and the bed are very poor;
● Group A particles have a small mean size and/or a lowdensity (less than about 1.4 g/cm3). Beds with suchparticles expand considerably before bubbles appear.When the air supply is turned off the bed collapsesslowly, typically at a rate of 0.3–0.6 cm/s. Bubbles, oncethey appear, generally rise faster than the interstitial gas;
● Group B contains particles with a mean size of40–500 µm and a density of 1.4–4 g/cm3. With Group Bparticles, bubbles start to form at or only slightly abovethe minimum fluidising velocity. Bed expansion is smalland the bed collapses rapidly when the air supply isturned off. Similarly, most bubbles rise faster than theinterstitial gas;
● Group D particles are generally larger and/or denserthan those in other groups. They require, due to theirweight, much higher fluidising velocities than particlesin groups A and B. It appears that bubbles do not formuntil about 5 cm above the air distributor. All but thelargest bubbles rise more slowly than the interstitial gas.
3.4 Bubble dynamics
Bubbling is a basic phenomenon of bubbling fluidised beds.The behaviour of bubbles is generally complicated and noteasily predicted or measured. However, it is crucial to severalaspects of the fluidised bed performance, including (Fueyoand Dopazo, 1995): ● mixing. The upward motion of bubbles enhances solids
mixing and hence promotes heat and mass transfer; ● bed expansion. The bubble-phase volume determines the
extent of bed expansion and hence affects the overallbed height;
● elutriation. The bursting of bubbles at the bed surface,which throws particles into the freeboard zone, increaseselutriation.
Bubble formationThe process of bubble formation has been investigated usingvarious computer simulations (Laux and Johansen, 2003;Ouyang and others, 2001). It has been shown that when gasenters the air distributor, bubbles are formed at the orificeswith small sizes. They then rise towards the bed surface witha growing size or converge into larger bubbles. Somebubbles coalesce horizontally but vertical coalescence isgenerally prominent. At the bed surface some bubbles breakup, leading to fluctuation of the bed surface. In the bubbleformation process, the bed expands gradually from a fixedbed at the beginning and finally reaches a roughly stablesurface during the fluidisation.
Bubble geometryIn a fluidised bed, the shape of a bubble is either nearlyspherical or a spherical cap with a rear indentation(see Figure 5). The indentation is usually filled with particlesthat move upwards with the bubble, forming the so-calledwake. It has been reported that the angle u increases withincreasing particle diameter and an increase in operatingpressure has the opposite effect (Fueyo and Dopazo, 1995).
2000
1000
100 1000
Dp µm
r p-r
gkg
/m3
30004000
C
AB
D
Figure 4 Geldart's (1973) particle classificationaccording to fluidisation behaviour(reproduced from Fueyo and Dopazo, 1995)
The size of a bubble, Db, can be predicted from arelationship suggested by Darton and others (1977):
Db = 0.54 (u-umf)0.4 (z + 4Ao0.5)0.8 g-0.2 (4)
where u is the gas velocity, umf is the minimum fluidisingvelocity, z is the height above air distributor, and Ao is thedistributor area per orifice.
Bubbles grow while rising towards the bed surface. Whenthe maximum size is reached, they become unstable and aresplit into smaller ones. Mori and Wen (1975) gave the initialbubble diameter, Db0, as:
Db0 = 0.347 [Ao(u-umf)]0.4 (5)
and the maximum bubble diameter, Dbm, as:
Dbm = 0.652 [Ab(u-umf)]0.4 (6)
where Ab is the bed cross-sectional area.
Bubble rising velocityThe velocity of an isolated bubble without wall effects, ub`,can be calculated using the following equation (Davidsonand Harrison, 1963):
ub` = 0.71 (gDb)0.5 (7)
It is indicated that smaller bubbles rise more slowly thanlarger ones. In the presence of walls, the isolated bubblevelocity needs to be corrected. Wallis (1969) gave the bubblevelocity as:
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Fluidisation
ub = 1.13 ub` e-Db/D (8)
where D is the diameter of the bed. Equation (8) only appliesto 0.125#Db/D#0.6. Outside this range, the bubble is eithersmall enough so that wall effects can be neglected, or largeenough to be considered as a slug.
Bubble interactions change the bubble size. When bubbleinteractions are considered, the bubble velocity is given bythe following equation (Davidson and Harrison, 1963):
ub = ub` + u-umf (9)
Stein and others (2000) compared bubble and averageupward solids velocities at various heights above thedistributor. The bubble velocity was calculated according toEquations (4) and (7). The average upward solids velocitywas measured, using a column with a 141 mm diameter and600 mm height at u-umf = 0.15 m/s and 0.3 m/s. It wasshown that the bubble velocity is about twice as much as theupward solids velocity except in the distributor region.
3.5 Solids mixing
The mixing of solids in a fluidised bed has significant effectson the mass and heat transfer processes. It can also affect thechemical reactions and therefore the combustion and sulphurretention efficiencies. A weak mixing may result inbypassing the fluids and formation of dead zones (Mostoufiand Chaouki, 2001). Solids mixing is affected by a numberof factors such as the arrangement of immersed tubes, gasvelocity, and particle characteristics. It behaves differently inbubbling and circulating fluidised beds, which are discussedbelow.
Bubbling fluidised bedIn bubbling fluidised beds, solids motion is driven by bubblemotion. Bubbles carry particles upward in their wakes anddrifts. This upward flow of solids is balanced by a downwardflow which occurs at places where there are no bubbles.Horizontal motion (lateral convective mixing) of particlesoccurs mainly at the bed surface where bubbles burst, andnear the air distributor where particles are carried away bybubbles. It is also seen in the regions with a high velocitygradient. However, vertical solids mixing is generally farmore significant and hence contributes more to the bedperformance (Stein and others, 2000).
Stein and others (2000) have shown that for relatively deepbeds, particles travel upwards in the central zone anddownwards in the wall zone. Both upwards and downwardsmotions are evident near the distributor. O’Brien andSyamlal (2001) have distinguished the solids flow patternsbetween a shallow bed and a deep bed. In a shallow bed of11.5 cm, the upward motion of bubbles induces a downwardflow of particles in the central zone. There is also a reversesolids flow in a small upper-outside region of the bed. In adeep bed of 23.1 cm, the downward flow region becomessmaller and higher up in the bed. However, the reverse solidsflow has intensified so that particles move upwards in thecentral zone and downwards in the wall zone.
u
Figure 5 Bubble with a spherical-cap shape(Fueyo and Dopazo, 1995)
Immersed cylinders can affect solids mixing in a bubblingfluidised bed (Ding and others, 2001). The cylinders act as abarrier to solids motion between the regions above andbelow the cylinder row. This results in a strong local particlecirculation in the lower region and hence a more extensivemixing. However, solids exchange between the two regionsis reduced with the degree of reduction depending on thenumber of cylinders.
Wang and others (2001a) studied the degree of solids mixingin bubbling fluidised beds using the Lacey index. The indexhas a value of unity when the components of a mixture arecompletely segregated. It decreases to 0 for a fully randommixture that has an identical probability of finding a particleof a component in every point. Figure 6a shows the effect ofgas velocity on the process of mixing. It shows for all thecases that as mixing proceeds the mixing index decreasesfrom unity to 0.28 at which an equilibrium of mixing isreached. This suggests that the degree of mixing achievableunder the conditions does not vary with gas velocity. It isnoted, however, that the time required to reach the mixingequilibrium decreases with increasing velocity. In otherwords, the rate of mixing increases with increasing gasvelocity.
Figure 6b shows the effect of particle density on the processof mixing. Note that the curve for rp=900 kg/m3 almostcoincides with that for rp=2650 kg/m3. It can be seen that atthe mixing equilibrium the mixing indexes are all around0.28 for the particle densities examined. This indicates thatparticle density has little effect on the degree of mixingachievable. However, the mixing rate is found to decreasewith decreasing particle density from 2650 kg/m3 to1590 kg/m3 (in a higher density range). It then increaseswhen particle density is further decreased to 900 kg/m3 (in alower density range).
Figure 6c shows how the process of mixing is influenced bythe particle diameter. It is interesting to note that the mixingindex at the mixing equilibrium increases from 0.28 to 0.46when the particle diameter increases from 1 mm to 2 mm. Inother words, the degree of mixing decreases. It is also shownthat the time required for the mixing index to decrease from1 to 0.46 increases from 2 s to 5 s. This implies that themixing rate decreases with increasing particle diameter.
Circulating fluidised bedSolids mixing in a CFB is intensive due to turbulentconditions. However, segregation may occur particularlywhen solids of a broad particle size distribution and/ordifferent densities are fluidised. Easily fluidised particlestend to be entrained by the fluidising gas while others tend tosink and remain in the lower part of the CFB. This results inspatial distributions of solids with different size and densityas well as mass concentration.
There have been extensive studies on solids mixingbehaviours in a CFB. It is shown that the CFB riser can becharacterised with different zones according to their solidsflow pattern. It is generally accepted that solids moveupwards in the core zone and downwards in the wall zone(Bruhns and others, 2001; Zhang and others, 2001). A
13Understanding fluidised bed combustion
Fluidisation
0.6
0.4
0.2
0.0
0 6 8
Time, s
Lace
y in
dex
2 4
0.8
1.0
u/umf=2.0, 1.75, 1.5, 1.25
(Dp=1 mm, rp=2650 kg/m3)
a)
0.6
0.4
0.2
0.0
0 6 8
Time, s
Lace
y in
dex
2 4
0.8
1.0
rp=900, 2650 kg/m3, rp=1590 kg/m3
(Dp=1 mm, u/umf=1.5)
b)
0.6
0.4
0.2
0.0
0 6 8
Time, s
Lace
y in
dex
2 4
0.8
1.0
Dp=1,2 mm
(u/umf=1.5, rp=2650 kg/m3)
c)
Figure 6 Effects of process parameters on theLacey index: (a) gas velocity; (b) particledensity; (c) particle size (Wang andothers, 2001)
number of studies have estimated the thickness of thedown-flow wall zone, measuring from the wall to a pointwhere the net particle flux is zero. For example, Zhou andothers (1995) reported an average thickness of around0.015 m for a 0.146 m diameter riser. Zhang and others(1995) studied a 1.72 m diameter bed and found an averagethickness of 0.09 m. They also analysed data from a numberof other sources. It was suggested that the wall zonethickness, d, is related to the bed diameter, D, according tothe following correlation:
d = 0.05 D 0.74 (10)
A number of models for solids mixing in a CFB have beenproposed (Bruhns and others, 2001; Kallio and other, 2001;Schlichthaerle and other, 2001). In the model by Bruhns andothers (2001), the CFB riser was vertically divided into threeregions, namely the dense bottom region, the upper diluteregion and the exit region. The bottom and exit regions wereassumed to be completely mixed. The upper dilute region(the main region) was further divided into an up-flowingcore zone and a down-flowing wall zone. Solids mixing inthis region was then modelled according to mass transferbetween the core and wall zones.
Figure 7 shows the simulated and measured axial profiles ofthe mean particle size. It can be seen that the mean particlesize decreases with increasing height above the distributorplate. Furthermore, higher values are found in the wall zonethan in the core zone. The axial profiles of particle densityare determined by mixing 60 kg quartz sand with 40 kg ironpowder (accounting for 40%). Figure 8a shows that the massfraction of higher density solids (iron powder) is higher inthe lower and wall zones. This together with the aboveindicates higher density and coarser solids tend to sink andmove to the wall zone. These findings generally agree withthose by other workers (Younis and others, 1999; Karri andKnowlton, 1998). It is also interesting to see that, when thegas velocity increases from 4.6 m/s to 7.5 m/s, the massfraction of iron powder falls to around 40% and becomesless distinctive both in the axial direction and between thewall and core zones (Figure 8b). This is due to the higherdegree of turbulence and hence the better mixing.
Winaya and others (2001) examined the axial profile ofsolids concentration for various gas velocities ranging5.35–7.68 m/s. The experiments were undertaken with sandwith a mean particle diameter of 245 µm and a density of2520 kg/m3. It was shown that for all the velocities examinedthe solids concentration is higher in the bottom zone anddecreases along the riser height. This is again attributed tothe sinking of solids. With increasing gas velocity, however,the carryover of solids from the bottom zone or the degree ofturbulence increases. As a result, the variation in the solidsconcentration becomes less distinctive.
3.6 Heat transfer
In fluidised beds, effective gas-solids contacting gives a highrate of heat transfer from burning fuel to immersed coolersurfaces/wall. The fraction of radiative heat transfer is
14 IEA CLEAN COAL CENTRE
Fluidisation
generally small compared with that of convective heattransfer. The transfer of heat by convection-enhancedconduction (including radiation) is expressed as:
Q = h DT (11)
where Q is the heat flux in W/m2, h is the overall averageheat transfer coefficient in W/m2K (typically in the order of200–500 W/m2K), and DT is a characteristic temperaturedifference between the bed and the wall.
400
350
300
250
0 6 8 10 12
Height above distributor plate, mu=4.6 m/s
Mea
n p
artic
le s
ize,
µm
2 4
450
500 wall zone (calculated)core zone (calculated)wall zone (measured)core zone (measured)
Figure 7 Axial profile of the mean particle size ina CFB riser (Bruhns and others, 2001)
50
45
40
35
0 6 8 10 12
Height above distributor plate, mu=4.6 m/s
Iron
pow
der
in m
ixtu
re, m
ass
%
2 4
55
60 wall zone (calculated)core zone (calculated)wall zone (measured)core zone (measured)
a)
50
45
40
35
0 6 8 10 12
Height above distributor plate, mu=7.5 m/s
Iron
pow
der
in m
ixtu
re, m
ass
%
2 4
55
60 wall zone (calculated)core zone (calculated)wall zone (measured)core zone (measured)
b)
Figure 8 Axial profile of the mass fraction of ironpowder in a CFB riser (Bruhns and others,2001)
The bed/wall heat transfer is a sequence of intermittentevents of either a dense phase or an almost particle-free gascoming close to the wall and exchanging energy with it. Thedense phase may be a highly loaded emulsion, cluster orpacket of particles, while the gas may be a bubble with asmall amount of particles. As a result, the overall heattransfer coefficient is found as
h = g hp + (1-g) hg (12)
where hp and hg are the average heat transfer coefficients dueto the particulate and the gas phase, respectively, g is theaverage fraction of the wall area occupied by the particulatephase. For dense beds, the first term on the right-hand side ofEquation 12 is important. For small solids concentrations orin the freeboard region, the second term may becomedominant (Fueyo and Dopazo, 1995).
The mean bed particle size has a significant effect on theheat transfer coefficient. It has been shown that theconvective heat transfer coefficient varies approximately asthe inverse square root of the mean bed particle size. Finebeds therefore require less heat transfer surface than coarsebeds. Typically, convective heat transfer coefficients in theorder of 200 W/m2K are found with a mean bed particle sizeof 1 mm (Merrick, 1984).
Figure 9 shows that the heat transfer coefficient increaseswith increasing solids concentration in a CFB (Winaya andothers, 2001). This is due to the presence of more particles,leading to greater heat interaction between the bed and thewall. The heat transfer coefficient is also found to increasewith bed temperature because of higher radiationcontribution. In addition, at higher bed temperature gasthermal conductivity is higher, which results in increased gasconvection.
3.7 Summary
Fluidising velocity is perhaps the most important parameterof fluidisation. Its choice affects most of the other processparameters. It typically ranges from 1 to 3 m/s for bubblingfluidised beds. In circulating fluidised beds, higher velocitiesof up to 8 m/s are used.
15Understanding fluidised bed combustion
Fluidisation
The mean diameter of the bed particles depends mainly onthe choice of fluidising velocity. It ranges from 0.5 mm to1.5 mm for bubbling fluidised beds. For circulating fluidisedbeds, much finer bed particles with a mean diameter about150 µm are used. The particle size, together with density andshape, determine its drag and terminal velocity in the air and,eventually fluidisation behaviour.
In bubbling fluidised beds, bubbles are formed at the airdistributor and rise towards the bed surface with a growingsize. When the maximum size is reached, they becomeunstable and split into smaller ones. In this process, the bedexpands and finally reaches a stage with a roughly stablesurface. Bubbles generally rise faster than the interstitial gasand upward solids except in the distributor region.
In fluidised beds good mixing is usually achieved. This givesa good heat distribution and, consequently, a uniformtemperature distribution. Solids mixing in bubbling beds isdriven by bubble motion. Its behaviour varies with thearrangement of immersed tubes, gas velocity, particledensity, and particle size. In circulating beds, solids mixingis caused by turbulent conditions. Segregation may occurparticularly when solids of a broad particle size distributionand/or different densities are fluidised.
Effective gas-solids contacting gives a high rate of heattransfer from burning fuel to immersed cooler surfaces/wall.The heat transfer coefficient varies as the inverse square rootof the mean bed particle size and therefore, fine beds requireless heat transfer surface than coarse beds. In contrast, thecoefficient is found to increase with solids concentration andbed temperature.
120
100
80
60
40
0
0 5 10 15 20 25
Solids concentration, kg/m3
Hea
t tra
nsfe
r coe
ffici
ent,
W/m
2 K
140
160
20
Bed temperature = 500°CBed temperature = 550°C
Figure 9 Effect of solids concentration on heattransfer coefficient (Winaya and others,2001)
4 Combustion
When coal is fed into a fluidised bed combustor, the particlesundergo a number of processes (see Figure 10) including:● drying (with or without shrinkage of the particles);● devolatilisation (with or without swelling and
fragmentation);● combustion of volatile matter; and● combustion of the residual char particles.
This chapter examines the mechanisms and chemicalreactions involved in each of the above processes. Theeffects of process parameters and fuel properties on thecombustion efficiency are then discussed.
4.1 Drying
The drying of a coal particle occurs once it enters a fluidisedbed. This process removes the fuel moisture which varieswidely from the typical range of 2–20% to extreme highvalues of up to 70%. It can be further divided into two stages(Anthony and Preto, 1995): ● the evaporation of surface moisture (which is present in
the coal naturally or, when coal is fed as a paste orslurry); and
● the loss of inherent moisture.
It appears that the evaporation of surface moisture does notaffect the coal combustion process directly. In the case ofslurried fuels, however, feeding such fuels may lead toagglomeration of particles.
Low rank coals can contain more than 40% inherentmoisture, and its evaporation may occur in conjunction withand slow down the devolatilisation and ignition processes. Inaddition, the loss of water can result in significantmorphological changes in low rank coals. A volumereduction of up to 40% was reported for an Australian browncoal during drying in a fluidised bed combustor (Agarwaland LaNauze, 1989).
4.2 Devolatilisation
Further heating releases the volatile matter of the coalparticle, that is, devolatilisation. Typically, devolatilisationstarts at about 300°C (Anthony and Preto, 1995) andproceeds at temperatures lower than that of the fluidised bed(Daki and others, 1997). It is a pyrolytic process whereby thelarge and heavy molecules of organic substances in the coalbreak up. The weaker carboxyl, hydroxyl and aliphatic bondsbreak up at lower temperatures, while at temperatures greaterthan 650°C the stronger heterocyclic components decompose(Ross and others, 2000; Strezov and others, 2002). Theprocess produces various gaseous species such aslow-molecular-weight hydrocarbons, hydrogen, carbonmonoxide, carbon dioxide, methane, and tars. It finallyleaves a solid carbonaceous residue (containing typically98% carbon) which is called char. The amount of char varies
16 IEA CLEAN COAL CENTRE
widely for different fuels, and is typically about 45 wt% forbituminous coal, 20 wt% for peat and sewage sludge, andapproximate 10 wt% for wood-based fuels and variouswaste-derived fuels (Kilpinen and others, 2002).
Devolatilisation phenomena During devolatilisation, coal particles are surrounded by acharacteristic luminous, plume-like flame. The flame isformed almost immediately or a few seconds afterintroduction into a fluidised bed combustor and for most ofthe time remains near the surface of the bed throughoutvolatiles evolution. The coal particles initially appear as darkspots inside the flame and eventually turn yellow towardsextinction of this flame (Urkan and Arkol, 1994).
The coal structure changes significantly duringdevolatilisation. This can be a sequence of events of coalparticle softening, resolidification and contraction (Strezovand others, 2002). Daki and others (1997; 2000) noted thatthe structural change is due to hydrocarbons transformationfrom solid to gaseous phase. In this process the volume ofhydrocarbons increases by approximately three orders ofmagnitude. This leads to an increase in the pressure insidethe coal particle, and subsequently changes in the coalstructure such as increase in porosity and appearance ofbubbles in the coal particle.
The structural change can vary from one coal to another(Anthony and Preto, 1995). It has been reported that for largeparticles of a bituminous coal with a free swelling index of5–5.5, the coal particle swelled, softened and developed ashell and eventually a cenospherical structure. However,another bituminous coal with a free swelling index of 1–2,devolatilised with little swelling and the creation of twozones (the char and unreacted coal), but without cenosphereformation.
Figure 10 Coal combustion processes (modifiedfrom Anthony and Preto, 1995)
Devolatilisation degreeThe volatile release may vary with coal rank, devolatilisationtemperature, and gas pressure (Basu, 1999; Wall and others,2002; Bencteux and others, 2002). Figure 11a shows therelationship between the volatile release and coal rank at atemperature of 900°C. It can be seen that the volatile releasedecreases from 37% to 26% due to an increase of the Ccontent in coal from 73% to 93%. Figure 11b illustrates howthe volatile release varies with temperature for an anthracitecoal. It shows that an increase in temperature from 700°C to1000°C leads to a linear increase in the volatile release from13% to 32% (Bencteux and others, 2002).
Okumura and others (2002) calculated the volatile release ofa high-volatile bituminous coal at various gas pressures.They found that, at a gas pressure of 0.1 MPa, the total tar
17Understanding fluidised bed combustion
Combustion
and gas release accounted for about 51% of the originalparticle weight. As the gas pressure increased, the tar vapourrelease decreased abruptly while the release of gases slightlyincreased. As a result, the total volatile matter releasedecreased with increasing gas pressure. It decreased toaround 35% of the original particle weight at a gas pressureof 1.5 MPa. However, the amount of char formed was foundto increase with increasing gas pressure.
Devolatilisation rateCoal devolatilisation is a relatively rapid process. Itcompletes in seconds to minutes, depending on the coalparticle size. It is generally acknowledged that thedevolatilisation rate, measured as devolatilisation time, maybe controlled by three main mechanisms: chemical reactionkinetics; heat transfer to and within the particle; and masstransfer of volatile products within the particle. A number ofmodels have been proposed assuming different ratecontrolling mechanisms (Ross and others, 2000; Agarwal,1986; Agarwal and others, 1984; Stubington and Sumaryono,1984; LaNauze, 1982). Agarwal and others (1984) suggestthe existence of three different regimes:● for very small particles less than 0.1 mm, the
devolatilisation time is independent of the particle size,indicating chemical reaction kinetics is therate-controlling factor. A similar finding was reported byJuntgen and van Heek (1979), where for particle sizesbelow 0.6 mm chemical kinetics is rate controlling;
● as the particle size increases, both chemical reactionkinetics and heat transfer control the rate ofdevolatilisation. The extent of the mixed regime dependson the type of coal and the operating conditions such asbed temperature;
● for large particles greater than about 1 mm, the rate isheat transfer controlled with the devolatilisation timebeing proportional to the square of the particle diameter.
Ross and others (2000) also noted that the rate controllingmechanism may change during the devolatilisation process.At lower particle temperatures, heat transfer is ratecontrolling. When the particle temperature reaches the bedtemperature, however, chemical kinetics plays an increasingrole.
The devolatilisation time has be defined or measured in anumber of different ways as described below:● flame extinction time (FET), which may be taken as a
valid measure of devolatilisation time for particlesreacting in air. It is the time elapsing between theappearance and disappearance of any visible flame (forexample, luminous, plume-like yellow flames) at the bedsurface, measured with a stopwatch (Urkan and Arkol,1994);
● CO2 evolution profile, which measures the CO2
concentration as a function of time after injection of theparticle into the bed. For smaller particles, thedevolatilisation time is given by a distinct end point ofdevolatilisation, after which the CO2 concentrationremains approximately constant for some considerabletime. For larger particles, however, the profile shows aslower rate of decrease before levelling out at a roughlyconstant value. The devolatilisation time is defined as
70
30
25
20
80 90 100
% C in coal
Vola
tile
rele
ase,
%
35
40
a)
25
20
0
1100
Pyrolysis temperature, °C
Vola
tile
rele
ase,
%
30
35
b)
15
10
5
1000900800700600
Figure 11 Effects of coal rank and temperature onvolatile release (Bencteux and others,2002)
Table 1 Reported values of A and n for fluidised bed combustion
Reference A n Coal used Particle size,mm
Bedtemperature, °C
Measurementtechnique
Stubington and others(1992)
1.84 1.5 – – 850 –
Stubington and others(1997)
1.78 1.5 Blair Athol >3 850 FET + CO2 data
1.67 1.5 Lemington >3 850 FET + CO2 data
1.84 1.5 South Bulli >3 850 FET + CO2 data
Urkan and Arikol(1994)
0.66–1.88 1.42–1.91 Various Turkish 1.0–11.2 860 FET
Ross and others(2000)
4.11–6.85 1.03–1.16 Australian Bowmans 6.5–14.5 750–950Temperatureresponse
4.45 1.19 Australian Morwell 6.5–14.5 850Temperatureresponse
the time when the slope first changes (Stubington andothers, 1997);
● the time taken from coal particle immersion into the beduntil the centre temperature of the particle equals thebed temperature. The particle temperature is measuredusing a sheathed type K thermocouple (Ross and others,2000; Heidenreich and Zhang, 1999).
Factors affecting devolatilisation rateThe effect of particle size on devolatilisation time for largecoal particles (generally >1mm) has been extensivelyreported in the literature. Results have typically beencorrelated by the following power-law relationship:
tv = A dpn (13)
where tv is the devolatilisation time, dp is the initial coalparticle diameter, and both A and n are a correlationconstant. Table 1 lists the data for A and n from a number ofinvestigations. It can be seen that there are largediscrepancies in the values of A and n. The variation may beattributed to the differences in measurement technique,definition of devolatilisation time, coal type, and bedtemperature (Ross and others, 2000; Stubington and others,1997). It is also noted that experimental values for ngenerally vary between 1 and 2. This does not agree with themodel prediction which suggests tv a dp
2 for heat transfercontrolled mechanism. However, the measurements of thecoal particle centre temperature for particle sizes between6.4–20 mm have shown a time lag before the particle centrereaches the bed temperature. This supports the theory thatthe devolatilisation rate for large particles is heat transfercontrolled (Ross and others, 2000).
The devolatilisation time is also influenced by several factorsincluding heating rate, oxygen concentration, bedtemperature, coal rank, and moisture content. Figure 12
18 IEA CLEAN COAL CENTRE
Combustion
shows the experimental and calculated time-resolveddevolatilisation history of Beypazari lignite particles at twodifferent heating rates (Selcuk and others, 2001). It can beseen that heating rate has very little effect on thedevolatilisation degree, that is, the final yield of volatiles(both about 35% of the original particle weight). However,the devolatilisation time is reduced by half when the heatingrate is increased from 10°C/min to 20°C/min.
Ross and others (2000) studied the devolatilisation times oftwo coals in different gas environments. They observed thatas the oxygen concentration was increased, thedevolatilisation time decreased. The devolatilisation times forBowman and Morwell coals in air were found to be 38 and24% less than in nitrogen, respectively. This generally agreedwith that reported by Stubington and others (1991), a 19%decrease in the devolatilisation time. Accordingly, thedevolatilisation times for the two coals in air were 10 and
8% less than in air/steam (15 vol% oxygen), respectively.This might be attributed to the formation of a volatilediffusion flame around the coal particle when changing gasenvironment, especially from nitrogen to air. The flameincreased the particle heating rate and therefore decreasedthe devolatilisation time.
Figure 13 compares the devolatilisation times for Bowmanscoal particles fluidised in air at 750°C, 850°C and 950°C(Ross and others, 2000). It is shown that the devolatilisationtime and correlation constant A decreases with increasingbed temperature. This is in general agreement with valuesreported in the literature. For example, Urkan and Arkol(1994) found that for 3.35–4 mm coal particles, thedevolatilisation time decreased by a factor of 1.3–1.6,depending on coal type, as the bed temperature increasedfrom 650°C to 920°C. This is because an increase in bedtemperature enhances external heat transfer and thereforedecreases the devolatilisation time.
It appears that coal moisture content also has a significanteffect. Urkan and Arkol (1994) measured the devolatilisationtimes for dried and moist coals, tvd and tv, respectively. Theratio tvd/tv was found to be less than one for all the moisturecontents, indicating a dried coal had a shorter devolatilisationtime than a moist coal. The effect became more pronouncedas the moisture content increased. This was because moreenergy was required to evaporate water while the coalparticle temperature remained constant until drying wascomplete. As a result, the internal heat transfer was retardedand hence the devolatilisation time increased. A similar trendwas found for the combustion of pine wood in a fluidisedbed (Diego and others, 2003). Figure 14 shows the results atvarious bed temperatures for particles of 10 x 16 x 15 mm3.It can be seen that the devolatilisation time increased almostlinearly with moisture content. In addition, the slope wasslightly higher for lower bed temperatures.
It is not clear whether or how coal type affects thedevolatilisation time. This may be because of the difficultyof identifying representative coal properties describing thedependence of devolatilisation behaviour on coal type.Agarwal (1986) used kinetic parameters and the temperatureat which devolatilisation is complete to modeldevolatilisation, but found negligible effect of coal type forparticles >1 mm. Stubington and others (1991) also did notfind any correlation between the devolatilisation time andcoal type.
There are contraventions to the above findings. Urkan andArkol (1994) used the volatile matter heating value (VMHV)as a parameter to characterise the influence of coal type. Thedevolatilisation time was then found to increase withincreasing VMHV, implying an increase in devolatilisationtime with coal rank. This was thought to be a consequence ofincreased times for the centre of a coal particle to reach thetemperature at which devolatilisation is complete or in thiscase the bed temperature. Eatough and Smoot (1996) alsofound that a bituminous coal has a longer devolatilisationtime than a lignite. This may be due to the greater volatilesyield and the softening nature of the bituminous coal, whichslows transport of volatiles to the particle surface. In
19Understanding fluidised bed combustion
Combustion
addition, the lignite produces more gaseous volatiles (forexample, CO, CO2, HC) prior to ignition.
4.3 Combustion of volatiles
The volatiles released from the coal particles then burn in thepresence of oxygen. Oxygen taking part in the reactionscomes mainly from the fluidising air, but also from the coal.In general, the combustion of volatile species is a rapidprocess. CH4 and H2 oxidations are particularly fast, whileCO oxidation is comparatively slow (Desroches and others,2002; Lawrence, 2001). The combustion contributes to bothheat production (as much as 50% of the total heat released )and emissions.
bedtemperature
750°C850°C950°C80
60
40
20
0
0 5 10 15
Particle diameter, mm
Dev
olat
ilisa
tion
time,
s 100
120A
6.854.424.11
n
1.031.161.14
Figure 13 Devolatilisation times for Bowmans coalparticles fluidised in air at different bedtemperatures (Ross and others, 2000)
bed temperature650°C750°C850°C950°C
100
80
60
40
0 20 30 40 50
Moisture, %
Dev
oliti
lisat
ion
time,
s
120
140
10
Figure 14 Effect of moisture content ondevolatilisation time for pine woodparticles (Diego and others, 2003)
Chemical reactionsIn fluidised beds, combustion of volatiles can be ahomogeneous or a heterogeneous reaction over char and bedsolids surfaces. Oxidation of CO, H2, CH4, C2H6 and C3H8 isa homogeneous process (Liu and Gibbs, 2002; Kulasekaranand others, 1999). However, HCN and NH3 oxidation can beeither homogeneous or heterogeneous (Desroches and others,2002). Chemical reactions for some volatile species are listedbelow.
Combustion of carbon monoxide:
CO + 1/2O2 ➝ CO2 (A)
Combustion of hydrogen:
2H2 + O2 ➝ 2H2O (B)
Combustion of methane:
CH4 + 3/2O2 ➝ CO + 2H2O (C)
Combustion of C2H6:
C2H6 + 5/2O2 ➝ 2CO + 3H2O (D)
Combustion of propane (stepwise oxidations):
C3H8 ➝ 3/2C2H4 + H2 (E)
C2H4 + O2 ➝ 2CO + 2H2 (F)
CO + 1/2O2 ➝ CO2 (G)
H2 + 1/2O2 ➝ H2O (H)
Combustion of HCN:
HCN + 5/4O2 ➝ NO + CO + 1/2H2O (I)
HCN + 3/2O2 + NO ➝ N2O + CO + 1/2H2O (J)
HCN + 5/4O2 ➝ 1/2N2 + CO2 + 1/2H2O (K)
Combustion of NH3:
NH3 + 5/4O2 ➝ NO + 3/2H2O (L)
NH3 + 3/4O2 ➝ 1/2N2 + 3/2H2O (M)
The formation/emissions of NO, N2O and CO are furtherdiscussed in details in Chapter 5.
Combustion modesIn bubbling fluidised beds, volatiles may burn either in asolid-free bubble phase or, in the spaces between the solidparticles, that is, in an emulsion phase (Bulewicz, 2002;Srinivasan and others, 1998). Gas can be exchanged betweenthese two phases by diffusion and convection (Lawrence,2001). In circulating beds, the distinction between thebubbles and the particulate phase is lost. Srinivasan andothers (1998) reviewed experimental observations in the
20 IEA CLEAN COAL CENTRE
Combustion
literatures. They found that the combustion mode depends onthe bed temperature as follows:● at relatively low bed temperatures (less than about
700–750°C), the volatiles may not burn within the bed;● at moderate bed temperatures (from about 830–900°C),
the volatiles are likely to burn only in the bubbles; and● at high bed temperatures (greater than about 900°C),
combustion may occur within the bubbles and in theparticulate phase of the bed.
It was noted that the temperatures defining the abovecombustion regimes vary with the bed particle size andfluidised velocity. However, understanding of therelationships remains incomplete.
Stubington and Chan (1990) have shown that volatilecombustion occurs in the emulsion phase for bed particlesizes of larger than 600 µm and temperatures of higher than800°C. Experimental investigations by Hesketh andDavidson (1991) showed that propane combustion isinhibited in the presence of particles for bed temperaturesbelow about 837°C. However, Hayhurst and Tucker (1990)have shown that CO does not oxidise in the emulsion phaseup to about 1000°C. Instead, its oxidation occurs only inbubbles and in the freeboard zone of a bubbling bed.Lawrence (2001) simulated the combustion of CH4 and COin a 20 cm deep bubbling bed (consisting of 300 µm sandparticles) at temperatures of 700°C and 950°C. Similarly, itwas found that the combustion did not take place in theparticulate phase even at 950°C.
Most of the mathematical models are based on the two-phasetheory. However, Srinivasan and others (1998) proposed athree-phase model which accounts for the presence of acloud/wake phase. This phase assumes the gas is well-mixed,in which the phase itself moves upward with the bubblephase, and the solids in the cloud phase are well-mixed withthe solids in the emulsion phase. Combustion of CO, CH4,and propane in a fluidised bed was then simulated accordingto the model and the results were compared withexperimental data. It was concluded that:● combustion of CO at 1000°C is inhibited by the presence
of bed particles, that is, in the emulsion and cloud/wakephases and thus can occur only in the bubble phase;
● oxidation of CH4 to CO at three different temperatures(850°C, 950°C, 975°C) might occur in all the threephases, that is, bubble, emulsion and cloud/wake phases;
● in the stepwise oxidations of propane, reaction (E)represents an endothermic pyrolysis which occurs in allphases; CO oxidation occurs only in the bubble phase;oxidation of C2H4 at 750°C is inhibited by bed particlesand hence occurs only in the bubble phase; however, themodel could not predict whether reaction (H) occurs inall phases or only in the bubble phase.
Bed/freeboard combustion splitUnder normal operating conditions, the volatile matterreleased will burn completely. In a bubbling bed, however,some of the combustion may occur above the bed becausepart of the volatile matter is inevitably released in a zonewhere oxygen is not available (near the coal feed point) andat the top of the bed.
Selcuk and others (2001) estimated the fractions of volatilesreleased in the bed and in the freeboard for burning a lignitewith a VM/FC ratio of 2.16. The fuel was fed at 0.22 mabove the distributor plate and the expanded bed height was1 m. The times spent by a devolatilising coal particle in thebed and over the bed surface were calculated. A comparisonof the time spent in the bed with the devolatilisation history(the cumulative yield of volatiles against time) gave thefraction of volatiles released in the bed. The remainingvolatiles were assumed to be released to freeboard while theparticle was at the bed surface. It was shown that 75% of thevolatile matter was released in bed and 9% to freeboard. Thelatter figure was expected to increase as the feeding pointapproaches the bed surface. The split of volatiles combustionbetween bed and freeboard was calculated from the oxygenconsumption in the respective regions. It was found to be80/20, which was comparable to the value 84/16 determinedfrom measured gas concentration profiles.
Chirone and others (1999) compared the volatiles combustionsplits for a low-volatile (23%) South African bituminous coaland a high-volatile (75.1%) Robinia biomass. Differentlyfrom the above, the combustor is divided into the bed, thesplashing zone and the freeboard. The splashing zone extendsbetween the average bed surface level and the maximum levelreached by bed solids due to bubbles bursting. It was shownthat for the bituminous coal, 35–70% of the volatiles isburned in the bed and the remaining in the splashing zone.For Robinia only 10% is burned in the bed and the remaining90% in the splashing zone.
In another study for two high-volatile fuels, Chirone andothers (2001) showed that the combustion of volatiles cancontinue into the freeboard. In the case of Robinia, 7% of thevolatiles are burned in the bed, 62% in the splashing zoneand 31% in the freeboard. For a tyre derived fuel with avolatile content of 63.4%, the corresponding fractions are18%, 61% and 21% respectively. However, it was noted thatthe estimates are based on various assumptions withuncertainties.
4.4 Combustion of char particles
Char burning is the final stage of the coal combustionprocess. It is much slower than the combustion of volatilematters. Char is normally present in the bed as less than 1%of the bed mass if crushed coal is used and less than 5% forlarger size grading (Merrick, 1984). The concentration mayvary according to the fuel activity (Anthony and Preto,1995). A burning char particle is normally hotter than itssurroundings. Its surface temperature is up to about 100°Chigher than the bed temperature (Merrick, 1984). Thisprovides the necessary rate of heat transfer from the charparticle to the bed.
Char comprises mainly carbon (fixed carbon), ash, nitrogenand sulphur. At elevated temperatures (above 750°C), itoxidises to gaseous products (CO, CO2, NO, NO2, N2O, andSO2) and leaves ash as a residual product. Here, discussionsare focussed on the carbon combustion while the pollutantformations are addressed in Chapter 5.
21Understanding fluidised bed combustion
Combustion
Combustion mechanismsIn the combustion of carbon particles, following three mainchemical reactions may occur:
C + 1/2O2 ➝ CO
CO + 1/2O2 ➝ CO2
C + CO2 ➝ 2CO
Figure 15 illustrates three possible combustion mechanismsas described below (Basu, 1999). Mechanism 1 assumes thatoxygen diffuses to the char particle surface and oxidises it toCO. The CO oxidises further to form CO2 in a gas-phasereaction. As the CO oxidisation occurs very close to thecarbon surface, CO2 may be considered as the primarycombustion product (single-film theory). This mechanism ispredominant for a low Reynolds number flow, for large charparticles (greater than 1 mm), or at high bed temperaturesaround 900–1300°C.
In Mechanism 2, oxygen diffuses to the char surface andproduces both CO and CO2. CO and CO2 then diffuse awayfrom the carbon surface. CO further reacts in a gas-phasewith oxygen arriving from the bulk gas, and forms CO2.Both CO and CO2 are the primary carbon combustionproducts. If the temperature is low or the particle size issmall, however, CO may escape into the free stream andleave unburned. The primary CO/CO2 product ratio dependson the type of char and the particle temperature (Kulasekaranand others, 1999). It is about 10 for typical CFBtemperatures (Desroches and others, 2002).
Mechanism 3 assumes that oxygen can not reach the carbonsurface. It reacts with the CO in a gas-phase reaction awayfrom the carbon surface. Part of the CO2 formed diffusesback to the carbon surface where it is reduced to CO
O2
CO2
CO2
CO
CO2
CO
CO2
O2
O2
CO2
reaction zone
carbon
model Idirect oxidation
model IIItwo film theory
model II
Figure 15 Combustion mechanisms for non-porouscarbon particles (Basu, 1999)
(two-film theory). However, it has been suggested that thereduction of CO2 on the carbon surface is too slow to besignificant.
The oxidation of char for all three mechanisms can berepresented as:
C + (1/f)O2 = (2-2/f)CO + (2/f-1)CO2
where f is the mechanism factor and equal to 1 for CO2
transport from the surface (Mechanism 1) and equal to 2 forCO transport (Mechanisms 2 and 3).
The above mechanisms all assume char particles arenon-porous. However, char particles are generally highlyporous. Hence, the combustion may not necessarily takeplace on the external surface of the particle. Three differentcases are considered. When the oxygen diffusion rate ismuch greater than the chemical reaction rate on the carbonsurface, most of the oxygen may penetrate into the pores ofthe char. As a result, the combustion occurs mainly on theinternal surface. This type of combustion is prevalent forfine char particles. When the diffusion and chemicalreaction rates are comparable, part of the oxygen penetratesinto the pores of the char, thus the combustion occurs onboth the external and internal surfaces. This type ofcombustion occurs with medium size char particles. Whenthe chemical reaction is much faster than the oxygendiffusion, the oxygen is entirely consumed before it canenter the char pores. This case generally occurs with largechar particles (Basu, 1999).
Combustion modelsIn principle, char combustion in fluidised beds may occur infollowing three different ways (Anthony and Preto, 1995):● char particle size decreases continuously while its
density remains constant (that is, a shrinking particlemodel);
● char particle density decreases continuously while itssize remains constant (a progressive conversionmodel);
● char particle density within a shrinking char coreremains constant, with the appearance of an ash layer (ashrinking core model).
Most coals chars seem to react in an intermediate modebetween models 1 and 2, but roughly approaching model 1.In contrast, a high ash Portuguese anthracite was found toburn according to model 3 (Anthony and Preto, 1995).
Burning rateWhen the oxygen diffuses from the free stream to thereaction site on the unburned carbon core, it undergoesresistance: through the inert particles around the charparticle; through the gas film layer around the char particle;and through the ash layer. The burning rate of individual charparticles therefore depends on both the mass transfer ofoxygen and the chemical reaction rate. It can be predictedusing the shrinking particle model (Basu, 1999; Selcuk andothers, 2001). Basu (1999) has suggested that:
q = Co/(1/km +1/kc) (14)
22 IEA CLEAN COAL CENTRE
Combustion
where q is the burning rate on the external surface expressedas kg of carbon/m2 s, Co is the oxygen concentration in thefree stream, km is the mass transfer coefficient of oxygen,and kc is the chemical reaction rate on the external surface ofchar.
The chemical reaction rate depends on the type of coalburned whilst the mass transfer rate is given by (Basu, 1999):
km = 12fShDo/dcR.Tm (15)
where Sh is the Sherwood number, Do is the moleculardiffusivity of oxygen, dc is the initial diameter of the charparticle, R is the universal gas constant, and Tm is the meantemperature of the diffusion layer around the carbon particle.The value of km varies with the fluidisation regime and ishigher in circulating beds than in bubbling beds. As a result,the burning rate in circulating beds is higher than that inbubbling beds under similar thermal and chemicalconditions.
The burning rate determines the burnout time of a charparticle. The burnout time, tb, can be correlated to processparameters by the following relationship (Anthony and Preto,1995):
tb = a(0)Tba(1) ua(2) dc
a(3) (16)
where a(i) values are constants, Tb is the bed temperature,and u is the fluidising velocity. In extreme cases, a(3) equalsto 1(chemical reaction controlled) and 2 (mass transfercontrolled). Typically, a(3) adopts a value between 1 and 2for a combined regime.
Combustion/elutriation of fine charIn fluidised bed combustion char particles continuouslyreduce in size. This is attributed to three different processes:char particles are burned out, which has been discussedearlier; char particles are fragmented; and char particles areattrited. As a result, a significant amount of fine chars can beproduced within the bed.
Fragmentation is normally attributed to thermal shock andthe over-pressures produced (Anthony and Preto, 1995). Itoccurs above some critical coal particle size, reportedly,between 13–19 mm. As well as on the particle size,fragmentation behaviour also depends strongly on themechanical properties of the coal. In a study of brown,bituminous and anthracitic coals with particle sizes of4–9 mm, fragmentation only occurred with the anthraciticcoal. However, another study showed that bituminous coalsfragmented on completion of devolatilisation. Urkan andArkol (1994) studied various Turkish coals ranging frombituminous to lignite. They observed a few coals fragmentedafter devolatilisation was virtually complete. However, themajority preserved their original particle shape and showedvery little change in particle size. Wang and Stubington(1998) tested an Australian black coal and observed charfragmentation at about 2 minutes after feeding coal particlesinto the bed. The ratio of the number of fragmented charparticles (with sieve sizes >1 mm) to the number of initialcoal particles (with sieve sizes of 4–4.75 mm) was 33/20.
Attrition is the process whereby asperities on the charsurface are broken off during interparticle collisions in thebed. It is generally acknowledged that combustionconsiderably enhances the attrition rate (Anthony and Preto,1995; Basu, 1999). This is because combustion createsragged surfaces, which are easily attrited. However, Walshand Li (1994) argued that attrition is mainly a mechanicalprocess with a rate proportional to collision frequency butunaffected by combustion.
Basu (1999) notes that the overall weight loss of a burningchar particle is given by:
dWc/dt = W + ka ud Wc/dav (17)
where Wc is the overall weight loss, W is the weight lossonly due to combustion, ka is the dimensionless attrition rateconstant, ud is the driving velocity which depends on thefluidisation regime, and dav is the average carbon particlediameter. In bubbling beds, ud = u-umf, indicating theattrition rate is proportional to the excess velocity overminimum fluidisation. In circulating beds, however, thedriving velocity is the slip velocity between larger charparticles and average bed solids. The rate constant formechanical attrition is independent of the fluidisation regimeand found in the range of (0.05–0.31) x 10-7. In the presenceof combustion, the rate constant increases to (0.5–5.0) x 10-7.
Chirone and others (1999) studied the combustion of abituminous coal and a biomass using a mathematical model.They found that for both fuels char combustion takes placeessentially in the dense bed of the combustor. Charconcentrations in the splashing and freeboard zones, andassociated burning rates, are negligible. Coarse charcombustion prevails in the case of the low-volatilebituminous coal. Fine char contributes to about 10% of thefixed carbon conversion. This is proportionally larger whencompared with its concentration, because of its greaterapparent reactivity. In the case of the high-volatile biomassfuel, fixed carbon conversion in the bed is equally sharedbetween the coarse and fine chars in spite of the muchsmaller concentration of fine char. Again, this is due to themuch greater apparent reactivity of the biomass char fines.
When the char particles reduce to certain sizes, mainlybetween 10 and 100 microns, their free-fall velocity may beless than the fluidising velocity. Some of the particlesbecome entrained in the gas stream and are removed fromthe combustor (that is, elutriation). This contributes toinefficient combustion (Stubington and Cui, 2002; Anthonyand Preto, 1995). Some loss of char can also occur with theremoval of excess bed solids to control the bed depth.However, because of the low char concentration in the bedthis loss is small compared with that by elutriation. The totalchar loss typically represents 2–15% of the thermal input ofcoal to the combustor. However, values above 4% are usuallyunacceptable commercially (Merrick, 1984).
4.5 Combustion efficiency
Combustion efficiency is defined as the ratio of heat released
23Understanding fluidised bed combustion
Combustion
to the energy supplied by the fuel. It is generally higher inCFBC than in BFBC due to the use of finer particles,turbulent gas-particle mixing, and a high solid recycle rate.PBFBC achieves higher efficiencies than BFBC due tosmaller and more frequent bubbles and hence better gas andsolid contact. Combustion efficiency is affected by severalfactors as detailed below.
Fuel type It appears that fuel type has a significant effect on thecombustion efficiency. This was observed for the combustionof two different mixtures in a bubbling fluidised bed(Armesto and others, 2001). For the lignite/alpeorujo (solidresidues from olive milling) mixture, the combustionefficiency was in the range 94–98.6%. However, lowercombustion efficiencies (88.1–88.5%) were found for theanthracite/alpeorujo mixture. This might be attributed to thedifference in the fuel volatile content. The lignite had ahigher volatile content (25.8% compared to 7.1% with theanthracite), hence a greater reactivity and a greater charcombustion rate.
Bed temperatureHigher bed temperatures usually lead to higher reactionrates, and hence higher combustion efficiencies. Adanez andothers (2001) carried out a study on the combustion of ananthracite coal in a CFB. A gas velocity of 4 m/s and anexcess air of 20% were used. They found that an increase intemperature from 800°C to 900°C increased the carboncombustion efficiency from 94% to 96%. A test by Armestoand others (2002a) has also shown, for rice husk combustionin a BFB, the combustion efficiency increased from 97% to98% due to an increase in temperature from 844°C to 877°C.In a study for co-combustion of lignite/25% foot cake (solidresidues from olives, 2–4% oil, and 60–70% water) in aBFB, the combustion efficiency was found to increase withbed temperature. However, for an anthracite/15% foot cakemixture the combustion efficiency changed little withtemperature (Armesto and others, 2003).
Gas velocityAn increase in gas velocity generally decreases thecombustion efficiency. This was observed for the combustionof rice husk in a BFB, where an increase in gas velocityfrom 1 m/s to 1.2 m/s led to a decrease in combustionefficiency from 98.4% to 97% (Armesto and others, 2002a).A similar trend was reported for the combustion of a ligniteand an anthracite coal in a CFB at a temperature of 850°Cand an excess air of 20% (Adanez and others, 2001). In bothcases, the combustion efficiency decreased slightly as the gasvelocity was increased from 3.2 m/s to 5 m/s. This might beattributed to an increase in coal feed, which reduced thecombustion efficiency by increasing the throughput per unitbed area. In addition, the solid circulation flow rate wasincreased, hence increasing the flow rate of solid losses bythe cyclone.
Excess airAn increase in excess air increases the mean oxygenconcentration in the bed, and therefore the combustionefficiency. Adanez and others (2001) studied the combustionof an anthracite coal in a CFB. A bed temperature of 850°C
and a gas velocity of 4 m/s were used. They observed that anincrease in excess air from 10% to 30% led to a constant butsmall increase of carbon combustion efficiency from 95% to96.5%.
Figure 16 shows how combustion efficiency changed withexcess air for the combustion of olive cake and lignite in aCFB. A wider range of excess air ratio (from 1.1 to 2.2) wasstudied. In both cases, the combustion efficiency increasedwith increasing excess air initially. This was believed toresult from a decrease in the combustion losses due tohydrocarbons and CO. The maximum efficiencies for theolive cake and lignite (98.7%, 98.3%) were reached at excessair ratios of 1.35 and 1.39 respectively. A further increase inexcess air, however, was found to decrease the combustionefficiency (Topal and others, 2003).
4.6 Summary
When coal particles enter a fluidised bed, they undergo asequence of events including drying, devolatilisation,combustion of volatiles, and combustion of the residual char.
The drying process removes both the surface and inherentmoisture of a fuel. The evaporation of surface moisture doesnot affect the combustion process directly. Low rank coalscontain high inherent moisture, and its evaporation mayoccur in conjunction with and slow down the devolatilisationprocess. Furthermore, the loss of water can result insignificant morphological changes.
Further heating releases the volatile matter of the coalparticle. It appears that the volatile release increases with thedevolatilisation temperature but decreases with coal rank andgas pressure. The devolatilisation rate is controlled bychemical kinetics or heat transfer, depending on particle size.It is argued that the controlling mechanism changes duringthe heating-up of a particle. At lower temperatures heattransfer is rate controlling. When approaching the bedtemperature, chemical kinetics become important. Thedevolatilisation time of a particle increases with increasingparticle size and moisture content. It decreases withincreasing heating rate, oxygen concentration and bedtemperature.
The volatiles released then burn rapidly in the air. It can be
24 IEA CLEAN COAL CENTRE
Combustion
homogenous or a heterogenous reaction over char and bedsolids surfaces. In bubbling beds, homogenous reactions canoccur either in bubbles or in a particulate phase. At low bedtemperatures volatiles may not burn within the bed. Atmoderate temperatures volatiles are likely to burn only inbubbles. At high temperatures, the combustion may occur inboth phases. In circulating beds, the distinction between thebubbles and the particulate phase is lost. Under normalconditions, the volatiles will burn completely. However, inbubbling beds some of the combustion may occur above thebed, particularly for high-volatile fuels.
Char burning is the final stage of coal combustion. Itproduces CO and/or CO2 as the primary product. Reactionscan occur both at the particle surface and internally. Mostchars appear to react in a shrinking particle model, that is,particle size decreases while its density remains constant.The burning rate depends on the coal type and other factorssuch as the fluidisation regime and char diameter. It isgenerally higher in circulating beds than in bubbling beds.Fines are produced in the bed due to char burning out,fragmentation and attrition. They account for small amountsbut contribute significantly to fixed carbon conversion due totheir greater apparent reactivity, particularly for high-volatilefuels. However, some of the fines may elutriate andcontribute to inefficient combustion.
In general, combustion efficiency is higher in CFBC than inBFBC due to the use of finer particles, turbulent mixing, anda high solid recycle rate. It is also higher in PBFBC than inBFBC due to smaller and more frequent bubbles and hencebetter gas and solid contact. It increases with fuel volatilecontent or bed temperature whereas gas velocity has theopposite effect. An increase in excess air increasescombustion efficiency initially. However, a further increasemay decrease combustion efficiency.
9694929088
80
2.4
Excess air
Com
bus
tion
effic
ienc
y, %
98100
2.32.22.121.91.81.71.61.51.41.31.21.11
868482
coalolive cake
Figure 16 Effect of excess air on combustionefficiency for olive cake and lignite(Topal and others, 2003)
5 Pollutant formation
In fluidised bed combustion, various pollutants aregenerated. These generally include nitrogen oxides, nitrousoxide, sulphur oxides, particulates, carbon monoxide, andtrace elements. This chapter addresses how these pollutantsare formed and what factors may affect their formation.Where appropriate, the controls of pollutants are alsodiscussed.
5.1 Nitrogen oxides and nitrousoxide
During coal combustion, nitrogen in both the combustion airand fuel may be converted into nitrogen oxides (NO andNO2 , together referred as NOx). At the relatively lowtemperatures of FBC, nitrogen in the combustion air does notoxidise to a significant extent. Thermal-NOx is thereforenegligible, thus virtually all NOx are generated fromfuel-nitrogen (Armesto and others, 2002b; Minchener andothers, 2000). As a result, NOx emissions are reducedcompared to pulverised coal combustion that takes place athigher temperatures. Under FBC conditions, the oxidation ofNO to NO2 is limited by a low equilibrium concentration(Johnsson, 1994). In practice, more than 90% of the NOxemissions are actually in the form of NO. Hence, NOxemissions from FBC can be taken as NO (Anthony andPreto, 1995).
It has become apparent that FBC produces significantquantities of nitrous oxide (N2O), again due to the lowcombustion temperatures. The N2O emissions range from20 ppm to 300 ppm compared to the 10 ppm typically foundin conventional coal combustion (Armesto and others,2002b). As the formation of NO and N2O are closely related,they are considered together as below.
Formation and reduction mechanismsThe formation of NO and N2O from coal-nitrogen duringcombustion is a complex process. It involves coaldevolatilisation, volatile species oxidation, char oxidation,and finally, partial reduction to N2. Figure 17 shows the NOand N2O formation and reduction pathways. Duringdevolatilisation, the organically bound nitrogen in fuel ispartitioned into volatile-nitrogen and char-nitrogen. Thevolatile-N compounds are mainly released as HCN and NH3.In the char, the nitrogen is bound in aromatic structures.
The volatile-N compounds are subsequently oxidisedthrough homogenous gas-phase reactions or heterogenouscatalytic reactions over char and other bed solids such aslimestone, ash and sand. HCN gives NO, N2O or N2, whileNH3 is essentially oxidised to NO or N2 (chemical reactionsare shown in Section 4.3). The reaction product depends onthe bed temperature and the catalyst activity of the bedmaterial (Desroches and others, 2002).
During char oxidation, nitrogen is oxidised mainly to NO,N2O and N2. However, small amounts of HCN and NH3 may
25Understanding fluidised bed combustion
also be formed. The main reactions are as follows (Liu andGibbs, 2002):
Char-N + 1/2O2 ➝ -CNO
-CNO + 1/2O2 ➝ NO + CO2
-CNO + NO ➝ N2O + CO2
The NO and N2O formed are then partially reduced to N2.
NO can be reduced by NH3 in the homogenous phase. N2Ocan be decomposed due to temperature effect. Both NO andN2O can be reduced either directly by char or throughreactions with CO catalysed by char and other bed solids.The reduction reactions are shown below (Desroches andothers, 2002; Lan and others, 2001):
NO reduction:
NO + NH3 + 1/4O2 ➝ N2 + 3/2H2O homogenous
NO + 2/3NH3 ➝ 5/6N2 + H2O char, bed solids
NO + Char ➝ 1/2N2 + CO char
NO + CO ➝ 1/2N2 + CO2 char, bed solids
N2O reduction:
N2O ➝ N2 + 1/2O2 homogenous
N2O + Char ➝ N2 + CO char
N2O + CO ➝ N2 + CO2 char, bed solids
The rates of the above heterogenous catalytic reactions canvary widely, depending highly on the activity of a catalyst
HCNFuel-N N2ON2O
N2
N2
Cha
r-N
NO NO
N2
Char
NH3O2
H, CO
Char,Limestone
O, OH
Volatile-N
Char-N
NO
O2
O2
Figure 17 NO and N2O formation and reductionpathways (Anthony and Preto, 1995)
(Johnsson, 1994). In general, char and calcined limestonehave a high catalytic activity. Quartz sand has a low catalyticactivity. However, the activity of coal ash very much dependson the type of coal. For example, ash from low-rank coalsshows a higher activity. The activity of limestone decreasesin the course of sulphation, and fully sulphated limestone hasa low activity in most reactions. The rate of a given reactiondepends very much on the type of catalyst. It should benoted, however, that different reactions have very differentrates over the same catalyst. For example, char can be10 times more active for the reduction of N2O than for thereduction of NO (Armesto and others, 2002b). This results ina much higher rate for the reduction of N2O.
Conversion efficienciesConversion of fuel-N to NOx can range from less than 5% toabout 40%, with typical conversions being normally muchless than 20% (Anthony and Preto, 1995). It appears thatfuel-N is converted to N2O to a lesser extent. Boavida andothers (1997) studied the combustion of South African andCimpor coals in a FBC at various temperatures. They foundthat the fuel-N to NO conversions varied within 7.5–15%and 11–20% respectively. Correspondingly, the fuel-N toN2O conversions for the two coals were found to be 3–6%and 5–9% respectively. A study by Armesto and others(2002b) has also shown that, for both a lignite and abituminous coal, the conversion of fuel-N to NO is higher incomparison with the conversion to N2O.
As discussed earlier, fuel-N conversion processes comprisecoal devolatilisation, volatile-N species oxidation, andchar-N oxidation. The above overall conversions thereforedepend on the fractional conversions at each step ofcombustion and from each of the nitrogen compounds, whichare detailed below.
Devolatilisation. The partition of fuel-nitrogen betweenvolatiles and char and the amounts of HCN and NH3 formedduring devolatilisation, are important to the formation of NOand N2O. It appears that the partition of fuel-nitrogendepends on the coal type and devolatilisation temperature.Bencteux and others (2002) reported that the release ofvolatile-N decreased with increasing coal rank. Boavida andothers (1997) studied the devolatilisation of two differentcoals. Cimpor coal was found to give higher fractions ofvolatile-N than South African coal. It was also shown thatfor both coals volatile-N increased with increasingdevolatilisation temperature.
HCN and NH3 releases appear to vary with the coal type,devolatilisation temperature as well as gas pressure (Desrochesand others, 2002; Löffler and others, 2001; Okumura andothers, 2002). A number of studies have been published in theliterature but do not provide conclusive evidence. However,most experimental results indicate the following general trends(Johnsson, 1994; Okumura and others, 2002):● NH3 release from lower-rank coals is greater than or the
same as from higher-rank coals;● NH3 to HCN ratio for lower-rank coals is greater than or
the same as for higher-rank coals;● HCN release from higher-rank coals is greater than or
the same as from lower-rank coals;
26 IEA CLEAN COAL CENTRE
Pollutant formation
● HCN release increases with increasing devolatilisationtemperature;
● HCN release decreases with increasing gas pressure.
A recent study was conducted by Bencteux and others(2002) on the devolatilisation of bituminous and anthracitecoals at temperatures of 700–1000°C. It was shown that NH3
yield was quite stable with both temperature and coal rank,around 10%. In contrast, HCN yield increased exponentiallywith temperature and was much higher for bituminous coal(see Figure 18). In a study of a Spanish lignite, Armesto andothers (2002b) found that NH3 yield fluctuated between5.8% and 17.9% but did not have a clear relation withtemperature. However, HCN yield increased from 0.98% to1.72% with increasing temperature from 609°C to 853°C.
Oxidisation of volatile-N. It is generally accepted that HCNis a very important precursor for N2O formation in fluidisedbeds. NH3 is mainly oxidised to NO and N2, and onlyproduces minor amounts of N2O (Johnsson, 1994; Löfflerand others, 2001). Oxidisation of volatile-N compounds canproduce higher levels of N2O in comparison with NO. In astudy by Boavida and others (1997), volatiles from South
20
15
10
5
0
1100
Pyrolysis temperature, °C
N c
onve
rsio
n, %
25
30
a) Bituminous coal
900700500
NH3
HCN
20
15
10
5
0
1100
Pyrolysis temperature, °C
N c
onve
rsio
n, %
25
30
b) Anthracite coal
900700500
NH3
HCN
Figure 18 NH3 and HCN yields as a function oftemperature for bituminous andanthracite coals (Bencteux and others,2002)
African and Cimpor coals were burned separately attemperatures of 700–1000°C. The results from the two coalswere similar. For South African coal, the conversion ofvolatile-N (related with % of coal-N) to N2O was 2–9%depending on the combustion temperature. In comparison,the conversion to NO was low, between 0.3 and 1.2%.
Johnsson (1994) reviewed a number of studies on thecontribution of volatiles to the total N2O formation influidised beds. Most studies showed that the volatilescontributed approximately 40–70% of the N2O formation fora bituminous coal and 70–85% for a lignite. However, onestudy found that the volatiles contributed less to N2Oformation than char for a bituminous coal.
Oxidisation of char-N. In contrast to volatile-N oxidation,char-N oxidation generally gives much higher levels of NOin comparison with N2O. There are many reported data onthe conversion efficiencies. It is suggested that about20–80% of char-N is converted to NO and 1.5–16% to N2O(Johnsson, 1994). However, these values do not representintrinsic conversions because the NO and N2O formed arepartly reduced to N2 by char. Char is usually found to bemore active for the reduction of N2O than for the reductionof NO, thus giving a higher degree of N2O reduction. Thismay partly contribute to the observed lower N2Oconversions.
The wide variation in the reported char-N conversions is dueto the use of different experimental techniques as well as thefollowing factors (Johnsson, 1994):● amount of char in the bed. A decrease in char content by
factors of 10 and 20 may increase the conversion to NOby 40–80% and 10–50%, respectively. However, afurther decrease in the amount of char does not changethe NO conversion;
● char particle size. A decrease in char particle size byfactors of 10 and 12 may increase the conversion to NOfrom 28 to 56% and from 33 to 51%, respectively;
● devolatilisation temperature. An increase in thedevolatilisation temperature of the char increasesconversion of char-N to NO and decreases conversion toN2O;
● combustion temperature. An increase in the combustiontemperature generally increases the conversion of char-Nto NO and decreases the conversion to N2O, with adegree depending on the coal type (Figure 19);
● coal rank. Conversion to NO is almost independent ofcoal rank or decreases with increasing coal rank.However, conversion to N2O is found to increase withrank (Figure 19);
● nitrogen content of the char. With increasing nitrogencontent of the char, conversion to NO decreases butconversion to N2O increases;
● oxygen concentration. A minor influence of oxygenconcentration has been found but with inconsistentresults.
Effects of process parameters on NO andN2O formations Table 2 summarises how fuel and operating parameters mayaffect NO and N2O formations from fluidised bed
27Understanding fluidised bed combustion
Pollutant formation
combustion. The effects of each parameter are detailedbelow.
Fuel-N content. Typically, coal has a nitrogen content
0.40
0.20
0.00
600 1100
Temperature, °C
Frac
tiona
l con
vers
ion
of c
har-
N to
NO
0.60
1000900800
a)
0.10
0.05
0.0
600 1100
Temperature, °C
Frac
tiona
l con
vers
ion
of c
har-
N to
N2O
0.15
1000900800
b)
700
lignite
bituminous
anthracite
semi-anthracite
subbituminous
lignite
bituminous
bituminous
bituminous
subbituminous
anthracite
semi-anthracite
700
0.20
Figure 19 Conversions of char-N to NO and N2O asa function of combustion temperaturefor different coals (Johnsson, 1994)
Table 2 Effects of fuel and operating parameterson NO and N2O formations
Increasing parameters NO N2O
Fuel-N content Increasing Increasing
Fuel-volatile content Increasing Decreasing
Temperature Increasing Decreasing
Excess air Increasing Increasing
Air staging Decreasing Decreasing
Gas velocity – Increasing
Limestone addition Increasing Decreasing
NH3 injection Decreasing Increasing
between 1and 2% (dry, ash free basis). In general, the higherthe fuel-N, the more NOx and N2O produced, although thedegree of conversion depends on fuel reactivity andcharacteristics. A very reactive fuel like wood, for example,gives a higher degree of fuel-N conversion. However, sincewood has a low fuel-N content it produces less NOx andN2O than would a typical coal (Anthony and Preto, 1995).
Figure 20 shows the effect of fuel-N content on NOemissions. It is suggested by both model prediction andexperimental study that NO emissions increase linearly withfuel-N content. For a fuel-N content of 0.1% (dry, ash-freebasis), the NO emission is about 50 ppm. However, thisincreases to approximately 150 ppm when fuel-N content isincreased to 0.3%. In a study by Leckner and Lyngfelt(2002), high-nitrogen (7.1%) sewage sludge was found togive NO emissions of about 1000 ppm. This wassignificantly higher than the 100 ppm produced fromlow-nitrogen (0.14%) wood. At the same time, aconsiderable reduction of NO was seen in the co-combustionof sludge with other fuels.
Fuel-volatile content. With increasing fuel-volatile content,NO emissions usually increase but N2O emissions decrease(Basu, 1999). This may be attributed to enhanced NH3
production as compared to HCN, which tends to bepreferentially oxidised to NO (Anthony and Preto, 1995).Fuels such as petcoke may result in N2O emissions as highas 150–200 ppm. However, high-volatile coals producerelatively low levels of N2O, often about 50 ppm. In the caseof biomass fuels levels of N2O are even lower. For instance,in BFBC de-inking sludge tends to produce levels of lessthan 10 ppm (Minchener and others, 2000).
In a study by Armesto and others (2002b), the SpanishPuertollano lignite (a volatile content of 26.3%) and theColombian Carbocol bituminous coal (a volatile content of38.6%) were tested. Figure 21 shows that the Carbocolyielded a higher fuel-N to NO conversion than thePuertollano (48–35% compared to 17–21%). However, theCarbocol was found to give a lower conversion to N2O thanthe Puertollano (5–9% compared to 9–18%).
Combustion temperature. Combustion temperature hassignificant effects on both NOx and N2O emissions. An
28 IEA CLEAN COAL CENTRE
Pollutant formation
increase in combustion temperature increases combustionrate and concentrations of radicals (Hamalainen and Aho,1994). At the same time, higher temperatures promote theoxidation of nitrogen radicals (such as NCO and NH) to NOaccording to the reactions (Armesto and others, 2003):
NCO + O ➝ NO + CO
NH + O ➝ NO + H
In addition, higher temperatures reduce the char and COconcentrations in the combustor due to increased combustionrate. This in turn decreases the heterogenous reduction ofNO to N2 on the char surface.
In contrast, an increase in combustion temperature decreasesthe following N2O formation (Lan and others, 2001):
NCO + NO ➝ N2O + CO
The formation of N2O from char-N is also accordinglydecreased. However, the reduction reactions of N2O withhydrogen radicals (H, OH) are significantly enhanced (Lanand others, 2001; Armesto and others, 2002b):
N2O + H ➝ N2 + OH
N2O + OH ➝ N2 + HO2
Hence, it can be expected that with increasing temperatureNO emissions increase but N2O emissions decrease.
250
200
150
100
50
0
0.5
Fuel-N, %
NO
em
issi
on, p
pm
v
300
350
0.40.30.20.10.0
NO modelNO experimental
Figure 20 Effect of fuel-N content on NO emissions(Liu and Gibbs, 2002)
40
30
20
10
0
900
Temperature, °C
Con
vers
ion
N-N
O, %
50
60 PuertollanoCarbocol
875850825800
a)
12
8
4
0
900
Temperature, °C
Con
vers
ion
N-N
2O, % 16
20 PuertollanoCarbocol
875850825800
b)
Figure 21 Conversions of fuel-N to NO and N2O forlignite and bituminous coals (Armestoand others, 2002)
Most studies in the literature support the above generalstatement (Leckner and Lyngfelt, 2002; Basu, 1999; Anthonyand Preto, 1995). Figure 22 shows how fuel-N conversions toNOx and N2O varied with temperature for an anthracite-footcake mixture. It can be seen that when temperature wasincreased from 800° to 850°C conversion to NOx increasedfrom 30 to 50% while conversion to N2O decreased slightlyfrom 11 to 8% (Armesto and others, 2003). Lan and others(2001) have also shown, for both bituminous and anthracitecoals, an increase in NOx emissions and a decrease in N2Oemissions when the bed temperature was increased from810°C to 960°C. In tests with South African and Cimporcoals, Boavida and others (1997) found that the fuel-N toNO conversions for both coals increased with temperature inthe range of 700–1000°C. It was also shown in a pilot-scaleCFBC that NOx increased almost linearly from 80 ppm to200 ppm for a high volatile subbituminous coal due to anincrease of temperature from 760°C to 880°C. Thisrepresented an increase of 1 ppm/°C (Anthony and Preto,1995). In modelling biomass combustion in a CFB boiler,Liu and Gibbs (2002) found a decrease in N2O emissionsfrom 75 ppm to 13 ppm with increasing temperature from750°C to 900°C. A study by Armesto and others (2002b)showed that for a Spanish lignite N2O emissions decreasedfrom 560 mg/m3 to 375 mg/m3 when temperature wasincreased from 810°C to 880°C.
However, a few results contradicting the general trend of NOemissions have been obtained. For wood-chips combustion inCFB boilers, bed temperature (between 740°C and 850°C)was found to have little effect on NO emissions (Lyngfeltand Leckner, 1999). In studying a Colombian bituminouscoal, Armesto and others (2002b) found a linear decrease infuel-N to NO conversion with increasing temperature from810°C to 885°C. A modelling on biomass combustion in aCFB boiler has also suggested that the NO emissionsdecreased with increasing bed temperature (Liu and Gibbs,2002). This might be because, with increasing bedtemperature, more volatiles are released in the bottom densezone where volatile-N is mainly oxidised to N2. As a result,less char-N is carried over to the oxygen-rich zone (abovethe secondary air inlet) where it is oxidised to NO. Inaddition, the rate for the reduction of NO by char increaseswith temperature. This adds to decreased NO emissions (Liuand Gibbs, 2002). In spite of the above interpretations, the
29Understanding fluidised bed combustion
Pollutant formation
mechanisms determining the temperature dependency of NOemissions appear to remain unclear.
Excess air. An increase in excess air (the air supplied inaddition to the stoichiometric air requirement, that is, the airrequired just to burn the fuel) generally leads to increases inboth NO and N2O emissions. This is because with increasingexcess air the combustion rate increases. As a result,homogenous reactions for oxidations of HCN and NH3 areenhanced but heterogeneous reactions for reductions of NOand N2O are weakened. At the same time, the reductions ofN2O by H and OH are also weakened due to a decrease inthe radical concentrations (Lan and others, 2001).
It appears that the degree of impact differs between fuel typeand operating pressure. For instance, the effect on NO issmaller for biofuels such as wood than for coals (Lyngfeltand Leckner, 1999). On the other hand, the effect on N2Oemissions is relatively small at PFBC conditions (Anthonyand Preto, 1995).
A modelling on biomass combustion in a CFB boiler hasshown approximately linear increases of NO and N2Oemissions with total excess air ratio (total air/stoichiometricair) (Liu and Gibbs, 2002). The NO prediction roughlyagreed with the experimental results with birch and fir chips.However, a deviation of up to 25% between the model andexperiments was found. It was also shown that, forwood-chips combustion in CFB boilers, an increase in totalair ratio from 1.13 to 1.23 increased NO emissions from70 ppm to 95 ppm (Lyngfelt and Leckner, 1999).
Air staging. Air staging can substantially reduce both NOand N2O emissions from FBC. In a BFBC, secondary air isintroduced to the freeboard at one point or different levels(Abelha and others, 2003). However, in a CFBC injection ofsecondary air to the combustion chamber at differentlocations has not proven to be very effective. It has beenshown that a significant reduction in the emissions can beachieved by adding secondary air in the cyclone outlet,where most combustion is completed (Lyngfelt and Leckner,1999; Leckner and Lyngfelt, 2002).
The decreases in the emissions may be attributed to anenhanced reducing atmosphere in the bed. Under stagedcombustion, char and CO concentrations in the bed increase,therefore enhancing the rates of NO and N2O reductions onchar. In addition, more fuel-N is decomposed as N2 (Lan andothers, 2001). The effectiveness of air staging appears todepend on the type of fuel. A large reduction in the NOemissions from a CFBC was found for coals. However, forhigh-volatile fuels such as wood and sludge only a smalleffect was observed because of the low char concentration inthe bed (Leckner and Lyngfelt, 2002).
Lan and others (2001) observed that, for a Chinesebituminous coal, NOx emissions decreased by 46% when thesecondary air ratio was increased from 0 to 25% under a bedtemperature of 870°C. However, in contrast to the generaltrend N2O emissions were found to increase with secondaryair ratio. This was because in the test the freeboardtemperature increased by 70–80%. As a result, N2O
40
30
20
10
0
860
Temperature, °C
N c
onve
rsio
n, %
50
60
850840830820810800790
NOxN2O
Figure 22 Conversions of fuel-N to NOx and N2Oas a function of temperature for ananthracite-foot cake mixture (Armestoand others, 2003)
formation reactions were significantly enhanced in spite ofan increase in N2O decomposition at the same time.
Table 3 shows the effect of air staging on NO and N2Oemissions from combustion of chicken litter and peat in aBFBC (Abelha and others, 2003). It clearly shows that airstaging reduced the emissions for both 100% chicken litterand 50% chicken litter/50% peat mixture firing. In the caseof 100% chicken litter, with a secondary air/fluidising airratio of 0.4 and one point injection, NO and N2O emissionswere reduced by approximate 40–50% and 20% respectively.Further reductions were observed when secondary air wasinjected into the freeboard at two different levels and with astrong turbulence.
Gas velocity. Armesto and others (2002b) reported that thesuperficial gas velocity had various degrees of effect on N2Oemissions depending on fuel characteristics. For a Spanishlignite, N2O emissions increased from 390 to 425 mg/m3 (byless than 10%) with an increase of gas velocity from 4 to 6m/s. For a Colombian bituminous coal, however, an increaseof almost 50% was found when the gas velocity wasincreased from 5.1 to 5.9 m/s. This might be due to adecrease in the residence time with increasing gas velocity.In the case of the bituminous coal, the high-carbon ashplayed an important role in the reduction of N2O. Withincreasing gas velocity, the contact time of the gas and ashparticles decreased. This decreased the reduction of N2O,hence giving rise to higher N2O emissions.
Limestone addition. Limestone has a complex role in NOformation and reduction. As noted earlier, it is a catalyst forboth NO formation from volatile NH3 and reduction of NOby CO. The effect of limestone addition seems to differ inBFBC and CFBC (Johnsson, 1994; Anthony and Preto,1995). In CFBC an increase in NO emissions is normallyobserved. However, in BFBC or PBFBC the effect issomewhat less clear and there are instances of increase, noinfluence, or decrease. An explanation of the differentobservations in BFBC and CFBC is given elsewhere(Johnsson, 1994; Anthony and Preto, 1995). In a BFBC thelimestone is confined to the dense phase, where the COconcentration may be high and the O2 concentration may below. Consequently, catalysed reduction of NO by CO maydominate over oxidation of NH3. However, in a CFBC thelimestone is distributed over the whole riser, and oxidation of
30 IEA CLEAN COAL CENTRE
Pollutant formation
volatiles to NO may occur in the upper dilute zone, wherethe CO concentration is lower. As a result, catalysedoxidation of volatiles like NH3 to NO may becomedominant, therefore increasing NO emissions.
Fuel type also appears to be an important factor. For reactivefuels such as bituminous coals, the oxidation of volatile-Ndominates, and limestone addition tends to increase NOemissions. Conversely, for low volatile fuels like petroleumcoke, the catalysed reduction of NO may dominate overoxidation of volatile-N. In this case, limestone addition candecrease NO emissions (Minchener and others, 2000;Anthony and Preto, 1995).
Limestone addition is usually found to decrease N2Oemissions (Basu, 1999; Lan and others, 2001). In some cases,however, no effect or an increase has been observed(Johnsson, 1994). The decrease is because the limestonecatalyses the decomposition of N2O, and the different trendsmay be attributed to the differences in limestone activity anddegrees of sulphation in FBC (Johnsson, 1994; Armesto andothers, 2002b). An additional explanation for decreased N2Oemissions is that limestone addition may change the oxidationpath for HCN from homogenous oxidation, selective for N2Oformation, to heterogenous catalytic oxidation which isselective for NO formation (Johnsson, 1994; Lan and others,2001). This agrees well with the decrease in N2O andsimultaneous increase in NO observed on addition oflimestone. In cases where limestone addition has no effect, theoxidation of HCN over char and bed material may alreadyfavour the heterogenous reaction path (Johnsson, 1994).
Ammonia injection. NH3 injection can also be effective inreducing the NO emissions from FBC. It was observed thatinjection of NH3 with the fluidising air increased the NOemissions, whereas injection in the splashing zone had littleeffect. However, a large reduction in NO emissions wasfound when NH3 was injected into the freeboard of a BFBCor the cyclone of a CFBC, in both cases with low solidsconcentration. This might be attributed to the concentrationof solids. At higher solids concentrations (for example, in theriser of a CFBC), the catalytic oxidation of NH3 to NOdominated over the reduction of NO by NH3 in the gasphase. In contrast, at lower solids concentrations thereduction of NO by NH3 became more important (Johnsson,1994).
Table 3 Effect of air staging on NOand N2O emissions from a BFBC (Abelha and others, 2003)
Conditions usedNO emissions, ppm at11% O2
N2O emissions, ppm at11% O2
100% chicken litter, no air staging 280–360 30–45
100% chicken litter, air staging
secondary air/fluidising air = 0.4, one point injection, little turbulence 160–180 24–36
secondary air/fluidising air = 0.4, two level injection, strong turbulence 120–140 18–22
50% chicken litter and 50% peat, secondary air/fluidising air = 0.4two-level injection, strong turbulence
80–120 12–26
Figure 23 shows the NOx emissions from a BFBC as afunction of NH3/NOx molar ratio (Lan and others, 2001). AtNH3/NOx molar ratio of 5, a 39% NOx reduction wasachieved without air staging. With NH3 injection and a 15%secondary air ratio, the NOx emissions were reduced by60%.
It is generally accepted that NH3 injection increases the N2Oemissions from FBC (Anthony and Preto, 1995; Basu, 1999;Lan and others, 2001). However, no explanation of thisobservation can be found in the literature.
5.2 Sulphur oxides
Sulphur in the fuel is oxidised to SO2 during the combustionprocess. The higher the fuel sulphur content, the more SO2
produced. This also applies to co-combustion in FBC. Forinstance, a lignite/foot cake mixture was found to producemuch higher SO2 levels than an anthracite/foot cake mixturedue to different sulphur contents of the mixtures (Armestoand others, 2003). The lignite had a higher sulphur contentthan the anthracite (0.84% compared to 0.27%). This gave ahigher sulphur content for the lignite/foot cake mixture thanfor the anthracite/foot cake mixture at the same mixing ratio.
In FBC, the SO2 formed is captured in situ by added sorbentmaterial which most commonly is limestone or dolomite.The sulphur retention can be greater than 95% but sorbentutilisation levels are relatively low. Here, the mechanisms ofsulphation, sorbent performance, and process parametersaffecting sulphur retention are discussed.
Sulphation mechanismsIn units operating at atmospheric pressure, limestone ordolomite first calcines to CaO or CaO.MgO. The CaOcomponent then reacts with SO2, producing CaSO4 as a solidresidue. In the case of dolomite, the MgO componentremains unreacted. The global chemical reactions are as
31Understanding fluidised bed combustion
Pollutant formation
follows (Podolski and others, 1995; Anthony and others,2003):
Limestone:
CaCO3 ➝ CaO + CO2
2CaO + 2SO2 + O2 ➝ 2CaSO4
Dolomite:
CaCO3.MgCO3 ➝ CaO.MgO + 2CO2
2CaO.MgO + 2SO2 + O2 ➝ 2CaSO4.MgO
In PFBC, the partial pressure of CO2 is much higher than inatmospheric FBC at the same temperature. Under thiscondition, the calcination of limestone is unfavourable butdolomite may partially calcine, that is, only the MgCO3
component calcines. As a result, SO2 reacts directly withcalcium carbonate in limestone or dolomite. Thus, thesulphation process can be represented by the followingglobal reactions (Podolski and others, 1995; Anthony andPreto, 1995):
Limestone:
2CaCO3 + 2SO2 + O2 ➝ 2CaSO4 + 2CO2
Dolomite:
2CaCO3.MgCO3 + 2SO2 + O2 ➝ 2CaSO4
.MgO + 4CO2
It is argued that reverse sulphation may occur when the spentsorbent enters the lower dense bed (Basu, 1999). In thisoxygen deficient zone the CaSO4 formed is reduced,releasing sulphur dioxide. Studies of this phenomenon inCFB show that both C and CO reduce the CaSO4, but CO isthe stronger reducing agent. The reduction reaction with COis given below:
CaSO4 + CO ➝ CaO + SO2 + CO2
However, no reduction of CaSO4 takes place at temperaturesblow 850°C.
Sorbent performanceFBC has sorbent utilisation levels of typically 30–50%(Anthony and others, 2003). This is significantly lowcompared with the stoichiometric calcium utilisationsachievable by conventional flue gas desulphurisation. Thelow level of utilisation increases limestone use for a givenSO2 reduction, hence the associated costs of transport,processing of solids and disposal.
Sorbent performance in FBC has been well detailed in theliterature (Podolski and others, 1995; Anthony and Preto,1995). It is generally accepted that sorbent performance hasno relationship to chemical composition, except for a weakcatalytic effect of iron oxide on reactivity. For example, ahigh-purity limestone (of >95% CaCO3) will not necessarilyperform better than a low-purity one (of <80% CaCO3).
240
200
160
120
80
40
0 3 4 5 6 7 8
NH3/NOx molar ratio
NO
x, p
pm
1 2
280
320
secondary air ratio = 0% bed temperature = 873°Csecondary air ratio = 15% bed temperature = 851°Csecondary air ratio = 15% bed temperature = 821°C
Figure 23 Effect of ammonia injection on NOxemissions (Lan and others, 2001)
It appears that pores in sorbent particles play an importantrole (Podolski and others, 1995; Anthony and others, 2003;Anthony and Preto, 1995; Basu, 1999). At atmosphericpressure, both limestone and dolomite calcine, resulting in asignificant increase in porosity. For limestone, the porositycan change from 0.3–12% to over 50%. A number of studiessuggest a correlation between the mean pore size/total porevolume of calcined sorbents and their sulphation capacity. Itis expected that the sulphation reaction takes place rapidlyinitially because of the large pore volume. As sulphationproceeds, the reaction product CaSO4 (which has a greatermolar volume than CaO) tends to block the original pores. ACaSO4 layer is thus built up particularly on the largerparticles. This tends to prevent further reaction between theunreacted, but no longer accessible CaO core and SO2.However, this sulphation mode does not necessarily apply tosmaller particles, perhaps less than 500 µm (the exact sizeprobably depends on the sorbent characteristics). For suchparticles, pore diffusion does not limit the penetration of SO2
to the centre and the sulphation occurs effectively across theparticle.
In PFBC, dolomite and limestone are expected to performdifferently according to the role of pores. As limestone doesnot calcine, it would be unreactive due to its relatively lowporosity. However, the MgCO3 component in the dolomitecalcines, creating sufficient pores into which SO2 can readilyenter to react with CaCO3. Early data suggested Cautilisation of about 50% for dolomite but only 30–35% forlimestone. In more recent studies, however, uncalcinedlimestone was shown to perform as well as or even betterthan the precalcined limestone in PFBC. It has beengenerally acknowledged that both dolomite and limestonecan retain their reactivity under conditions where the sorbentwould not be expected to calcine.
Effects of process parameters on sulphurretentionIt has been observed that both pressure and excess air levelhave little effect on sulphur retention (Podolski and others,1995). However, other factors such as bed temperature, gasvelocity, Ca/S molar ratio, and sorbent particle size can playimportant roles as detailed below.
Bed temperature. It is well established that at atmosphericpressure the maximum sulphur retention is achieved attemperatures between 800°C and 900°C, typically at about850°C. At either side of the temperature range, sulphurretention can fall rapidly. There have been severalexplanations for this phenomenon (Anthony and Preto, 1995):● higher temperatures enhance the reverse sulphation
reaction, therefore decreasing sulphur retention;● higher temperatures enhance the sintering of sorbent
particles. This reduces their porosity and surface area,hence decreasing the reaction with SO2;
● the SO2/SO3 equilibrium plays an important role, withhigh temperatures reducing the availability of SO3 forreaction with CaO;
● higher temperatures increase the sulphation rate.However, the reaction product tends to block small poresmore quickly at the particle surface, thus preventing thepenetration of SO2/SO3 to the unreacted CaO core;
32 IEA CLEAN COAL CENTRE
Pollutant formation
● oxygen depletion in the dense bed, due to increasedvolatile combustion at higher temperatures, inhibits thesulphation reaction above 900°C.
For PFBC, however, there is no pronounced maximum forsulphur retention as a function of temperature (Anthony andPreto, 1995). Within the expected range of operatingconditions, sulphur retention is found to increase slightlywith bed temperature. When dolomite is used as a sorbent,the increase with bed temperature can continue up to as highas 950°C (Podolski and others, 1995).
Gas velocity. An increase in fluidising gas velocitydecreases the residence time of gas in the bed. This reducesthe effective contact between the SO2 and sorbent particles,and hence sulphur retention. Adanez and others (2001)studied sulphur retention in a CFBC using limestone for twodifferent coals. It was shown that for the lignite sulphurretention decreased from 85% to 77% with increasing gasvelocity from 4 m/s to 5 m/s. For the anthracite, sulphurretention decreased from 60% to 45% when the gas velocitywas increased from 3.2 m/s to 5 m/s.
Ca/S molar ratio. For both atmospheric and pressurisedFBC, sulphur retention increases with increasing Ca/S molarratio (Adanez and others, 2001; Wang and others, 2001b;Podolski and others, 1995). It has been shown that for PFBCsulphur retention in excess of 95% can be achieved at Ca/Sratios less than 2.0 and 90% sulphur retention is possiblewith Ca/S ratios between 1 and 1.5 (Podolski and others,1995). However, this can vary from case to case. Forinstance, for a CFBC with limestone addition the sulphurretention was found to be generally below 90% for Ca/Sratios between 2 and 3 (Adanez and others, 2001).
Sorbent particle size. A decrease in the sorbent particlesize increases sorbent reactivity, hence enhancing sulphurretention (Adanez and others, 2001). The degree of impactcan be very different. In a study for PFBC at 0.8 MPa,sulphur retention only increased slightly when themass-mean sorbent particle size was reduced from 0.74 mmto 0.37 mm. However, another study for PFBC at pressuresof 0.5–1 MPa showed an increase in sulphur retention fromjust over 50% to greater than 90% with decreasing mass-mean particle size from 0.99 to 0.23 mm (Podolski andothers, 1995).
5.3 Particulates
Fine particles are predominantly formed from the mineralmatter in the coal during the combustion process. Themineral matter present in the coal can be regarded as aninert. However, some chemical changes do take place. Forexample, water of crystallisation is removed and carbonatesand sulphides decompose to form oxides. This produces aresidue which is referred to as ash (Merrick, 1984). A largeproportion of these ash particles will be sufficiently small.They will therefore not remain in the bed and be carried outin the flue gas stream. In addition, attrition of the bedmaterials (including sulphur sorbent if added) also producesa significant amount of fine particles.
Winter and Liu (2000) described the ash formation in CFBC.In this process, small areas of the mineral matter (which isincluded in the coal matrix and well dispersed) form fine ashparticles. These fine ash particles are removed from the charsurface due to external forces. They are subsequentlyelutriated or circulated by the cyclone, depending on theparticle size. Usually, they can be found in the flyash. On theother hand, larger areas of the mineral matter (which is notnecessarily included in the coal matrix) are liberatedgradually and form larger ash layers. These ash layersusually stay connected to the char particle. When thecombustion is completed, the ash layers form to larger ashparticles which are fully subject to attrition. As a result, morefine ash particles are generated.
It is argued, however, that the flyash is formed due mainlyby the following two mechanisms: a small fraction of theash-forming compounds in the fuel vaporises and formsnew particles by nucleation; the non-volatile ashcompounds are then transferred to the char surface wherethey agglomerate and coalesce with each other (Lind andKauppinen, 2000).
Particulates from FBC differ from those generated frompulverised coal combustion. In general, they exhibit thefollowing characteristics:● smaller sizes. In CFBC larger flyash particles are
recycled back to the combustor, thus the particlesescaping from the cyclone are relatively small;
● FBC operates at temperatures below the ash fusionpoint. As a result, it produces irregularly shaped flyashparticles compared to the spherical particles formed bymelting and condensation in pulverised coal combustion(Takeshita, 1994; Buhre and others, 2002; Latva-Somppiand others, 1998);
● with limestone addition in FBC, the flyash containssignificant levels of CaSO4, CaO and CaCO3, and ismore alkaline and porous than that from pulverised coalcombustion (Armesto and Merino, 1999; Takeshita,1994);
● the flyash from atmospheric FBC has much higherlevels of unreacted CaO than PFBC flyash. This isbecause CaCO3 calcines to CaO at atmosphericpressures whereas this does not occur in PFBC(Armesto and Merino, 1999).
5.4 Carbon monoxide
Carbon monoxide is the product of the incompletecombustion of carbon or carbonaceous material. It is formedwhen the oxygen supplied is less than the amount required toconvert all of the carbon in the fuel to CO2 or, when themixing of air with fuel is not efficient (Podolski and others,1995; Abelha and others, 2003). CO emissions from FBCcan be less than 100 ppm due to good mixing and hence highcombustion efficiency (Minchener and others, 2000).Generally, CFBC produces lower levels of CO than BFBCdue to the higher degree of mixing in the bed.
CO emissions from FBC may be influenced by a number offactors including coal type, bed temperature, fluidising air
33Understanding fluidised bed combustion
Pollutant formation
velocity, excess air, and air staging. The effect of eachparameter is discussed below.
Coal typeCoal type (reactivity) has a strong effect on CO emissions.Armesto and others (2001) studied the combustion of ananthracite and a lignite in a bubbling bed. It was found thatthe CO emissions from the anthracite were much higher thanthose from the lignite. This might be attributed to the lowerfreeboard temperature for the anthracite although the densebed temperature was the same for the two coals.
Bed temperatureIt can be expected that CO emissions decrease withincreasing bed temperature due to the increased combustionrate. This trend was found in the combustion of wood-chipsin a CFBC, where an increase in bed temperature from740°C to 850°C resulted in a decrease in CO emissions from85 ppm to near zero (Lyngfelt and Leckner, 1999; Lecknerand Lyngfelt, 2002). However, the situation may be differentfor a BFBC. Armesto and others (2002a) studied thecombustion of rice husk in a BFBC at a fluidising airvelocity of 1.2 m/s. It was found that the CO emissionsincreased with bed temperature initially but droppeddramatically with a further increase in temperature.
Fluidising air velocityAn increase in the fluidising air velocity decreases theresidence time of CO in the bed. This reduces the conversionof CO to CO2, therefore increasing the CO emissions.Figure 24 shows the CO emissions for the combustion of ricehusk in a bubbling fluidised bed at a temperature of 840°C.It shows that the CO emissions increased from 1085 mg/m3
to 1688 mg/m3 with increasing the fluidising air velocityfrom 1 to 1.2 m/s (Armesto and others, 2002a).
1500
1000
500
1.3
Fluidising air velocity, m/s
CO
em
issi
ons
(mg
/Nm
3 at
6%
O2)
2000
2500
1.21.110.9
Figure 24 Effect of fluidising air velocity on COemissions (Armesto and others, 2002)
Excess airExcess air also has a significant impact on the CO emissions.This was observed for the combustion of olive cake and coalin a CFBC (see Figure 25) (Topal and others, 2003). In thecase of olive cake, the CO level was about 3000 mg/m3 at anexcess air ratio of 1.25. It decreased rapidly with increasingexcess air ratio and reached a minimum at an excess air ratioof 1.35. When the excess air ratio exceeded 1.6, however, theCO level was found to increase. A similar trend was
observed for coal combustion. At an excess air ratio of 1.25,the CO level was around 1000 mg/m3. It decreased sharplywith increasing excess air ratio due to the increasedcombustion rate. At an excess air ratio of 1.36, the minimumCO level was reached. However, a further increase of theexcess air ratio resulted in an increase in the CO level. Thiswas because the effect of residence time became dominant.An increase in excess air increased the gas velocity. This inturn decreased the residence time of CO in the bed, resultingin incomplete combustion.
Air stagingAir staging is shown to significantly reduce the CO level of aBFBC. Abelha and others (2003) studied the combustion ofchicken litter in a BFBC. They observed that without airstaging the CO level was considerably higher and fluctuatedbetween 1500–6000 ppm. However, with a secondaryair/fluidising air ratio of 0.4 and one point injection, the COlevel decreased to 360–540 ppm. When secondary air wasinjected to the freeboard at two different levels and with astrong turbulence, the CO level further decreased to about50–120 ppm.
5.5 Trace elements
Coals and other solid fuels may contain various traceelements in the form of either elements or compounds. Theyare usually measured in parts per million although theconcentration of sodium and chlorine can approach 1% insome coals. During the combustion, these trace elements arereleased into the flue gas and solid waste streams. There arecurrently 23 trace elements of interest to environmentalregulators and plant operators, including antimony, arsenic,barium, beryllium, boron, cadmium, calcium, chlorine,chromium, cobalt, copper, fluorine, lead, manganese,mercury, molybdenum, nickel, phosphorus, potassium,selenium, sodium, vanadium and zinc. Most of these areclassified as hazardous air pollutants. Mercury, which isfound in virtually all coals, has been of particular interestand heightened perhaps due to its toxicity (Podolski andothers, 1995).
34 IEA CLEAN COAL CENTRE
Pollutant formation
At FBC temperatures, most of these trace elements (such asheavy metals) in the fuel vaporise and pass to the gas phase(Abelha and others, 2003; Oikari and others, 2003). Forinstance, mercury is highly volatile and present almostexclusively in the vapour phase, in the form of eitherelemental mercury or mercury salts such as HgCl2, HgO,HgS and HgSO4 (Ho and others, 2001). Some trace metalsmay react with the bed material (such as sand and ashparticles) and reside in the bed (Lind and others, 1999). Therest are carried out of the bed by the combustion gases. Asthe gas temperature drops, these trace elements maycondense onto cooler surfaces and flyash particles which arethen captured by the cyclones or fabric filters (Abelha andothers, 2003; Podolski and others, 1995). However, thoseremaining in the gas phase or enriched on sub-micronparticles will escape from the particulate control unit. Thefinal emissions of trace elements will therefore depend notonly on their concentrations in the fuel but also on how theypartition to the bed material, flyashes and flue gases (Attallaand Nelson, 2002).
Mojtahedi and others (1990) studied the partitioning ofmercury, arsenic and seven other trace elements in a PFBC.Peat and coal containing respectively 0.3 and 2 ppm ofmercury were burned. In some runs, limestone containing4.7–4.9 ppm of mercury was added. The combustion gas wascleaned by a cyclone and a ceramic candle filter before itwas cooled and depressurised. Ashes from the bed, cyclone,filter and cleaned flue gas were analysed. They surprisinglyfound that only 8–16% of the mercury in the coal andsorbent left as vapour. Of the remainder, about 75% wascaptured by the bed ash, 20% by the cyclone ash and 5% inthe filter catch. As with mercury, each of the other elementswas found in the ashes at the four locations with the quantitydiminishing (but more notably than for mercury) along thegas stream from bed to filter exhaust. The level in the filteredgas was only 0.5–5% of the element input in the fuel andsorbent.
A study by Lind and others (1999) has also shown, duringthe combustion of forest residuals in a CFB, 65–68% of zincin the fuel was captured by the bed ash. However, lower
5000
4000
3000
2000
1000
0
1 2.4
Excess air
CO
em
issi
ons
(mg
/m3
at 7
% O
2) 6000
7000
2.22.01.81.61.41.2
CoalOlive cake
Figure 25 Effect of excess air on CO emissions from olive cake and coal combustion (Topal and others, 2003)
levels of enrichment were observed for copper (28–30%) andlead (14–16%). They assumed that the enrichment was dueto the reactions of zinc, copper, and lead with SiO2 in thebed sand and ash particles under the bed conditions.
Kouvo and Backman (2003) investigated co-combustion ofbiomass, peat, and refuse-derived fuels in a bubblingfluidised bed. They reported that the bed material capturedsubstantial amounts of lead, copper, zinc and manganese.However, these metals were released from the bed when theprocess parameters or the fuel characteristics changed. Forinstance, high bed temperatures were found to lead to Pbrelease from the bed material. In the case of zinc, the releasewas strongly dependent on the chlorine content of the fuelwhile the temperature had less effect. This might be becausethe ZnO (presumably) present in the bed reacted with Cl,forming ZnCl2 which was subsequently released from thebed material.
In contrast, a recent study by Abelha and others (2003) hasshown much lower levels of enrichment in the bed ash thanin the flyash leaving the combustor. In this study, chickenlitter was burned alone or with peat in a bubbling fluidisedbed. Ashes collected from the bed and the two cyclones wereanalysed. It was found that only 6–12% of the potassium inthe fuel was enriched in the bed ash. Another 56–62% wascaptured in the first cyclone ash and 22–32% in the secondcyclone ash. The amount emitted as vapour was thereforevery small. For chlorine, about 70% of the total input wascaptured in the cyclone ashes. A similar trend was observedfor most of the heavy metals such as copper, nickel, lead,manganese and chromium. In fact, the first cyclone ash wasfound to contain higher levels of heavy metals than thesecond cyclone ash. This suggested that most of thecondensation occurred in the first cyclone.
Another example was found for the combustion of biomassand waste in a bubbling fluidised bed (Coda and others,2001). Chlorine was observed to completely volatilise fromthe bed ash. It was recovered only in the flyash collected bythe cyclone and filter.
5.6 Summary
Nitrogen oxides and nitrous oxide. At the relatively lowtemperatures of FBC, nitrogen in the combustion air does notoxidise to a significant extent, thus virtually all NOx and N2Oare generated from fuel-nitrogen. The formation of NO andN2O is a complex process. During devolatilisation, nitrogen inthe fuel is partitioned into volatile-N (mainly as HCN andNH3) and char-N. Both the volatile-N and char-N aresubsequently oxidised to NO and N2O. HCN is an importantprecursor for N2O formation while NH3 is mainly formed toNO and N2. Char-N oxidation generally gives much higherlevels of NO compared to N2O. The NO and N2O formed arethen partially reduced to N2 through homogenous reactions orheterogenous reactions over char and other bed solids. Theformations of NO and N2O are affected by several factors:● the higher the fuel-N, the more NO and N2O produced;● an increase in fuel-volatile content increases NO
formation but decreases N2O formation;
35Understanding fluidised bed combustion
Pollutant formation
● with increasing temperature NO formation increases butN2O formation decreases;
● an increase in excess air generally leads to increases inboth NO and N2O formations;
● air staging can substantially reduce both NO and N2Oformations;
● an increase in gas velocity increases N2O formation;● limestone addition decreases N2O formation. It normally
increases NO formation in CFBC, but the effect issomewhat less clear in BFBC or PBFBC;
● NH3 injection can reduce NO formation but increasesN2O formation.
Sulphur oxides. The SO2 is formed due to oxidisation ofsulphur in the fuel. It is then captured in situ by addedsorbent (limestone or dolomite). At atmospheric pressure, thesorbent calcines to CaO or CaO.MgO. The CaO componentthen reacts with SO2, producing CaSO4 as a solid residue. InPFBC, limestone does not calcine but the MgCO3
component in dolomite does. Thus SO2 reacts directly withCaCO3 in limestone or dolomite. The calcination createssufficient pores into which SO2 can readily enter to reactwith sorbent. As sulphation proceeds, the reaction producttends to block the original pores. This prevents furtherreactions, resulting in low sorbent utilisation. However, thistheory does not explain why uncalcined limestone can retainits reactivity in PFBC. At atmospheric pressure, sulphurretention has a maximum at bed temperatures of 800–900°C.In PFBC, however, it increases slightly with bedtemperature. It is also found to increase with increasing Ca/Sratio and with decreasing gas velocity or sorbent particlesize.
Particulates. Fine particles are formed predominantly fromthe mineral matter in the coal. In this process, water ofcrystallisation is removed and carbonates and sulphidesdecompose to form oxides (ash). A large proportion of theseash particles will be sufficiently small and elutriated. Inaddition, attrition of the bed materials also produces asignificant amount of fine particles. Particulates from FBCdiffer from those generated from pulverised coal combustion.They are generally smaller in size, irregularly shaped, andmore alkaline and porous.
Carbon monoxide. CO is the product of the incompletecombustion of carbon or carbonaceous material. It is formedwhen the oxygen supplied is less than the amount required toconvert all of the carbon in the fuel to CO2 or, when themixing of air with fuel is not efficient. The CO level varieswith coal type (reactivity). It generally decreases withincreasing bed temperature, decreasing gas velocity, or theuse of air staging. It is also found to decrease rapidly withincreasing the excess air ratio initially, but increase with afurther increase of the excess air ratio.
Trace elements. At FBC temperatures, most of the traceelements in the fuel vaporise and pass to the gas phase.Certain amounts of trace elements are captured by the bedmaterial, depending on the type of element, bed temperatureand fuel characteristics. The rest is carried out of the bed bythe combustion gases. As the gas temperature drops, theycondense onto cooler surfaces and flyash particles which are
then captured by the cyclones or fabric filters. However,those remaining in the gas phase or enriched on sub-micronparticles will escape from the particulate control unit. Thefinal emissions will therefore depend not only on theirconcentrations in the fuel but also on how they partition tothe bed material, flyash and flue gas.
36 IEA CLEAN COAL CENTRE
Pollutant formation
6 Conclusions
Fluidised bed combustion differs from conventionalpulverised coal combustion. It uses a continuous stream ofair to create turbulence in a mixed bed of fuel, inert materialand coarse fuel ash particles. Combustion occurs attemperatures typically between 800°C and 900°C. A constantmixing of particles encourages rapid heat transfer andcomplete combustion. Fluidisation, combustion and pollutantformation constitute the fundamental issues of FBC.
Fluidisation● fluidising velocity is a key parameter of fluidisation. Its
choice affects most other process parameters. It rangesfrom 1 m/s to 3 m/s in bubbling beds and can be up to8 m/s in circulating beds;
● the bed particles are present in a wide size range. Inbubbling beds, the mean diameter ranges from 0.5 mmto 1.5 mm. In circulating beds much finer particles areused. The particle size, together with density and shape,determine its terminal velocity and finally fluidisationbehaviour;
● in bubbling beds, bubbles are formed at the airdistributor and rise with a growing size. When themaximum size is reached, they become unstable andsplit into smaller ones. In this process, the bed expandsand finally reaches a stage with a roughly stable surface.It would be worthwhile to understand better how thebehaviour of bubbles may affect solid mixing;
● in fluidised beds, solids mixing is driven by eitherbubble motion or turbulence. Good mixing is usuallyachieved. However, segregation may occur particularlywhen solids of a broad size distribution and/or differentdensities are fluidised;
● effective gas-solids contacting gives a high rate of heattransfer. The heat transfer coefficient decreases with thebed particle size and hence, fine beds require less heattransfer surface than coarse beds. In contrast, thecoefficient increases with increasing solids concentrationand bed temperature.
CombustionWhen coal particles enter a fluidised bed, they undergodrying, devolatilisation, combustion of volatiles, andcombustion of the residual char:● the drying process removes the surface and inherent
moisture of a fuel. Low rank coals contain high inherentmoisture, and its evaporation may occur in conjunctionwith and slow down the devolatilisation process. Also,the loss of water can result in significant morphologicalchanges;
● further heating releases the volatile matter of the coal. Itappears that the volatile release increases withtemperature but decreases with coal rank and gaspressure. The devolatilisation rate is controlled bychemical kinetics or heat transfer, depending on particlesize. It is argued that the mechanism changes during theheating-up of a particle. In general, the devolatilisationtime increases with particle size and moisture content. Itdecreases with increasing heating rate, oxygen
37Understanding fluidised bed combustion
concentration and bed temperature. However, the effectof coal type is unclear;
● combustion of volatiles can be a homogenous or aheterogenous reaction. In bubbling beds, homogenousreactions can occur either in bubbles or in a particulatephase, depending on the bed temperature. In circulatingbeds, the distinction between the bubbles and theparticulate phase is lost. Under normal conditions, thevolatiles will burn completely. In bubbling beds some ofthe combustion may occur above the bed, particularlyfor high-volatile fuels;
● char burning is the final stage of coal combustion. Mostchars appear to react in a shrinking particle model. Theburning rate depends on the coal type, fluidisationregime and char diameter. It is generally higher incirculating beds than in bubbling beds. Fines areproduced in the bed due to char burning out,fragmentation and attrition. They account for smallamounts but contribute significantly to fixed carbonconversion. However, some of the fines may elutriateand contribute to inefficient combustion.
In general, combustion efficiency is higher in CFBC than inBFBC and higher in PBFBC than in BFBC. It increases withfuel volatile content or bed temperature whereas gas velocityhas the opposite effect. An increase in excess air increasescombustion efficiency initially. However, a further increasemay decrease combustion efficiency.
Pollutant formationAt FBC temperatures, virtually all NOx and N2O aregenerated from fuel-nitrogen. During devolatilisation,nitrogen in the fuel is partitioned into volatile-N (mainly asHCN and NH3) and char-N. Both the volatile-N and char-Nare subsequently oxidised to NO and N2O. HCN is animportant precursor for N2O formation while NH3 is mainlyformed to NO and N2. Char-N oxidation generally givesmuch higher levels of NO compared to N2O. The NO andN2O formed are then partially reduced to N2 throughhomogenous or heterogenous reactions. The formations ofNO and N2O are affected by several factors:● the higher the fuel-N, the more NO and N2O produced;● an increase in fuel-volatile content increases NO
formation but decreases N2O formation;● with increasing temperature NO formation increases but
N2O formation decreases;● an increase in excess air generally leads to increases in
both NO and N2O formations;● air staging can substantially reduce both NO and N2O
formations; ● an increase in gas velocity increases N2O formation;● limestone addition decreases N2O formation. It normally
increases NO formation in CFBC, but the effect issomewhat less clear in BFBC or PBFBC;
● NH3 injection can reduce NO formation but mayincrease N2O formation.
SO2 is formed due to oxidisation of sulphur in the fuel. It is
then captured in bed by added limestone or dolomite. Atatmospheric pressure, the sorbent calcines to CaO whichthen reacts with SO2. In PFBC, limestone does not calcinebut the MgCO3 component in dolomite does. Thus SO2
reacts directly with CaCO3 in limestone or dolomite. Thecalcination creates sufficient pores into which SO2 canreadily enter to react with sorbent. As sulphation proceeds,the reaction product tends to block the pores and henceprevent further reactions. However, this theory does notexplain why uncalcined limestone can retain its reactivity inPFBC. At atmospheric pressure, sulphur retention has amaximum at bed temperatures of 800–900°C. In PFBC,however, it increases slightly with bed temperature. It alsoincreases with increasing Ca/S ratio and with decreasing gasvelocity or sorbent particle size.
Fine particles are formed predominantly from the mineralmatter in the coal. In this process, water of crystallisation isremoved and carbonates and sulphides decompose to formoxides (ash). A large proportion of these ash particles will besufficiently small and elutriated. In addition, attrition of thebed materials also produces fine particles. Particulates fromFBC are generally smaller in size, irregularly shaped, andmore alkaline and porous.
CO is formed when the oxygen supplied is less than theamount required to convert all of the carbon in the fuel toCO2 or, when the mixing of air with fuel is not efficient. TheCO level varies with coal type (reactivity). It decreases withincreasing bed temperature, decreasing gas velocity, or theuse of air staging. It also decreases rapidly with increasingexcess air ratio initially, but increases with a further increaseof the excess air ratio.
Most trace elements in the fuel vaporise and pass to the gasphase. They are subsequently captured by the bed material,depending on the type of element, bed temperature and fuelcharacteristics. The rest leaves the bed, and condenses ontocooler surfaces and flyash particles which are then collectedby a particulate control unit. However, those remaining in thegas phase or enriched on sub-micron particles will escape inthe flue gas.
As seen above, considerable advances in understanding FBCprocesses have been achieved over the last decades.However, there are a number of areas requiring further work.These include, in particular, how the behaviour of bubblesmay affect solid mixing, the mechanisms controllingdevolatilisation rate, the effect of coal type ondevolatilisation time, the effect of limestone addition on NOformation in BFBC and PBFBC, and the sulphationmechanisms in PFBC.
38 IEA CLEAN COAL CENTRE
Conclusions
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