Powder Technology 171 (2007) 81 – 95 www.elsevier.com/locate/powtec
Numerical modelling of flow and transport processes in 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 Navier–Stokes 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–ε) 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 cement specifications. After cooling, the clinker is fed into grinding 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 carbonate and silicon dioxide. During its heating/drying at temperatures 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. ☆
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
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 of the raw-mix according to the chemical reaction: 1160 K
CaCO3 Y CaO þ CO2 þ 178 kJ=mol
ð1Þ
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
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Fig. 1. Schematic of cement production.
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 calcination 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 medium–low temperatures (b1400 °C) in the kiln reduces the production of NOx, although combustion control and kiln burner design is still significant, 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 production 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 economically advantageous, requires particular attention in order to avoid undesirable emissions of polluting and erroding gases. From the above, it becomes apparent that control of calcination is important because it affects fuel consumption, pollutant emissions and the final cement quality. Therefore, understanding the mechanisms of flow and transport phenomena in the calciner may contribute to more efficient production and better quality of cement. Recently, calciners have been studied with different geometries and operational conditions in 2D and 3D CFD simulations. Huanpeng et al [2] studied the influence of various physical parameters on the dynamics of gas–solid two-phase flow in a precalciner using kinetic theory of granular flow to represent the transport properties of the solid phase 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 decomposition ratio during the simultaneous injection of two types 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 aerodynamic 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 Navier–Stokes equations for
Fig. 2. Calciner device.
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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–ε) 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 A 1A 1 A W AU ðqUÞ þ ðqU UÞ þ ðqrV UÞ þ q At r Ar Ax r Ah r Ah A AU 1A AU 1 A 1 AU CU CU r CU ¼ þ þ þ SU Ax Ax r Ar Ar r Ah r Ah ð2Þ where U, V, W are the time-averaged velocities in the axial, radial and circumferential direction, respectively, ΓΦ the transport coefficient, and Φ any time-averaged transported fluid property. 2.2. Particle dynamics The particle trajectories are calculated from their corresponding motion equation: qp −q dUp ¼ FD ðU−UpÞ þ gi þfi ð3Þ dt qp where, the subscript “p” denotes particle. For spherical particles, FD in the drag force term is: 3lCD Re FD ¼ 4qp Dp2
ð4Þ
where the drag coefficient is calculated from: CD ¼ a1 þ
a2 a3 þ Re Re2
ð5Þ
and α1, α2 and α3 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
MD ¼ e−½ðD=D o Þ ¯
lnðlnM D Þ ¯Þ lnðD=D
Particle heat transfer is due to convection, radiation and devolatilization, as follows: dm h Ap Tl þ dtp hfg þ Ap ep rH4R Tp ðt þ ΔtÞ ¼ h Ap þ Ap ep rTp3 0 1 dm h Ap Tl þ p hfg þ Ap ep rH4R dt A þ @Tp ðtÞ− h Ap þ Ap ep rTp3 e
−
Ap ðhþep rTpe Þ t mp Cp
2.5. Devolatilization model The devolatilization model of Kobayashi [6] is used: E1 R1 ¼ A1 exp ; RTp
E2 R2 ¼ A2 exp RTp
ð8Þ
ð9aÞ
ð9bÞ
where, R1 and R2 are competitive volatilization rates at different temperature ranges. These yield an expression for devolatilization: Z t Z t mv ðtÞ ¼ ða1 R1 þ a2 R2 Þexp ðR1 þ R2 Þdt dt ð10Þ mpo −mash 0 0 The Kobayashi model requires known kinetic parameters (A1, E1) and (A2, E2) and the contribution of the two reactions via the factors a1 and a2. More specifically A1 = 2.0e + 07 s− 1 and A2 = 1.0e + 07 s− 1 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 a1 should 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:
ð6Þ
ð7Þ
mo Tp qg dmp ¼ −4pDp Dim dt Sb ðTp þ Tl Þ
where n is calculated from: n¼
2.4. Particle heat transfer
2.6.1. Diffusion model The reaction rate is determined by the diffusion of the gas oxidant into the particle surface:
The particle sizes follow a Rosin–Rammler distribution: n
83
Each size interval is represented by an average diameter for which the trajectory calculations are performed.
ð11Þ
In this model the particle diameter is assumed constant and, as its mass decreases, the active density decreases resulting in
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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. The model proposed by Baum and Street [7] and Field [8] is used, in which the diffusion rate is:
The calculation of σp is repeated in the entire trajectory for n particles. Then, the source term that is introduced into the energy equation is: rT 4 −jqr ¼ −4p a þ Ep þ ½a þ ap G p
ð19Þ
2.8. Chemical reaction models
The particle size is kept constant, until a significant reduction in its mass leads to a new size estimation.
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 the explicit intermediate calculation of chemical compound forming 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. Moreover, 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:
2.7. Particle radiation
f ¼
The radiation from the coal particles into the gas is incorporated via the P-1 model [9–10]:
The value of f is calculated from the solution of a timeaveraged transport equation: A A A lt A f¯ ðq ¯f Þ þ ðqui f¯ Þ ¼ ð21Þ þ Sm At Axi Axi rt Axi
R1 ¼ C1
½ðTp þ Tl Þ=20:75 Dp
ð12Þ
and the kinetics rate: E R2 ¼ C2 exp − RTp
ð13Þ
The kinetics rate incorporates the effects of chemical reaction in 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 R1 R2 ¼ −pD2p P0 dt R1 þ R2
ð14Þ
rT 4 jdðCjGÞ þ 4p a þ Ep −½a þ ap G ¼ 0 p
ð15Þ
where, Ep and αp are calculated from: Ep ¼ lim
V Y0
ap ¼ lim
V Y0
N X
epn Apn
n¼1 N X n¼1
epn
4 rTpn pV
ð16aÞ
Apn V
ð16bÞ
The quantity Γ in Eq. (15) is: C¼
1 3½a þ ap þ rp
ð17Þ
and σp is calculated from: rp ¼ lim
V Y0
N X n¼1
ð1−fpn Þð1−epn Þ
Apn V
ð18Þ
Zk −ZkO ZkF −ZkO
ð20Þ
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, ¯ f 2V , describing the interaction between chemistry and turbulence, is solved: ! 2V lt A f¯ rt Axi !2 A¯f e 2V þ Cg lt −Cd q f¯ Axi k
A ¯ A A 2V ðq f 2V Þ þ ðqui f¯ Þ¼ At Axi Axi
ð22Þ
where, σt, 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 fragmentation (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 devolatization gases take place as follows [1,13–23]: Heterogeneous reactions CðsÞ þ O2ðgÞ ⇒CO2ðgÞ
ð23aÞ
2CðsÞ þ O2ðgÞ ⇒2COðgÞ
ð23bÞ
CðsÞ þ 2H2ðgÞ ⇒CH4ðgÞ
ð23cÞ
CðsÞ þ CO2ðgÞ ⇒2COðgÞ
ð23dÞ
CðsÞ þ H2 OðgÞ ⇒COðgÞ þ H2ðgÞ
ð23eÞ
Homogeneous reactions 2COðgÞ þ O2ðgÞ ⇒2CO2ðgÞ
ð24aÞ
COðgÞ þ H2 OðgÞ ⇒CO2ðgÞ þ H2ðgÞ
ð24bÞ
COðgÞ þ 3H2ðgÞ ⇒CH4ðgÞ þ H2 OðgÞ
ð24cÞ
CH4ðgÞ þ 2O2ðgÞ ⇒CO2ðgÞ þ 2H2 OðgÞ
ð24dÞ
HCðgÞ þ 1:5O2ðgÞ ⇒CO2ðgÞ þ H2 OðgÞ
ð24eÞ
The decomposition and polymerization reactions of the superior and unsaturated hydrocarbons are also added: fragmentation
Superior HCðgÞ Y Inferior HCðgÞ þ CðsÞ Unsaturated HCðgÞ YSaturated HCðgÞ Polymerization
Unsaturated HCðgÞ þ H2ðgÞ Y Superior HCðgÞ 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
85
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 percentage of volatiles (∼ 1.5%) like H2 and N2, requiring a temperature 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 temperature 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 temperature 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 jets from 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 [21–29]. The heating rate of coal particles depends on their size and contact with the thermal source. For example, the heating rate of coal 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 experiments concerning rapid pyrolysis is not sufficient to determine the composition and distribution of intermediate products for various coals [30–32]. Thus, the mathematical models developed for devolatilization 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 volatiles–char–ash to char–ash 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 þ 1=2O2 →CO2 –283:2kJ=mol
ð25aÞ
H2 þ 1=2O2 →H2 O–242kJ=mol
ð25bÞ
CH4 þ 2O2 →CO2 þ 2H2 O–802:86kJ=mol
ð25cÞ
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 mathematical models. These were developed in order to describe the solid coal combustion and have found important application in reactions 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 CO2 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 (b 100 μm) are considered to burn according to the first model and larger (up to N 2 mm) according to the second. However, the two models constitute only the two extreme cases of char combustion and cannot, therefore, establish a general theory [29–36]. 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 of temperature 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, experimental data by Field [8] and Borghi [38] showed that the reaction of char–CO2 is very slow in comparison with the reaction of char–O2. Therefore, the latter can be considered as the main reaction on the char surface when the essential quantity of oxygen is available. However, the presence of CO2 cannot be ignored and, thus, there always exists the probability of parallel reactions [26,27,29–31]. Based on a comparative analysis of existing data
for coal combustion and on the constitution and granulometry of particles (average char diameter ≪100 μm), 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 coefficient Dim of oxidant in the porous char used in the present model was 5.0e–05 m2/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 interface 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 μm 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 the power 0.6. The value of the exponent (b 1) 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 endothermic. Its thermodynamic state is defined by the reaction enthalpy ΔH and the equilibrium pressure PCO2,eq,: ΔH ΔS þ PCO2;eq ¼ exp − ð26Þ RT R 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 temperature–equilibrium 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 [40–46] for the present endothermic reactions
D.K. Fidaros et al. / Powder Technology 171 (2007) 81–95
87
Fig. 3. Variation of calcination chemical kinetics with temperature.
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 ΔÇ900 = 1660 kJ/kg CaCO3 = 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 CO2, 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 temperature 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 decisive role [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]: 19; 680 7 Pe ¼ 1; 826 10 exp − ð27Þ T
Fig. 4. Calciner side view.
The reaction rate at the interface is expressed as follows, Borgwardt [39]: ð28aÞ
Rate ¼ −ks ACaCO3 where,
Ea ks ¼ Aexp − RT
ð28bÞ
The activation energy Ea of the decomposition reaction is in the range 165–205 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 decomposition rate by modifying it as proposed by Darroundi and Searcy [48]: ksV ¼ ks for Pb10−2 Pe
ð29Þ
ksV ¼ ks ðPe −PÞ=Pe for 10−2 Pe bPbPe
ð30Þ
The effect of temperature on calcination chemical kinetics is shown in Fig. 3. During calcination, the thermal conductivity of Table 1 Mass flow rates at the inlets of the calciner Case 2 — pet coke
Kind of mass flow
Case 1— coal Quantity [kg/s]
Percentage
Quantity [kg/s]
Percentage
mCaCO3 mCoal mTertiary Air mAir Coal Total
52.47 3.78 39.36 0.97 96.59
54.3% 3.9% 40.7% 1.0% 100.0%
52.47 3.17 39.36 0.97 95.98
54.6% 3.3% 41.0% 1.0% 100.0%
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Table 2 Ultimate analysis for the solid feeds Components
Humidity Volatiles Ash Fix carbon Total
Case 2 — pet coke
Case 1— coal As used (%)
Dry basis (%)
As used (%)
Dry basis (%)
1.45 27.70 12.05 55.80 97.00
0.0 28.11 12.23 56.62 96.96
1.45 13.21 0.14 81.70 96.50
0.0 13.40 0.14 82.90 96.44
lime depends on the state of the solid material and differs considerably 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 percentage of mass rate is that of CaCO3, followed by tertiary air, coal and finally the coal feeding air. The Rossin–Rammler distribution of the raw-mix size had an average value of d = 16.6 μm and a spread parameter of n = 0.822, while the coal had d = 34.5 μm 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
Fig. 5. Velocity distribution in vertical symmetry plane (left) and at 90° (right) for Case 1.
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Fig. 6. Temperature field in the vertical symmetry plane (left) and at 90° (right) for Case 1.
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. 7. Concentration distributions of CO2 (left), H2O (middle) and O2 (right) for Case 1.
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limiting calcination. The CO2 concentration decreases progressively towards the exit, mainly because the available CaCO3 is also decreasing. The amount of CO2 produced by coal combustion 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 where intense 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 combustion). 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 of Figs. 8 and 9, show the parts of the trajectories where the 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 ash by the gases. The average passage length of 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 1 reaches 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
Fig. 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).
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 5–6 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 are observed in the opposite side of the rawmix inlets, mainly due to the trapping of small coal particles while the concentration of CaCO3 particles is exceptionally small. The main body of the calciner is kept at temperatures well above the effective calcination temperature, so that large CaCO3 quantities 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 released. Also at the exit, higher temperatures are observed, because 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 calciner height 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 combustion of solid coal remains). Also, in the semi-cylindrical part, where intense calcination takes place, the concentration of O2 remains at low levels because of high raw-mix and CO2 concentrations, while in regions where coal particles have been trapped, still lower concentrations are observed. Higher concentrations 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 s
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Fig. 10. Velocity distribution in vertical symmetry plane (left) and at 90° (right) for Case 2.
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 temperatures 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 than that of the pet coke. The longer residence time corresponds to the particles that collide in the upper conical part (before the exit) and
Fig. 11. Temperature field in the vertical symmetry plane (left) and at 30° (right) for Case 2.
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Fig. 12. Concentration distributions of CO2 (left), H2O (middle) and O2 (right) for Case 2.
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 rawmix 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 those of 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 decomposition along the calciner height, starting from where calcification
Fig. 13. Details of fuel particle 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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is initiated until the device exit. The active calcination height for Case 1 is approximately 27.5 m. An amount of 50% of CaCO3 is calcinated in the first 6 m, while the calcification rate is stabilized to lower values for the next 16 m and is reduced in the last part, where gas acceleration is also observed. The active calcination height for Case 2 is also 27.5 m with differences observed mainly in the evolution of the calcification. An amount of 70% of CaCO3 is calcinated in the first 7.5 m, while the calcination rate gradually decreases with height. In the last 5 m, the calcination rate decreases further, due mainly to the increasing gas speed, despite the fact that the temperature is well above the calcination threshold. Fig. 14 shows also the suppressive role of increased CO2 concentration, which is of great importance for the performance of the calciner. In the region defined by the lower conical part and the raw-mix inlets, the temperature should be maintained at a specific range in order to avoid locally high CaCO3 production. This is related to the CO2 partial pressure that constitutes the basic controlling parameter of calcification. Although there are no detailed measurements to compare with the present numerical results, it should be noted that the model is capable to predict the range of gas temperatures (850– 900 °C) and the asymmetric distribution of the particles (coal and raw-mix) in the exit of the calciner, as observed in the Olympus cement plant. In Case 1 (coal fuel), the velocities at the exit of the calciner are in the same range as those measured (17–21 m/s) with higher values observed in regions of higher temperatures. There were no data available to compare the numerical results for Case 2 (pet coke fuel). 5. Conclusions A numerical model was presented for the prediction of the velocity, temperature and concentration fields of gases and of particle trajectories in an industrial low NOx calciner for cement production. Models were also included for radiation, chemical reactions and turbulence effects. From the results of the parametric study the following main conclusions can be drawn: The small recirculation region observed near the raw-mix and tertiary air inlets, increases the active length of the device and the particle residence time. The rapid calcination near the raw-mix inlet produces high local CO2 concentrations, which limits calcination. Higher temperatures are observed near the coal inlet where combustion of volatiles occurs. The high temperature regions observed along the calciner are due to coal particles with orbits that are not intermingled with the raw-mix, suggesting that more attention should be paid to its granulometry than to that of coal. The high unevenness of the Rossin–Rammler particle distribution used is mainly responsible for the small percentage of the non-calcinated CaCO3. Despite the energy surplus observed (more in Case 2 than Case 1) calcination could not be completed because of the high CO2 concentration released immediately after the entry. The upper conical part of the calciner causes an acceleration of the gases and particles, which reduces calcination. The resulting high temperatures can cause problems of thermal stresses, and can lead to erosive by-products from SiO2
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chemical activation and NOx formation. Small recirculations are also formed and act as a bumper wall for the particles, resulting to local high temperatures and erosion of the walls. Finally, the model was capable to predict the range of velocities and gas temperatures and the asymmetric distributions in the calciner exit, as observed in the Olympus cement plant of AGET Hercules. Nomenclature A pre-exponential factor Apn area of projection of n-particle Ap area of projection of a particle ACaCO3 CaCO3 concentration a1, a2 devolatilization coefficients or evaporation factors CD drag coefficient C1,2 constants Dp particle diameter D0 size constant Dim diffusion coefficient of the oxidant d diameter ¯ D mean diameter E, Ea activation energy Ep equivalent particle brightness epn brightness of the nth particle FD coefficient for drag force term fi additional force fpn scattering factor of nth particle f mixture fraction G incident radiation gi gravitational acceleration hfg evaporation latent heat h thermal convection coefficient k thermal conductivity kS surface reaction rate MD mass fraction of particles with diameter larger than D mv(t) sum of volatiles evaporated up to time t mpo initial particle mass mash mass of ash in the particle mo local fraction of gas oxidant mp particle mass N total number of particles in a volume V n size distribution parameter P pressure Pe local pressure P0 partial pressure of oxidant in the gas environment of particle PCO2,eq equilibrium partial CO2 pressure qr radiation flux R universal gas constant Re Reynolds number R1 surface reaction rate R2 surface devolatilization rate Sb stoichiometric coefficient of the reaction Sm source term T temperature Tp particle temperature Tpn temperature of n-particle
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t U Up U,V,W zk zKF zKO ΔS ΔH
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time gas velocity particle velocity time-averaged axial, radial and circumferential velocities mass fraction for the chemical element k mass fraction for fuel stream mass fraction for the oxidizer stream change of grammolecular entropy reaction enthalpy
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