Mathematical Modeling of a Single-Stage, Downward-Firing

Article pubs.acs.org/IECR

Mathematical Modeling of a Single-Stage, Downward-Firing, Entrained-Flow Gasifier Job S. Kasule,†,‡ Richard Turton,*,†,‡ Debangsu Bhattacharyya,†,‡ and Stephen E. Zitney† †

National Energy Technology Laboratory, U.S. Department of Energy, Morgantown, West Virginia 26507, United States Department of Chemical Engineering, West Virginia University, Morgantown, West Virginia 26506, United States

‡

ABSTRACT: Gasifiers are the centerpieces of coal-fired integrated gasification combined cycle (IGCC) plants. Mathematical models of gasifiers have been developed in recent literature to describe the physical and chemical processes taking place inside the reactor vessels. These models range from simple one-dimensional (1D) steady-state equilibrium models to higher-order, sophisticated, dynamic 2D and 3D computational fluid dynamics (CFD) models that describe coupled gas−solid hydrodynamics, heat and mass transfer, and reaction kinetics over the complex gasifier geometry. In the current work, a 1D steady-state model of a single-stage, downward-firing, oxygen-blown, slurry-fed, entrained-flow gasifier has been developed for use in the context of IGCC process simulation. In this mathematical model, mass, momentum, and energy balance equations for solid and gas phases are considered. The model includes a number of heterogeneous and homogeneous chemical reactions along with devolatilization and drying of the slurry feed. The solid−gas heterogeneous reaction rates are calculated using the unreacted shrinking-core model. A detailed model of the radiative heat transfer has been developed considering interactions between the solids and all internal gasifier surfaces (side wall, top, and bottom surfaces), as well as interactions between the surfaces themselves. No a priori wall temperature profile is assumed in this model. The heat loss from the gasifier wall to the environment is also considered in the energy balance equations. In slurry-fed gasifiers, recirculation near the inlet of the gasifier is promoted by rapid mixing of the slurry feed with a portion of the hot reaction products. This violent mixing results in a significant rise in temperature that helps in evaporating the water and devolatilizing the coal. The recirculation is achieved by appropriately designing the feed burner and feeding the oxygen through a swirling annular injector. In the current gasifier model, a heuristic recirculation model has been developed and the conservation equations have been appropriately modified. The equations describing the gasifier are formulated as a set of ordinary differential equations (ODEs) in Aspen Custom Modeler (ACM). The ODEs are discretized using finite differences, and the resulting highly nonlinear system of algebraic equations is solved using a Newton-type method. The gasifier model is then validated using pilot plant and industrial data. This paper presents a number of parametric studies that have been performed using the 1D steady-state gasifier model to provide insight into the gasifier performance as the inlet and operating conditions change. Results are presented as profiles for species concentration and gas, solid, and wall temperatures. The effect of coal feed types on composition are also presented. In addition, a radiant syngas cooler (RSC) model has been developed in Aspen Plus and coupled with the gasifier model, thereby enabling the RSC exit stream composition to be compared to available industrial data.

1. INTRODUCTION The integrated gasification combined cycle (IGCC) process is a promising option for power generation because of its higher efficiency and environmental advantages over conventional coal utilization technologies.1−4 The gasifier plays a key role in the IGCC process by converting solid carbonaceous fuels, such as coal, petcoke, or biomass, into synthesis gas or syngas (a mixture of mainly CO and H2). The hot raw syngas is cooled in a water quench or radiant syngas cooler (RSC) to recover heat for producing steam in a steam turbine. The synthesis gas is then cleaned and subsequently combusted in a gas turbine (GT) to generate electricity. It is critical that optimal operation of the gasifier is understood for efficient operation of the IGCC plant. The extremely intense conditions and low residence time inside the gasifier cause rapid heating and reaction rates, giving rise to multiple time scales for the physical and chemical processes. Numerical simulations can help in gaining insight and developing a deeper understanding of the optimal operation of the gasifiers. © 2012 American Chemical Society

Several mathematical models of entrained-flow gasifiers have been developed, ranging from simple one-dimensional (1D)5−10 and equilibrium11 models to sophisticated dynamic 3D computational fluid dynamics (CFD) models12−16 that describe coupled gas−solid hydrodynamics, heat and mass transfer, and reaction kinetics over the gasifier geometry. The higher-order models generally include additional details such as turbulence but are too computationally expensive to be used directly in operability and controllability studies. Thus lowerdimensional models are required for these applications. Varying degrees of simplifications have been made while developing steady-state 1D gasifier models. Some have considered only mass and energy balances while neglecting momentum balances.5−10,17 Most of these models have also assumed an arbitrary wall temperature profile. 6,7 The Received: Revised: Accepted: Published: 6429

September 16, 2011 April 2, 2012 April 11, 2012 April 11, 2012 dx.doi.org/10.1021/ie202121h | Ind. Eng. Chem. Res. 2012, 51, 6429−6440

Industrial & Engineering Chemistry Research

Article

Figure 1. Schematic of a control volume (CV) in the gasifier.

assumption of plug flow within the gasifier with no mixing and recirculation has been made in almost all the models except those by Ubhayakar et al.5 and Monaghan et al.18 Most of these studies assume that the feed enters the gasifier at temperatures greater than 550 K so that all the slurry makeup water has already evaporated and enters the gasifier as steam. Therefore, the devolatilization is assumed to be instantaneous. However, in reality, the slurry feed has to travel a finite length inside the gasifier before attaining the high temperature needed to initiate the volatilization. In this length, mixing and recirculation play a key role. This length depends upon the feed inlet conditions and can affect the overall residence time of the reactants inside the gasifier. Therefore, neglecting this phenomenon is not appropriate in models intended for studying the effect of disturbances in feed inlet conditions and for analyzing dynamics and control. The major objective of the current study is the development of a detailed model of a downward, entrained-flow, slurry-fed, oxygen-blown (GEE-Texaco type) gasifier that will be eventually used for dynamic studies. The model includes detailed energy balance equations for the reacting phases and the gasifier wall. Energy loss to the environment is also considered. No assumptions about the wall temperature profile are made. A heuristic mixing and recirculation model, simpler than that of Smith and Smoot17 and similar to that of Monaghan et al.,18 is included in the gasifier model to capture the initial energy transfer that promotes a stable flame zone within the initial section of the gasifier. This gasifier model is developed with the aim of incorporating it into an existing plant-wide simulation of an IGCC plant.4

• While carrying out the energy balances, potential and kinetic energies of the system are considered to be negligible in comparison to the thermal energy due to the high temperature in the gasifier. • No particle attrition is considered in the model. In the two-phase model, conservation equations for mass, momentum, and energy for each phase are developed and solved in conjunction with the required constitutive closure equations. These balance equations are written for a control volume (CV) in the gasifier as shown in Figure 1. A heuristic recirculation model is incorporated in the gasifier model during the initial processes of evaporation and devolatilization. This involves recirculation of a predetermined fraction (α) of the gas from the higher temperature combustion zone to the colder region (mixing zone) just after the gasifier entrance. The fraction (α) is the ratio of the recirculated gas to the inlet gas stream flow rates. The effect of the recirculation is accounted for on the mass and energy balances in the gas phase, while its effect on the momentum balance is neglected. Details of the model equations for the recirculation zone are given in the Appendix. 2.1. Continuity Equations. The continuity equations for the solid and gas phases respectively, are obtained as

2. GASIFIER MODEL DESCRIPTION The following assumptions are made in the gasifier model development: • The radial dispersions of mass, momentum, and energy are neglected. • The entrained-flow system is assumed to be very dilute in the solid phase such that the interparticle interactions may be neglected. The ash layer formed due to the reaction of the coal particle is assumed to remain on the particle surface, and consequently a shrinking-core model is assumed. • The ideal gas law is assumed to hold for the gas phase. • The temperature inside the solid particle is assumed to be uniform. • The ash is assumed to be inert, but its effect as a catalyst has been considered in the relevant kinetic equations.

where ρs and ρg are the densities of the solid and gas phases, respectively, ε is the void fraction in the gasifier, and Γs−g is the net rate of consumption of the solid phase (coal) by the heterogeneous reactions. The last two terms on the right-hand side of eq 2 account for the mass recirculated from the hotter combustion region to the colder inlet region as illustrated in the Appendix. The term mrg represents the mass that enters a CV, while the term mmg represents the mass that leaves a CV. Species conservation equations are written as

d(ρs (1 − ε)Us) dx

d(ρg εUg) dx

= −(1 − ε)Γs − g

(1)

= (1 − ε)Γs − g − mrg + mmg

(2)

d (ερ Ugy ) = R gi − mrg ygi + mmg ϖgi dx g gi

(3)

d ((1 − ε)ρs Ux s sj) = R sj dx

(4)

where eqs 3 and 4 are the species balances in the gas and solid phases, respectively, and ϖgi is defined in the Appendix. 6430

dx.doi.org/10.1021/ie202121h | Ind. Eng. Chem. Res. 2012, 51, 6429−6440


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Syamlal and Bisset.25 This simple phenomenological model predicts the yields of tar and some major gas components while preserving a strict elemental balance. This model is based on data such as proximate and ultimate assays, tar composition, etc., obtained from certain lab-scale experiments that characterize the coal. A number of assumptions are made in the model to determine the stoichiometric coefficients of the devolatilization and cracking reactions. For example, all the sulfur in the coal is converted to H2S while all the nitrogen is converted to NH3. The kinetic parameters of the above reactions/processes are given by Syamlal and Bisset.25 2.2.4. Heterogeneous Char−Gas Reactions. After coal devolatilization, the residual char, represented as carbon (C), can undergo any of the following heterogeneous reactions: char combustion, char−steam gasification, char−carbon dioxide, and char−hydrogen. These reactions are shown in Table 1.

The gas phase can be modeled as a compressible gas obeying an ideal gas law such that its density is calculated from the following equation: RTg 1 = ρg Pg

N

∑ (yi /MW)i i=1

(5)

where yi and MWi are the mass fraction and molar weight of the ith gaseous species, respectively, and N is the total number of gaseous species. 2.2. Chemical Reactions. The coal slurry mainly undergoes two broad reaction stages: the initial stage reactions (processes) that occur when the fresh coal feed is first heated and the subsequent combustion and gasification reactions. In some studies,6,7,19 the reactor is divided into several reaction zones, such as drying, devolatilization, combustion, and gasification zones, in which various sets of reactions are considered. No such restriction is made in the current model, thus eliminating the need to determine the interface between these zones that would otherwise result in a multipoint boundary value problem.20 Various reactions are considered to take place simultaneously within the gasifier with their rates determined solely by the conditions at each point in the reactor. 2.2.1. Initial Stage Processes. The initial stage reactions/ processes occurring when coal is heated are complex in nature and lead to a wide variety of products whose composition depends not only on the type of coal but also on the processing conditions. The three initial stage reactions are postulated as drying/water evaporation, devolatilization, and heterogeneous char−gas reactions. moisture → steam

Table 1. Solid Phase Reactions reaction

VM → αd tar + βdCOCO + βdCO2CO2 + βdCH4 CH4 + βdH2 H 2 + βdH2OH 2O + βdH2SH 2S + βdNH3 NH3 (7)

C+

steam gasification CO2 gasification H2 gasification

C + H 2O ↔ CO + H 2

C + CO2 ↔ 2CO

C + 2H 2 ↔ CH4

diff

where VM is obtained from the proximate analysis of the coal. The tar produced during the devolatilization reaction undergoes further cracking reactions:

Wen and Chaung6 Wen and Chaung6 Wen and Chaung6 Wen and Chaung6

ash

(

)

ksY

(9)

where Y = rc/Ro; rc is the radius of the unreacted core; Ro is the original radius of the particle; and kdiff, kash, and ks are the gas film diffusion coefficient, ash diffusion coefficient, and the surface reaction constants, respectively.6 The ash diffusion constant is obtained using the correlation given by Syamlal and Bisset.25

tar → αc FC + βcCOCO + βcCO2CO2 + βcCH4 CH4 + βcH2 H 2 + βcH2OH 2O + βcH2SH 2S + βcNH3 NH3 + higher hydrocarbons

⎛ ⎛2 ⎞ 1 2⎞ O2 → ⎜2 − ⎟CO + ⎜ − 1⎟CO2 ϕ ϕ⎠ ⎝ ⎝ϕ ⎠

char combustion

reference

In the char combustion reaction, ϕ is a mechanism factor that gives the ratio of CO2 to CO in the reaction products. This factor varies significantly with the temperature of the reaction.26 For the high temperature environment prevailing in slagging gasifiers, carbon monoxide is favored at higher temperatures while carbon dioxide is favored at lower temperatures.27 It has been generally accepted that, in the high temperature environment within the entrained gasifier (T > 1000 °C), the heterogeneous, char−gas reactions can be considered as surface reactions. Due to the dilute nature of the entrained gasifier, particle−particle collisions are less frequent and the ash layer formed may be assumed to remain on the reacting particle. Thus, it is reasonable to apply the shrinking-core model28 (SCM) to estimate the char−gas reaction rates. In this formulation, it is also assumed that the temperature is uniform throughout the particle. According to the SCM model, the overall reaction rate can then be written as 1 rate = 1 (Pi − Pi*) 1 1 1 + 1 − + 2 k k Y

(6)

2.2.2. Drying/Water Evaporation. The moisture entering the gasifier is assumed to be the combination of the makeup water in the coal slurry feed and the original moisture in the coal, as specified in the proximate analysis of the coal. The water evaporation model adopted in this study assumes that the two types of water exist as a single water phase and a single evaporation rate is used. The evaporation rate is based on the work of Rao et al.21 2.2.3. Devolatilization. During devolatilization, the dry coal is thermally decomposed to release the volatile matter (VM), leaving behind a high carbon residue generically known as char. This process is shown below:

+ higher hydrocarbons

stoichiometry

kash = kdiff (εash 2.5)

(8)

where FC is the fixed carbon. Various devolatilization models, with varying degrees of complexity, exist in the literature.22−24 In the current study, the devolatilization model used is taken from the MGAS model of

(10)

where εash is the voidage of the ash layer, Pi − Pi* is the effective partial pressure of the ith component (O2, H2, H2O, or CO2) in the gas participating in the gasification reactions and takes into account the reverse reaction effect, Pi is the partial 6431



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pressure of component i, and Pi* is the equilibrium pressure6 of reactant i. 2.2.5. Homogeneous Gas−Gas Reactions. A number of homogeneous reactions are possible in the gasifier, but this study is limited to the reactions shown in Table 2.

where dp and μg are the particle diameter and dynamic viscosity of the gas phase, respectively. 2.4. Energy Balance Equations. 2.4.1. Gas and Solid Phase Energy Balances. Equations 16 and 17 are written for energy conservation of the gas and solid phases: d(Ugερg Cp ,gTg)

Table 2. Gaseous Reactions reaction

CO +

k4 1 O2 → CO2 2 k5

CH4 + 2O2 → CO2 + 2H 2O k6 1 O2 → H 2O 2

H2 +

dx

reference

+ hw − g [Tw − Tg ]} − (1 − ε)

Westbrook and Dryer29

{egFg − sσ(Tg 4 − Ts 4) + hg − s(Tg − Ts)} gas‐phase reactions

Wen and Onozaki,31 Karan et al.32

keq2

CH4 + H 2O ←→ CO + 3H 2 keq3 1 3 N2 + H 2 ←→ NH3 2 2

d((1 − ε)ρs Us 2)

Friedrichs and Wagner34

dx

= −(1 − ε)

dx πDi 6 Fw − sσ[Tw 4 − Ts 4] + (1 − ε) = AR dp {egFg − sσ[Tg 4 − Ts 4] + hg − s[Tg − Ts]} solid‐phase reactions

+

(11)

Nu = 2.0 + 0.6Re1/2Pr1/3

(13)

where the drag coefficient, CD, is given by Rowe and Henwood39 as Re < 1000 (14)

The particle Reynolds number, Re, is given by Re = (1 − ε)ρg d p

|Ug − Us| μg

(17)

(18)

where Nu, Re, and Pr are the dimensionless heat transfer coefficient, particle Reynolds number, and Prandtl number, respectively. Because coal particles in a slurry feed are usually very small, the Reynolds number, Re, can be neglected and as a consequence Nu may be assumed to be equal to 2. 2.4.2. Wall Energy Balance. In order to determine the wall temperature, Tw, a detailed energy balance across the wall is carried out. This balance considers radiation between wall and solids, convection between wall and gas, radiation between wall and top and bottom ends of the gasifier, and the energy loss to the surrounding environment. At each wall element l, all the energy contributions to the CV are summed up and must be equal to 0 at steady state. The skin temperature of the wall is also one of the calculated variables in the following study. The gasifier wall is assumed to be composed of an inner layer comprised mostly of refractory, a middle insulation layer, and an outer steel layer of given thicknesses.18

3C Dρg (1 − ε)−2.65 (Ug − Us)|Ug − Us|

⎧ 24 ⎪ [1 + 0.15Re 0.687], C D = ⎨ Re ⎪ 0.44, Re ≥ 1000 ⎩

(1 − ε)( −ΔHk)rk

where the term 6/dp is the ratio of the surface area of a particle to its volume assuming spherical particles. The view factors Fg−s and Fw−g, used in the radiation terms, must correctly account for the change in area between the emitting and receiving bodies. Radiation transfer to the gas phase is neglected as the gas is assumed to be transparent. The various view factors between the surfaces that exchange radiative heat are calculated accordingly from Siegel and Howell.40 The last two terms in the gas balance equation (mrghrg and mmghmg) are the enthalpies of the recirculated gas. These terms are equal to 0 outside the recirculation zone. The Nusselt number, Nu, for the gas−solid interphase heat transfer is calculated according to the equation due to Ranz and Marshall41 as

(12)

4d p

∑ k

where fs is the drag force per unit volume of particles, Us and Ug are the solid- and gas-phase velocities, and Pt is the total pressure in the system, taken to be the same as the gas-phase pressure. The drag force per unit volume, fs, is given by the correlation of Arastroopour and Gidaspow38 as fs =

(16)

d(Us(1 − ε)ρs Cp ,sTs)

dPt + (1 − ε)ρs g dx

+ (1 − ε)fs

ε( −ΔHj)rj − mrg hrg

+ mmg hmg

Lindstedt et al.33

dPt + ερg g − (1 − ε)fs dx

∑ j

The kinetic parameters for the above reactions were obtained from the references in Table 2. The kinetics for the water gas shift reaction (WGS) were modeled as a combination of the catalytic rate of Wen and Onozaki31 and a noncatalytic rate. The latter is a slightly modified form suggested by Karan et al.32 2.3. Momentum Balance. The momentum balance equations for gas−solid systems have been developed previously35−37 and have been modified in the current work using some simplifications such as neglecting the shear stress and particle−particle interaction forces, to give the following balance equations: =−ε

6 dp

Westbrook and Dryer29

+

keq1

dx

πDi {egFw − gσ[Tw 4 − Tg 4] AR

Peters30

CO + H 2O ←→ CO2 + H 2

d(ερg Ug 2)

=

(15) 6432



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Table 3. Proximate and Ultimate Analysis (As Received) of the Various Coal Feeds proximate

ultimate

coal

moisture

VM

FC

ash

C

H

O

N

S

Pittsburgh No. 8 Illinois No. 6 PRB lignite

1.00 11.12 17.89 14.20

33.52 34.70 36.24 43.40

57.69 44.19 40.27 41.40

7.79 9.99 5.60 1.00

76.83 63.75 58.37 62.07

5.49 4.50 3.85 4.49

6.03 6.88 13.20 17.55

1.40 1.25 0.80 0.68

1.46 2.51 0.29 0.08

∑ (qconv,w− g + qrad,w− s + qrad,w − w l

l

all p ≠ l

study of Bhattacharrya et al.4 and correspond to the feed conditions given in a U.S. Department of Energy report42 on the Tampa Electric Company (TECO) gasifier. The results are then compared to the TECO data and to other results in the literature. The feed conditions to the gasifier indicate that the slurry makeup water enters the gasifier in the liquid state and that this water must be evaporated from the coal before any gasification can take place. Other model parameters such as the heat capacities for the solid and gas phases are obtained from the METC Gasifier Advanced Simulation (MGAS) model.25 The wall thermal properties including the thicknesses of the refractory, insulation, and outer steel layers are obtained from Monaghan et al.18 A schematic diagram of the GEE-Texaco gasifier including the radiant syngas cooler is shown in Figure 2.

+ qloss + qrad,w − top

+ qrad,w − bot) = 0

(19)

where qconv = h(Tw − Tg),

qloss = heff (Tw − Tsurr),

qrad,w − s = Fw − sσ(Tw 4 − Ts 4), qrad, w − w l

all p ≠ l

=

∑ ewFw − w σ(Tw 4 − Tw 4) l

p

l

p

(20)

p

2.5. Model Solution and Parameters. The system of ordinary differential equations for the gasifier model described above is discretized using the first-order backward finite difference method and then simultaneously solved in the Aspen Custom Modeler (ACM) environment using a Newtontype method with appropriate boundary conditions and a good initial guess. The system of equations was solved for a small length of the gasifer which was then increased stepwise until the entire length of the gasifier was reached. Since the gasifier model with coupled fluid flow, heat and mass transfer, and complex reactions proved difficult to converge when the discretization mesh was changed, obtaining a good initial guess for the newly introduced nodes was very crucial. The initial guess of the variables for these new nodes was taken as the values at the last grid points of the previous domain length. A uniform grid was used throughout the modeling process even though multiple grid sizes are possible in the solver environment. The system of equations was then solved iteratively until convergence was realized. The gasifier model is solved for the four different coal types shown in Table 3, and additional parametric studies are performed to gain more insights on the model behavior. Table 4 gives some additional model input parameters that were selected to closely match the feed conditions in the IGCC Table 4. Additional Sample Model Parameters and Input Conditions Used in the Simulation parameter gasifier length (cm) gasifier inside diameter (cm) particle diameter (μm) emissivity of gas, particle emissivity of wall gas−wall heat transfer coeff (kcal/h·m2·°C) input conditions gasifier pressure (atm) coal feed rate (g/s) particle diameter (μm) all feed temperature (K) oxidant composition

value

Figure 2. Schematic of the GEE-Texaco gasifier with RSC considered in the study.

350 152 100 0.9 0.78 122

3. RESULTS AND DISCUSSION The main results of the current gasifier modeling work are summarized below. These results are in the form of model validation results, temperature profiles, species concentration profiles, and additional sensitivity studies. 3.1. Model Validation. Using Illinois No. 6 coal type, the gasifier model results are validated by comparing two sets of experimental data: Texaco pilot plant data42 and industrial data43 from the Tampa Electric Company (TECO). The

54 60 000 100 303 95% O2, 5% N2 6433



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Bhattacharrya et al.4 that certain reactions, most importantly the WGS, continue to take place within the RSC. In order to reconcile this inconsistency, the operation of the RSC was also accounted for in this work as illustrated in Figure 2. The RSC was modeled as a plug flow reactor in Aspen Plus with the inlet conditions identical to those exiting the gasifier. The dimensions and configuration of the RSC were obtained from Robinson and Luyben.45 Similar to the study by Robinson and Luyben,45 a multiple tube reactor configuration with a constant coolant temperature (336 °C) was used and only the WGS reaction was modeled in the RSC. The exclusion of other reactions in the RSC is reasonable because other gas species, such as CH4, have very low partial pressures for reaction rates such as the methane reforming to have any significant effect on the overall gas composition. In fact, the results remained unchanged when both the WGS and methane reforming reactions were modeled within the RSC. The results from the RSC modeling are shown in Figure 4, and it is clearly seen that the WGS reaction continues to take

validation runs were carried out with the gasifier dimensions and operating conditions shown in Table 5. Table 5. Conditions for the Validation Runs conditions gasifier configuration outer diameter (cm) internal diameter (cm) length (cm) operating conditions coal feed rate (kg/s) oxygen/coal ratio water/coal ratio pressure (atm)

Texaco pilot plant

TECO

152 15.8 330

400 179 662

0.1875 0.90 0.61 24

40 0.82806 0.4108 26

The Texaco pilot plant gasifier is internally divided into two sections: a partial oxidation zone and a quench section for cooling the syngas stream. During the validation run, only the partial oxidation zone was considered and its length7 was taken to be 330 cm as shown in Table 5. The inner diameter could not be found in the open literature but was back-calculated based on the assumption that the residence time in the pilot gasifier was similar to that of the TECO gasifier, whose dimensions were available.43,44 The validation is shown in Figure 3.

Figure 4. Species composition profile along the RSC.

place within the initial 20−25% length of the RSC before being quenched by the cooling fluid. The RSC exit compositions are then compared to the TECO data as shown in Figure 5. Figure 5 shows good qualitative and Figure 3. Comparison of current model with pilot plant data (dry basis).

The comparison shows a general qualitative agreement of the model’s results with the pilot data. Quantitatively, the model’s results compare fairly with the pilot data. The methane concentration shows some mismatch, but the two values are very low and are well within generally acceptable values. The underprediction in the methane is attributed to the faster kinetics of the methane destruction reactions. Slower kinetics rate schemes are also available6 which end up overpredicting the methane concentration. Tuning of the reaction kinetics could have been done to match the pilot plant data, but no such attempt was made in this study. Other apparent mismatches in species concentration could be possibly attributed to the differences in the temperature profiles in the two gasifiers and consequently the carbon conversion. However, even without any tuning of the kinetic parameters, the model predictions showed an acceptable agreement with the pilot plant data. When comparing the current results with those from TECO, it should be noted that these industrial data are from the clean syngas stream downstream of the radiant syngas cooler (RSC), while the current results are the conditions exiting the gasifier prior to entering the RSC. Moreover, it is reported by

Figure 5. Comparison of TECO data with current model results at RSC exit (dry basis).

quantitative general agreement of the model predictions with the industrial data. The CO and CO2 concentrations are slightly higher in the current model than in the industrial data, and this is attributed to the higher carbon conversion in the current study (almost complete conversion for the conditions considered) than in the industrial case (∼98%). The lower methane as explained above is attributed to the faster methane destruction kinetics. Again, no tuning of reaction kinetics was 6434



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slightly leveling off at the point when the entire water is evaporated but the temperature is still below the devolatilization temperature. The volatile matter is then released almost instantaneously as the temperature reaches the devolatilization temperature. Figure 8 shows the profiles of the major gas species along the length of the gasifier. Previous studies6−8 have shown that there

carried out in the current study other than using reaction schemes as obtained from the open literature. 3.2. General Model Predictions. This section presents general model predictions in the form of temperature profiles, species concentration, and carbon conversion profiles and the effect of coal types on the exit gasifier product distribution. These runs were obtained with gasifier configuration and operating conditions shown in Table 2. For the use of Pittsburgh No. 8 coal, Figure 6 shows the gas and solid temperature profiles along the reactor length. At the

Figure 8. Major gas species concentration profiles along the dimensionless gasifier length (water/coal, 0.4; oxygen/coal ratio, 0.8; Pittsburgh No. 8 coal).

Figure 6. Temperature profile and carbon conversion along the gasifier length (water/coal, 0.4; oxygen/coal ratio, 0.8; Pittsburgh No. 8 coal).

is no coexistence of the combustible gas species components such as CO and H2 when sufficient oxygen is present. This is also observed in the current study. The small accumulation of CO and H2 at the front of the reactor is due to recirculation. However, this occurs at locations before the combustion temperature is reached in the gasifier. The results also show peaks in the CO2 and H2O concentration profiles at the point when the maximum temperature occurs in the gasifier. This is also consistent with previous results.6−8 As discussed previously, the current model assumes no a priori wall temperature profile. The temperature of the wall is calculated and its profile is shown in Figure 9. It should be

beginning of the gasifier, the solid temperature gradually increases and then levels off slightly at the point when all of the water is evaporated (at a scaled reactor length of approximately 0.05). Following another gradual increase, the solid temperature rapidly increases as the volatile matter, consisting of CO, H2, and CH4, evolves and is subsequently combusted, releasing additional energy. This temperature behavior is further illustrated in Figure 7. The corresponding carbon conversion

Figure 7. Profiles of volatile matter, moisture content, and temperature of the solids along the gasifier length (water/coal, 0.4; oxygen/coal ratio, 0.8; Pittsburgh No. 8 coal).

Figure 9. Calculated wall temperature profile (Tw) compared to the linear wall temperature profile.6

profile shows that there is no significant consumption of carbon until all the devolatilization takes place and the temperature rapidly increases due to the combustion of volatiles. However, after the volatile products have combusted, there is appreciable carbon conversion mainly due to the carbon gasification reactions. Figure 7 shows the profile of the mass fraction of the slurry water and volatile matter (VM) of the Pittsburgh No. 8 coal and the solid temperature profile. As the water starts to evaporate with the initial increase in temperature, the mass fraction of the volatile matter increases in the solid phase before

mentioned that, apart from the commonly assumed linear wall temperature profile,6,7 which is also plotted in Figure 9, there is no other known temperature profile in the literature to which the current profile could be compared. The wall is seen to experience an initial increase in temperature, attaining a maximum value and then steadily decreasing. This characteristic is remarkably different from the assumed linear temperature profile. In addition, the temperature gradient along the wall is 6435



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calculated to be much less steep than the assumed slope of the linear temperature profile. 3.3. Sensitivity Studies. A series of sensitivity studies was performed to gain more insight into the response of the gasifier to changes in certain key parameters. The studies include the effect of coal feed properties on the product distribution, the effect of water-to-coal and oxygen-to-coal ratios on the maximum temperature attained in the reactor, the effect of the recirculation ratio on temperature, and the effect of the water-to-coal and the oxygen-to-coal ratios on coal conversion. Figure 10 shows the product distributions obtained when different coal feeds are used at the same feed conditions. The Figure 12. Effect of water-to-coal ratio on maximum phase temperatures at oxygen/coal ratios (oc) of 0.65 and 0.8 (Illinois No. 6 coal).

Figure 10. Product distribution of exit gas for different coal feeds (water/coal, 0.4; oxygen/coal ratio, 0.8).

results show a marked variation in the gaseous product distribution but with CO, CO2, and H2 consistently produced as the major dominant species and CO having the highest composition in all cases. It is also interesting to note that the maximum phase temperature, particularly that of the solid phase, varies markedly with coal type as seen in Figure 11. These results highlight the strong dependence of the gasifier response on the coal type.

Figure 13. Effect of oxygen-to-coal ratio on maximum phase temperatures at water/coal (wc) ratios of 0.4 and 0.5 (Illinois No. 6 coal).

from 0.4 to 0.5. As expected, increasing the water-to-coal ratio decreases the maximum temperatures while increasing the oxygen-to-coal ratio increases the maximum temperatures. In addition to the above observations, the position of the maximum temperature is seen to shift downstream from the inlet of the gasifier with increasing water-to-coal ratio. This result is similar to that shown by Vamvuka et al.9 The higher this ratio is the longer it takes to evaporate the water and hence the ignition point tends to drift away from the gasifier inlet as seen in Figure 14. A shift in the position of the maximum temperature seen in Figure 14 is similarly observed when the model is run at different recirculation ratios as seen in Figure 15. However, the recirculation ratio has no appreciable effect on the maximum values of the temperature.

Figure 11. Maximum gas and solid temperatures for different coal feed types.

Figures 12 and 13 show the co-effect of the water-to-coal and oxygen-to-coal ratios on the maximum temperatures attained by the solid and gas phases inside the gasifier for Illinois No. 6 coal. In Figure 12, as the water-to-coal ratio increases, the maximum temperatures of both solid and gas phases decrease at any fixed oxygen-to-coal ratio. In addition, by increasing the oxygen-to-coal ratio (for example from 0.65 to 0.8), the maximum temperature of each phase increases. This is further illustrated in Figure 13, in which the maximum phase temperatures increase with increasing oxygen-to-coal ratio. However, at a fixed oxygen-to-coal ratio, the maximum temperatures decreased as the water-to-coal ratio increased

Figure 14. Effect of water-to-coal ratio (wc) on solid temperature (oxygen/coal, 0.8; coal, Illinois No. 6). 6436



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oxygen-to-coal ratio significantly affects the carbon conversion more than the water-to-coal ratio. Figures 18 and 19 show the co-effects of water-to-coal and oxygen-to-coal ratios on the composition of the major gas

Figure 15. Effect of recirculation ratio (α) on solid temperature.

Additionally, the effect of the recirculation ratio on the composition of the exit gas species is shown in Figure 16. The results show that this ratio has a negligible effect on the gas composition. Figure 18. Product gas composition as a function of oxygen-to-coal ratio (broken line, water-to-coal ratio = 0.55; solid line, water-to-coal ratio = 0.4).

Figure 16. Effect of recirculation ratio on the main gas composition.

Figure 17 shows the effect of the water-to-coal and oxygento-coal feed ratios on carbon conversion. At a fixed oxygen-toFigure 19. Product gas composition and carbon conversion as a function of water-to-coal ratio (solid line, oxygen-to-coal ratio = 0.8; broken line, oxygen-to-coal ratio = 0.9).

products. In Figure 18, CO and CO2 concentrations are seen to increase while the H2 concentration decreased with increasing oxygen-to-coal ratio. However, at any fixed oxygen-to-coal ratio, the concentrations of CO2 and H2 increased while that of CO decreased with increasing water-to-coal ratio as shown by the two plots with water-to-coal ratio values of 0.55 and 0.4. This is due to mainly the competing char-gasification and WGS reactions. The decrease in CO and increase in H2 and CO2 are consistent with the direction of the WGS reaction equilibrium. In Figure 19, increasing the water-to-coal ratio at fixed oxygento-coal ratio leads to an increase in H2 and CO2 concentrations and a decrease in CO concentration. Again this is consistent with the direction of the WGS reaction equilibrium. Increasing the oxygen-to-coal ratio (for example from 0.80 to 0.90) at any fixed water-to-coal ratio shows a decrease in H2 and an increase in CO2 and CO concentrations. An increase in oxygen content results in more combustion of H2 and carbon, hence the observed reduction in H2 and increased CO and CO2, the major products of the char-combustion reaction. These results are due to the competing reactions (mainly char combustion, char gasification, and WGS) taking place in the gasifier as reported by Wen and Chaung6 and Govind and Shah.7 It is also

Figure 17. Effect of water-to-coal feed ratio on carbon conversion as a function of oxygen-to-coal ratio.

coal ratio, the carbon conversion is seen to decrease with increasing water-to-coal feed ratio. This is largely attributed to the decrease in temperature as the water-to-coal ratio is increased. Conversely, conversion significantly increases with increasing oxygen-to-coal feed ratio. This is due to the increase in temperature as a result of the additional energy added from the exothermic combustion reactions. With the water-to-coal and oxygen-to-coal ratios used in the model, carbon conversion of at least 99% is achieved with an oxygen-to-coal ratio of 0.8 and a water-to-coal ratio of less than 0.4. It is evident that the 6437



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apparent that the water-to-coal ratio has a larger effect on the gas product composition than the oxygen-to-coal ratio.

In the model, a fraction of the hotter gas phase stream from the recirculation section (2) of the gasifier is recirculated back to a colder region at the entrance of the reactor (region 1). The total mass recirculated (ṁ recir) is determined from the recirculation ratio (α), which is a parameter in the model, and the inlet gas feed flow rate (ṁ in):

4. CONCLUSIONS A steady-state model of an entrained, GEE-Texaco type, downward-flow gasifier is presented. In addition to mass, momentum, and energy balance equations for the solid and gas phases, the gasifier model includes heterogeneous char−gas and homogeneous gas−gas reactions and equations describing the drying and devolatilization processes for the slurry feed. The gasifier wall temperature profile is not assumed a priori, but rather is calculated from the detailed multisurface, multimechanism energy balance model. A heuristic recirculation model is considered to capture the initial energy transfer to the slurry feed within the “mixing” zone at the gasifier entrance. For validating the gasifier model, results are compared with the pilot plant data and the available industrial data from the TECO IGCC plant. The model predictions compared appreciably with the pilot plant data even without any tuning of the reaction kinetics. The model results compare satisfactorily to the TECO results when an RSC model is included at the gasifier outlet. The residual mismatch in the syngas composition is attributed to the higher carbon conversion in the current model. The gasifier model is also used to simulate the gasification of different coal types. The results show a strong dependence of the product composition and maximum phase temperatures on the type of the coal fed to the gasifier. A recirculation ratio of greater than 10.4% affects only the position at which the ignition occurs in the gasifier without having any major effect on the maximum phase temperature and the product gas composition in the gasifier. The optimum ranges for the water-to-coal and oxygen-tocoal ratios necessary for achieving at least 99% carbon conversion for the cases considered in the study are 0.3−0.4 and 0.8−0.9, respectively.

α = ṁ recir /ṁ in

(A.1)

The mass and energy balances in region 1 (mixing zone) and region 2 (recirculation zone) are modified to account for the recirculation as shown in eqs A.2 and A.3. ∂(ρg ε)

∂(ρg εUg)

+

∂t

∂x

∂(ερg Cp ,gTg)

= (1 − ε)Γs − g − mrg + mmg

(A.2)

∂(Ugερg Cp ,gTg)

+

∂x ∂t πDi 4 = {egFw − gσ[Tw − Tg 4] + hw − g [Tw − Tg ]} AR 6 − (1 − ε) {egFg − sσ(Tg 4 − Ts 4) + hg − s(Tg − Ts)} dp gas‐phase reactions

∑

+

ε( −ΔHj)rj − mrg hrg + mmg hmg (A.3)

j

where mrg and mmg denote the mass of the gas recirculating from the recirculation zone (section 2) and that recirculating into the mixing zone (section 1), respectively. The corresponding terms for the energy associated with the recirculating streams into or out of regions 1 and 2 are mmghmg and mrghrg, respectively, for which the specific enthalpy terms are calculated as hmg =

■

APPENDIX: RECIRCULATION MODEL This section briefly describes the heuristic recirculation model incorporated in the current study. In an industrial gasifier, two regions, namely mixing and recirculation, are known to exist at the immediate entrance of the gasifier. These two regions are crucial in maintaining stability of the ignition front as well as aiding in the initial drying and devolatilization processes. The recirculation model is thus used in the current model to mimic these two sections, shown in Figure 20 as regions 1 and 2 of lengths L1 and L2, respectively. This model is based on the mixing and recirculation model seen in previous studies by Smith and Smoot17 and Monaghan et al.18

m

1 m

∑ hk k=1

(A.4)

N

h=

∑ yh i i

(A.5)

i=1

hi = h0 +

∫T

T

Cp dT

o

(A.6)

where m represents the total number of control volumes/cells in the recirculating region (section 2) and i represents the gas species. hrg is the same as h of the control volume under consideration within section 2. mmg and mrg are calculated as mmg = ṁ recir /(AR L1)

(A.7)

mrg = ṁ recir /(AR L 2)

(A.8)

where AR is the cross-sectional area. The corresponding species balance equation in the above regions is written as ∂ ∂ (ερ y ) + (ερ Ugy ) = R gi − mrg ygi + mmg ϖgi ∂t g gi ∂x g gi

(3a)

where

ϖgi = Figure 20. Illustration of the recirculation model. 6438

1 n

n

∑ ygi ,k k=1 dx.doi.org/10.1021/ie202121h | Ind. Eng. Chem. Res. 2012, 51, 6429−6440


Article

Γg−s = rate of char consumption (g/cm3·s)

ygi is the mass fraction of the gas species in control volume k within the length L1−L2 and n is the total number of control volumes in length L1−L2. The lengths L1 and L2 are parameters in the model. A recirculation ratio of 12−16% was found to be sufficient for initiating the ignition within 10% of the gasifier length for all the coals studied here, and a recirculation ratio of 14.8% was used in most of the runs in this study.

■

Subscripts

w, g, s = wall, gas, and solid, respectively rg = recirculated gas flow out of recirculation gas mg = recirculated gas flow in mixing zone Superscript

■

AUTHOR INFORMATION

Corresponding Author

* = equilibrium value

REFERENCES

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*E-mail: [email protected]. Tel.: (304) 293-9364. Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This technical effort was performed in support of the National Energy Technology Laboratory’s ongoing research in Process and Dynamic Systems Research under RES Contract No. DEFE0004000.

■

NOTATION At = cross-sectional area of reactor (cm2) Cp,g, Cp,s = specific heats of gas and solid, respectively (cal/ g·K) CD = drag coefficient Di, Do = inside and outer gasifier diameters respectively (cm) eg = emissivity of gas Fw−s = view factor of solid from the wall Fw−g = view factor of gas from the wall Fg−s = view factor of gas from solid ΔHk = enthalpy of kth reaction (cal/g) ΔHj = enthalpy of jth reaction (cal/g) fs = drag force on solids per unit volume of particles (N/ cm3) g = acceleration due to gravity (cm/s2) hg−s = convective heat transfer coefficient between gas and solid (cal/cm2·K·s) hw−g = convective heat transfer coefficient between wall and gas (cal/cm2·K·s) Pg = total pressure of gas (atm) R = gas constant (cal/g-mol·K) Rgi = net rate of generation or consumption of gas species i due to chemical reactions (g/cm3·s) Rsi = net rate of generation or consumption of solid species i due to chemical reactions (g/cm3·s) xi = mass fraction of solid species i yi = mass fraction of gas species i MWi = molar weight of gas species i (g/mol) Tw = wall temperature (K) Ts = solid-phase temperature (K) Tg = gas-phase temperature (K) Ug = gas-phase velocity (cm/s) Us = solid-phase velocity (cm/s) Kash = ash film diffusion constant (g/cm2·atm·s) Kdiff = gas diffusion constant (g/cm2·atm·s) Ks = surface reaction constant (g/cm2·atm·s)

Greek Symbols

ε = void fraction σ = Stefan−Boltzmann constant (cal/cm2·s·K4) ρ = density (g/cm3) α = recirculation ratio 6439



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