The fluidised combustion of coal

THE FLUIDISED

COMBUSTION

OF COAL

JANOS M. BEI~R Department of Chemical Engineering, Massachusetts Institute of Technology, Cambridge, Massachusetts 02139 Fundamental and engineering aspects of fluidised combustion are reviewed, The main areas discussed include: --the application of the two phase theory of fluidisation to the combustion of coal, --the relevance of available information on the kinetics of volatile evolution, volatile combustion and of the carbon-oxygen and carbon-carbon dioxide reactions to the combustion of coal in the fluidised bed, --the retention of sulphur by the addition of sorbents to the bed solids and the conversion of fuel nitrogen to NO followed by its partial reduction by char to Nz, --the rates of heat transfer to tubes immersed in the bed, and --the effects of operation at elevated pressures. Finally research needs arising from the practical application of fluidised combustion are discussed.

1. Introduction

In the process of meeting the growing demand for electric power and heat, coal will play an increasingly important role world wide but particularly in the United States because of the very large coal reserves. Fluidised combustion offers a clean and relatively inexpensive method for coal combustion, a method insensitive to fuel quality, capable of reducing sulphur and nitric oxide emission and providing the best hope yet for the development of a coal fired gas turbine. A coal fired gas turbine will then enable further improvements in energy conservation to be obtained by way of combined gas turbine-steam cycles and a convenient combination of combined power and heat generation. Fluidised combustion of coal involves the burning of coal particles--normally in the millimeter size range--in a hot fluidised bed of inert solids. The burning particles account for less than 2% weight of the total bed solids, more than 98% being coal ash and limestone or dolomite, the latter of which may be added for sulphur retention in the bed. The bed {emperature is usually 750~176 where the lower temperature is a combustion stability 439

limit while the higher temperature is the ceiling for clinker free operation of the bed. Fluidisation-like processes have been used for coal combustion and gasification for more than half a century although the fluidisation process as we know it today was developed in the late 1930's and has become a general tool of chemical engineering only after World War II. Some of the early examples included the Winkler gas generator, 1the Szikla-Rozinek gasifier-combustor 2 which had a s p o r e d bed slagging gasifier, Stratton spouted bed boiler system, a the Ignifluid boiler combustion system 4 and the Two-Stage Fluidised bed gasifier-Cyclone Combustor System. ~ An impetus was given to the development of fluidised combustion by the recognition that heat exchange surfaces immersed in the bed enable good control of the bed temperature to be achieved without any detrimental effect upon mixing, i.e. solids reeirculation in the bed. Experimental studies carried out in the early 1960's at BCURA 5 and CRE in Britain and at Pope, Evans and Robbins in the USA under contract to the Office of Coal Research to develop a coal fired package boiler 6 showed that coal burning fluidised beds can be operated with 18-20% excess air at temperatures

440

ENERGY PRODUCTION FROM COAL

l

~._I

0

~l

:2=

w

z

t~ uJ c::

0

o

:::)

m

~ra

0 @

n~

-t I 0 0

0 0

0

I

u

0

(atoms ~., o",!qJol

9

llNfl

..I0 3 Z I S o

0

v)

~. ,.=,

;,,

=Lk. o

UJ

~"

"a nQl

r Z

! 0 0 0

0 0

!

0

(aloes ,(jo~,tqJo) 11Nfl ~0 3ZIS

THE FLUIDISED COMBUSTION OF COAL

441

unpleasant deposit formation.) W h e n limestone or dolomite is added to the b e d material the emission of sulphur dioxide can be reduced. While there have b e e n extensive studies made and a n u m b e r of texts p u b l i sh ed on the fluid dynamics an d heat transfer in fluidised

lower than those at w h i c h ash clinkering occurs. O w i n g to the relatively low temperature in the bed the emission of nitrogen oxides is lower and little or no vapourisation of alkali salts and of silica occurs. (The condensation of these vapours on heating surfaces causes

TABLE I Fluidised-bed coal combustion facilities*

Combustion and regenerator size

Organisation National Coal Board (NCG) and British Coal Utilisation Research Association (BCURA) Pope, Evans & Robbins (PER)

U.S. Bureau of Mines

Gas Pressure, velocity, Temperature, arm fps F

6" diam, 12" • 12" 36" • 18" 48" • 24" (36" • 24") 27" diam 12" • 16" (FBC) 20" x 6' (FBM) 20" x 20" (CBC) 28' 8.2" x 12' (Rivesville) 18" diam

1 1 5 1 1 1 1 1 1

2-3 2-8 2 6-14 8-14 12-15 12-15 9-12 2-3

1240-1680 1420-1680 1470-1745 1420-1680 1450-2000 1600-2050 1450-2000 1550 1450-1650

Consolidation Coal

4" diam

1.5

2-3

1700-1900

CSIRO

9" diam

1

2-7

750-1850

6" diam 3" and 6" diam

1 8

3-7 2-5

1400-1700 1450-1650

3" diam 3" diana 12.5" and 5" diam 18" diam

1 10 10

2-4 5 5-10

1500-2000 1500-2000 1700-2000

20" diam 7" diam

1 4

4-10 5-10

1400-1800 1400-1800

6" diam

5

5-10

1400-1800

Bergbau-Forschung

16" x 31"

1

4-8

1470-1560

Lurgi

5" diam 40" diam 6" diam 24" diam

1

1

> 10 > 10 1-5 2-10

2000-2200 2000-2200 1400-2500 1400-1700

6" diam

1

3-30

1400-1700

36" diam

6

10

2100

12" x 12" diam

1

2-4

1382-1632

Argonne National Lab.

EXXON R & D

Foster-Wheeler Combustion Power

Battelle-Columbus

Sheffield University

*(After Kiang, K. D. et al 4~

1 1

Status/purpose 62-73/Combustion

66- 75 / Combustion and regeneration, development plant 66-72 / Combustion heat transfer 68-70/Combustion and regeneration 67-68 / Combustion 68-75 / Combustion, regeneration 68-75 / Combustion regeneration Operating/Combustion 73-75 / Combustion (adiabatic) 74-75 / Combustion 74-75/Combustion 62-70 / Fast bed 49 / Bloating 60-75 / Combustion 74-75/Combustion 75/Combustion gasification 72-76 / Combustion

442

ENERGY PRODUCTION FROM COAL

beds, 8,9a~ it was not until the NCB's publication of their research results 6 and the publication of reports of research sponsored by NAPCA, EPA, the Office of Coal Research and the NCB H-16 that information on the combustion of coal and on pollutant emission from fluidised combustion became available. The progress of the development of the size of experimental combustors, atmospheric and pressurised, is illustrated by Fig. 1.69 Table I contains data on experimental fluidised combustion facilities.

2. Fluidisation Characteristics of Gas-Solid Systems As the flow rate of gas through a bed of solids is increased a limit is reached where the bed starts to expand and the particles become rearranged to offer a reduced flow resistance. On a further increase of flow rate, the particles separate and become freely supported in the gas; the bed is fluidised. At this point the frictional drag of the particles is equal to the apparent weight of the bed, i.e. the pressure drop is equal to the total weight of the bed per unit distributor area. In a gas-fluidised bed, as a result of a further increase in fluidisation velocity two "phases" are formed: a continuous, dense or "particulate" phase and a discontinuous, lean or "bubble" phase. The bubbling fluidised bed resembles a boiling liquid. It is on this physical model that Davidson and Harrison have based their two phase theory of fluidisation. 17 Their basic assumptions are: a) the bed consists of a particulate phase in which the gas flow rate is equal to the rate of incipient fluidisation, i.e. the voidage fraction is essentially constant and b) the bubble phase which carries the additional flow of fluidising gas. From a number of detailed experimental studies, ~8-2zresults of which supported the two phase theory, information is now available on bubble formation and bubble rise velocity, on the flow of the fluidising gas in the vicinity of the rising bubble and on estimated values of the coefficient for mass exchange between the bubble and the surrounding particulate phase. 23 3. Models of Fluidised Combustion of Coal For the successful modelling of combustion in a fluidised bed quantitative descriptions

are needed of the motion of the fluidising gas and of the solids, the devolatilisation of the coal particles, and the combustion of the volatile matter and of the char. The problem is further complicated by the attrition of coal particles in the bed and by the reactions taking place in the free board, above the fluidised bed. In the following the two most important steps in the combustion process are discussed, the transport of the oxidant introduced with the fluidising gas to the coal particle, and the pyrolysis and combustion of the coal particle in the fluidised bed. 3.1. Gas Interchange between Bubble and Particulate Phase

The original two phase theory of fluidised beds was first published by Orcutt and Davidson and a full description is given in Davidson and Harrison's 17 book. Since this elegant model was published studies have been carried out to refine and extend it. s Because of the relative simplicity of the model and its success in describing the movement of gas and solids in fluidised beds, most mathematical descriptions of fluidised combustors have adopted Davidson's two phase theory for calculating gas interchange between bubble and particulate phase. The two phase theory implies that the flow in excess of that represented by the minimum fluidising velocity Uo is going to form bubbles and the particulate phase remains at incipient fluidisation. If the flow rate of gas to the bed is A r Uthen A r Uo flows through the particulate phase and Ar(U-Uo) flows through the bed in the form of bubbles. With the absolute bubble rise velocity assumed by Davidson and Harrison as UA

=

(U-

Uo) + 0.711(gDe) '/2

(1)

where the second term on the RHS represents the natural rising velocity of a bubble, the material balance on a single rising bubble can be given as (q + kcS)(C P - Cb)

= jc.

dt

= v vdc A

dy

(See Fig. 2 for the scheme and nomenclature of the analysis) and by integration with respect to y and for the boundary conditions of C b = C o at y = 0 w e h a v e C b = Cp + (C o - Ce)e-Ou/vAv

(3)

THE FLUIDISED COMBUSTION OF COAL

(5) Oxidant that is consumed by reaction within the particulate phase. The rate of consumption is kHCp (1-NV) where k is the reaction rate constant based on unit volume of the particulate phase.

CH

ttttt

The oxygen concentration in the free board, CH, can be obtained by the sum of terms representing the oxygen leaving the particulate and bubble phases respectively as

cb

H

0

tt

443

q u

)

UC a = (U-

+ UoCemoleO2/mZS

(6)

where C b, can be substituted for from Eq. (3). This reads further to the expression for the fraction of oxidant leaving the bed unreacted as

ttt [../

Uo)CbH

CH

Co

Co

FIG. 2. Assumed bubble pattern in fluidised bed showing concentrations. (After Davidson and Harrison. 17)

= I~e-x +

1 - 13e-x

(7)

k' + 1 - f3e - x

where 13 = 1 - U o / U is the fraction of inlet flow passing through the bed in the form of bubbles, k' = k H o / U

where Q is an equivalent flow rate to and from the bubble Q = q + kcS

(1) Oxidant that enters the particulate phase directly at the rate of U o C O (2) Oxidant that enters the particulate phase through the interphase transfer from bubbles. If N is the number of bubbles per unit volume of bed the flow rate of the oxidant can be given as C bdg

X = OH/UAV

(4)

In their analysis Davidson and Harrison assumed that the particulate phase is perfectly stirred while the bubble concentration changes with height in the bed. The material balance of the oxidant in the perfectly mixed particulate phase can be assumed to consist of the following terms:

NQ

and

(5)

0

where C b can be substituted for from

Eq. (3). (3) Oxidant that leaves at the top of the bed at the rate of Uo C r, (4) Oxidant that leaves the particulate phase and enters the bubble phase at the rate of NQHC p.

Xis the transfer factor, i.e. the number of times a bubble is flushed out while it rises to the top of the bed. 3.2 T h e R e a c t i o n s o f C o a l in the F l u i d i s e d Bed

The particulate phase of the bed where the coal particle undergoes pyrolysis and dombustion is maintained at nearly constant temperature because of the effective heat transfer between reacting and nonreacting particles on one hand and particles (mainly inert) and heat exchangers immersed in the bed, on the other. The coal, particle size normally 1-3 mm (but sometimes with some fines below lmm,) is injected in to the bed, usually near the distribution plate. The gas residence time is approximately 1 sec, while the coal particle will reside in the bed much longer, for about 100-300 secs. The bed temperature, the particle size and the proximate volatile matter in the coal will determine the volatile yield in the fluidised bed. Results of isothermal decomposition of coal obtained by Badzioeh24 and Kimber and

444

ENERGY PRODUCTION FROM COAL

Gray,2s are relevant to the pyrolysis of coal particles in the fluidised bed. It can be estimated that the time needed for the devolatilisation of a 1 mm particle is 0.5-1 sec which is of the same magnitude as the time scale of mixing in the fluidised bed. It follows that assumptions in some models that the volatile release is instantaneous on coal injection are not realistic. It is more likely that the volatiles are rather uniformly dispersed in the bed by the recirculation of the coal particles undergoing devolatilisation. When a large proportion of the feed coal consists of fine particles < < 1 mm, then a fast release of volatile matter near the coal injection point can be expected. The Combustion of Volatiles At temperatures above 650~ and in an oxidising atmosphere the rate of burning of volatiles is fast relative to the time required for the volatile evolution. If it is assumed that the volatiles form a uniform spherical shell around the coal particle the thickness of this shell can be calculated from the amount of volatiles evolved. It is then possible to calculate the rate of shrinkage of this film as a function of the rate of influx of oxygen, the latter being calculated from the equation for diffusion of 0 2 through combustion products and nitrogen. The total burning time is then calculated as the time for the shrinking film to vanish completely 26 as:

tb

0.305 TFg

(18;o)0

,o Rdr

(8)

where R o = 0.5d(41/MOo T + 1)1/3 tb is burning time of volatiles (sec) R is outer radius of spherical annulus occupied by volatiles (cm) d is particle diameter (cm) is mass fraction of oxygen in main stream gas T is temperature (~ Po is initial density of coal particles (g/cm a) M is mean molecular weight of coal volatiles; subscript o refers to initial condition Among the least known details of the combustion process are those concerning the combustion of volatiles in the fluidised bed. Firstly, there is doubt that the volatile film surrounds the particle in a uniformly thick film.

Visual observation suggests that some of the volatile matter is ejected in a jet-like manner beyond the stagnant boundary layer surrounding the particle. Secondly, the oxygen concentration in the particulate phase where most of the volatiles burn, depends on the kinetics of evolution of volatiles, their burning and also upon the interphase gas transfer between particulate and bubble phase. The visual observations made on a 3" bed 27 suggest that when the interphase transfer rate is high because of small bubble size, the volatile burns completely in the bed and the volatile evolution-combustion process can be described by means of the isothermal decomposition-volatile film combustion models as discussed above. When bubbles grow and coalesce as in larger and deeper beds, it can be expected that part of the volatile combustion will take place at the bubble boundary or in the freeboard of the fluidised bed. The Combustion of the Solid Residue Much of the early work by Nusselt, 28 and by Hottel and coworkers 29,a~ on the burning of solid carbon spheres in an oxidising atmosphere is relevant to the problem of char burning in fluidised combustion. These investigators, finding that boundary layer type diffusion was the rate determining step in the combustion process, were joined later by Spalding and by Thring and Essenhigh, the latter of whom were able to verify the Nusseltian "square law" of burning time (i.e. burning time is proportional to the square of the initial diameter of a spherical coal particle) by a well-controlled series of experiments on particles larger than 300 Ixm. Hottel and Stewart 31 were the first to suggest that the rate of chemical reaction at the coal surface might not be neglected when burning pulverised coal. Later Essenhigh and Perry 3z and Be~r and Essenhigh 33 suggested that for small particles (<100 ixm), and when the char reactivity was low, the chemical reaction, most likely an activated adsorption step, was controlling the overall rate. (Figure 3) Be6r, Marsden, Lee and Thring 34 were able to show from their experiments that the rate of combustion of a South Wales Anthracite of 30 txm mean particle size was controlled by the rate of the chemical reaction. The low temperature in the fluidised bed represents a borderline case. It is generally found by experimenters that it is not easy to maintain stable operation of a fluidised bed below 650~ bed temperature. It is at around this temperature that even for large particles

THE FLUIDISED COMBUSTION OF COAL Captive p:,rticle

I'.F. range

range

10'

.[ EXCC.S,~ ~,lt"

1 I'1j

* set_ r / 100% --

oO

Set B

v

-

t0%

-

100%,

o~

445

the diffusion control gives way to rate control by the desorption step in the chemisorption reaction. This desorption reaction has a high temperature coefficient. At this low temperature limit the combustion process is unstable; a small drop in temperature can result in extinction. Avedesian and Davidson 35 have assumed boundary layer type diffusion control for their 0.25-0.26 mm size particles. They also assumed that the carbon surface is oxidised by CO z and the resulting C O diffuses to a reaction zone surrounding the particle where it burns with oxygen to C O z, Figure 4 illustrates the model of the char particle burning and shows the concentration profiles. Assuming steadystate conditions for a shrinking spherical particle while it burns they integrate the equation = 0

r2

dr

(9)

for the boundary conditions

r=d/2 r=R r=~

1

10

100

Particle d i a m e t e r

1,000

1 era.

Cco 2=0

Cco=C o~=0 Cco ~=0,Co~=CP

where C e is the 0 2 concentration in the dense phase. The analysis shows that for R = d, the molar flow of Co2 towards the particle

In microns (d,)

FIG. 3. Comparison of calculated burning times of coal particles in a flame assuming boundary layer type diffusional control (Set A) and chemical reaction control (Set B) respectively. (After Be6r and Essenhigh. 33)

n = 4arRZE

[d(C02) l =4,rrdECp dr n

(10)

where E is the effective diffusion coefficient having a value lower than the molecular diffu-

Fro. 4. Model of char particle reacting in a fluidised bed. (After Avedesian and Davidson. 35)

ENERGY PRODUCTION FROM COAL

446

siGn coefficient D c because of the inert particles surrounding the coal particle. For the case when O~ can reach the surface and there is direct COz formation the rate equation is n o

=

3~ort

~

-

T

r

,
(10/a)

2~rdEC

~

-

[

-

-r

---~

'

By defining a transfer coefficient kg by no

= kg~dzCe

so that

kg =

9

2E/d

the Sherwood n u m b e r can be written as

g I

d Sh= kg--=2E/D Dc

c o PO,mtSnet ;~cLodeJ ,~ teos~ sqJ~'eS c:cu ct;_~

80

and the ra{e equation rewritten as n = 2~rShDcdC p

po

(10/b)

This equation is then used by the authors 35 to calculate the rate of combustion of a char particle throughout its lifetime with Sh = const. For char particles > ash particles Sh - 2% and for char particles < ash particles Sh---> 2. 9 o is the void fraction of the particulate phase. They obtained the following expression for the burn-out time of a batch of coal particles 35,36

i

40 [ I I I I ! ! i 0 1 2 3 ~5 6 7 (d~l) (INITIAL AVERAGE CHAR PARTICLE DIAMEIERi2(mm;t}

d, = 0 . 3 9 t u r n ;

rn = 3.6g c a r b o n ; U o = 4 6 t u r n # ; 7 2 0 k,g c a r b o n / m ~ 2 1 4 r a m / s ; (b) U = 3 0 0 m r n / s ; (c) U = 383 m m / s Pc =

(a) U :

FIG. 5. Burn out time as a function of initial average particle size. (After Avedesian and Davids o n . 3s )

Nm i

tb

400

1 2 C o A [ U _ ( U _ U o)e - x ]

+

pod~

48

ShD c

Co

(11)

The first term represents mass transfer from bubble to particulate phase and the second is the diffusion limited m i n i m u m burning time of a single particle when the charge Nm{ is small. Here N is the n u m b e r of particles fed each of mass m,, A is the cross sectional area of the bed and X is the interphase transfer factor. Avedesian and Davidson found good agreement between their predicted and measured batch type experimental results. (Figs. 5, 6). Typical burn-out times for 2.5 m m char particles varied from 180 to 300 sec depending upon the conditions of fluidisation. More recently Campbell and Davidson have extended the above theory to a continuously operated system. 36 Assuming constant initial particle size they obtained a linear bed particle size distribution function P = ad

1-- ~r---

1

9 .......

d~= leTmm

(12)

,oo /./

F// .

/o•

y

Z/./

g 1,00 / ~ L __t

V--

di= Zs

1>

,oo

Go

l-- - T

Q /

-

A ,

I

2

~

7.

o

__~t__s

,

2

~

I __

,

INITIAL CHARO[-. OF CAROON m (ql d. = 0 " ] 9 ram; Uo = 46 m m / s ; Pc = 720 kg carbon/rn ~ (a) U == 214 ram/s; (b) U ~ 300 m m / s ; (c) U == 383 mm/$

Fic. 6. Burn out time as a function of the initial charge of the char (batch process). (After Avedesian and Davidson. 35)

THE FLUIDISED COMBUSTION OF COAL where P is d e f i n e d as

N(d, d') =

f

d'

Pdd

(13)

d

w i t h N the n u m b e r of particles in the size range (d, d ' ) and a in Eq. 12 d e f i n i n g the b e d particle size distribution, given as

447

C a m p b e l l and D a v i d s o n tested their theory against experiments in an electrically heated tubular quartz reactor at 900~ and found good agreement b e t w e e n predicted a n d experimental results. Basu, Broughton and Elliott a7 reexamined the combustion m e c h a n i s m and set up experiments to test the conclusions about the primary product of the carbon oxidation u n d e r condi-

Om i

6/=

4~rShDc.da~ {C O - Q m , / 1 2 A [ U - ( U - U o ) e - X ] }

tions of fluidised combustion. Their model I is based on the a s s u m p t i o n that both 0 2 and CO 2 can reach the carbon surface. Model II is identical to that of Davidson. They p o i n t out that the difficulty of the Avedesian-Davidson model is that the two film theory (CO oxidising to CO 2 at a distance from the coal surface), i.e. the e n d o t h e r m i c reaction CO z + C ~ 2CO, requires 7.1.10 -4 KW heat flow to the surface of a particle of 3 m m dia. O n the other hand, the heat balance on the reacting particle with excess temperature over that of the surroundings shows that the particle is loosing more heat b y radiation than it gains b y conduction. It follows therefore that an excess temperature of the reacting particle surface can be m a i n t a i n e d only if the reaction is at least partly exothermic, i.e. O z is reaching the surface (Model I). The correct m o d e l l i n g of the c o m b u s t i o n process is important not only for the p r e d i c t i o n of the b u r n i n g time a n d particle size distribution in the b e d b u t also to enable calculations to be made of the species concentration distribution in the particulate phase w h i c h in turn is required for p r e d i c t i n g CO, NO x and SO x emission from the f l u i d i s e d bed. Song and Sarofim as have c o m p a r e d the simple Avedesian D a v i d s o n model with results of P.C. Wu 39 who s t u d i e d the kinetics of the carbon-CO 2 reaction. W u showed that the initial specific rate of reaction ro can be correlated with a L a n g m u i r type rate equation of the form ro =

K1Pco2

it was not possible to have reached the high rate of combustion w i t h o u t 0 2 reaching the surface directly. T h e y show also from values of r A v z / K 1 which they f o u n d to be of the order of 10~2-10 3 that it is u n l i k e l y that the C O z concentration fell to zero at the particle surface as assumed by Avedesian and Davidson in their c o m b u s t i o n model. It seems therefore that while their m o d e l has made a pioneering contribution to the calculation of c o m b u s t i o n of coal b u r n i n g rates in fluidised beds, a more sophisticated model is n e e d e d for the general representation of the c o m b u s t i o n kinetics. Song and Sarofim 38 a s s u m e d that 0 2 diffuses from the b u l k gas flow in the particulate phase to the carbon surface where it reacts to form CO. The rate of oxidation of the carbon monoxide with oxygen to form carbon dioxide is given b y Hottel et al.'s semi-empirical equation, 6s d C co -

-

-

3

x

101~

dt

exp - - RT

(~P H2O~

e,o~

e, c o

(15)

where Cp, H2O, C e.o2 and Ce,co are concentrations in the particulate phase averaging from particl e surface to effective radius. In their equations of mass transfer to the surface of the shrinking particle they take into account the gas phase reaction as:

(14)

1 + K2Pco + KaPco ~

--= Ot

T h e y have calculated the average rate from Avedesian and D a v i d s o n ' s experimental data a n d found it to be c o n s i d e r a b l y in excess of that reported b y Wu. 39 T h e y c o n c l u d e d that

D ---[_r2 Or

r2 Or - 0.5

and

k' CcoC~

448

ENERGY PRODUCTION FROM COAL

Cco oIl (r2 Cco)l Ot

k r2 Or

of bed and temperature. 4~ The SO 2 can react directly with limestone b y the exothermic reaction

Or

- k ' C c o C c o 2 o.a 1 C a C O a + SO2 + 2 0 2 = CaSO4 + C O 2

where the subscript p has been omitted for simplicity. By using some of Davidson's experimental data and the steady state solution of the above differential equations Song and Sarofim 38 have c o m p u t e d the concentration profiles a r o u n d the particles and the escape of CO from the bed. Figure 7 shows their results c o m p a r e d with experimental data reported b y Skinner. 6 Their combustion m o d e l has also the potential of being extended to the treatment of volatile combustion.

In case the temperature exceeds the e q u i l i b rium value ( d e p e n d e n t u p o n C O S concentration) for the e n d o t h e r m i c calcination reaction CaCO 3 = CaO + CO 2 it will then be lime that will react with SO2: 1 CaO + SO 2 + -2-02 = CaSO4 2 Calcination of limestone is favoured b y atmospheric fluidised b e d operation. It does not occur at elevated pressures. Calcination is instrumental to the formation of porous particles with large internal surface area. C a S O 4 on the other h a n d tends to block the pores and to form a skin on the particle impervious to further reaction. As has been p o i n t e d out b y Hsieh et al, 49 one of the primary p r o b l e m s of obtaining high sorbent utilisation is the fact that the reaction tends to be self-choking i.e.,

4. Pollutant Emission Control 4.1 SO 2 Removal by the Addition of Sorbent W h e n lime or limestone is a d d e d to a coal b u r n i n g fluidised b e d , the sulphur in the coal reacts with the a d d i t i v e and the SO 2 e m i s s i o n can be reduced to an extent d e p e n d e n t u p o n the feed rate, chemical nature and particle size of the sorbent a n d u p o n operational variables of the bed such as f l u i d i s i n g velocity, d e p t h

0.2

.Sglnner~t ;o

0.1

.

.

.

.

.

.

.

.

.

.

.

.

"'~..

ProposedModel

Ic :0 ~U

i

- tO

I

I

I

I

0

tO

20

3O

i

40

E , Exce F1c. 7. Comparison of predicted CO emission with experimental data. (After Song and Sarofim.38)

THE FLUIDISED COMBUSTION OF COAL lOC

~.~o~pp~'go7

9

9~

Dolomite

R e d u c t i o n 80 If)

sulphur emission: 70 %

Limestone/ "

the SO z retention b y sorbents such as limestone or dolomite. Figure 8 illustrates the potential of sulphur removal from a b e d in w h i c h coal with 3% S is b e i n g burned. 5~ The 700 and 100 p p m SO 2 marked in the graph are emission limits for sparsely and densely p o p u l a t e d areas respectively.

-- "E~i'EEio~4.2 NO~ Emissions from Fluidised Bed Combustors

6o

i

5~

'2

--3

Ca/S n;ol ratio

Fic. 8. Reduction in sulphur emission as a function of the Ca/S mol ratio for a coal with 3%S. (After Hoy and Roberts cited by Locke.5~ the reaction products block pore diffusion b y gaseous reactants into the b o d y of the stone. There seems to be agreement b y a n u m b e r of research workers r e g a r d i n g the following: The C a / S mole ratio in the feed stream, the combustion temperature and the mean gas residence time are the most important variables of SO z removal in the bed. At C a / S ratio near 4 almost 100% of SOz removal has been achieved. The o p t i m u m temperature d e p e n d s on the C a / S ratio. T h e o p t i m u m temperature range is 790~176 Outside this temperature range the SO 2 removal is reduced. SO 2 removal decreases with increasing f l u i d i s i n g velocity (at a rate of about 5% per 30 c m / s e c ) a n d increases slightly with b e d height (probably a contact time effect). There is some evidence that the c o m b u s t i o n reaction occurring in the b e d is s u p p o r t i n g

1 200 .

0 CENTRE A WALL

~

C o m p a r e d with the literature on SO 2 emission, relatively little information has been p u b l i s h e d on NO xemission from fluidised bed combustors. The m a i n publications include those of Jonke et al at the Argonne Res. Lab., 51 Robinson et a152 at Pope, Evans and Robbins, a n d Skopp and H a m m o n s 53 at Exxon Res. and Eng. More recently a n u m b e r of publications have resulted from the research at the University of Sheffield. ~4-57 T h e general conclusions can be given as: a) The emissions of NO x from fluidised bed combustion have the potential of meeting the existing standards for coal fired steam generating plant (<0.7 l b / 1 0 6 BTU). b) Under the operating conditions of the fluidised b e d c o m b u s t o r up to 90% of the NO x formed is " F u e l NO". c) NO is generated both during the volatile and the char combustion. The nitrogen left in the residual char decreases as the devolatilisation temperature rises. d) NO is formed preferentially near the coal inlet point where the oxygen concentration is high, irrespective of the input stoichiometry. (Fig. 9) NO emission is

~z

~

O CENTRE ~ WALL

1 200 --

I 000

I 00(]

800

800

>D

449

"600

600

~

; 804~

400

'~ 400 o.

EX.AIR ; 710 C

d

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200 --

200

:-~

"----BED

FREEBOARD

I

[

0"5

t'0

I

BED

H

.h N

I

1"5 3"0 HEIGHT ABOVE

0

0

:~"

FREEBOARD

I

l

I

0-5

1-0

15

DISTRIBUTOR,

pl

.IN

1 3'0

m

F[c. 9. Spatial distribution of NO measured in a 0.3 x 0.3 m fluidised combustor. (After Pereira, Gibbs and Be~r.~)

450

ENERGY PRODUCTION FROM COAL

therefore a weak function of input stoichiometry. In the temperature range between 660-850~ the NO emission increases linearly with temperature at an approximate rate of 2.3 ppm/~ (Fig. 10). e) NO is reduced by reaction with nitrogenous fragments in the volatiles and also by the heterogeneous reaction with char. 55 The latter reaction is strongly temperature dependent and becomes significant only above 800~ f) Staged combustion has a beneficial effect on NO x emission, 10% secondary air inject/on resulting in a 20% reduction in NOx emission. NO emissions from a 12" square combustor could be interpreted with good approximation by assuming that the "volatile NO" formation and its destruction takes place in the particulate phase outside the particle stagnant boundary layer. 55 (Fig. 11) For the further refinement of an NO~ emission model, more information is needed on the char-NO reaction and on the oxygen concentration profiles in the particulate and bubble phases, and on the dispersed coal particle concentrations in the bubble phase as affected by fluidised combustion operational variables.

2.2

REGRESSIONLINES

2.1

E 104=-5-016~'.007 T *.81/. N ~

2-0

5. Heat Transfer to Immersed Surfaces The temperature control of the bed can be achieved by cooling tubes immersed in the bed. For avoiding large local departures from the mean temperature of the bed it is essential that heat transfer surfaces should not obstruct solids circulation in the bed. Little difference is reported between heat transfer coefficients of vertically or horizontally arranged tubes, although there may be some advantage of the vertical tubular inserts in fluidised beds of low diameter to height ratio for the control of any slugging tendency. In general the heat transfer coefficient increases with the reduction of particle size and of particle density and with increasing fluidising velocity and bed temperature. Although peak heat transfer rates are somewhat lower than those in pulverised coal fired boilers, the overall heat transfer rates are considerably higher in fluidised beds because the heat transfer surfaces are more uniformly loaded. A further advantage is that a larger fraction of the total heat exchange surface can normally be immersed in the bed than can be exposed to the high temperature medium in other combustion systems. For example in fluidised combustors burning 3 m m - - 0 coal, heat transfer coefficients to water cooled tubes of approximately 554

N =I .3

1.0 i.TI~

1-8

,8

1.6 1.4 qr.=.

LU 1-2 1.0

-6

.i, EXCESS

KEY ~-S S-10n-,sls-2oz~-2szs.3o 3t.~s

.4

AIR

92

:OAL9 9 9 $ 9 < J 700

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O

Z~

[3

850

0

V

<~

p

T ,~

FIG. 10. NO emission as a function of bed temperature and input stoicbiometry. (After Pereira, Gibbs and Be~r.s6)

THE FLUIDISED COMBUSTION OF COAL

!

" \

I

\

i

\

COAL

N~

)

ij 9

~..~,.

!

9

\/,.o,.o.~,

/

~

" ~ (.CNAR.

COl

"

I

/

PARTICLEliOUNOAItY- LAYER

~'x

I i

451

/

",/ \\

C~R

PARTICLE

I \<..o,~ ,.o..o..o~/ !

- NH i ~

NM ~

N

NO-dV..vv~~ o

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* N l

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/

i

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Flc. 11. Physical model of the NO formation during the volatile evolution and combustion of a coal particle. A. the reaction of nitrogen bearing volatile matter in oxygen rich atmosphere outside the stagnant boundary layer of the particle. (Volatile matter ejected during pyrolysis). B. the reaction of nitrogenous compounds contained in the char, within the fuel rich stagnant boundary layer. (After Pereira and Be~r.~s)

452

ENERGY PRODUCTION FROM COAL W m - 2 K - 1 (80 B T U ft -~ h- 1oF - l) are o b t a i n e d with tubes 50 m m apart, rising further as the pitch to diameter ratio of the tube b a n k arrangement is increased. F i g u r e 12 represents b e d to immersed t u b e heat transfer coefficients as a function of particle size and for two temperature levels of the tube. 57 The variation of the heat transfer coefficient a r o u n d the circumference of an i m m e r s e d tube is s h o w n in Fig. 13. Also s h o w n are local particle circulation patterns close to the tube a n d the pressure field t h r o u g h the bed. 58 Heat transfer m e c h a n i s m s in fluidised beds are discussed a n d quantitative information is given in two c o m p r e h e n s i v e texts b y Zab r o d s k y 59 and Botterill 6~ respectively.

12~176 J

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"~..~.. 100,~ .o v,

80.

~E

60.

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;.

.

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Tube tern

0 0 0 4 0-006

0.01

Mean

0.02

paJticle size:

004

0"06

0.1

in.

6, C o m b u s t i o n E f f i c i e n c y

Flc. 12. Fluidised bed heat transfer coefficients. (After Locke et alY)

Conditions for the b u r n o u t of char in fluidised beds are favourable because of the good

280

270

-~ L~60

a

ZSO

240

230

3O

2O

I0

0

(ram)

Flc. 13. Variation in local heat transfer coefficient, h, fron an immersed horizontal tube. Also shown are the local particle circulation patterns close to the tube and the pressure field through the bed, (After Noaek.58)

THE FLUIDISED COMBUSTION OF COAL

I I

$

I

%C0 I| ~

-10BS$ MESH COAL -101,.'t0 BSS MESH COAL

-

- %OI

~-

I-J

453

~

~

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~r

I ',

.i

X~'101.30 BSS ESH CO t IS'I.E / & - ; O B S S MESH CHAR 'S".EA

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~1

' I/ - #

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!1 I OlS

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I I O'SO O?S I-0 HEIGHT ABOVE Ols/rllOOTOR (m)

I I-S

I I

0

FIG. 14. The effect of coal feed size and volatile content on the CO2, CO and 02 concentrations in the bed and the freeboard. (After Gibbs and Be~r.6~)

solids recirculation and long residence time. The largest loss of carbon from the bed is caused by the elutriation of combustible particles. Larger particles carried away by the fluidising gas and precipitated from the gas can be recycled into the bed, while small particles can be introduced into so called carbon burn up cells with lower fluidising velocities as recommended by Bishop 61 and reported by Robinson et al. 52 The amount of carbon burning in the free board above the bed is small, less than 5%, due to the relatively short residence time, low temperature and low oxygen concentration. Conditions for the combustion of volatile products in the bed are less favourable than those for solids because of the chance of unburned gas--mainly CO--escaping into the free board. The combustion of CO will go to completion as the bed temperature is raised to about 800~ and the excess air level reaches 20%. Because of the relatively poor lateral mixing in the free board a small amount of air (e.g. 10% of fluidising gas), or recirculated combustion products, used for the stirring of the free board by jets, can be effective in

reducing CO emission. The fraction of very fine coal particles in the feed coal has an important effect on the combustion efficiency. Fine particles will increase the "carry over" losses and also the CO emission from the bed. Figure 14 shows the CO2, CO and 0 2 concentrations measured in a 0.3 • 0.3 m experimental fluidised bed and in the free board for different coal feeds and for a bed tdmperature of 800~ and a fluidising velocity of 0.91 m/s, obtained by Gibbs and Be6r. 62 7. Pressurised Fluidised Combustion of Coal Because at present gas turbines need premium fuels, they are only used for peak load power generation. Previous attempts to use cheaper fuels such as coal or heavy fuel oils have been unsuccessful. In pulverised coal fired gas turbines the abrasive sintered ash caused serious erosion of the turbine blades and the alkali salts and SiO vapourised in the high temperature flame condensed on the blades, causing deposit formation and corro-

454

ENERGY PRODUCTION FROM COAL

sion. Ash from low temperature fluidised combustion is not sintered but soft and friable. The alkali salts are not vapourised to any substantial degree but are retained in the ash. There are further advantages of fluidised combustion at elevated pressure but there are also some problems that arise from the greatly increased volumetric heat release rates obtained in pressurised combustors. Although the rate of the boundary layer type diffusion which is assumed to control the rate of burning of carbon particles in the bed is unaffected by pressure, the interphase mass transfer rate is higher at elevated pressure because the size of the bubbles is smaller. 6~ (The exchange factor X (Eq. 7) is proportional to the bed height and inversely proportional to bubble volume.) Figure 15 illustrates the effect of pressure upon the heat release rate per unit distributor area. At 1500 KNm -2 (15 atm) the heat release rate increases by more than an order of magnitude compared with atmospheric pressure operation. The high values of heat release rates require more heat exchange surface to be immersed into the bed, as the heat transfer coefficient is virutally uneffected by pressure. As the

Cool

size

~ ~. j

J

6mm-O

3rnrn - 0

I

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~1 o =i ,ua!u. Oi

o,

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2

3

Velocity: m/sFIc. 15. Area of bed per unit power output vs. fluidising velocity and pressure. (After Roberts et al.64)

economy of pressurised operation is not as sensitive to the pressure drop across the bed as are atmospheric fluidised combustion processes, the bed height can be increased to accommodate the extra heat exchangers. However, the problem remains of absorbing sufficient heat from bed areas close to the distributor plate without obstructing solids recirculation and thus avoiding hot spots in the bed. Because of the smaller bubble size and better interphase transfer, pressurised operation improves combustion kinetics in the bed. Material withdrawn from the pressurised combustor contains typically less than 0.1% carbon. As the particles become smaller in size and their "effective" boundary layer diminishes, their burnout process can become kinetically controlled which in turn implies that the rate of combustion becomes dependent upon pressure and temperature. At a bed temperature of 900~ at 500-600 KNm-2 (5-6 atm) pressure with 2 sec residence time in the bed and 15-20% excess air, BCURA reported combustion efficiencies in excess o f 99%. 64 The variation in sulphur retention with Ca/S ratio and with temperature at elevated and atmospheric pressure is shown for dolomite and limestone in Fig. 16. With limestone, pressurised combustion gives lower sulphur retention. Limestone does not calcine at 800~ and sulphation is restricted to the outer layers of the limestone particle at higher partial pressures of CO 2. In the case of dolomite the calcination of the magnesium carbonate produces a large internal reaction surface. In addition the better gas/solid contacting at high pressure brings some improvement to the performance of dolomite at elevated pressures. The effect of temperature (Fig. 16b) is different from that at atmospheric operation. With the temperatures taken to the limit of slagging free operation no reduction in sulphur retention was found. This is important because gas turbine cycles require the bed to be operated at this maximum temperature. Results of experiments at Exxon, ea BCURA 64 and ANL 67 are in agreement about significant reductions in N O emission as the total pressure in the fluidised combustor is increased. Figure 17 shows the effect of pressure upon the NO concentration of the flue gas as determined by Vogel et al ~7 in ANL's bench scale pressurised fluidised combustor. NO is preferentially formed near the bottom of the bed where the 0 2 concentration is high. The lower NO emission for pressurised operation may be ascribed to a reduction of the formed NO through the bed by reactions with char and

THE FLUIDISED COMBUSTION OF COAL

I00

,

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I ~

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2.5

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oil

of Ca/S

m:~| r a t i , ~

on sulphur

retention

850

900

950

Temperature (~ Effect.

of temperat.re on sulphur r e l r t , t ion

Fie. 16. The effect of the Ca/S tool ratio and of temperature upon sulphur retention in atmospheric and pressurised fluidised beds. (After Roberts et al.64)

volatile products made more effective by the improved interphase mass transfer.

8. The State of Practical Application

There have been a number of pilot plant flui'dised combustors in operation at various research centers for some time now (Fig. 1, Table I) but the development has been taken -to a new stage, nearer industrial application, by the construction of two demonstrationplant-size boilers. The first of these is a water tube boiler at Babcock and Wilcox Ltd. in Renfrew Scotland, an existing boiler has been originally equipped with a spreader stoker which was replaced by a fluidised bed approximately 3m x 3m in size using CSL technology (Fig. 18). The converted boiler in operation

since June 1975. It burns coal crushed to 95% smaller than 3 mm. Under the initial operating conditions the fluidising velocity is limited to 4 ft/s and the boiler generates about 23,000 l b / h r of steam. The fluidised bed boiler is designed to enable the fluidising velocity to be increased up to 8 ft/s with the corresponding increase in steam generating capacity. ~8 The other demonstration plant was built by the Foster Wheeler Co. using Pope, Evans and Robbins technology. This boiler which has a capacity of 30 MW (e), is located at the Rivesville, West Virginia plant of the Monangahela Power Co., and is to be commissioned in the near future. (Fig. 19) The principles of design of the Rivesville boiler are discussed in a recent paper by Ehrlich. 66 The development of pressurised fluidised combustors will be greatly assisted by a new

456

ENERGY PRODUCTION FROM COAL

240O 2200

,,,,

,

,

1800 lEO0 NOX-I SERIES: ARKWRIGHT COAL ALUNOUM BED 1550 "F flED TEMP 3-5 fl/ser GAS VELOCITY

1400 1200 I000

NOX-,~ SERIES:

800

ARK~,TRIGHT COAL TYMOCHTEE COLOMITE CoJS RATIO ' 1.5 1550 ~ BED TEMP 3.S fl/sr162 GAS VELOCITY

GO0

400

~.x

~

~'~'~ ~. ~~, ~ = _

2OO

_ _ _ . f NOX- I/ SERIES ~'NOX'[ SERIES

%

i

l

I

i

I

I

15 30 45 60 90 120 PRESSURE, psio

FIe. 17. The effect of pressure on NO emission from a fluidised combustor. (After Vogel et al.67)

experimental plant to be built in England under the aegis of the International Energy Agency with the participation of Britain, West Germany and the USA and operated by the National Coal Board. These new experimental and demonstration plants hold out the promise of accelerated development and the long awaited industrial application of atmospheric and pressurised fluidised combustors. 9. General Conclusions Research Needs Fluidised combustion is the most promising new method of direct coal utilisation for steam raising, for process heat generation or for power generation in gas turbine cycles. The special advantages include high combustion efficiencies and high heat absorption rates at low combustion temperatures, effective removal of SO 2 at low capital and operating costs, reduced NO xemission, and the possibility of passing the combustion products from low grade fuels through gas turbines. The development of fluidised combustion

FIc. 18. The BCSL Rcnfrew fluidised combustion boiler. (After Locke et al.sT)

THE FLUIDISED COMBUSTION OF COAL

457

and burn as a diffusion flame at the bubble boundary? b) What are the processes that control particle elutriation from the bed? c) What proportion of the volatiles and particles burn in the freeboard? C. on the emission o f pollutants a) What is the role of the SO 2 diffusion through the particulate phase to the sorbent surface and to the bubble respectively and the role of the heterogeneous reaction at the surface of the sorbent particle in the retention of sulphur in the bed? b) What are the mechanisms of alkali and trace metal retention by limestone and dolomite in the bed? c) What are the detailed mechansims of fuel nitrogen evolution, oxidation and of the NO reduction by nitrogeneous fragments and by char?

Ftc. 19. The 30MW (e) multicell fluidised-bed. Rivesville, W. Va. (After Ehrlich. 66)

has reached the stage of industrial application and there is now an urgent need to obtain answers to some of the questions arising from design requirements and also to generalise information in terms of mathematical models that give representative, quantitative description of the processes of combustion, heat transfer and pollutant emission in fluidised beds. Some of these questions are: A. on interphase m a s s transfer: a) How does bubble size, frequency of bubble shedding, bubble growth and coalescence depend upon distributor type, fluidising velocity, bed and immersed tube geometry, and total pressure? What is the role of bubbles in oxygen bypassing? b) What is the role of bubbles in particle mixing? What is the relative magnitude of particle mixing time and char burn-out time? This ratio may be important for determining the spacing of coal feed points. B. on the kinetics o f combustion: a) What are the processes that govern the escape of the volatile matter? If particle mixing time is commensurable with volatile evolution time, does the volatile matter burn in the vicinity of the particle or does it react with the bubble oxygen

These questions together with many others concerning the engineering development of fluidised combustors, such as start up, load following, deposit on turbine blades, corrosion of superheater tubes immersed in the bed etc., will have to be studied in closely coordinated programmes if we wish to see the goal of industrial application of fluidised combustion accomplished in the 1980"s.

Nomenclature

Ar

bed cross sectional area constant in Eq. 12. a inlet concentration Co Ce concentration in the particulate phase particle dia. d De diameter of sphere that has the same volume as the bubble molecular diffusion coefficient Dc effective diffusion coefficient E Fg mass fraction of 0 2 in main gas stream (Eq. 8) bed height H initial bed height (at incipient Ho fluidisation) K 1 , K 2 , K ~ Langmuir equation constants (Eq. 14) pseudo first order reaction rate k constant mass transfer coefficient kc mean molecular weight of coal M volatiles mass of particles charged to the m bed (batch process)

458

ENERGY PRODUCTION FROM COAL

N n

P

Q

R

ro rnvg t

T U

Uo

CA V X E

0 Pc

n u m b e r of particles molar flux of oxygen to the particle partial pressure effective volumetric interchange b e t w e e n b u b b l e and bed radius of reaction zone around particle; outer radius of spherical a n n u l u s occupied by volatiles initial reaction rate of carbon with CO~ average reaction rate Of carbon with C O 2 time; t b = b u r n i n g time temperature superficial gas velocity to the bed superficial gas velocity at incipient fluidisation absolute b u b b l e rise velocity b u b b l e volume transfer factor = Q H / U A V particulate phase voidage fraction density carbon particle density

REFERENCES 1. WINKLER, F.: U.S. Patent No. 1.687118 (Oct. 1923) and DRP437.970 (1922). 2. SZIKLA,G. AND ROZINEr, A.: Feuerungstechnik 4, 1938. 3. STRATrON,J. F.: Power, 68, 486, Sept. 1928. 4. GODEL, A. A.: Revue General de Thermique 1966, 5(52) 349-359. 5. NOVOTNY,P.: Sb. Prednasek 50 Vyroci Ustavu Vysk. Vyuziti Paliv, 104-11 (1950). 6. SKINNER,D. J.: Fluidised Combustion of Coal, 1971, Mills and Bonn Monograph CE/3. 7. COUTANT, R. W., McNULTV, J. S., BARRETT, R. E., CARSON, J.

8. 9.

10. 11.

12.

13.

J., FISHER, R. AND LOUCHER, E.

H.: NAPCA Aug. 1968. P.B. 179907. KUNII,D. ANDLEVENSPIEL,O.: Fluidisation Engineering, John Wiley and Sons Inc., 1969. ZARRODSKY, S. S.: Hydrodynamics and Heat Transfer in Fluidised Beds. Translation, MIT Press, 1966. BOTTERILL,J. S. M.: Fluid Bed Heat Transfer, Academic Press, 1975. Proc. First Intern. Conf. on Fluidised Bed Combustion NAPCA Hueston Woods State Park, Oxford, Ohio, Nov. 1968. ARCHER,D. H., ET AL.: Evaluation of the Fluidised Bed Combustion Process Volumes I, II, and III. A report to EPA by Westinghouse Res. Labs. (1971). KEAIRNS,D. L. ETAL.:Evaluation of the Fluidised Bed Combustion Process. Volumes I, II, and III. Ibid. (1973).

14. Proc.2ndlnternationalConferenceonFluidised Bed Combustion, EPA, Hueston Woods, Ohio, Oct. 1970. 15. Proc. 3rd International Conference on Fluidised Bed Combustion, EPA-650/2/73-053. 16. Proc. Fluidised Combustion Conference, Inst. Fuel London, Sept. 1975. 17. DAVIDSON, J. F. AND HARRISON, D.: Fluidised Particles, Cambridge Univ. Press, 1973. 18. PARTRIDGE, B. A. AND ROWE, P. N.: Trans. Inst. Chem. Eng., London, 44, T-335 (1966). 19. SZ~r~LV,J.: Trans. Inst. Chem. Eng., London, p. 197 (1962). 20. CniaA, T., AND KOBAYASHI, H.: Chem. Eng. Sci. 25 1375 (1970). 21. DAVIES, L. AND RICHARDSON, J. F.: Trans. Inst. Chem. Eng., London, 44, T-293 (1966). 22. PEREIRA, J. R, AND CALDERBANK, P. H.: Proc. Fluidised Combustion Conference, Inst. Fuel, London, Paper B.2, Sept. 1975. 23. YOSHIDA,K." Rapporteur's report to session B., Ibid. 24. BADZIOCH,S.: BCURA, Mon. Bull. 25, 285-301 (196i). 25. IOMBER,G. M. AND GRAY, M. D.: Combustion and Flame, 1967. 26. FIELD, M. A., GILL, D. W., MORGAN,B. B. AND HAWRSLEY,P. G. W.: "Combustion of Pulverised Coal", BCURA (1967). 27. PEREI~, J. F.: Ph.D. Thesis, University of Sheffield, England, 1975. 28. NUSSELT,W.: Z. Ver. Dt. Ing., 68, 124-128 (1924). 29. DAVIES,H. ANO HOTTEL, I-I. C.: Ind. Eng. Chem. 26, 889/892 (1934). 30. PARKER,A. S. ANDHOTTEL, H. C.: Ind. Eng. Chem. 28, 1334/1341 (1936). 31. HOTTEL, H. C. AND STEWART,I. McC.: Ind. Eng. Chem., 32 719 (1940). 32. ESSENHIGH, R. H. AND PERRY, M. G.: Inst. Fuel Conference; Science in the Use of Coal, Paper D.1. (1958). 33. BEtR, J. M. AND ESSENmGH,R. H.: Nature 187, No. 4743 pp. 1106-1107 (1960). 34. BEI~R,J. M., LEE, K. B., MARSDEN,C. ANDTHRING, M. W.: Journees de la Combustion et de la Conversion de L'Energie, Paris, May, 1964. 35. AVEDESIAN, M. M. AND DAVIDSON, J. F.: Trans. Inst. Chem. Eng., London, 51, 121 (1973). 36. CAraPBELL,E. K. ANDDAVIDSON,J. F.: Fluidised Combustion Conference, Paper A2, London, Sept. 1975. 37. BAsu,P., BROUGnTON,J. ANDELLIOTT,D. E.: Ibid., Paper A3. 38. SONG,Y. H. AND SAROVXM,A. F.: Fluidised Bed Head Start Program, Phase 1. M.I.T. Cambridge, Mass., Sept. 1975. 39. Wu, P. C.: Sc. D. Thesis, M.I.T. (1949). 40. K1ANG,K. D., NACK,H., OXLAY,J. H. AND REID, W. T.: 1975 Lignite Symposium, Energy Res.

THE FLUIDISED COMBUSTION OF COAL Center, Grand Forks, N. Carolina, May 1975. 41. HARRINGTON,R. E., BORGWARDT,R. H., ANDPOTTEa, A. E.: Amer. Ind. Hyg. Assn. J., 29, (1968). 42. FALKENBERBY,H. L., AND SLACK, A. V.: Chem. Eng. Progr. 65 (12), 61(1969). 43. COUTANT, a. W., BARRETT, R. E., AND LOUGHEB,

44. 45.

46.

47.

48. 49.

50. 51. 52.

53. 54.

E. H.: "SO 2 Pickup by Limestone and Dolomite Particles in Flue Gas", American Society of Mechanical Engineers Preprint No. 69WA/APC-1 (1969). POTTER,A. E.: Amer. Ceram. Soc. Bull., 48(9), 855 (1969). ATTIC,R. C., ANDSEOOR, PAUL:"Additive Injection for Sulfur Dioxide Control", Babcock and Wilcox Co. Research Center Report 5460, Research Center, Alliance, Ohio, 1970. DAVIDSON,D. C., SMALL,A. W.: Second International Conference on Fluidised Bed Combustion, Hueston Woods, Ohio, October 4-7, 1970. HARVEY,RICHARDD.: "Petrographic and Mineralogical Characteristics of Carbonate Rocks Related to Sorption of Sulfur Oxides in Flue Gases", Interim Report to the National Air Pollution Control Administration, Contract Number CPA 22-69-65, june 22, 1970. BOBGWARDT,R. H., AND HARVEY, R. D.: Environ. Sci. Technol., 6 (4), 350 (1972). O'NEmL, E. P., I~AmNS, D. L., ANDIQVrLE, W. F.: "Kinetic Studies Related to the Use of Limestone and Dolomite as Sulfur Removal Agents in Fuel Processing", Third International Conference on Fluidised Bed Combustion, Hueston Woods, Ohio, November 1972. LOCKE,H.B.: Achema, Frankfurt 1973, Dechema Monograph. Vol. 73. JONKE,A. A. ET AL.: Argonne Natl. Lab., Publ. No. ANL/ES-CEN-1001 (1969). ROBINSON, E. B., GLENN, R. D., EHRLICH, S., BISHOP,J. W., AND GORDON, J. S.: EPA Contract CPA 70-10, February 1972, PB 210 828. SKOPP,A. ANDHAMMONS,G.: ASME Winter Ann. Meeting, Nov./Dec. 1971. PEREIRA,J. F., BE~R, J. M., GIBBS, B. M. AND HEDLEY, A. B.: Fifteenth Symposium (Interna-

55.

56.

57.

58. 59. 60. 61. 62.

459

tional) on Combustion, p. 1149, The Combustion Institute, 1974. PEBEIRA,F. J. ANnBE~B,J. M.: Second European Symposium on Combusion, p. 339, Orleans, France, 1975. PEREIRA,F. J., GIBBS, B. M. AND BEI~R, J. M.: Combustion Institute Central States Section Meeting, Apr. 1975. LOCKE, H. B., LUNN, H. G., HoY, H. R. AND ROBERTS,A. G.: Fourth International Conference in Fluidised Bed Combustion, MITRE Corp., McLean, Virginia, Dec. 1975. NOACK,R.: Chemie Ing. Technik 42, 371 (1970). ZABBOI~SKY, S. S.: Hydrodynamics and Heat Transfer in Fluidised Beds, M.I.T. Press (1966). BOTTERILL,J. S. M.: Fluid Bed Heat Transfer, Academic Press (1975). BisHoP, J. W.: U.S. Patent No. 3,508,506 (Apr. 28, 1970). GIBBS, B. M. ANn BEER, J. M.: Symposium on High Temperature Reaction Engineering. Inst. Chem. Eng. and VDI. June 1975; I. Chem. E. Symposium Series No. 43, Paper 23.

63. HOKE, R. C. AND BERTRAND, R. R.: Fluidised Combustion Conference, Inst. Fuel, London, Paper D5, Sept. 1975. 64. ROBERTS,A. G., STANTAN,J. E., WILKENS, D. M., BEACHAM,B. AND HOY, H. R.: Institute of Fuel Series No. 1, Fluidised Combustion London, Sept., 1975, Paper D4. 65. GUEDESDE CABVALHO,J. R. F. ANDHARRISON,D.: Ibid. Paper No. 31. 66. EHRLICH,S. H.: Ibid. Paper C4. 67. VOGEL, G. I., Swirl W. M., MONTAGNA,J. C., LENC, J. F. ANDJONKE, A. h.: Fluidised Combustion Conference, Inst. Fuel, London, Paper D3, Sept. 1975. 68. HOTTEL, H. C., WILLIAMS, G. C., NERNHEIM, N. M. ANDSCHNEmER:Tenth Symposium (International) on Combustion, p. 111, The Combustion Institute, 1965. 69. THUBLOW,G. G.: Rapporteur's Report. ~luidised Combustion Conference, Inst. Fuel, London, Sept. 1975.

COMMENTS I. Mahawili, E. I. du Pont de Nemours, USA. Could you please comment on the possible backmixing in a fluidized bed and how it affects plug flow kinetics modeling and residence time calculations. Also, is the time function of local in-bed concentration measurement significant? Author's Reply. When the dense phase solids

descend faster in a bubbling bed than the gas can percolate through the solids, the flow will be reversed through the dense phase. This gas recirculation occurs when the fluidizing velocity is high ( U o / U,~f-> 6 to 11),1and its effects upon the kinetics of reactions in the bed can be significant. For example, it was found that the NO emission from the FBC will reach a peak as the fluidizing velocity

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ENERGY P R O D U C T I O N FROM COAL

is increased and then decrease as the velocity is further increased, z This was interpreted as b e i n g due to the reduction of NO by char as the high NO concentration gas was recirculated through the dense phase. Under these condictions plug flow kinetics are clearly not adequate for modeling the NO emission from the FBC. Time resolved concentration measurements made in a coal b u r n i n g fluidized bed a showed that there are significant O z concentration variations at a point in the b e d - - t h e frequency of the variations being equal to that of bubbles passing at the point of sampling. Such variations will have to be taken into account in combustion modeling, particularly when calculating CO or NO emission from the FBC.

REFERENCES 1. KUNII, DAVID, AND L~.VE.~SPmL, O.: Fluidization Engineering, John Wiley & Sons, 1969. 2. PEREIRA,F. J. ANn BE~r~, J. M.: NO Formation

from Coal Combustion in a Small Experimental Fluidized Bed, 2 n d European Symposium on Combustion, Orleans, France, pp. 339-345, Sept. 1975. 3. GIBbs, B. M., PE~EIrU, F. J., A~D BE~R, J. M.:

Coal Combustion and NO Formation in an Experimental Fluidized Bed, Institute of Fuel Symposium Series No. 1., Fluidized Combustion Paper No. D6 1975.

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B. L. Vollerin, Battelle Geneva Research Center, Switzerland. For utility boilers, do you think that the optimum operation point of such fluidized beds will result in an uneconomical pressure requirement for the f l u i d i z a t i o n / c o m b u s t i o n air?

Author's Reply. It seems that economic operation of atmospheric pressure FBC's will limit the bed height to about 120-150 cm. These bed heights, however, permit already high heat release rates per unit distributor area to be obtained (>1 M W t h / m 2 ) , as well as advantages such as simple and inexpensive coal preparation compared with that for pulverized coal combustion. Also, the advantage of a reduction of required heat exchange surface outweighs the disadvantage due to the higher pressure drop across the fluidized bed. At elevated pressure, these advantages tilt the balance even more in favor of fluidized combustion.

Sridhar K. I~la, Union Carbide Corp., USA. I would like to mention that fluidized bed combustion of coal can be carried out successfully in the temperature range 1000~176 This requires a careful control of the bed temperature and the fluidization velocity to avoid defluidization. A major advantage of operation in this temperature range is that the ash released from the b u r n i n g coal particles can be accreted by the bed particles, thus minimizing fly ash emission from the bed. Such a combustor is called an ash agglomerating fluidized bed combustor and finds application in coal gasification technology.

Bimal K. Biswas, Foster Wheeler Energy Corp., USA. It is interesting to note that the author achieved NO~ reduction substantially above the bed from the already produced NO x within the bed, due to a heterogeneous reaction of carbon and NO~ at around 800~ In our own experience, we obtained similar results when b u r n i n g "coke breeze" and other carbonaceous materials in a rather large fluid bed with e m b e d d e d tube bundles (boiler type). However, onset of the heterogeneous reaction seemed to vary from 750~ to 850~ We do not know if these variations are due to different types of carbons that are produced after devolatilization, or the effective contribution of volatile and carbons on NO x followed by the reduction of NO xby a heterogeneous type of reaction with carbon. I wonder if Dr. Be6r might have additional comments on the subject of heterogeneous NOx and carbon reaction in the fluid bed.

Author's Reply. Batch type experiments in w h i c h a 3" dia. coal ash-fused alumina bed was fluidized with gas mixtures containing NO showed that a batch of char particles dropped into the bed can reduce significantly the NO the reduction being dependent upon bed temperature and char particle size. There was evidence found that in the absence of oxygen and at temperatures above 800~ the carbon is gasified by NO with reaction products of CO and CO~ in the exit gas. In the presence of 0 2 the NO-char reaction continued to reduce the overall emission of NO but to a reduced extent?

REFERENCES 1. B. M. GIBBS, F. J. PEREI~, ANn J. M. BEEa: This Symposium.