handbook of combustion (online) || fluidized beds

11Fluidized BedsZbigniew Bis

11.1Introduction

The experiences of the last two decades have clearly demonstrated that the safety ofpower supplies cannot be guaranteed without the production of electricity from theconversion of chemical energy of solid fossil fuels. Bituminous and sub-bituminouscoals, as well as lignite, are the reservoirs of primary energy with the longest provenand documented perspective of use. However, the requirements of sustainabledevelopment and the limitation of the emission of the greenhouse gases have forceda search for other modern power generation technologies. Societies in most of thedeveloped countries have solved the problems associated with the formation of the�acid rains� and the emission of other gaseous and solid pollutants in order to meetthe required emission standards. The efficiency of the conversion of energy of solidfuels has significantly increased and the process has become more �environmental-friendly.� However, commercialization of the power production market has forcedthe necessity to reduce power generation costs. That need has, for example, beenmetby the use of some fuels that have been so far been treated as waste fuels (e.g.,municipal and industrial waste, etc.). The combustion of such more problematicfuels must be realized, however, with high efficiency and their availability and mustbe economically profitable to compete with other advanced technologies in the freemarket. Such goals have been achieved by implementation of the Clean CoalTechnology Program, the excellent product of which is the technology of fluidizedbed combustion. Currently, fluidized bed boilers are designed either as atmosphericor pressurized units. The furnaces of both types of those boilers can be designedeither as bubbling or circulating fluidized beds. The present chapter gives basicinformation on the requirements and conditions needed to fluidize the solids as wellas the characteristic features and know how, along with industrial achievements andapplications and future perspectives for development.

It is commonly known that coal is the world most abundant fossil fuel. In theUnited States alone, for example, coal reserves exceed the oil reserves of the rest of the

Handbook of Combustion Vol. 4: Solid FuelsEdited by Maximilian Lackner, Franz Winter, and Avinash K. AgarwalCopyright � 2010 WILEY-VCH Verlag GmbH & Co. KGaA, WeinheimISBN: 978-3-527-32449-1

j399

world. In someEuropean countriesmost energy is supplied from coal, like in Poland,for example, where it coversmore than 90% of the primary energy supply. Geologicalinvestigations have indicated that proven coal reserves are large enough to provideenergy for the next 200–300 years.

However, coal is not an excellent fuel since its many impurities, like sulfur andnitrogen, are �trapped� inside its particles [1]. In a typical PC (pulverized combustor),the coal is crushed into very fine particles, and then blown into the furnace where it isignited to form a long and lazy flame. In other types of the combustors, like stokerboilers for example, the burning coal is placed on various kinds of grates. Thetemperature inside the furnaces of those boiler types can be very high and oftenexceed 1000 �C. When the coal is burned under such conditions the impurities areeasily released into the air. Therefore, the coal was treated as a �dirty� fuel in the1970s. Unfortunately, such a conviction is still dominant in many countries.However, the subsequent energy crisis and blackouts made painfully clear to thesocieties of many countries that oil and gas are �strategic� fuels and very powerfulpolitical tools. Therefore, the end of the twentieth and the beginning of the twenty-first century has been the right time for an intensive search for solutions andtechnologies that provide a safe supply of power from coal, and alternative orrenewable fuels. Such activities also arose from the present requirements andenvironment protection standards, and the need to rationalize the use of primaryenergy sources, and the use of primary energy sources for sustainable developmentand minimization of CO2 emission. The implementation of the fluidized bedcombustion technology has been an excellent support for realization of thoseambitious goals.

In the early 1990s, the magazine POWER called the development of fluidized bedcoal combustors �the commercial success story of the last decade in the powergeneration business.� The success, perhaps themost significant advance in coal-firedboiler technology in half a century, was achieved largely through the US Departmentof Energy Office of Fossil Energy�s technology program called The Clean CoalTechnology (CCT) Program that began in 1985 [2].

The development of the fluidization technology has been dated back to the 1920swhen the first commercial Winkler fluidized bed reactor was constructed in Ger-many [3]. In the UK the development of fluidized bed coal combustion technologybegan in the early 1960s. At almost the same time (mid-1960s) decision-makers in theindustrial sector throughout the world recognized that fluidized bed boilers are notonly a potentially lower-cost andmore efficient tools for coal combustion but can alsoconvert the coal chemical energy in amuch cleaner way. Furthermore, the turbulent,that is, �fluidizing,� mixing of solids, coal, and oxygen can effectively improve thecombustion process and heat transfer rate. Since the fluidized bed combustiontechnology also evolved from efforts to develop a combustion process where thepollutant emission could be easily managed without any external emission controlsystems (such as e.g., scrubbers) the technology provides conditions for burning thefuel at 850–950 �C, that is, temperatures well below the threshold where nitrogenoxides form [4]. The excellent mixing in fluidized bed furnaces provides theconditions for intensive contact of the flue gases and sulfur-absorbing agents, such

400j 11 Fluidized Beds

as, for example, limestone or dolomite, so that over 95% of coal�s sulfur can becaptured by the sorbent inside the FBC (fluidized bed combustor).

Coal, like the other fossil fuels and all living creatures on Earth, is mainlycomposed of carbon molecules [5]. Accordingly, during the combustion processthose molecules react with oxygen and form carbon dioxide. CO2 is a colorless andodorless gas that contributes significantly, as many scientists believe, to the so-called�greenhouse effect,� that is, the process associated with increase of the Earth�stemperature and altering our planet�s climate [6]. To decrease the costs of reducingcarbon dioxide emission new research and investigation are going on at somesophisticated FBC technologies, such as, for example, the oxygen-fired FBC. Thelatter process represents a �flagship project� of carbon capture and storage (CCS)technology [6].

11.2Theory

The crucial components for stable operation of afluidized bed are a sufficient batch ofsolid loose bulkmaterial, thefluidized apparatus, and thefluidizing agent [7–11]. Theloose solids bulk material is the population of solids (usually of mineral origin) formeddue to natural or mechanical grinding process. A batch (bed) of loose material is acertain amount (kg or m3) of the solids that is placed in a container or on the grate(fluid distributor) of the fluidized bed apparatus. The fluidizing agent is fluid (gas orliquid) used to initiate and maintain stable operation of the fluidization process.The fluidized bed apparatus is a separate systemor part of a plant (Figure 11.1) used forindustrial realization of the fluidization process for various chemical and physical

Figure 11.1 Schematics of two main fluidized bed systems: bubbling fluidized bed reactor (left)and circulating fluidized bed reactor (right).

11.2 Theory j401

technological processes. To achieve the fluidization and operate the system in astable way the operator must know the particle size distribution (PSD) of the bedsolids. Themaintenance and stability of the PSD is crucial for proper operation of theindustrial fluidized bed system (e.g., the boiler) with respect to both load andemission [7, 11–15].

Since practically all the solids (either natural or artificial ones) are of different sizeand shape they are described by the term poly-dispersive to be distinguished from themono-dispersive solids. The latter are seldom found in practice.

The solids particles are forming a statistical set where the particle diameter, d, is therandom variable. The distribution of the diameters forms a spectrum. The spectrumis, according to statistical mathematics, the relationship between the number ofsolids particles Dni that are within the interval:

di�Dd2

; di þ Dd2

� �

and the diameter di, where di is the median diameter of each i-th range, and Dd iswidth of each of the ranges (Dd is a constant value). The spectrum is made based onmeasurements of the diameters of solids particles. The measurement relies onclassifying each diameter within the corresponding diameter range. The totalnumber of the particles (strength of the set) is:

NT ¼Xmi¼1

Dni

where m is the number of diameter ranges.The particle numberDni in each diameter range divided by the total number,NT, of

particles in the set represents the contribution (histogram; Figure 11.2) of each of thediameter ranges:

D�ni ¼ DniNT

Figure 11.2 Histogram of particles of a loose material.


The total sum of heights of each of the contributions is 1 (or 100%). The histogramcan be replaced by a solid curve that is, the probability density function (PDF):

fnðdÞ ¼ D�niDd

¼DniDdPm

1Dni

ð11:1Þ

Assuming the diameter d is continuous in the range (dmin, dmax) we may write:

fnðdÞ ¼ d�ndd

¼dnddÐdmax

dmin

dndd dd

ð11:2Þ

The distribution of the variable d, is thus:

FðdÞ ¼Xm1

D�ni ð11:3Þ

or:

FðdÞ ¼ðd0

f ðdÞdd ¼ðd0

d�ndd

dd ¼

Ðd0

dndd dd

Ðdmax

dmin

dndd dd

ð11:4Þ

All the above relationships are applicable not only to particle diameters but also tosurface and volume (mass) distributions. The particle mass distribution may be, forexample, described as:

d�Vdd

� d3d�ndd

ð11:5Þ

where d�V is the volumetric contribution of solids particles for a given range.Based on that rule we obtain:

faðdÞ ¼dndd d

a

Ðdmax

dmin

dndd d

add

ð11:6Þ

and:

FaðdÞ ¼

Ðd0

dndd d

add

Ðdmax

dmin

dndd d

add

ð11:7Þ

For a¼ 0 we have a number distribution, a¼ 2 a surface distribution, and a¼ 3 avolumetric (mass) distribution.

11.2 Theory j403

To describe the functions f(d) and F(d) various equations have been formulated toapproximate the experimental distribution data. Some of the commonly used are theequations of:

. Rosin–Rammler

. Nukiyama–Tanasawa

. Log–normal distribution.

The Rosin–Rammler equation describes the total volumetric (mass) solids dis-tribution as:

F3ðdÞ ¼ 1�exp � diX

� �n� �ð11:8Þ

where

X ¼ scale parametern ¼ shape parameter

When the particle diameter is equal to the scale parameter we obtain:

diX¼ 1; F3ðdÞ ¼ 1�e�1 ¼ 0:632

The above value of 0.632 means that 63.2% of the volume (mass) of the solids iscontributed by particles of sizes less than X, while the remaining 36.8% correspondsto the contribution of solids of sizes >X.

The Rosin–Rammler equation may be differentiated. As a result we obtain thevolumetric distribution function:

f3ðdpÞ ¼ dF3ðdpÞddi

¼ nXn

dn�1i exp � di

X

� �n� �ð11:9Þ

Figure 11.3 shows example functions f3(d) and F3(d).

11.2.1Average Particle Size

The average particle diameter is the diameter of a set of uniform spherical particles ofthe same size. The term was introduced in order to conduct the calculations andanalyses, as well as to predict the behavior of the fluidized bed systems. Severalvarious average diameters are distinguished (Table 11.1); all of them can be calculatedfrom the following relationship:

dpq ¼

ffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiPm1dpi Dni

Pm1dqiDni

p�q

vuuuuut ð11:10Þ


11.2.2Parameters of a Bed of Solids

While the solids particles seem to form a solid body, a SEM (scanning electronmicroscopy) picture brings to light their complicated structure (Figure 11.4). A well-developed network of pores, chinks, and splits in conjunctionwith an irregular shapemakes, on one hand, the geometric interpretation of individual particle sizes verydifficult, but, on the other hand, it is a fundamental factor intensifying the hetero-geneous reaction on particle surface, as it is for the processes of, for example, coalcombustion or SO2 sorption by limestone or dolomite.

Figure 11.3 Example solids volumetric (mass) distribution functions: probability densityfunction (PDF) – f3(d) (Equation 11.9) and cumulative distribution function (CDF) – F3(d)(Equation 11.8).

Table 11.1 Exponents �p� and �q� used to calculate average diameters.

p q Average diameter Application

1 0 d10 Arithmetic Comparison of various systems2 0 d20 Surface Surface control, surface phenomena

(e.g., absorption, evaporation)3 0 d30 Volumetric Volume control, volumetric phenomena2 1 d21 Surface, relative Disintegration of particles, adsorption3 1 d31 Volumetric relative (P Robert) Evaporation, interparticle diffusion,

combustion3 2 d32 Volumetric-surface (Sauter) Heat and mass transfer, fluid –

particle systems4 3 d43 Mass (de Brouckere) Solids handling, combustion

11.2 Theory j405

The density of solids is defined as the quotient of solids mass mp and theirhydraulic volume Vp:

rp ¼mp

Vpð11:11Þ

Since the picnometric liquid does not penetrate the inside of the pores the value ofthe hydraulic volume exceeds the real volume of the solids (Figure 11.4), and thus thedensity determined from (11.11) is often called the apparent density. The absolute(true) density of solids can be calculated from the relationship:

rap ¼mp

Vm¼ mp

ð1�epÞVpð11:12Þ

where:

ep ¼ Vpor

Vp¼ Vp�Vm

Vp¼ 1�Vm

Vp

is the porosity of the solids particle determined by BET (Brunauer–Emmet–Teller)sorption or by mercury porosimetry.

For a fixed bed of solids the total bed volume is the sum of the volumes of allindividual particles Vp and the volume of the void space (�pores�) between theparticles. The total volume of the poresmay be determined as the difference betweenthe bed volume and the volume of solids (Vbed � Vp). By analogy to the porosity ofparticles, we may introduce the term �bed porosity:�

e ¼ Vbed�Vp

Vbed¼ 1� Vp

Vbed¼ 1�s ð11:13Þ

By similarity to the true density of the particles wemay also define the bulk densityof solids:

rb ¼mp

Vbed¼ mp

ð1�eÞVpð11:14Þ

Figure 11.4 SEM cross-section of an individual solid particle.


Using (11.11), the relationship (11.14) may be rearranged as:

rb ¼mp

Vbed¼ rp

ð1�eÞ ð11:15Þ

Equation 11.15 is commonly used to determine the bed porosity since from it wemay easily obtain the relationship:

e ¼ 1� rbrp

ð11:16Þ

The density of solids depends on the degree of compression of their particles, andthus to determine the solids density the requirements contained in the correspond-ing standardsmust be obeyed [10]. As an example, for the bulk densities of silica sandofrb¼ 1450 kgm�3 andrp¼ 2650 kgm�3

, the porosity of the bed solids composed ofthat sand yields:

e ¼ 1� rbrp

¼ 1� 14502650

¼ 0:45

11.2.3Critical Fluidization Velocities

When the population of fine solid particles formed a packed bed in a vessel, such as,for example, shown in of the reactors in Figure 11.1, the solids exert thrust on thebottom of the reactor:

Pp�Pg ¼ Dp ¼ AHð1�eÞðrp�rgÞgA

¼ Hð1�eÞðrp�rgÞg ð11:17Þ

where Pp and Pg are the gravity force of solids and the fluid (gas), respectively.When a gas orfluid is pumped upward through a bed offine solid particles at a very

low flow rate the fluid percolates through the void spaces (pores) without disturbingthe bed and, accordingly, the bed remains in a fixed (or packed) state. For such a statethe pressure drop through a packed bed of spherical particles with mean Sauterdiameter d32 may be determined from the well-known Ergun equation [10]:

�Dp ¼ H 150ð1�eÞ2

e3mUd232

þ 1:75ð1�eÞe3

rgU2

d32

" #ð11:18Þ

The minimum fluidization velocity of the bed of solids is, according to thedefinition, �the minimum velocity of filtration of gas through the bulk bed of loose;for which the bed of solids behaves like the liquid since the gravity of the solids equalsthe drag force of the upward flowing filtrating gas.� According to the above definitionthe fluidization condition for the bed may be for U¼Umf, and e¼ e0, written as:

�DpH

¼ DPH

ð11:19Þ

11.2 Theory j407

By using the relationships (11.16) and (11.17) in (11.18) we obtain:

150ð1�e0Þ2

e30

mUmf

d232þ 1:75

ð1�e0Þe30

rgU2mf

d32¼ ð1�e0Þðrp�rgÞg ð11:20Þ

The solution of the Equation 11.20 is presented in Figure 11.5.Introducing the non-dimensional Reynolds and Archimedes numbers:

Remf ¼rgUmf d32

mð11:21Þ

and:

Ar ¼ grgd332ðrp�rgÞm2

ð11:22Þ

and modifying Equation 11.20 we obtain:

150ð1�eÞe3

Remf þ 1:751e3

Re2mf ¼ Ar ð11:23Þ

The solution of Equation 11.23 is:

Remf ¼ 42:857ð1�e0Þ 1 þ 3:1� 10�4Are0

3

ð1�e0Þ2 !0:5

�1

24

35 ð11:24Þ

When the upward flow rate is very high the drag force of the gas is higher thanparticle gravity and the bed solids may be elutriated from the apparatus.

Figure 11.5 Graphical solution of Equation 11.20; the abscissa of the intersection point of thefunctions Dp¼ f(U,d32,rg,m) and Dp¼ f(H,rp, rg, e) represents the minimum fluidization velocity.


If the distances between the solid are long enough so that mutual collisionsbetween the particlesmay beneglected, the behavior of the particlesmay be describedby a simple momentum equation:

mpdwp

dt¼ Fd�Fg ¼ pd232

4CDrg

U�wp

�� ðU�wpÞ2

�pd3326

rpg ð11:25Þ

The terminal velocity is, according to the definition, �the lowest gas velocity forwhich a single solids particle can be �suspended� by the upward flowing fluid.� Theterminal velocity is also often treated as the minimum velocity of the pneumatictransport.

According to the above definition the condition for suspension of a single solidsparticle may be written as follows.

For U¼Ut:

dwp

dt¼ 0 and wp ¼ 0 ð11:26Þ

Using the Ar and Re numbers and Equation 11.26 we obtain from Equation 11.25:

34CDRe

2t ¼ Ar ð11:27Þ

where:

Ret ¼rgUtd32

mð11:28Þ

CD ¼ 24Re

ð1þ 0:15 Re0:687Þ for Ret < 1000 ð11:29Þ

CD ¼ 0:44 for RetY1000 ð11:30ÞSolution of Equation 11.27 yields:

Ret ¼ffiffiffiffiffiffiffiffiffiffi43ArCD

rð11:31Þ

Figure 11.6 gives an example of the relationship (11.24) for three assumed values ofporosity for packed beds and Equation 11.31.

For a given loosematerial the fluidization phenomenonmay only be observed for acertain range of the velocity of the fluidizing agent, Umf<U<Ut. Accordingly, theinformation regarding the value of the velocity of the fluidizing agent is crucial forproper design of thefluidized bed system and for the selection of the appropriate fansto maintain the required gas flow rate and control, for example, of the operation of afluidized bed boiler. In addition, maintenance of the proper solids particle sizedistribution is also important, particularly for the case of a FB (fluidized bed) boiler.

An increase in size of the solids particles brings about not only a significantincrease of Umf and terminal velocity Ut but also a narrowness of the range of gasvelocity for which a stable fluidization may be maintained. With coarse solids it is

11.2 Theory j409

difficult to initialize the fluidization since high gas velocities are required (significantenergy consumption), and it is quite difficult to maintain fluidization since at suchhigh gas velocities the pneumatic conveying state may easily be achieved. Suchremarks are the practical reasons why stones (pebbles) are not transportedpneumatically.

The above discussion applies predominately to beds of narrow particle sizedistribution. In commercial plants (e.g., the furnaces of atmospheric or pressurizedfluidized bed boilers) the fluidized beds consist of polydispersive mixtures of solids.An example composition and particle size distribution of such solids are shownbelow in Figure 11.19.

In such a bed the minimum fluidization velocity must be determined for thereal particle size distribution in the bed, which may differ significantly from theparticle size distribution of the feed material due to various in-bed processes, suchas, for example, combustion, attrition, elutriation of fines, or agglomeration ofparticles.

The fluidization velocity may be estimated for an average (e.g., median) particlesize; for such a case the permeability average is themost appropriate.However, as theflow rate of the upward flowing fluid is increased, the fine bed particles, that is, thosein the voids between the larger particles, will fluidize much earlier than the largerones, and the bed will partially fluidize at a gas velocity significantly lower than thepredicted value (average). To fluidize the whole bed the minimum fluidizationvelocity should be estimated for the largest particle. However, particular attentionmust be also paid to the terminal velocity of the finest particles to avoid anyentrainment of the fines from the top of the bed.

For a bimodal distribution of the bed solids (as it is very often found in industrialplants) one may distinguish two size ranges as, for example, shown in Figure 11.7.For such case several fluidization conditions may exist. Some basic ones are:

Figure 11.6 Non-dimensional minimum fluidization velocity (Remf) for different porosities (e0)of the bed at minimum fluidization state, and terminal velocity (Ret) for particles of different loosematerials, characterized by the Archimedes number (Ar).


. Complete segregation, depending on formation of a layer offines on the surface ofthe coarse particles (the coarse solids are a certain grate for fine particles).

. Partial segregation, depending on partial mixing of the coarse and fine solidsat the boundary of the two layers (the boundary between both beds is �washedout�).

. Complete mixing – as a result both solids types are uniformly spread within thebed.

When the diameters d1M and d2M are significantly different, it is easy to segregateboth solids types, particularly if the gas velocity is increased aboveU1

mf (Figure 11.6).For simplicity of further analysis, such materials may be called as fine (d1M ¼ df ) andcoarse (d2M ¼ dc). A further increase of the gas velocity aboveU1

t allows separation ofthe fine particles from the bed of coarse ones. During the �separation� a significantdifference between the values of the velocities of both particle types can be observedand unavoidablemutual impacts between both particle types can occur (Figure 11.8).

The effect of the momentum exchange of both solids populations is the mutualinteractions force [16]. Its vertical component may be expressed as:

ðFfcÞz ¼p

4Gs0d2c

dfdc

þ 1

� �2

ðwf Þz ð11:32Þ

where Gs is the solids mass flux (kgm�2 s�1).In this case the condition of suspension of coarse particle in the upward flowing

suspension of fine particles and gas is [according to (11.26)] [16]:

ðFdÞz þFg þðFfcÞz ¼ 0 ð11:33Þ

Using the relationships (11.21) and (11.22) the solution of the above equation forthe coarse and fine solids may be written in a non-dimensional form:

Figure 11.7 Bimodal particle size distribution (PDF) function with two peaks (modes d1M; d2M).

11.2 Theory j411

Rect ¼ðGs0Þ2c2Dc

þ 2Gs0dcReftc2Dcd

f32

þ 43ArccDc

" #0;5�Gs0

cDcð11:34Þ

where according to Reference [16]:

Gs0 ¼ kðGsf ÞGsf ; kðGsf Þ ¼ 2�exp � GsfGs�f

� �2" #

Gs�f ¼ 13 kgm�2 s�1

For Gs0 ¼ 0 Equation 11.34 is the same as Equation 11.31. Figure 11.9 shows thesolution of Equation 11.34 for the solids in Figure 11.7.

Figure 11.8 Impacts between fine and coarse solids during a vertical gas–solids flow; df¼ 67mm,dc¼ 8mm, U¼ 4ms�1.

Figure 11.9 Terminal velocity of coarse particles versus solids mass flux of the fine particlessuspended in the gas. The solid line corresponds to Equation 11.34.


The above equation shows, contrary to the results in Figure 11.6, that pneumatictransportation of the pebbles may be easily conducted by the suspension of fineparticles and the gas. Such results are of great practical importance since they mayexplain, for example, the behavior of polydispersive solidsmixtures in the furnaces offluidized bed boilers (particularlyMSFB:multi solidsfluidized bed [16]), the behaviorof coarse coal particles in the CFB (circulating fluidized bed) furnace, or thephenomenon of high separation efficiency of the cyclones of CFB boilers [13].

11.2.4Structure of Fluidized Bed

The type of solids, its particle size distribution and density, as well as the densityand viscosity of the fluidizing gas, are important parameters affecting not only theminimum fluidization conditions but also the structure of the bed [7, 8, 10, 16, 17].

Accordingly, the Geldart classification of solids powders (Figure 11.10) is a usefultool to predict the fluidization structure of real solid materials. Geldart�s diagramdescribes the relationship between particle diameter and the density differencebetween the solids and the gas. Derek Geldart [9] identified four groups of loosematerials (four particles types: A, B, C, and D); each type behaves quite differentlywhen they reach the state of fluidization.

Group A particles are characterized by:

. the bed expands considerably before bubbling occurs;

. bubbling bed fluidization (Figure 11.11), gas bubbles rise more rapidly than therest of the gas;

Figure 11.10 Geldart classification of loose solids. Particle properties are related to the type offluidized beds [9].

11.2 Theory j413

. bubbles spit and coalesce frequently through the bed, maximum bubble size isless than 10 cm;

. gross circulation of solids occurs.

Group B particles are the most common. These beds:

. are made of coarser particles than group A particles and are more denser,

. form bubbles as soon as the gas velocity exceeds Umf,

. form small bubbles, at the distributor, that grow in size throughout the bed.

Group C particles:

. difficult to fluidize and tend to rise as a slug of solids,

. form channels in large beds with no fluidization, and tend to be cohesive.

Group D particles:

. very large, dense particles;

. form bubbles that coalesce rapidly and grow large;

. form bubbles that rise more slower than the rest of the gas phase;

. form beds whose dense phase surrounding the bubbles has low voidage;

. cause slugs to form in beds when the bubble size approaches the bed diameter,and spout from the top of the bed easily.

Figure 11.12 presents the schematics of the evolution of the bed structure for thesolids of group B for increased superficial gas velocity U. The parts of the figuresreferring to the following states:

Figure 11.11 Structure and mixing of loose solids (Geldart�s group A) in a bubbling fluidizedbed.White solids are themain bedmaterial, while the black solids are carried by the gas bubbles andtrack the bubble passages.


. a – filtration,

. b– start of fluidization (minimum fluidization) (incipient),

. c and d – bubbling and turbulent fluidization, respectively,

. e – fast (circulating) fluidization,

. f – pneumatic transport.

Filtration is the state before the fluidization and its proper maintenance isimportant for various industrial processes, such as, for example, combustion ofsolid fuels in the stoker boilers. As can be seen in Figure 11.13 the pressure drop of apacked bed depends, according to the relationship of Ergun (11.18), not only on thegas velocity and solids particle size but also on the bed voidage. In Figure 11.12 [8]there are two lines (solid anddashed) in region a. The solid one represents the value ofthe pressure drop for increased gas velocity (beginning of fluidization). For a �loose�bed, for example, (or a decrease of gas velocity) the pressure variations are repre-sented by the dashed line. The characteristic feature of the fluidization is a constantvalue of the pressure drop of the bed for various gas velocities (Figure 11.5). Duringturbulent fluidization (region d) the measured values of the pressure drop of the bedfluctuate intensively (standard deviation of pressure drop D(Dp) increases) due tointensive coalescence and breakup of the gas bubbles. The result is the characteristicfeature of turbulent fluidization, that is, very intensive mixing of solids. The numberand size of the gas bubbles are proportional to the gas velocity. Since the cross-sectionof the fluidized bed is constant the increase in number and size of the bubblesshortens the distance between the bubbles and increases their coalescence. As aresult, the size of the bubbles increases and larger bubbles are formed. At the bed

Figure 11.12 Comparison of the bed structure and bed pressure drop for increasing velocity of thegas: (a) filtration (packed bed), (b) incipient of fluidization (minimum fluidization), (c) bubblingbed, (d) turbulent fluidization, (e) fast fluidization, (f) pneumatic transport.

11.2 Theory j415

surface the bubbles �erupt� like a volcano and eject significant amounts of solids intothe freeboard.During the �eruption� the height of thefluidized bed (aswell as the bedpressure drop) are instantaneously changed, bringing about the fluctuations of thegas pressure according to Equation 11.17. In the case of large and high fluidized bedsthose dynamic phenomena may interfere with the fluidized bed apparatus and beresponsible for its low frequency vibrations that have to be taken into considerationduring the reactor design phase.

The number of bubbles increases at higher gas velocity and so does the mixing ofsolids and their elutriation from the bed. However, during the �eruption� of bubbles,solids of all particle sizes are elutriated, while only some of them are �returned� to thebed – the finest solids are carried over by the gas phase. During the eruption ofbubbles (Figure 11.11) the structure of the fluidized bed is similar to that of boilingwater. Accordingly, if the height of the reactor is insufficient, stable operation ofthe reactor can only be maintained if a special solids separation device is installed inthe freeboard and the separated solids are returned to the bed. The value of theterminal velocity of individual solids particles decides whether the particles arecarried over by the gas or not. The terminal velocity of coarse particles depends on theconcentration of the fines according to the relationship (11.34) – see also Figure 11.9.Consequently, the change from bubbling to turbulent fluidization brings about asignificant increase in elutriation, so that the stable maintenance of the bed and thedetermination of the bed height become very difficult. It is also difficult to identifythe individual gas bubbles since their coalescence is so rapid that one practicallyobserves well-developed channels connecting the gas distributor with the bedsurface. Under such conditions the amplitude of pressure fluctuations becomes amaximum, and, in fact, further increase in gas velocity brings about transition to fast

Figure 11.13 Comparison of bed structures and pressure gradient versus gas velocity; Upt is theminimum pneumatic transportation velocity; the labels a–f are as described in Figure 11.12.


fluidization. In fast fluidizationmode the solids are carried over by the gas and fill thereactor (region e in Figure 11.12). Furthermore, a specific �inversion� of the basicelements of bed structure is observed; the instabilities that were discrete so far, like,for example, upward flowing large bubbles, are replaced by clusters, that is, theagglomerates of solids that are falling against the gas–solids flow. During suchoperation of the reactor a significant part of the bed may be carried out of the vessel.For steady-state operations the particles must be separated from the gas (e.g., by acyclone) and recirculated back into the bed. Under such conditions the hydraulicresistance of the bed strongly depends on the recovering intensity, that is, the massflow of the recirculated solids (kg s�1), or on the solidsmassfluxGs (kgm�2 s�1). Thefast fluidization process cannot be maintained in a reactor designed for bubblingfluidization and a specific type of apparatus must be used (Figure 11.1). In com-mercial operation the fast fluidization mode is maintained in a circulatingfluidized bed (CFB). Increasing the gas velocity in a CFB reactor above the terminalvelocity for the coarsest solid particles brings about the transition to the pneumatictransportation mode. It is practically impossible to use one type of reactor for allfluidization modes.

Measurement of the pressure drop of the bed shown schematically in Figure 11.12cannot determine straightforwardly the mode of fluidization because its valuedepends on the bed height (i.e., the mass of solids in the reactor). The pressuregradient measured for two ports located not far from each other is a much betterparameter. In commercial reactors of large cross-sections the pressure drop mea-sured according to the schematics shown in Figure 11.13 is practically equal to thebed hydraulic resistance since the resistance of the gas flow is so small that it can beneglected. The pressure drop may be described by Equation 11.17.

If we replace the bed height, H, in Equation 11.17 by the distance Dh betweenthe neighboring pressure measurement ports then Equation 11.17 can be trans-formed into:

DpDh

¼ ð1�eÞðrp�rgÞg ð11:35Þ

or:

DpDhðrp�rgÞg

¼ ð1�eÞ ¼ s ð11:36Þ

An example of relationship (11.35) is shown in Figure 11.13.The main difference compared to Figure 11.12 is the monotonically decreasing

pressure gradient for increased velocity of the fluidizing gas. According to therelationship (11.36), starting from the region b, the line in Figure 11.13 alsorepresents average concentrations of solids in the reactor between the pressuremeasurement cross-sections, the distance between which isDh. That information isof great practical importance since it allows us to identify the mode of fluidizationinside the reactor [7] and also to determine the solids concentration along the reactorheight by measurements of the bed pressure, which is performed by pressuremeasurement ports located at one of the reactor�s walls (e.g., at the wall of a CFB

11.2 Theory j417

boiler [12, 14]). An example pressure distribution along the furnace of a CFB boiler isshown below in Figure 11.23. By multiplying the two sides of Equation 11.17 by thereactor cross-section area, A, and dividing them by the gravity acceleration, g, weobtain the relationship allowing us to estimate themass of solids inside of the reactor:

ADpg

¼ AHð1�eÞðrp�rgÞ ¼ mbed ð11:37Þ

The above relationship is commonly used by CFB staff to control the operation ofcommercial CFB boilers [12, 14, 18].

11.2.5Heat Transfer in Fluidized Bed

The intensive heat transfer is another advantage of the fluidization process [19–21].In a CFB combustor the heat transfer is put into practice by a complicatedmechanism and so several parameters are taken into consideration, such as, forexample, physical and geometrical properties of the bed solids, the bed structure, gasflow, radiation, and so on. It may be assumed that in a CFB boiler the heat istransferred in three ways:

1) gas convection2) solids convection3) radiation.

The equivalent bed-to-surface (ab-w) heat transfer coefficient may be expressed as:

ab-w ¼ ag-w þap-w þarad ðWm�2 K�1Þ ð11:38Þ

where:

ag-w¼ gas-to-surface heat transfer coefficient (gas convection),ap-w¼ solids-to-surface heat transfer coefficient (solids convection),arad¼ radiation heat transfer coefficient (radiation).

The values of the components of the sum (11.38) are plotted in Figure 11.14 versusrelative solids concentration N. The component of particle convection, ap-w, thatrepresents the solids-to-surface heat transfer varies considerably. In accordance withthe results of References [19, 20]:

ap-w ¼ 34

aplgd32

ð1�e0ÞNffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffið1�N

23Þ

qðWm�2K�1Þ; N ¼ 1�e

1�e0ð11:39Þ

where a is a coefficient.The value of the solids-to-surface heat transfer coefficient rapidly increases

after the bed starts to be fluidized since the solids begin to move rapidly and thenumber of mutual interactions between the solids, as well as between the solids andthe heat transfer surface, is also increased similarly to the well-known Brownianmovements. As long as the frequency of interactions increases so does the heat


transfer coefficient. However, an increase of gas velocity brings about a decrease ofthe solids concentration (Figure 11.13) and despite the increase of the average solidsvelocity the frequency of the mutual solids–surface collisions decreases. From theheat transfer point of view two ranges of fluidization may then be distinguished(Figure 11.14):

1) where the solids concentration is high and the heat transfer rate depends mainlyon solids velocity;

2) where the solids velocities are quite high and the heat transfer rate dependsmainly on solids concentration.

In thefirst range, just after the state ofminimumfluidization is reached (areaBB inFigure 11.14), an increase in gas velocity brings about an increase of heat transferrate, while in the second range [areas CFB and PT (pneumatic transport) inFigure 11.14], for which an increase of gas velocity brings about a decrease of solidsconcentration, the rate of heat transfer by solids convection also decreases.

The relationship between the gas and solids mixing in a fluidized bed is crucialfrom the point of chemical processes, such as, for example, combustion, SO2

sorption, and so on [7, 11, 22]. As shown in Figure 11.11, two components of themixing processmay be distinguished. The first component (along the reactor height)gradually increases versus the number and size of the bubbles, particularly for theturbulent fluidization regime where the mixing occurs throughout the whole bed.The second component (lateral mixing, i.e., across the reactor) is similar to thephenomenon of heat transfer by solids convection –the only difference is that themaximum mixing is achieved for a bubbling fluidization regime.

Figure 11.14 Components of the heat transfer mechanism in two-phase systems. BB – bubblingbed, CFB – circulating fluidized bed, PT – pneumatic transport.

11.2 Theory j419

11.3Application in Industry

11.3.1Introduction

In a fluidized bed boiler, crushed coal particles are fed into boiler�s furnace, and thenmixed with other, mostly inert, solids (see Figure 11.19) against upward-blowing jetsof air. The mass of burning coal in the �hot fluidized bed� would bubble and tumblehere and there just like the boiling lava inside a volcano (Figure 11.15). As a result ofthis turbulent mixing of gas and solids, muchmore effective chemical reactionsmayrun and a uniform temperature of the bed may be maintained.

The first generation of fluidized bed combustion technology was developed asatmospheric bubbling fluidized bed combustors (ABFBCs) [23]. The early ABFBboilers were designed similar to the simple reactors, like the one on the left-hand sidein Figure 11.1. Such boilers were characterized by significant cross-sections to keepthe gas velocity low enough to minimize the elutriation of unburned coal fines andcontrol themass and temperature of the bed. Since the heat transfer rate is very goodat the bubbling fluidization (Figure 11.14) the maintenance and control of the bedtemperature below 950 �C (i.e., below the ash softening temperature) can be put intopractice by immersion of a heat transfer surface in the bed. Heating of the bed above950 �C usually brings about agglomeration of ash and coal particles, the formation oflarge agglomerates, and, finally, the melting of the bed and stoppage of fluidization.The lack of fluidization and, as a result, the lack of mixing brings about a further

Figure 11.15 Fluidized bed of hot solids inside the furnace of a commercial FB boiler. Black areascorrespond to just-introduced coarse coal particles, while light areas are the burning volatile flamessurrounding the coal particles.


increase of local temperature until the boiler is totally blocked and an emergencyshutdown has to be executed.

Apart from their many advantages, ABFBC boilers were from the beginning noteasy units for commercial operation due to, for example, the difficulties associatedwith proper control of the bed temperature in the range above a stable combustiontemperature (>600 �C) and below the ash softening (>900 �C), control of bedmass inventory due to intensive elutriation, control of the unburnables in the flyash at low level, quite low cross section heat duty (<2MWm�2), and complicatedcoal feeding systems due to large cross section area of industrial combustors [23].Accordingly, a rapid development of atmospheric circulating fluidized bedcombustors (ACFBCs) (Figure 11.17) took place from the end of the 1970sand from the beginning of the 1980s. These boilers are designed with thefundamental elements essential for circulating fluidization (cf. Figure 11.1) andthe combustion technology is based on coupling the advantages of turbulentand fast fluidization regimes (Figure 11.13). Accordingly, the technology hasovercome most of the difficulties mentioned above and as a result a rapid increasein the number of new units and their capacity can nowadays be observed(Figure 11.16).

Both types of atmospheric fluidized bed combustion technologies (ABFBC andACFBC) cross the commercial threshold, and most of boiler manufacturers arecurrently offering fluidized bed boilers as a standard package. In parallel to theACFBC, pressurized fluidized bed combustion (PFBC) technology has beendeveloped. The idea of the fluidized bed combustion systems was based on earlierdevelopment of the atmospheric fluidized bed combustion technology, but thefurnace was pressurized and the energy of the flue gases was used to drive a gas

Figure 11.16 Development of the US and world fluidized bed boiler market.

11.3 Application in Industry j421

turbine and operate the whole system as a combined cycle [24]. The first generationof pressurized fluidized bed combustors is operated based on the bubbling bedtechnology. A stationary fluidized bed in such boiler types can be established byusing of low air velocities to fluidize the bed material, and immersion of a heatexchanger (tube bundle) into the bed to generate steam and control bed tempe-rature. For such units standard cyclone separators and sophisticated tube filters areused to remove the particulate matter from the flue gas prior to entering a gasturbine.

The second generation of pressurized fluidized bed combustor is designed tooperate in a CFB regime technology with several efficiency enhancement measures.CFB technology has the potential to improve the operational characteristics by usingof higher air flow rates to fluidize and carry over the bed solids, and it provides theconditions for recirculating nearly all the bed solids by the application of high-volumehot cyclone separators.

The relatively cleanflue gas can thenflow to the heat exchangers. Such an approachtheoretically simplifies the feeding system and extends the contact time between thesorbent and the flue gas, but also reduces the probability of tube erosion, andimproves the combustion efficiency and SO2 capture.

A major efficiency enhancing measure for the second-generation PFBCs isthe integration of a coal gasifier (carbonizer) to produce a fuel gas [24]. The fuelgas is combusted in a topping combustor and then fed into the gas turbine,which is the more efficient part of the combined cycle. The topping combustormust be characterized by flame stability during the combustion of low-calorificgas, and must also have low-NOx emission characteristics. To take maximumadvantage of the increasingly efficient commercial gas turbines, the high-energygas outletting the topping combustor must be nearly free of particulate matterand alkali/sulfur content. All the effluents from the PFBCs must nowadays beessentially free of mercury, a hazardous air pollutant [1] soon-to-be regulated inEurope.

Fluidized bed combustion is now a well proven and popular technology. It is fuelflexible and almost any combustible material (coal, municipal waste, biomass, etc. –see Figure 11.17) may be burned without any problems associated with the emissionof sulfur dioxide and nitrogen oxides. The required emission standards are metwithout any need for expensive add-on control systems.

11.3.2Fundamental Rules in Designing and Operating the Fluidized Bed Boilers

A fluidized bed boiler, shown schematically in Figure 11.18, must be designed,constructed, and operated to meet at least three – often competing – requirements:

. easy fuel combustion and effective heat transfer to the working fluid;

. stable bed fluidization for all the bed loads;

. maintenance of thermal and flow conditions inside the furnace to maintainemissions of the flue gases below the standards.


Figure 11.18 View of an industrial large-scale CFB boiler.

Figure 11.17 Fuel range applicable for fluidized bed combustion. Fuels on the right-hand side canbe used with standard boiler design, while those on the left cause more challenges for multi-fueloperation and boiler design. (Figure reproduced with permission from VTT Finland.)


The most effective combustion temperature is 850 �C, since it is optimal for SO2

capture by fine sorbent particles [22] according to:

CaCO3 �!850�CCaOþCO2 ð11:40Þ

CaO þ 12O2 þ SO2 �!850�C

CaSO4 ð11:41Þ

Tomeet the requirements regarding furnace heat transfer theflue gas temperatureand the bed particle size distribution must be selected to maintain the solidsconcentration along the furnace, in order to maintain a uniform temperaturedistribution so as to meet the emission standards for SO2 and NOx at low consump-tion of limestone [12, 14, 22]. One has to keep inmind thatmaintaining too low a bedtemperaturemay bring about poorer calcination of sorbent (11.40) and in fact a lowerreaction rate. The NOx formation rate from the fuel nitrogen is decreased butincreased consumption of sorbent to maintain the required emission of SO2 mayoffset the positive effect of the reduction of NOx emission due to the catalytic effect ofthe surplus CaO on the formation and emission of NOx.

Too high a bed temperature brings about an increase of NOx emission (decreasingN2O [4]) and sorbent consumption due to blockage and sintering of the surface ofsorbent particles [12, 14, 22] that limits the sorption of SO2. Furthermore, theprobability of ash sintering is increased at higher temperatures [22]. Figure 11.19compares particle size distributions of the bed solids in a commercial CFBC(circulating fluidized bed combustor).

To maintain the required particle size distribution of the fluidized bed at least tworequirements must be met:

Figure 11.19 Cumulative distribution function (CDF) of solids particle size in a commercialCFBC (circulating fluidized bed combustor) [13].


1) The feed solids (coal and sorbent) must be ground to the designed sizes(Figure 11.20).

2) The solids separator located at furnace outlet must guarantee the optimumseparation efficiency.

Unfortunately, due to unstable coal grinding susceptibility and variations of its ashcontent and ash properties, the first requirement is very difficult to fulfill, as can beseen in Figure 11.20. The same applies to the technical aspects of proper preparationof sorbent. Fulfillment of the second condition requires an explaination of thephenomenon of high separation efficiency of large cyclones, which has been done inReference, [13], and the definition of the criteria of optimum separation efficiency ofthe cyclones of CFB boilers was presented in Reference [12].

To give the reader an impression of the size of CFB cyclones Figure 11.21 shows aninlet to a commercial CFB cyclone. The boiler is equipped with two cyclones of thatsize. It is commonly known that the fundamental solids separation mechanism incyclones is a force balance between the centrifugal and drag forces that act on solidsparticles. To increase the separation efficiency the centrifugal force has to beincreased by, for example, a decrease of the cyclone diameter. This is a commonindustrial practice.

However, the diameters of the cyclones of commercial CFB boilers are in the 10mrange and still their separation efficiency exceeds 99%. The investigations presentedin Reference [13] indicated that the mutual interaction force between the coarse andfine solids, described by the Equation 11.32, is an additional factor contributingconsiderably to the increase of separation efficiency since it changes the fundamentalforce balance acting on a solid particle and two forces (a centrifugal and particle

Figure 11.20 Comparison of the required and real CDF of solids particle size of coal and sorbentfor a large-scale commercial CFBC [13].


collision force) are acting to counterbalance the particle drag force. By their activity acertain �continuouslymoving layer of granular filter of coarse particles� is formed onthe internal wall of the cyclone [13]. Such a granularfilter captures the solidsfines andcarries them back into the cyclone discharche pipe. Figure 11.22 shows the overalleffect of the action of both separation mechanisms.

The concept of optimum separation efficiency is a complicated phenomenondirectly related to the distribution of solids concentration along the height of theCFBC furnace [12] as shown in Figure 11.23. The solids concentration along thefurnace height is a superposition of the concentration of two solids types – fine andcoarse; the fine solids (elutriable particles in Figure 11.23) are carried over by the gasinto the cyclone, while the coarse ones (non-elutriable particles in Figure 11.23) arefluidized in the lower part of the furnace.

The increase in cyclone separation efficiency of aCFBCbrings about an increase insolids mass flux [12, 25] since it is directly associated with a decrease of the mass ofsolids that are not separated by the cyclone and, accordingly, decrease of the d50 of thesolids that are circulating within the CFB loop. Since for such a case the mass of fine(elutriable) particles in the bed is increased the boiler operators must increase the

Figure 11.21 Inlet window to one of the cyclones of a commercial CFBC [13].


mass flow of the bottom (bed) ash that is removed from the lower part of the furnaceto maintain the mass inventory constant for a given boiler load. As a result, thecontribution of fines in the bed is increased while themass of coarse (non-elutriable)solids is decreased.

Figure 11.22 Function of grade cyclone separation efficiency for a commercial CFBC cyclone [13].

Figure 11.23 Distribution of pressure (directly proportional to solids concentration, seeEquation 11.17) along the furnace of a commercial CFBC [12], showing the influence of differentcyclone separation efficiencies.


According to Equation 11.39 the particle-to-wall heat transfer (main componentin Equation 11.38) in the furnace is increased and the bed temperature isdecreased. As a result, the emission of NOx is also decreased and the desulfurizationconditions are also improved (Figure 11.24). Further improvement of thecyclone separation efficiency is, however, not welcomed since the dense bed in thelower part of the furnace disappears (Figure 11.23), and, accordingly, the bedtemperature in that section is increased and the emission of NOx and SO2 isincreased again (Figure 11.24). The relationship shown in Figure 11.24 staightfor-wardly defines the optimum separation efficiency of the cyclones for a givenCFB boiler.

11.4Outlook

So far, the development of ACFBC has indicated significant potential of the fluidizedbed combustion technology and its suitability to realize the tasks associated with theCCT. Similar to the situation in the 1970s, associated with the replacement of thebubbling FBCs by the circulating FBCs, in the 1990s new Compact CFBCs haveappeared. Such boilers have been produced by Foster Wheeler [26, 27]. Apart fromtheir compact structure (Figure 11.25) these boilers are also equippedwith anoriginalexternal heat exchanger (EHE), called INTREX�; up to that time the EHE had beendesigned in Lurgi CFBCs only. Figure 11.26 shows the schematics of an INTREX [27].

Figure 11.24 Effect of modification of the cyclone separation efficiency of a commercial CFBC onthe emission of NOx and SO2 [12].


Apart from the INTREX the Compact CFBCs are also equipped with octagonalsolids separators that are integrated with the furnace. The application of INTREXincreased theflexibility and operation reliability of newfluidized bed boilers since thetemperature in the bed canbemore efficiently controlled and the processes of erosionand corrosion of the heat transfer surfaces may be limited.

This is very important since erosion is themost serious weak point of CFB boilers.The largest number of brown coal-fired Foster Wheeler Compact CFBCs is operatedat the Turow Power Plant (Poland) [27].

The experience and know-how gathered during operation of those units haveallowed Foster Wheeler to design a pioneering supercritical once thorough (OT)460 MWe CFBC [26]. The unit of such size has been in commercial operation at theLagisza Power Station, Poland, since June 2009. Figure 11.27 shows the schematicsof the boiler. The boiler is highly integrated and designed to burn various fuels withhigh efficiency and low emission of SOx and NOx. Its successful start up will be a keymilestone for further development of OT CFBCs.

Another advanced alternative for fluidized bed combustion is the technologyof oxy-combustion where the solid fuels are burned in atmospheres of oxygen

Figure 11.25 Schematic of a Compact CFB boiler.

11.4 Outlook j429

Figure 11.26 Schematics of the lower part of a Compact CFB boiler with INTREX EHE [27].

Figure 11.27 Isometric view and furnace cross-section of the OT CFBC at PKEs Lagisza PowerStation, Poland [26].


concentration >21% [6, 28]. The target goal is the development of the technologywhere the combustion process would run at atmospheres of 100%. oxygen In such acase the oxygen-fired CFBC system would produce flue gases composed of mainlycarbon dioxide and water that would be ready for sequestration or use for otherpurposes. In such a case the main advantage of the CFB over PC seems to be theeliminationof thefluegas recirculation to control the combustion temperature [6], andthe use of the circulating solids for such purpose in CFB systems (cf. Figure 11.23).Current research and engineering investigation are focused on (i) optimization of thesolids andflue gas recycle process to control the bed temperature, (ii) determinationofhow to place the heat transfer surfaces inside a CFBC system tomaximize the benefitsof increased solids recycle, and (iii) identification of appropriate methods for stagedadditionof the oxygen/recycledfluegas streams tomaintain thedesigned temperatureprofile along the combustor.

11.5Summary

As a summary of the information given in this chapter, Table 11.2 compares theadvantages and disadvantages of the fluidized bed combustion technology. The

Table 11.2 Advantages and disadvantages of fluidized bed combustion technology.

Disadvantages Advantages

. Large thermal inertia due to largemass of concreteused for the construction of the furnace and thecyclones, and, accordingly, extended start up from a�cold reserve� period

. Large thermal inertia, and, accordingly,short start up after relatively long �hot-reserve shut down� period

. Danger of erosion of the heat transfer surfaces . Lowminimum load, practically close to20–25%

. Additional power consumption to maintain thefluidization

. Low power consumption for fuelpreparation

.Need to control the particle size distribution of thesolids (fuel, sorbent, bed)

. Relatively low (<5%) concentration offuel in the bed and, accordingly, thepossibility of burning various fuels, in-cluding low-quality or waste fuels

. Some difficulties in managing and reusing thesolid combustion by-products

. Separation of the processes of com-bustion and heat transfer

. Significant consumption of additives to control theemission of sulfur compounds, particularly if highdesulfurization efficiencies (>90%) are required

. Easy flue gas emission control

. Limited size of the boilers, so far . Absorption of toxic combustion by-products by addition of cheap sorbents. Compact structure, easy to replace theexisting units since little free space isrequired

11.5 Summary j431

comparison includes the main industrial experiences regarding the fluidized bedcombustion and its industrial application.

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handbook of combustion (online) || fluidized beds

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