Cement Model

7/27/2019 Cement Model

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Numerical modelling of flow and transport processesin a calciner for cement production

D.K. Fidaros, C.A. Baxevanou, C.D. Dritselis, N.S. Vlachos

Department of Mechanical and Industrial Engineering, University of Thessaly, Athens Avenue, 38334 Volos, Greece

Received 19 October 2005; received in revised form 1 September 2006; accepted 7 September 2006

Available online 29 November 2006

Abstract

Controlling the calcination process in industrial cement kilns is of particular importance because it affects fuel consumption, pollutant emission

and the final cement quality. Therefore, understanding the mechanisms of flow and transport phenomena in the calciner is important for efficient

cement production. The main physico-chemical processes taking place in the calciner are coal combustion and the strongly endothermic

calcination reaction of the raw materials. In this paper a numerical model and a parametric study are presented of the flow and transport processes

taking place in an industrial calciner. The numerical model is based on the solution of the NavierStokes equations for the gas flow, and on

Lagrangean dynamics for the discrete particles. All necessary mathematical models were developed and incorporated into a computational fluid

dynamics model with the influence of turbulence simulated by a two-equation (k) model. Distributions of fluid velocities, temperatures and

concentrations of the reactants and products as well as the trajectories of particles and their interaction with the gas phase are calculated. The

results of the present parametric study allow estimations to be made and conclusions to be drawn that help in the optimization of a given calciner.

2006 Elsevier B.V. All rights reserved.

Keywords: CFD; Coal combustion; Calcination; Calciner modeling; Cement production

1. Introduction

The main processes of cement production include raw-mix

preheating and calcination, clinker formation and cooling to

achieve a crystalographic structure that meets the required ce-

ment specifications. After cooling, the clinker is fed into grind-

ing or finish mills and is mixed with plaster and ameliorating

additives. The mills consume a very large amount of the total

energy required for cement production.

The raw-mix consists mainly of pulverized calcium carbon-ate and silicon dioxide. During its heating/drying at tempera-

tures from 100 C to 500 C the moisture evaporates and at 850

to 890 C the endothermous calcination reaction begins, where

CaCO3 is converted into CaO and CO2. The activation energy

for the calcination is provided by the combustion heat of the

fuel.

Dry heating of raw-mix in vertical suspension preheaters (see

Fig. 1) is mostly used, where calcination also takes place. The

innovation in the entire pyroprocess in modern cement plants is

the use of an additional calcining vessel, in which the raw-mix

undergoes calcination to a level of 90 to 95%. In this way, the

calcined raw-mix enters the rotary kiln at a higher temperature,

thus reducing the energy demand and the thermal load on the

kiln. After being heated to the appropriate temperature, it enters

the calciner together with the fuel and the hot tertiary air, Fig. 2.

The combustion heat released by the fuel causes calcination ofthe raw-mix according to the chemical reaction:

CaCO3 Y1160 K

CaO CO2 178 kJ=mol 1

The high fineness of the raw-mix and the good turbulent

mixing cause uniform and fast coal combustion and calcination

reactions. The products of the calciner are fed to the last cyclone

that feeds the rotary kiln. The placement of calcination outside

the cement kiln results in better quality of CaO and energy

savings. For example, in the Olympus plant of AGET Hercules

in Greece calcination takes roughly 60% of the total heat

Powder Technology 171 (2007) 8195

www.elsevier.com/locate/powtec

Dedicated to the late Professor Shao-Lee Soo, for his pioneering work in

multiphase dynamics. Corresponding author. Tel.: +30 2421074094; fax: +30 2421074085.

E-mail address: [email protected] (N.S. Vlachos).

0032-5910/$ - see front matter 2006 Elsevier B.V. All rights reserved.doi:10.1016/j.powtec.2006.09.011
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absorbed in the system, while 35% is spent for preheating and

5% for clinkering [1]. This ratio of 60:40 is reversed in the case

where the calcination is taking place inside the rotary kiln. In

addition, the good mixing of fuel, air and raw-mix in the calciner

results to faster calcination with good efficiency at relatively low

temperatures.The advantages of using calcination devices are: a) The

addition of a burner in the calciner increases the capacity of the

rotary kiln in comparison to using simple preheaters, b) The

reduction of thermal load and the increased rotational speed of

the kiln (to achieve better mixing at increased capacity) extends

the lifetime of the firebricks and, thus, the operational life of the

kiln, c) The reduction of energy demand and the minimal calci-

nation in the kiln reduce considerably the exhaust gases and the

kiln heat losses to the environment because the exhaust gases

absorb most of the radiation, d) The combustion at mediumlow

temperatures (b1400 C) in the kiln reduces the production of

NOx, although combustion control and kiln burner design is stillsignificant, e) The lower temperature required in the calciner

allows the use of fuels with relatively low thermal capacity

(usually bituminous coal), f) The reduction of the thermal load of

the rotary kiln decreases the condensing of vapours (SO3, Na, K

and Cl) in the combustion area. However, the volatile cycle is

still a concern because now it will take place in the preheater/

precalciner tower itself as opposed to the kiln), and g) The

reduced calcification percentage in the rotary kiln, decreases its

thermal load and improves its functional stability, as the kiln

burners are now used only for clinkering.

Calciners have become essential devices in cement produc-

tion but have also disadvantages: a) The lower temperatures of

the exhaust gases may cause condensation of volatile alkalis,

while the higher rotational speeds can increase the quantity of

alkaline dust in the kiln, b) Reduction of NOx emissions is not

common in all cement production systems using calciners,

mainly due to geometric and operational differences, as well as

to different quality and quantity of raw-mix and fuels, and c)

The utilisation of fuels with low energy value, although eco-

nomically advantageous, requires particular attention in order toavoid undesirable emissions of polluting and erroding gases.

From the above, it becomes apparent that control of cal-

cination is important because it affects fuel consumption,

pollutant emissions and the final cement quality. Therefore,

understanding the mechanisms of flow and transport phenom-

ena in the calciner may contribute to more efficient production

and better quality of cement.

Recently, calciners have been studied with different geom-

etries and operational conditions in 2D and 3D CFD simulations.

Huanpeng et al [2] studied the influence of various physical

parameters on the dynamics of gassolid two-phase flow in a

precalciner using kinetic theory of granular flow to represent thetransport properties of the solidphase in a 2D model. Hu et al. [3]

used a 3D model for a dual combustor and precalciner using a

Eulerian frame for the gas phase and a Lagrangean one for the

solid phase in order to predict the burn-out and the decompo-

sition ratio during thesimultaneous injection of twotypes of coal

and raw material into the device. Iliuta et al. [4] investigated the

influence of operating conditions on the level of calcination,

burn-out and NOx emissions of an in-line low NOx calciner, and

made a sensitivity analysis of their model with respect to aero-

dynamic and combustion/calcination parameters.

In the present work a numerical model is described for the flow

and transport processes taking place in an industrial calciner. The

model is based on the solution of the NavierStokes equations for

Fig. 1. Schematic of cement production.

Fig. 2. Calciner device.

82 D.K. Fidaros et al. / Powder Technology 171 (2007) 8195


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the gas flow and on Lagrangean dynamics for the discrete

particles, using a commercial CFD code. All necessary flow, heat

and mass transfer and chemical reaction models are presented

with the influence of turbulence simulated by a two-equation

(k) model. Limited available measurements from the Olympus

cement plant of AGET Hercules are used to verify the model.

2. Mathematical models

2.1. Gaseous phase

The general form of the time-averaged transport equation for

momentum, heat and mass of the gases is:

A

AtqU

A

AxqUU

1

r

A

ArqrVU

1

r

A

Ahq

W

r

AU

Ah

A

AxCU

AU

Ax

1

r

A

ArCUr

AU

Ar

1

r

A

AhCU

1

r

AU

Ah

SU

2

where U, V, Ware the time-averaged velocities in the axial, radial

and circumferential direction, respectively, the transport

coefficient, and any time-averaged transported fluid property.

2.2. Particle dynamics

The particle trajectories are calculated from their corre-

sponding motion equation:

dUp

dt FDUUp gi

qpq

qp fi 3

where, the subscriptp denotes particle.For spherical particles, FD in the drag force term is:

FD 3lCDRe

4qpDp2

4

where the drag coefficient is calculated from:

CD a1 a2

Re

a3

Re25

and 1, 2 and 3 are constants proposed by Morsi and

Alexander [5].

The additional force term fi

in Eq. (3) may be due to pressure

gradients, thermophoretic, Brownian or Saffman lift forces.

2.3. Particle size distribution

The particle sizes follow a RosinRammler distribution:

MD eD=Do

n 6

where n is calculated from:

n lnlnMD

lnD=D7

Each size interval is represented by an average diameter for

which the trajectory calculations are performed.

2.4. Particle heat transfer

Particle heat transfer is due to convection, radiation and de-

volatilization, as follows:

Tpt t

h ApTl dmpdt

hfg ApeprH4

R

h Ap ApeprT3p

Tpth ApTl

dmpdt

hfg ApeprH4

R

h Ap ApeprT3p

0@

1A

e

Ap hep rTe

p

mp Cpt

82.5. Devolatilization model

The devolatilization model of Kobayashi [6] is used:

R1 A1expE1

RTp ; 9a

R2 A2expE2

RTp

9b

where, R1 and R2 are competitive volatilization rates at different

temperature ranges. These yield an expression for devolatiliza-

tion:

mvt

mpomash

Zt0

a1R1 a2R2exp

Zt0

R1 R2dt

dt 10

The Kobayashi model requires known kinetic parameters

(A1, E1) and (A2, E2) and the contribution of the two reactionsvia the factors a1 and a2. More specifically A1 =2.0e +07 s

1

and A2 =1.0e + 0 7 s1 are the pre-exponential factors, and

E1 =1.046e +05 J/mol and E2 =1.67e + 05 J/mol are the

activation energies. It is recommended that the value of a1should be equal to the fraction of volatiles that is determined by

the proximate analysis, because this rate represents the volatile

evaporation at low temperatures. The value of a2 should be

equal to 1, as it expresses the contribution of the evaporation

rate of volatiles at very high temperatures.

2.6. Surface/coal combustion models

After devolatilization is completed, there starts the surface

chemical reaction of the coal particle which may be modelled as

follows:

2.6.1. Diffusion model

The reaction rate is determined by the diffusion of the gas

oxidant into the particle surface:

dmp

dt 4pDp Dim

moTpqg

SbTp Tl11

In this model the particle diameter is assumed constant and,

as its mass decreases, the active density decreases resulting in

83D.K. Fidaros et al. / Powder Technology 171 (2007) 8195


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a more porous particle. Eq. (11) proposed by Baum and

Street [7] ignores the contribution of kinetics to the surface

reaction.

2.6.2. Kinetic/diffusion model

The reaction rate is determined by the diffusion of gas

oxidant into the particle surface or by the reaction kinetics. Themodel proposed by Baum and Street[7] and Field [8] is used, in

which the diffusion rate is:

R1 C1Tp Tl=2

0:75

Dp12

and the kinetics rate:

R2 C2exp E

RTp

13

The kinetics rate incorporates the effects of chemical reactionin the internal surface of a coal particle and the epidermic

diffusion. The rates R1 and R2 are combined to give the

combustion rate of the coal (char) particle.

dmp

dt pD2pP0

R1R2

R1 R214

The particle size is kept constant, until a significant reduction

in its mass leads to a new size estimation.

2.7. Particle radiation

The radiation from the coal particles into the gas isincorporated via the P-1 model [910]:

jdCjG 4p arT4

p Ep

a apG 0 15

where, Ep and p are calculated from:

Ep limVY0

XNn1

epnApnrT4pn

pV16a

ap limVY0

XN

n1

epnApn

V16b

The quantity in Eq. (15) is:

C 1

3a ap rp17

and p is calculated from:

rp limVY0

XN

n1

1fpn1epnApn

V18

The calculation ofp is repeated in the entire trajectory forn

particles. Then, the source term that is introduced into the

energy equation is:

jqr 4p arT4

p Ep a apG 19

2.8. Chemical reaction models

The present modelling of mixture fraction [11,12] with the

method of probability density function (mixture fraction/PDF)

requires the solution of transport equations for one or two

conservative scalar properties. The effect of turbulence is also

considered. The method of mixture fraction with PDF has been

developed specifically for turbulent chemically reacting flow

simulations. The chemical reaction is determined by turbulent

mixing, which controls the limits of the kinetic rates. The PDF

method offers many advantages compared to the method of

finite reaction rate. The method of mixture fraction allows theexplicit intermediate calculation of chemical compound form-

ing and the interlacing of turbulence and chemistry. The method

is economic, because it does not require the solution of a large

number of transport equations for each chemical species. More-

over, it allows precise determination of auxiliary variables such

as density, and it does not use average values, in contrast to the

method of finite reaction rate.

For a binary system such as fuel and oxidant, the mixture

fraction can be formulated in terms of elemental mass fractions:

f ZkZkO

ZkFZkO20

The value of f is calculated from the solution of a time-

averaged transport equation:

A

Atqf

A

Axiqui f

A

Axi

ltrt

Af

Axi

Sm 21

The source term Sm is present only when particle mass

transport to the gaseous phase takes place.

Simultaneously with the solution of Eq. (21), a conservative

equation for the variance of mixture fraction, fV2, describing the

interaction between chemistry and turbulence, is solved:

A

AtqfV2

A

Axiquif

V2 A

Axi

ltrt

AfV2

Axi

!

CgltAf

Axi

!2Cdq

e

kfV2 22

where, t, Cg and Gd are constants equal to 0.7, 2.86 and 2.6,

respectively.

2.8.1. Coal reaction mechanisms

Coal combustion The most important physico-chemical

change in the coal particle during heating is thermal frag-

mentation (pyrolysis) at high temperatures. During this stage an



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important loss of weight occurs, because of dissolution of

volatile matter, the quantity and composition of which depend

on the ingredients of coal, its grain size and temperature. During

dissolution of volatiles, a number of parallel reactions occur,

with chemical combinations of reacting components or even

species such as, for example, CH4, CHOH, C2H6, H2, and S2.

After devolatization leading to production of water vapour, CO,CO2 etc, a series of progressive reactions of char and de-

volatization gases take place as follows [1,1323]:

Heterogeneous reactions

Cs O2gCO2g 23a

2Cs O2g2COg 23b

Cs 2H2gCH4g 23c

Cs CO2g2COg 23d

Cs H2OgCOg H2g 23e

Homogeneous reactions

2COg O2g2CO2g 24a

COg H2Og

CO2g H2g 24b

COg 3H2gCH4g H2Og 24c

CH4g 2O2gCO2g 2H2Og 24d

HCg 1:5O2gCO2g H2Og 24e

The decomposition and polymerization reactions of the

superior and unsaturated hydrocarbons are also added:

Superior HCgY

fragmentation

Inferior HCg Cs

Unsaturated HCgYSaturated HCg

Unsaturated HCg H2gYPolymerization

Superior HCg

Pyrolysis As temperature increases, the humidity and the

gases enclosed in the coal particles are released. The larger

percentage of the non-chemically combined water is evaporated

at temperatures below 105 C while the chemically combined at

temperatures exceeding 350 C. At pyrolysis temperatures,

certain types of coal melt, forming an intermediate product

called metaplast. With the increase of temperature the metaplast

is split, shaping the basic volatile products and semicoke,

causing the coal particles to swell. This is described by a factor

that depends on the composition of volatiles and the heating

rate. The increase of particle volume does not influence the

activity of pyrolysis, while the semicoke formed initially, is

decomposed as temperature increases.The rate of thermal decomposition increases with increasing

temperature up to a maximum value. Many researchers (for

example, [15,17,23]) have found that pyrolysis ends around 850

to 1000 C, while its duration is limited to a few seconds

depending on the particle size. After the volatiles have been

released, the remaining solid (char) still retains a small per-

centage of volatiles (1.5%) like H2 and N2, requiring a tem-

perature near 2000 C to be removed completely.

Experiments show that the determination of volatiles in coal

is demanding and time-consuming. Many measurements of

volatiles based on the ASTM standard, present large differences

in the percentage of volatiles depending on the rate of tem-perature increase and on the experimental method [10,15,16,

18,19]. The solid remains of the particles formed during thermal

decomposition are mainly fixed carbon, with high porosity and

large internal surface, and the inorganic part is ash. The tem-

perature varies between 1200 and 1800 C causing ash melting.

The composition and the nature of ash as well as its properties

(melting point, viscosity, etc) depend to a large extent on the

pyrolysis conditions.

In cases where the gaseous phase consists mainly of air, the

pyrolysis and the combustion of char proceed simultaneously.

However, in general, char combustion follows pyrolysis, with

only a very small time overlap. In ordinary coal particles, volatiles

tend to be emitted in concentrated but randomly distributed jetsfrom their surface. The larger jets reject volatiles during thermal

decomposition while smaller jets begin and end during this

period. When the gaseous phase is hot enough and rich in oxygen,

the jets of volatiles ignite to form jet flames. In relatively large

particles, the emission and combustion of volatiles can keep the

char surface free of oxygen. When the surface of hot char is

accessed by oxygen, there begins a heterogeneous combustion

reaction with longer duration, lasting 15 to 20 times than the

thermal decomposition of volatiles, depending on its evolution

and combustion conditions [2129].

The heating rate of coal particles depends on their size and

contact with the thermal source. For example, the heating rate ofcoal powder by a surrounding flame is 1000 C/s, but when the

flame is from powder coal particles, the rate may increase to

10000 C/s. The pyrolysis results in a number of products with

large differences in molecular weight, from gaseous hydrogen

up to heavy organic species (tar). The data provided by exper-

iments concerning rapid pyrolysis is not sufficient to determine

the composition and distribution of intermediate products for

various coals [3032].

Thus, the mathematical models developed for devolatiliza-

tion are based on the initial coal particle composition. Many

researchers, assume that the coal is considerably homogeneous,

so it is possible to be assumed as a heated mass and altered

gradually from volatilescharash to charash and finally to



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ash. From tables of ultimate analyses of coal and pet coke [1,2], it

appears that the main components of volatiles are CO, CH, H2O,

and H2. Given that the atmosphere of the calciner is oxidant and

assuming that all these components react with oxygen, the main

reactions considered as taking place are:

CO 1=2O2CO2283:2kJ=mol 25a

H2 1=2O2H2O242kJ=mol 25b

CH4 2O2CO2 2H2O802:86kJ=mol 25c

Char combustion The mechanism of char combustion has

been investigated more than pyrolysis, without definitive

answers to questions concerning the quantitative origin of

some constituents after the end of transformation. Qualitatively,

however, it has been modelled satisfactorily by various mathe-

matical models. These were developed in order to describe the

solid coal combustion and have found important application inreactions of porous solids with gases.

Two simple mathematical models describe the reaction of

coal grain with oxygen: the simple film and the double film

model. In the first model the oxygen is diffused via a constant

boundary layer in the surface of the char particle, where it reacts

to form CO and CO2. The CO is then diffused in the well-mixed

environment. In the second model, char reacts with CO 2 and not

with oxygen, in order to produce CO that is burned in a thin

flame inside the boundary layer. The CO reacts with oxygen

inside the boundary layer, and thus the oxygen never approaches

the char surface. Small particles (b100 m) are considered to

burn according to the first model and larger (up to N2 mm)

according to the second. However, the two models constituteonly the two extreme cases of char combustion and cannot,

therefore, establish a general theory [2936].

The real mechanism of combustion is more complicated,

because of many factors involved such as particle size, local

temperature, local oxygen concentration and reaction controlling

mechanism. Generally, the oxygen and CO are readily available

on the coal surface and can, therefore, react simultaneously with

coal and also between each other. The situation becomes more

complex when the char porosity is taken into consideration

(intrinsic model).

A more complex model proposed by Essenhigh [37] describes

better the above processes. In this model the distribution oftemperature and concentrations are extended to the center of the

particle. The more usual diffusion controlled combustion of CO

can be extremely fast, consuming all the local oxygen before it

reaches the char surface and reacting only with CO2. In the

chemically controlled combustion of CO2 and O2, these have

equal probability to react with the char surface. Moreover, ex-

perimental data by Field [8] and Borghi [38] showed that the

reaction of charCO2 is very slow in comparison with the reaction

of charO2. Therefore, the latter can be considered as the main

reaction on the char surface when the essential quantity of oxygen

is available. However, the presenceof CO2 cannot be ignored and,

thus, there always exists the probability of parallel reactions

[26,27,2931]. Based on a comparative analysis of existing data

for coal combustion and on the constitution and granulometry of

particles (average char diameter100 m), the selected model

for these particles was that of the kinetic/limited diffusion rate.

This is similar to that of shrinking-reactant particle core adopted in

the general theory of surface heterogeneous chemical reaction.

The diffusion coefficientDim of oxidant in the porous char used in

the present model was 5.0e05 m 2/s.

2.8.2. Calcination mechanisms

The calcination of limestone particles includes several stages,

with each one imposing different chemical kinetics rates: a) Heat

transfer from the gases to the particle surface and from it to the

reaction interface, b) thermal decomposition of CaCO3 in the

reaction interface, c) mass flux of CO2 from the reaction in-

terface to the gases.

For small limestone particles moving in high temperatures

gases, the internal and external heat and mass transfer rates are

high. Specifically, for particles with diameter between 1 and

90 m and gas temperatures between 748 and 1273 K, Borgwardt[39] has reported that the calcination is chemically controlled and

its rate is proportional to the surface area of the particle as

determined by the BET method (nitrogen absorption at 77 K).

Because, the limestone microstructure is not completely crystalic

and has a diverse form of porosity, the surface determined by the

BET method, is the sum of the porous surfaces accessed by

nitrogen. Under these conditions, the calcination happens on the

total available surface, giving pseudo-volumetric characteristics

to the reaction.

From the analysis of calcination data of high fineness

limestone in isothermal reactors, it is concluded that, for a better

description of the reaction evolution, the model of shrinking core

should be selected, with the size diameter raised to thepower 0.6.The value of the exponent (b1) is explained by the fact that the

calcination proceeds radially to the particle core, without

inhomogeneities in the reaction interface. When the raw-mix

particles are small, the reaction interface of the calcination is not

easy to determine. The internal thermal gradients and the partial

pressure of CO2 are also difficult to estimate. For these reasons,

most calcination models of high fineness particles consider that

the surface temperature is equal to the gas temperature,

neglecting the internal thermal gradients [40].

The decomposition reaction of CaCO3 is strongly endother-

mic. Its thermodynamic state is defined by the reaction enthalpy

H and the equilibrium pressure PCO2,eq,:

PCO2;eq exp H

RTS

R

26

These values depend on temperature and are influenced by the

nature of limestone, its degree of cleanliness and mainly by its

structural mesh. The lower the degree of cleanliness of raw

material, the lower is the reaction enthalpy. Also, the function of

temperatureequilibrium pressure PCO2,eq = f(T) develops to

lower temperatures because of the chemical kinetics of the

recently formed CaO and the impurities in the reaction

environment. The values of reaction enthalpies provided by the

open literature [4046] for the present endothermic reactions



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vary. Particularly, in the high interest range for calcination (800 to

1000 C), the reaction enthalpy is not linearly dependent on

temperature. Thus, for practical calculations 900 =1660 kJ/kgCaCO3 =396 kcal/kg CaCO3 can be assumed a constant value for

this specific temperature range [1,41].

The decomposition of limestone takes place in a reaction

zone, where the core of unreacted CaCO3 and the newly formed

CaO meet. This front moves from the perimeter to the center

with a certain speed, while heat is transferred simultaneously to

the core and CO2 is emitted to the outside. This reaction

proceeds in the following stages: a) Heat is transferred from the

surroundings to the particle surface, b) heat is conducted

through the reacted layer to the reaction zone, c) chemical

reaction occurs in the reaction zone, CO2 emission, nuclei

creation and reforming of CaO, and d) CO2 is diffused through

the CaO layer to the particle surface and the surroundings.The final reaction speed is a function of the rates of the

above stages. Because these rates are of the same order of

magnitude, a balance is achieved in the decomposition front,

under the prevailing temperature and partial pressure of CO 2,

so that the rates of the above stages become equal. If large

limestone particles exist, diffusion of mass and conduction of

heat will dominate, especially when the surrounding temper-

ature is high and the partial pressure low. In the case of low

temperatures and high partial pressures of CO2, the material

transformation occurs by the diffusion of CaO. For a fine

granulometry of ground limestone or raw-mix in ordinary

calcination conditions, the chemical kinetics play a decisiverole [1].

Thus, the proposed model, calculates the rates of particle

calcination and heat transfer by considering: a) the heat transfer

by convection from the gases to the particle and by conduction

to the particle interior, b) surface decomposition of CaCO3, and

c) mass transport of CO2 from the reaction interface via the

porous particle to the gaseous environment.

The calcination is a heterogeneous reaction and occurs at the

lime surface when the local pressure exceeds the criterion of

Baker [47]:

Pe 1; 826 10

7

exp

19; 680

T

27

The reaction rate at the interface is expressed as follows,

Borgwardt [39]:

Rate ks ACaCO3 28a

where,

ks Aexp Ea

RT

28b

The activation energy Ea of the decomposition reaction is in

the range 165205 kJ/mol.

The calcination of small limestone particles dispersed in the

gaseous phase, can proceed at temperatures up to 1600 C. The

effect of CO2 partial pressure is incorporated in the decompo-

sition rate by modifying it as proposed by Darroundi and Searcy

[48]:

kVs ks for Pb102Pe 29

k

V

s ksPe

P=Pe for 102

Peb

Pb

Pe 30

The effect of temperature on calcination chemical kinetics is

shown in Fig. 3. During calcination, the thermal conductivity of

Fig. 3. Variation of calcination chemical kinetics with temperature.

Fig. 4. Calciner side view.

Table 1

Mass flow rates at the inlets of the calciner

Kind of

mass flow

Case 1 coal Case 2 pet coke

Quantity [kg/s] Percentage Quantity [kg/s] Percentage

mCaCO3 52.47 54.3% 52.47 54.6%

mCoal 3.78 3.9% 3.17 3.3%

mTertiary Air 39.36 40.7% 39.36 41.0%

mAir Coal 0.97 1.0% 0.97 1.0%

Total 96.59 100.0% 95.98 100.0%



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lime depends on the state of the solid material and differs con-

siderably for non-calcinated, partially calcinated or fully

calcinated particles. This is mainly due to the different structure,

but also to the change of the specific surface area. In the reaction

region, the thermal conductivity of the particle is a linear

function of specific surface area and temperature. Thus, the

thermal conductivity of CaCO3 was 1.646W/(m.K) and of CaO

0.860 W/(m.K). The mass fraction of CO2 is determined from a

diffusion equation assuming a spherical particle.

3. Computational details

3.1. Calciner geometry

The modeled calciner, Fig. 4, consists of a cylindrical and a

conical section having three kinds of inlets at the bottom part

and an outlet at the top, from where the products such as

calcined raw-mix, CO2, and other gases exit. Raw-mix is fed

into the calciner via two 0.6 m diameter pipes inclined at 60

to the horizontal. The tertiary air enters axially from the bottom

via a concentric 2.6 m diameter duct and the coal is fed at the

lower conical part via two 0.2 m pipes at 30 to the horizontal.

The physico-chemical processes take place in the main volume

of the calciner, consisting of a 6.6 m diameter cylinder with

20 m height. The upper conical part has 1.1 m height and leads

to a cylindrical part with 4.3 m diameter and 5 m height. The

total calciner volume is 850 m3. The coal entries are at 2.4 m

height from the start of the cone and at 2.68 m from the

calciner axis.The computational domain consists of a hybrid mesh of

67.104 cells. Because of symmetry, the calculations were carried

out for one half of the calciner using the FLUENT code. Two

fuels (coal and pet coke) were considered and the total rate of

mass (raw-mix, coal and air) fed into the three kinds of inlets was

aproximately 100 kg/s. As shown in Table 1, the larger per-

centage of mass rate is that of CaCO3, followed by tertiary air,

coal and finally the coal feeding air.

The RossinRammler distribution of the raw-mix size had

an average value of d=16.6 m and a spread parameter of

n =0.822, while the coal had d=34.5 m and n =1.248. The

analysis of the raw meal and coal particles is given in Table 2.The tertiary air entered with a velocity 24 m/s, coal with 11.5 m/s

and the raw-mix with 1.5 m/s. The coal was fed pneumatically

while the raw-mix entered by gravity.

All the geometric data and the initial and boundary conditions

were supplied by Olympus plant of AGET Hercules in Volos,

Greece.

4. Results and discussion

4.1. Case 1 (Good quality coal)

Fig. 5 shows the calculated velocity distribution of the gaseous

phase for Case 1 (good quality coal) in two vertical diametral

Table 2

Ultimate analysis for the solid feeds

Components Case 1 coal Case 2 pet coke

As used (%) Dry basis (%) As used (%) Dry basis (%)

Humidity 1.45 0.0 1.45 0.0

Volatiles 27.70 28.11 13.21 13.40

Ash 12.05 12.23 0.14 0.14Fix carbon 55.80 56.62 81.70 82.90

Total 97.00 96.96 96.50 96.44

Fig. 5. Velocity distribution in vertical symmetry plane (left) and at 90 (right) for Case 1.



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planes normal to each other. The gases undergo an abrupt

deceleration at the beginning of the lower conical part due mainly

to the entry of the coal and secondly of the raw meal. In the main

cylindrical part, the velocity remains at 7 to 8 m/s, with regions of

higher velocity in front of the two raw-mix inlets and in the upper

conical part. At the exit, a region with higher velocities is

observed, a fact due to the relative absence of particles.In Fig. 6 higher temperatures are observed in the opposite

side of the raw-mix entries. This is due to the trapping of small

coal particles while the concentration of CaCO3 particles is low.

The main body of the calciner is maintained at temperatures

little above the threshold for calcination, so that calcification

takes place almost in the whole device. There are no spots of

high temperature but rather regions of low temperature because

of intense calcination, resulting to high heat absorption. The

high temperatures in regions where higher velocities prevail are

mainly due to the high concentration of burning coal and to the

absence of CaCO3 particles.Fig. 7 shows concentration distributions of CO2, O2 and H2O

in various horizontal cross-sections. Higher concentrations are

observed a little after the raw-mix inlet. It should be noted that

high CO2 concentrations result to high heat absorption, thus

Fig. 6. Temperature field in the vertical symmetry plane (left) and at 90 (right) for Case 1.

Fig. 7. Concentration distributions of CO2 (left), H2O (middle) and O2 (right) for Case 1.



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limiting calcination. The CO2 concentration decreases progres-

sively towards the exit, mainly because the available CaCO3 is

also decreasing. The amount of CO2 produced by coal com-

bustion is much smaller than that due to the calcination (ratio

1:6). Smaller O2 concentrations are observed in regions of high

temperature, mainly due to pyrolysis and the start of combustion

of the coal remains (char). In the part of the cylinder whereintense calcination takes place, the O2 concentration remains at

low levels because of the high raw-mix and CO2 concentrations,

while in regions where coal particles have been trapped, lower

CO2 concentrations are observed (mainly from coal combus-

tion). Higher H2O concentrations are recorded after the fuel inlet

in the regions of pyrolysis and coal combustion. These

concentrations decrease along the height of the device and

away from the combustion region.

The trajectories of coal particles are presented in Fig. 8 and

those of CaCO3 in Fig. 9. The average passage length covered by

all particles is about 54 m, while the longest exceeds 75 m. The

right parts ofFigs.8 and 9, show the parts of the trajectories wherethe particles are activated thermochemically. Thus, the blue colour

in the start of the trajectory corresponds to coal warming up, the

green and yellow to evaporation and combustion of volatiles,

respectively, and the red to the combustion of fixed carbon.

Finally, the blue colour in the end of the trajectory shows the

cooling of the ashby thegases. Theaverage passage lengthof coal

particles is 32 m, while the longest exceeds 45 m. Their average

residence time is 5 s, mainly because the air moves faster

sweeping the smaller coal particles. The average particle

residence time is 10 s while the longest 15 s. Calcination is

noticeable in the semi-cylinder defined by the raw-mix feeding

pipes and the coal inlets, which agrees with the temperature field

and the CO2 concentration. The predicted calcination for Case 1reaches 96.5%, and is realised in all the active calciner height.

4.2. Case 2 (Pet coke)

Fig. 10 shows the gas velocity field for Case 2 (pet coke

fuel). In the start of the conical part, the gases undergo a strong

deceleration, again for the same reasons as in Case 1. In the

main cylindrical section of the calciner, a region with higher

velocities opposite to the two raw-mix inlets and in the upper

conical part is observed. The gas velocity reaches 8.5 m/s,

encouraged by the relatively small particle load. In contrast, the

velocities decrease to 56 m/s in the remaining part. This is

mainly due to the higher particle concentration, which

influences the exiting speed of the gases.

In Fig. 11 higher temperatures relative to Case 1 are observed.

This is attributed to the better quality and utilization of the fuel.

Higher temperatures areobserved in the opposite side of the raw-

mix inlets, mainly due to the trapping of small coal particles

while the concentration of CaCO3 particles is exceptionallysmall. The main body of the calciner is kept at temperatures well

above the effective calcination temperature, so that large CaCO3quantities are calcinated in short times. The high temperatures

seen in regions where higher speeds prevail are attributed to the

absence of CaCO3 particles that would consume the heat re-

leased. Also at the exit, higher temperatures are observed, be-

cause calcification there is exceptionally limited.

Fig. 12 shows that higher concentrations of CO2, O2 and

H2O are present at a small distance after the raw-mix inlet.

Again it should be noted that high CO2 concentrations result to

high heat absorption, as far as it does not limit the calcination.

CO2 concentration decreases gradually along the 16.5m calcinerheight until the exit, where the available quantity of CaCO3 is

also decreased. Smaller O2 concentrations are observed in the

regions where the temperatures are high (pyrolysis and com-

bustion of solid coal remains). Also, in the semi-cylindrical part,

where intense calcination takes place, the concentration of O2remains at low levels because of high raw-mix and CO2concentrations, while in regions where coal particles have been

trapped, still lower concentrations are observed. Higher con-

centrations of H2O are seen little after the coal entry and in the

pyrolysis and coal combustion regions. These concentrations

are decreasing along the calciner height.

From the predicted trajectories of pet coke particles (not

shown), most calculated residence times did not exceed 10 sFig. 8. Trajectories of coal particles for Case 1 (combustion is marked in red).

Fig. 9. Trajectories of CaCO3 particles for Case 1 (calcination is marked in red).



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with the longest reaching 16.2 s. The average trajectory length

for all particles (raw-mix and pet coke) was roughly 52 m, while

the longest exceeded 73 m. The average residence time of pet

coke particles was smaller (4.5 s) than in Case 1. The average

trajectory length of these particles was 30 m while the longest

42 m. It should be noted that the combustion of most pet coke

particles is completed inside the device. The higher tempera-tures observed are due to the intense and fast char combustion,

and are accompanied by low concentrations of O2 and high

water vapour levels. This results from the volatiles and

hydrogen compounds in the char. The fast combustion of pet

coke forces the ash to abandon the calciner faster than other

particles.

From the predicted trajectories of CaCO3 particles (not

shown), their average residence time reached 10 s and the longest

15.6 s, indicating that the speed of these particles is higher thanthat of the pet coke. The longer residence time corresponds to the

particles that collide in the upper conical part (before the exit) and

Fig. 10. Velocity distribution in vertical symmetry plane (left) and at 90 (right) for Case 2.

Fig. 11. Temperature field in the vertical symmetry plane (left) and at 30 (right) for Case 2.



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are trapped by the rising particles from the central part of the

calciner. Calcination is realised fast and in a region close to the

calciner axis. The calcination rate for the particular fuel (Case 2)

reaches 98.7%, without taking advantage of the total active height

by the majority of CaCO3 particles. The reason that a small raw-

mix quantity is not being calcinated is the large diameter of

CaCO3 particles and the rapid acceleration observed near the

device exit.

Fig. 13 shows details of the fuel particle trajectories. It is

evident that the pet coke particles (Case 2) react faster than thoseof coal (Case 1), as soon as they enter the calciner.

Finally, the evolution of CaCO3 calcination for Cases 1 and 2 is

depicted in Fig. 14. The differences relate to the energy prevailing

levels observed, the aerodynamics (mass density, particle load)

and the CO2 concentrations that suppress calcination. The

evolution of calcination is represented by the CaCO3 decompo-

sition along the calciner height, starting from where calcification

Fig. 12. Concentration distributions of CO2 (left), H2O (middle) and O2 (right) for Case 2.

Fig.13. Details of fuelparticle trajectories for Case 1 (upper) and Case 2 (lower).(The pet coke particles (Case 2) react faster than the coal particles (Case 1)). Fig. 14. Evolution of calcination with height for Cases 1 and 2.



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t time

U gas velocity

Up particle velocity

U,V,W time-averaged axial, radial and circumferential

velocities

zk mass fraction for the chemical element k

zKF mass fraction for fuel streamzKO mass fraction for the oxidizer stream

S change of grammolecular entropy

H reaction enthalpy

Greek letters

absorption gas factor

p equivalent absorption factor

transport coefficient

p particle brightness

R particle temperature produced by the intensity of

thermal radiation

gas molecular viscosityt turbulent viscosity

p particle mass density

g gas density

Boltzmann's constant

p particle scattering equivalent factor

time-averaged transported fluid property

Acknowledgements

This work was partially supported by the General Secretariat

for Research and Technology of Greece and the cement

company AGET Hercules (EPET-II/96SYN121). The authors

wish to thank Dr. I. Marinos and Mr. T. Pissias of AGET for thedata and for useful discussions.

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