~ Pergamon Energy Convers. Mgmt Vol. 35, No. 12, pp. 1049-1060, 1994

Copyright © 1994 Elsevier Science Ltd 0196-8904(94)00031-X Printed in Great Britain. All rights reserved

0196-8904/94 $7.00 + 0.00

A MATHEMATICAL MODEL FOR EXIT GAS COMPOSITION IN A 10 MW FLUIDIZED BED COAL COMBUSTION

POWER PLANT

G. V. REDDY and S. K. MOHAPATRA Department of Fuel & Mineral Engineering, Indian School of Mines, Dhanbad-826 004, India

(Receh~ed 25 August 1993; received for publication 26 July 1994)

Abst rae t - -A mathematical model has been developed for the oxygen mass balance for a 10 M W fluidized bed coal combustion power plant operated at Jamadoba (TISCO, India). Assuming the three phase theory of fluidization, the fluid bed is considered to consist of a number of equivalent stages in series. Within each stage, an exchange of gas takes place between the bubble, cloud-wake, and emulsion phases. An effective chemical reaction rate of char combustion has been derived considering the single film theory of char combustion for shrinking particles. The model has been used to predict the consumption of oxygen in the fluidized bed combustor, the outlet gas composition, variation of oxygen concentrations in different phases and also the variation of average oxygen concentration along the bed height. Model predictions were compared with plant data, and reasonable accuracy was obtained.

Exit gas composition Three phase model Fluidized bed combustion Coal combustion process

N O M E N C L A T U R E

A~ = Cross-sectional area of bed (cm 2) A 0 = Distributor area per orifice (cm 2) C O = Initial oxygen concentration (g-mol/cm 3)

C~,, = Oxygen concentration at stage n in bubble phase C¢~,, = Oxygen concentration in cloud-wake phase at stage n

Cp = Oxygen concentration consumed in bed in combustion (g-mol/cm 3) C~ = Oxygen concentration at surface of burning char (g-mol/cm 3) Cs = Carbon concentration per unit dense phase (g-mol/cm 3) D b = Bubble diameter (cm)

Dbo = Initial bubble diameter (cm) Obmax = Fictitious maximum bubble diameter (cm)

D~ = Gas phase diffusivity of oxygen (cm2/s) D R = Diameter of bed d o = Average char particle diameter (cm)

dp~ = Average diameter of bed material (cm) fc~ = Ratio of cloud-wake volume to bubble volume

HLF = Expanded bed height (cm) //mr = Bed height at minimum fluidisation (cm)

K K~ Ko

(K~)b (K~)b

ND P~ Sh T~ V~ Uo

U~r ~'¢oal

XH, XW, XN

Greek letters

Eb Emf Pp Ps

= Overall chemical reaction rate constant (s -~) = Reaction rate in cloud-wake phase (s-I) = Reaction rate in emulsion phase (s ~) = Gas interchange coefficient between bubble and cloud (s L) = Gas interchange coefficient between cloud and emulsion (s -I) = Number of holes in distributor plate = Average pressure in combustor (atm) = Sherwood number = Bed temperature (K) = Surface temperature of char particle (K) = Superficial gas velocity (cm/s) = Minimum fluidisation velocity (cm/s) = Coal feed rate (gm/s) = Ultimate hydrogen, moisture, and nitrogen of feed

= Bubble fraction of bed = Mean voidage under minimum fluidisation conditions = Density of coal (g/cm 3) = Density of fluidising gas (g/cm 3) = Viscosity of gas

1049

1050 REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION

INTRODUCTION

Fluidized bed coal combustors are usually modelled as multiphase systems consisting of two or three distinct phases. According to the two phase theory of fluidization proposed by Davidson and Harrison [1], a gas fluidized bed is considered to be composed of two phases, a dense or emulsion phase consisting of solid particles and interstitial gas, and a dilute or bubble phase consisting of rising voids, essentially free from particles. It also assumes that all the gas in excess of the minimum fluidization flow rate passes through the bed as bubbles. On the other hand, the three phase theory, as proposed by Kunii and Levenspiel [2], assumes an additional phase consisting of a cloud-wake region. Bulk flow of gas through the emulsion and cloud-wake phases is assumed to be negligibly small.

There exist a number of mathematical models to predict the performance of fluidized bed coal combustor. The Avedesian and Davidson [3], Campbell and Davidson [4], Becket et al. [5], Gibbs [6], Bukur and Amundson [7] and De Souza-Santos [8] models have been based on two phase theory. These models differ substantially in the assumptions they have made regarding the flow behaviour of each phase, the extent of gas mixing and the mode of interphase gas transfer.

Some of the models have been developed in order to examine the relative importance of bed operating conditions and the model parameters, e.g. Avedesian and Davidson [3], Campbell and Davidson [4], and Becker et al., [5] while others have attempted to simulate the performance of small scale experimental units.

The Kunii and Levenspiel [2], Chen and Saxena [9] and Gordon and Amundson [10] models have been based on the three phase theory of fluidization. The majority of the models, whether they are based on two phase or three phase theory of fluidisation, have been validated from laboratory scale units.

Due to the complicated nature of a fluidized bed combustor, a model developed to describe a specific system may not represent well other systems differing in design and operating conditions. Therefore, more and more models should be introduced to represent large scale units in order to overcome the deficiencies of the existing models. In the present work, a three phase model has been presented, and it has been validated by the data collected from a small scale commercial plant which uses coal washery rejects.

BASIC ASSUMPTIONS OF THE MODEL

(1) The FBC is assumed to consist of a number of equivalent stages in series, similar to the approach suggested by E1-Halwagi and EI-Rifai [11] for fluid bed catalytic reactors. The height of each stage is equal to the average equivalent bubble diameter in the bed.

The following additional assumptions relevant to the FBC are made:

(2) It is assumed that the bubbles are carbon free, uniform in size and equally distributed throughout the bed. The gas flowing through the bubble phase is considered to be in plug flow.

(3) The reaction is isothermal, first order and does not involve changes in the number of moles.

FORMULATION OF THE MODEL

The oxygen that is present in the bubble cannot take part in combustion. The bubble phase oxygen is transported across the bubble cloud interface into the more carbon rich cloud-wake phase, where the combustion reaction takes place. Unreacted oxygen in the cloud-wake is further transported to the very carbon rich emulsion phase. Most of the oxygen is consumed in the emulsion phase. The net result of these events is that the observed combustion rate depends upon the transfer coefficients and chemical reaction rate constant.

Oxygen balance around stage 'n' for bubble phase

Oxygen in by ] Oxygen out by Oxygen transfer convection ] - convection - to cloud-wake = 0.

REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION 1051

It is represented symbolically as;

fjn Ub Cb. _, -- Ub Cb, -- (Kb~)bEb (Cb -- Cc,~ ) dZ = 0. (1) n I

Oxygen balance around stage 'n' for cloud-wake phase

IOxy en in IOxygen coming rom IOxyg n go'ng rom I convecnon + bubble phase by transfer - wake to emulsion by transfer

I Oxygen out from I Oxygen consumed - cloud-wake by convection + in cloud-wake phase = 0.

Symbolically, it can be represented as;

U¢wCcw._,+(Kbc)beb (Cb-Ccw.)dZ '4-K¢eCe. - (Ucw+Kcw+KcE)Ccw. = 0 (2) n t

where

K~ = (Kce)bEbAZ (3)

Kcw = (K) (f~w) (eb) (AZ) (4)

A Z = Z n - z " _ , = O b. (5)

Oxygen balance around stage 'n' for emulsion phase

I Oxygen in by [ Oxygen out by I Oxygen transfer from = [ Oxygen consumed by [ convection - convection + cloud-wake phase combustion reaction

which can be represented symbolically as:

U m f C e n - I "~ (Kc~)(Ccw. ) - (Umf + K~ + K~e)Ce. = 0 (6) where

Ke = K[I - Eb(l +f~w)]AZ (7)

Kce = (Kee)bEbAZ. (8)

Now the oxygen balance over a differential element of height (dZ) in the bubble phase is written as:

UbCb -- Ub(Cb + dCb) - (K~)bEb(Cb -- Cow, ) dZ = 0. (9)

Rearranging and integrating the above equation, we can get

Cb - Cow, = (Cb, ,-- Ccw,)exp[{--(Kb~)bEb(Z -- Z, I)}/Ub]. (10)

At the bottom of the bed (n = 0), the concentration of oxygen fed to each phase is the same as that of the incoming feed oxygen. Hence, the boundary conditions for solving the above equations are:

at n=O, C b n - ~ - C c w n ~ - C e n = C o . (11)

Equations (1)-(10), together with the boundary conditions, make up a complete mathematical description of the system.

E S T I M A T I O N O F M O D E L P A R A M E T E R S

In order to apply the proposed model in the calculation, it is necessary to present mathematical equations for the estimation of the various hydrodynamic parameters in the model.

The average equivalent bubble diameter: Db

Various correlations can be found in the literature for the estimation of bubble diameter in a fluidized bed. One of the widely used correlations was proposed by Mori and Wen [12], taking into account the effect of bed diameter and distributor type on bubble diameter, which is given below.

Ob = Dbm -- (Obm - - Db0)exp( -- 0.15/O R ) (12)

1052 REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION

where

and

Dbm = 1.6377[At(U0- Umr)] °4

Dbo = 0.8716[A t (Uo - - U m f ) / N D ] 0"4

(13)

(14)

(for perforated distributor plates with number of orifice openings = No) and

Db0 = 0.376(U0- Umf) 2 for porous distributor plates. (15)

However, the Mori and Wen [12] correlation, so far, has been used for laboratory type FBC having porous or perforated distributors where the bubble growth was restricted by the walls of the combustor. However, Rowe [13] presented an expression to estimate the bubble diameter for industrial scale combustors, where such restrictions are not expected due to the large cross-section of such combustors, which is given below;

Ob = (U0 - U m f ) 0"5" Z°75 "g-°25. (16)

Stubington et al. [14] modified the above expression for tuyer cap type distributors as follows:

Db = 0.43(U0 - Umr)°"(Z + 4x/~0)°'Sg-0.2. (17)

The distributor plate used in the 10 MW FBC power plant at Jamadoba is of the tuyer cap type and is made of carbon steel as shown in Fig. 1. Equation (17) has been used for calculating D b. Since the gas enters the bed horizontally with this type of distributor plate, weeping and damping of solids through the distributor are much reduced.

Volume ratio of cloud-wake phase to bubble phase: fcw

Rowe and Partridge [15] studied the behaviour of bubbles in a fluid bed by an X-ray method and found that the size of wake (ratio of wake volume to volume of bubble, fw) averages one-quarter of the total sphere volume and tends to increase as the particle size decreases. Hence, the value offw was taken to be 0.25.

12 holes T

3.5 mm 0

T 140 mm

I00 mm 1

~ ~ 5 mm

42 mm

Fig. 1. Schematic representation of the distributor plate.

REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION

Table 1. Hydrodynamic parameters

1053

Parameters Empirical correlation (all are in C.G.S. unit) Reference

Minimum fluidising velocity (Umf) Gas viscosity (gg) Gas density (pg) Rise velocity of a

crowd of bubbles (u b) Rise velocity of an

isolated bubble (ubr) Gas velocity in bubble

phase (Ub) Expanded bed height

(HLF) Total bed weight (W b) Number of equivalent

stages (N,t~e) Gas interchange coefficients

(K~) and (K.)b

/ , , \ Umf = 1SC~][{(33.7) 2 + 0.0408d~, .pg(pp- p,)g/#~},/2 33.71

\appg} #g = 1.4(10-s) • (Tb) °-5

p~ = 353.2(10-3)/Tb

ub = Uo - Umr + Ubr

Ubr = 0.711x~-' D b

U b = (U 0 - Umf)/(1.0 +few' (~mf)

HL F = Hmf/(l.0 - ~b)

Wb = Hmf' At(l - Cmf) " Pp

Nstag e = HLF/D b (K~:)b = 4.5(Umr/Db) + 5.85(D~/2" gt/4/Dg/4)

(K~) b = 6.781%r" Dg" Ub/D~,] I/2

Wen and Yu [231

Bird et al. [24] Bird et al. [24] Davidson and Harrison[l]

Davidson and Harrison[l]

E1-Halwagi and EI-Rifai [11]

EI-Halwagi and El-Rifai [11]

Kunii and Levenspiel [2l E1-Halwagi and EI-Rifai [I 1]

Kunii and Levenspiel [2]

For estimating the size of the cloud (fc), the Davidson and Harrison [1] correlation is widely used:

fc = 3Umf/(Eb " U b r - Umf). ( 1 8 )

But, in the present case, it gives absurd values due to which we have taken (fc) as an adjustable parameter, the value of which is adjusted between 0.025 and 0.15. Therefore, the ratio of the volume of the cloud-wake phase to the volume of bubble is finally obtained as:

f~w =fw +f~. (19)

Other parameters that have been used in the model are presented in Table 1.

D E R I V A T I O N OF R E A C T I O N RATE C O N S T A N T

The chemical reaction rate in fluidised bed combustion plays an important role in the formulation of the oxygen mass balance model. In this context, two basic mechanisms have been postulated for char combustion, e.g. the single film model [16] and the double film model [17]. According to single film theory, oxygen reaches the carbon surface and reacts to form varying amounts of CO and CO2. The CO formed is then oxidised to CO2 by incombing oxygen. The double film model postulates that oxygen never reaches the carbon particle surface. Only CO2 reacts with the carbon to produce CO, which, in turn is oxidised by incoming oxygen.

The 10 MW fluidized bed coal combustor at Jamadoba normally operates at 1173 K. Park [18] has shown that the temperature difference between burning char particles and the bulk of the bed can be a maximum of 200 K. However, Campbell and Davidson [4] have pointed out that the combustion mechanism on which the double film theory is based requires a particle temperature in excess of 1373 K, as a result of which the endothermic carbon dioxide reduction at the particle surface cannot be sustained due to the low particle temperature. Therefore, in the present case, the single film theory of char combustion for shrinking particles has been considered where carbon reacts with oxygen to produce CO2.

The model assumptions for char combustion are summarised below:

• Char and volatile matter burn at the same rate. • Release of volatile matter from coal is instantaneous. • Reaction between carbon and oxygen takes place only on the outer surface of the particles,

i.e. inner pores are inaccessible.

The rate of mass transfer is defined as the normal flow of material to a unit surface:

Qtraa~fer = I / S . d N / d t gmol oxygen/cm2-s. (20)

1054 REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION

Hence, the reaction step must similarly be defined as:

Qreaction = 1/S • d N / d t gmol oxygen/cm2-s. (21)

Oxygen diffuses through a film of thickness (AX) onto a plane surface of carbon particle where it reacts with carbon to yield the gaseous product CO2 which diffuses back through the film into the main gas stream. The intrinsic rate of chemical reaction will be proportional to the oxygen concentration at the carbon surface C~.

Considering the carbon oxygen reaction to be of first order,

1IS. d N / d t = KsC~. (22)

Now, at steady state, the intrinsic rate must be equal to the rate at which the oxygen is supplied to the surface via, the gas film. Hence,

Kg(Cp - C~) = Ks- C~ (23)

i.e. mass transported--surface reaction rate. The surface oxygen concentration C~ is derived from equation (23) as follows:

C~ = Kg. Cp/(Kg + Ks). (24)

Substituting equation (24) in equation (22), we get

1IS . d N / d t = Cp/(1/Ks + l/Kg). (25)

The chemical reaction rate constant (K) at any instant, which is controlled by surface reaction and gas diffusion coefficients, is obtained from the above equations as:

1/K = 1/Ks -I- 1/Kg (26)

where Ks, the surface reaction rate constant for carbon particles, is adopted from the expression given by Parker and Hottel [19] which is as follows:

K s = 4.32 x 10". x/-~p, exp[-44,000/RTp]. (27)

The mass transfer coefficient Kg is expressed by the dimensionless Sherwood number (Sh) as:

Kg = Sh . Dg/dp. (28)

There are many correlations available for estimating the Sherwood number, but in our work, the La Nauze et al. [20] correlation has been adopted, which was developed on the assumption of frequent renewal of gas at the particle surface. It was assumed that the particle environment changes often, as a result of which a steady state concentration boundary layer is never achieved, and the correlation is given as:

Sh = 2Emf+ [4.Emf.dp(Umf/Emfq- Vb)/l~Og] '/2 for dp/dpb~d >~ 4.0 (29)

and

Sh ~-- 2" Emf "}- [4" dp" Umf/7[Og] I/2 for dp/dpbed < 4.0. (30)

In the calculation, dp/dpbed = 4.0 has been assumed as the transition point.

SOLUTION PROCEDURE

After deriving the model parameters and the chemical reaction rate constant, the proposed model has been solved by using the calculus of finite difference equations. Manipulating equation (10), the value of Cow, in terms of Cb, is obtained. Substituting the value of Cew, in equation (1) and simpli- fying, the value of Ce, in terms of Cb, is obtained. Further, substituting the values of Cew, and Ce. in equation (6) and after simplification, equation (6) is expressed only in terms of Cb, which is a third order homogeneous linear finite difference equation with constant coefficients. The solution of that equation gives the oxygen concentrations in each phase leaving any stage n(1 ~< n ~< Ns~,g¢) as:

Cb. =f t G~' +fz G~ +f3 6~ (31)

Cow. = g, f~ G7 + g2f2 G~ + g3f3 G~ (32)

REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION 1055

Table 2. Feed coal analysis of the 10 MW F.B.C. power plant at Jamadoba (TISCO), India

Proximate Ultimate

Feed coal% Feed coal% Components by wt Component by wt

VM 15.18 C 26.99 FC 22.0 H 1.92 Moisture 2.86 S 0.5 Ash 59.96 N 0.48

O 7.29

Hence, the reaction step must similarly be defined as:

Co,, = r~f, G7 + r2f2G~ + r3f3G~ (33)

where f , gi, G , and r; are constants. Further details can be seen in Ref. [11]. The average gas composition leaving the nth stage, i.e. at the top of the bed is determined as

follows: Oxygen:

Carbon dioxide:

Moisture:

Nitrogen:

Cav~ = ( G G ° + Ucw G,,,,, + Um,-G,, )/Uo (34)

CO., = Co - C~.,g (35)

XH(1- XW)Wcoal (XW) Wcoal H20 = + (36)

4UoA, 18UoAt

N2 = (0.79/22,400)(273/Tb) + [XN(! - XW)]/(28. UoAt). (37)

From equations (34)-(37), the percentage gas composition can easily be calculated, which is representative of the flue gas composition according to the model assumptions.

The overall fractional consumption of oxygen (X) is given by

x = l - G v d C o . (38)

R E S U L T S A N D D I S C U S S I O N

The input data required for the proposed model was taken from a 10 MW FBC power plant at Jamadoba (TISCO, India). The plant uses the washery rejects from the adjacent coal washing plant as feed. The proximate and ultimate analysis of the feed coal is given in Table 2. The cross-sectional area of the bed is 26.6 m 2 and the bed height at minimum fluidization is 36 cm. The heat generated in the bed is removed by means of 51 mm o.d. stainless steel coils immersed in the bed. Gases leaving the combustor are dedusted by an electrostatic precipitator, There is no provision for returning the carry-over to the bed. Further to test the validity of the model, the data reported by De Souza-Santos [8], Mukherjee et al. [21] and Chen and Saxena [9] have been used.

The data reported by De Souza-Santos [8] are from a Babcock & Wilcox (U.S.A.) unit. The cross-sectional area of the bed is 0.9817 m 2, and the bed height is 0.7 m. The rate of coal feed to the combustor is 50 g/s. The heat generated in the bed is removed by means of 48.3 mm o.d.

Table 3. Flue gas analysis (Jew = 0.3)

Components

Percentage by volume

TISCO CFRI De Souza- Chen and Jamadoba Dhanbad Santos [8] Saxena [9]

Real Simul. Real Simul. Real Simul. Real Simul.

CO: 11-13 15.1 14.2 17.3 13.8 17.36 13-16 15.77 02 6 7 5.72 3.3 3.6 3.9 3.62 4~5 5.2 N 2 76-79 78.4 81.7 79.02 81.2 79.01 78,81 79.05

1056 R E D D Y and M O H A P A T R A : P O W E R P L A N T E X I T G A S C O M P O S I T I O N

stainless steel coils immersed in the bed, and the flue gas leaving the combustor is cleaned by a cyclone. There is provision for returning the carry-over to the bed. On the other hand, the data reported by Mukherjee et al.[21] is from an experimental plant having an octagonal bed cross-sectional area of 0.67 m 2 and a square disengaging chamber. The plant is designed to feed coal at the rate of 20 g/s and overflow is located at 0.75 m above the distributor plate. The data reported by Chen and Saxena [9] is from a pilot plant having a bed cross-sectional area of 0.9875 m 2. The plant is designed to feed monosized feed coal at the rate of 30 g/s. The gas distributor contains 100 orifices, and the bed height is 0.457 m.

The model has been used to predict the consumption of oxygen in the fluidized bed combustor, which makes it possible to predict the outlet gas composition and variation of average oxygen concentration along with the height of the bed. The simulated results are compared with real plant values in all the cases and presented in Table 3. In almost all cases, the agreement between the real plant data and the simulated results is very good.

As mentioned earlier, the function of cloud-wake phase (few) has been used as an adjustable parameter. For a value off~w = 0.3, the predicted flue gas analysis along with the plant data are given in Table 3. From the table, it is clear that the model predictions are in good agreement with plant results, thereby validating the model. Thef~w value has been varied, and the model predictions at different values off~w has been plotted in Fig. 2. The variation off~w between 0.275 and 0.4 influences the oxygen consumption in the bed, indicating the presence of a separate cloud-wake phase even in a small scale commercial plant. For the present case, a few value of 0.3 seems close to reality because the simulated values match the real values (Table 3).

Figure 3 shows the variation of oxygen concentration in the different phases and also the variation of the average oxygen concentration along with bed height. The rate of fall of oxygen concentration in the bubble phase is gradual throughout, whereas it is very steep in cloud-wake and emulsion phases at lower levels of the bed. As expected, the oxygen concentration in the bubble phase is highest followed by the cloud-wake and emulsion phases. This is because the main

1 . 0 -

0.9

0.8

0.7

o 0.6

~ 0.5

(~ 0.4

0.3

0.2

0.1

6. fcw 0.275

• fcw 0.3

~" T b = 1173 K

Exair = 0.3

Sh = 2.36

I I I I I I I 0.0 10 20 30 40 50 60 70

Bed h e i g h t (cm)

Fig. 2. Var ia t ion of oxygen convers ion with bed height at different f¢,~ values.

REDDY and MOHAPATRA:

2 . 0 -

POWER PLANT EXIT GAS COMPOSITION

A Bubble phase (Cb)

• Cloud wake phase (Ccw) 1.8 -- t::l Emulsion phase (Ce)

1,6 -- z ~ • Average concentration (Cavg)

\ 1.4 -- \ Tb = 1173K

fcw = 0.3

\ \ ,~ ~ ~ Tp= 1273K

.~ 1.0 - - ~ "• ~ E 9p= : : 8 3

= 0 . 8 - -

A

~ 0.6 --

0.2 -- ~ •

I I ~ I 0.0 10 20 30 40 50 60 70

Bed height (cm)

Fig. 3. Variation of oxygen concentrations in bubble, cloud-wake and emulsion phases with bed height.

2 . 0 - -

zx Exair 0.1

[] Exair 0.3

• Exair 0.2

e~

E

o

¢-

c)

x o

) .

<

1 . 8 - -

1 .6 - -

1.4 --

1.2 --

1 .0 - -

0 . 8 - -

0 . 6 - -

0.4 --

0.2 --

Tb= 1173 K

Sh = 2.36

pp = 1.68

fcw 0 3

I I I I 1 I 0.0 10 2 0 30 40 50 60 70

Bed height (cm)

Fig. 4. Variation of average oxygen concentration with bed height at different excess air factors.

1057

1 0 5 8 R E D D Y a n d M O H A P A T R A : P O W E R P L A N T E X I T G A S C O M P O S I T I O N

2 0

18

16

¢~ 14

" 12

~ ~o

e~

e "

o 6

0.0

14 - -

12 - -

I0 - -

8 - -

6 - -

4 - -

2 --

.0 --

~-0 -

0 .0

8 -

6 - -

4 - -

2 - -

0 - -

8 - -

6 - -

4 - -

J

A Db Tb = 1 1 7 3 K

[] C O 2 pp = 1 .68

• O x y g e n c o n v e r s i o n (x) Sh = 2 . 3 6

0 2 d = 0 . 1 0 9 2 c m pay

. m ~ nu

. n n ~

I I I I I I 0. I 0 . 2 0 . 3 0 . 4 0 .5 0 . 6

E x c e s s a i r f a c t o r

F i g . 5. V a r i a t i o n o f b u b b l e d i a m e t e r , o x y g e n c o n v e r s i o n , p e r c e n t a g e C O 2 a n d 0 2 w i t h e x c e s s air f ac tor .

combustion reaction is taking place in the emulsion phase, as a result of which the oxygen composition in this phase is the highest, therefore the concentration is the lowest.

In the cloud-wake phase, the oxygen concentration is intermediate between the bubble and emulsion phases. This is because, in the cloud-wake phase, the oxygen consumption is due to combustion as well as transfer to the emulsion phase. However, the combustion reaction in the cloud-wake phase is slow compared to that in the emulsion due to the presence of less particulate matter. The oxygen consumption in the emulsion phase is higher than that in the cloud-wake phase. The variation of the average oxygen concentration with bed height is similar to that of the oxygen variation in the bubble phase, indicating the dominating influence of the bubble phase oxygen concentration on the average oxygen concentration in the bed.

Chen and Saxena [9], with the help of their three phase mathematical model, also predicted the

c;

1.0 - -

0 . 9 - -

0 . 8

0 . 7 - -

0 . 6 0 . 0

De S o u z a Santos [8]

• C F R I

a Chen and S a x e n a 191

I J I I I 0.1 0 . 2 0 .3 0 . 4 0 . 5

E x c e s s a ir f a c t o r

F i g . 6 . V a r i a t i o n o f o x y g e n c o n v e r s i o n w i t h e x c e s s air f ac tor .

I 0 . 6

REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION 1059

oxygen concentration profiles in three phases along with bed height. The variation of the oxygen concentration in the three phases, per their model, is almost similar to our model predictions.

Assuming a two phase theory of fluidization, Gibbs et al. [22] have also reported the exist gas composition and oxygen profile in an experimental bed. The bed was constructed from stainless steel and was 1.83 m high and 0.3 m 2 in cross-section. Combustion gas samples were obtained from the bed, freeboard and exit flue by means of sampling ports located throughout the bed for continuous gas analysis. The bubble and particulate phase oxygen concentrations were ascertained by using a fast response mass spectrometer, the 'peaks' showing the oxygen concentration in the bubble phase and the 'lows' showing the particulate phase oxygen concentration.

Gibbs et al. [22] reported experimental results for the exist gas composition at 10% excess air as CO2 = 13.8%, CO = 1.0%, 02 = 4.0%, N2 = 81.2%, which are similar to our model predictions. They have also reported the variation of the average oxygen concentration in their experimental unit, which is 9% at the coal feeding point and 4% at the top of the bed. As per our predictions, the oxygen concentration in the Jamadoba plant varies between 21% (at the distributor) and 5.72% (at the top of the bed). However, Gibbs et al. [22] have measured the oxygen concentration well above the distributor plate.

Figure 4 shows the variation of the average oxygen concentration with bed height at different excess air factors as a parameter. As expected, the model predicts the decrease of oxygen concentration as the bed height increases at all excess air factors. The oxygen concentration is higher at all bed levels for higher excess air factors. Figure 5 shows the variation of the exit gas concentrations, oxygen conversion and bubble diameter with excess air factor. From Fig. 5, as the excess air factor increases, the CO2 percentage in the flue gas decreases because the amount of combustibles available in the bed remains the same. But, at the same time, the percentage of oxygen in the exit gas increases, leading to excessive loss of sensible heat in the flue gas. However, if the percentage conversion of the oxygen is considered, it decreases as the excess air factor increases. Figure 5 also shows the gradual increase of the bubble diameter with excess air factor. This is due to the fact that the superficial gas velocity increases as the excess air factor increases, leading to the gradual increase of the bubble diameter. The increase in the bubble diameter causes a decrease in the gas interchange coefficients, consequently the oxygen conversion decreases.

Figure 6 shows the variation of the oxygen conversion with excess air factor for data taken t¥om the literature [8, 9, 21]. In all cases, the oxygen conversion decreases with the increase of excess air factor. The oxygen conversion obtained from the data of Mukherjee et al. [21] and De Souza- Santos [8] are very close to each other, whereas the oxygen conversion obtained from Chen and Saxena [9] is lower.

C O N C L U S I O N S

(1) The three phase model predicts exit gas compositions from the 10 MW FBC power plant with reasonable accuracy.

(2) Oxygen concentrations in the bubble, cloud-wake and emulsion decrease with bed height. (3) The oxygen conversion in the FBC plant decreases as the excess air factor increases.

Acknowledgements--The financial aid given by the Council of Scientific and Industrial Research, India (Vide letter No. 2(332)/91 EMR II 14 April 1991) is gratefully acknowledged. The help extended by the staff of the 10 MW Power Plant TISCO, Jamadoba has been appreciated.

R E F E R E N C E S

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Fluidizied Combustion, A 2-1 (1975). 5. H. A. Becker, J. M. Beer and B. M. Gibbs, A model for fluidized bed combustion of coal. Proc. Inst. Fuel Syrup. Series

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1060 REDDY and MOHAPATRA: POWER PLANT EXIT GAS COMPOSITION

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