1

Modeling of Volatiles Condensation and Re-evaporation in Coal Gasification in a Low-Temperature,

High-Pressure Entrained-Bed Reactor

Ming-Hong Chen and Yau-Pin Chyou*

Chemistry Division

Institute of Nuclear Energy Research

Atomic Energy Council

Longtan, Taoyuan, Taiwan (R.O.C.)

Ting Wang

Energy Conversion & Conservation Center

University of New Orleans

New Orleans, Louisiana 70148, USA

ABSTRACT It has been reported that a newly introduced Entrained

Slagging Transport Reactor (E-STR) gasifier can reduce

capital and operating costs, while improve operability. E-STR

is a product modified from the existing E-Gasifier design. The

major modifications include increased operating pressure and

reduced temperature, aiming for producing higher H2/CO ratio

syngas for SNG production. Due to the reduced operating

temperature, catalyst is used to convert tar or heavy volatiles

to lighter gases.

Considering the lower operating temperature in an

E-STR gasifier, the objective of this study is to modify the

existing E-gasification simulation model by incorporating a

tar condensation and re-evaporating sub-model. In the

present study, the actual thermal property values of the

volatiles are not known, so different condensation and

re-evaporation tempertaures are used for paramtric study to

test the functionality of the newly implemented volatile

condesation and re-vaporation sub-models. The result shows

that incorporating the tar condensation model leads to a

formation of about 6.47 % liquid volatiles and an exit

temperature increase about 135 K, due to the release of latent

heat. When the re-vaporization model is implemented, the

liquid volatiles evaporate with an amount depending on the

setpoint of the evaporation temperature. It is concluded that

the volatile condensation and re-evaporation sub-modles have

been successfully developed and implemented. These

sub-models will be useful in the condition that the gasifier

temperature is intentionally kept low like the E-STR gasifer.

The results of E-GasTM

and the preliminarily E-STR

model are compared. Similar syngas compositions and outlet

temperature are observed between E-GasTM

and E-STR

without char recycling. For E-STR, higher interior

temperature is obtained near the second-stage injector, since

all coal slurry is fed at the 2nd

stage rather than a separate

injection as in E-GasTM

. This also leads to longer particle

resident time at the 2nd

stage for E-STR. Preliminary

modification on gemoetry and fed configuration results in

completed comsumption of char. However, for E-STR,

due to the changed gaisifer configuration and operating

condition, some char is unreacted and recycled. Therefore, the

reaction mechanisms of char are modified.

Several attempts have been made to reduce the reactor

temperature by chaning the feeding mass distribution between

the first and second stages, the oxygen/coal ratio, and the

operation pressure. The case using higher operation pressure

(70 atm) results in higher temperature at the 2nd

stage as well

as higher mole fraction of H2 and CO in exit syngas and

similar outlet temperature. It can be concluded that reducing

the oxygen/coal ratio is an effective way to reduce the reactor

temperature. The results indicate that high pressure and less

oxygen fed produce a more favorable higher H2/CO ratio for

synthetic natural gas (SNG) production.

Key words: gasification CFD modeling, entrained-bed reactor,

synthetic natural gas

1. INTRODUCTION

1.1 Background

The massive utilization of fossil fuel led to a significant

increase of carbon dioxide in atmosphere and expedites global

warming. Therefore, controlling and suppressing the

production of greenhouse gas are crucial issues for the

development of a sustainable environment. Coal is still the

most abundant source in traditional fossil fuels, and it is

expected to be employed for over 100 more years [1]. Due to

its low cost and stable supply, the USA and other industrial

countries have proposed the concept of Clean Coal

Technology (CCT). The Integrated Gasification

Combined-Cycle (IGCC) is an environment friendly coal

utilization technology, and it is also one of the most important

methods in the CCT. This technology has been gradually

commercialized, and more advanced developments continue

globally. With a multi-fuel feeding feature, IGCC can be

employed to produce power more efficiently with less

pollution than a traditional coal- burning power plants.

Combining the carbon capture and storage (CCS)

technology will make it the Sustainable IGCC (SIGCC), and it

will also be an essential option for the base-loaded power. If

integrated with hydrogen technologies, i.e., hydrogen turbine

and fuel cell, the SIGCC can achieve the vision of zero

emission.

There are five major IGCC demo plants globally, i.e.,

Wabash River and Tampa in USA, Buggenum (Nuon) in

Netherlands, Puertollano (Elcogas) in Spain, and CCP in

Japan. Abundant operating experience has been accumulated

from these five demo plants. All the gasifiers used in these

demo plants are entrained-bed gasifiers, and each has its

individual pros and cons. Besides the air-blown methods used

in MHI, Japan, the oxygen-blown method is employed for the

rest of the four plants. The gasifiers of Wabash River (E-Gas

gasifier) and Tampa (GE gasifier) employed coal-slurry

feeding, while those of Buggenum (Shell gasifier),

Puertollano (Prenflo gasifier) and CCP (MHI gasifier) all used

dry-fed. The downdraft flow condition is only employed in the

GE gasifier in Tampa. The E-Gas and MHI gasifiers feed coal

Proceedings of the 30th Annual International Pittsburgh Coal Conference, Beijing, CHINA, September 15 - 18, 2013

2

in two-stages, while the Shell, GE, and Prenflo use only

one-stage feed.

As mentioned earlier, coal gasification is an important

technology that can produce not only clean energy from

conventional coal resources, but it can also produce various

fuels and chemicals (Fig. 1). Popular coal-derived fuels

include gaseous fuels such as CO, H2, substitute natural gas

(SNG), methanol and liquid fuels such as diesel, jet fuels, and

methanol. The coal-derived chemicals, in turn, are as broad as

petroleum-derived chemicals. The most popular chemicals are

ammonia, urea, dimethyl ether, ethylene, and propylene (Fig.

1).

The Institute of Nuclear Energy Research (INER) in

Taiwan has been interested in developing CFD capability to

simulate the gasification process. Previous studies

successfully established the CFD capability to simulate

gasification process in an E-Gas type entrained-flow gasifier

to produce syngas for power generation [2]. To further

investigate different products that can be produced by coal

gasification, the gasification model needs to be continuously

expanded and improved by incorporating more complex and

detailed sub-models.

Coal gasification is an extremely complex process. It

will take many years of continuous effort to achieve

meaningful and practical results, and many years afterward for

continuous improvements. The object of this study is to model

and simulate the low-temperature and high-pressure Entrained

Slagging Transport Reactor (E-STR) introduced by the E-Gas

group [3]. The CFD gasification model developed in the

previous study [2] will be modified to simulate the newly

developed E-STR gasifier.

Carbon Source

Gasification

Synthesis Gas

Iron Reduction

H2 Ammonia

Dimethyl Ether

Ethylene &

Propylene

Oxo Chemicals

Polyolefins

Methanol

Fuel/Town Gas

Methyl Acetate

Acetic Anhydride

Power & Steam

Synthetic Natural Gas

Syngas

to

Liquids

Naphtha

& Waxes

Alcohols

Diesel/Jet/Gas Fuels

Acetic Acid

VAM Ketene

PVA Diketene

Acetate

Esters

Figure 1 Products (power, fuels, and chemicals) that can be

derived from gasifying coal or other hydro-carbon sources.

Therefore, modeling the entire gasification process in

E-STR, including the catalytic tar-cracking process, is

challenging.

E-STR (Fig. 2) is a product modified from an existing

E-Gasifier design. The major modifications include the

following features: (a) the horizontal cylinder in the

E-Gasifier is removed, (b) the oxidant is entirely fed at the 1st

stage with only the recycled char (no coal slurry), (c) the

coal/slurry mixture is entirely fed at the 2nd

stage, (d) the

operating pressure is much higher, and temperature is lower

than the counterparts in E-GasTM

gasifier, and (e) since the

second-stage temperature is not be sufficiently high enough to

crack the tar or heavy volatiles, catalytic material is needed

for tar/volatiles cracking. The main goal of E-STR is to

produce higher H2/CO ratio syngas suitable for SNG

production. Compared to the cross-type configuration of

E-GasTM

, the design of E-STR excludes the first-stage

cylinder. Therefore, several critical issues under high-pressure

operation can be simplified, such as welding, sealing,

maintenance and operation cost. The savings can be employed

to compensate the cost of the catalyst for the low-temperature

tar cracking process.

Figure 2 Configuration of the Entrained Slagging Transport

Reactor (E-STR) [3]

Choosing the E-STR gasifier in this study is (a) in our

previous studies [2], the E-Gasifier was investigated, and it is

more productive to continue to study the new product derived

from E-Gasifier; (b) E-STR is a new concept that is still in the

development stage. Investigating this new concept is

strategically important to keep INER in a lead position when

it comes to understanding and applying E-STR; (c) E-gasifier

and E-STR gasifier have certain advantages in producing

SNG over other commercial gasifiers. The result of this study

will pave the road for developing once-through SNG

production in the gasifier of the future.

1.2 Literature Review

In 1979, Wen et al. [4] investigated the Texaco (Now,

GE) down-draft entrainment pilot plant gasifier through their

mathematical model and assessed the influences of the input

operating parameters. They concluded: (1) an increase of

oxygen/fuel ratio significantly increases the carbon

conversion, resulting in slightly decreased H2 concentration

and increased CO concentration in the product gas; (2) an

increase in the steam/fuel ratio increases the H2 and CO2

concentration, but it decreases the CO concentration; (3) an

increase in the operating pressure can increase the degree of

carbon conversion particularly at high steam/fuel ratios; (4)

small fuel particles cause higher carbon conversion

efficiencies for a given mass feed rate.

3

Over the past few years, comprehensive research has

been devoted to the gasification process and performance via

CFD modeling [5-17]. Chen et al. [18] implemented a

three-dimensional model to simulate a 200-ton, two-stage air-

blown entrained type gasifier developed for an IGCC process.

The numerical analysis showed turbulent fluctuations

affected the volatiles and char-oxygen reactions and

significantly influenced the temperature and gas composition.

Furthermore, Chen et al. [19] also found carbon conversion is

independent of the devolatilization rate and less sensible to

coal particle size. However, it is sensitive to the

heterogeneous char-oxygen, char-CO2, and char-steam

reaction kinetics.

Since 2005, numerous attempts have been made by the

research team of the Energy Conversion & Conservation

Center (ECCC) at the University of New Orleans [20-26] to

develop a gasification numerical model through

ANSYS/FLUENT. The model is based on a two-stage,

entrained-flow gasifier following the geometry and input

parameters of Bockelie et al. [8] and Chen et al. [18]. The

ECCC team investigated various operating conditions, and

the results are summarized as follows [21]: (1) coal powder

feedstock tends to generate more CO and a higher exit

temperature than coal slurries with equivalent coal mass flow

rate; (2) the exit gases of oxygen-blown gasifiers have

higher mole concentrations of CO2 and CO than air-blown

gasifiers; (3) one-stage operation yields higher H2, CO, and

CH4 than a two-stage operation but with a lower syngas

heating value.

Silaen and Wang [20] found that the water-gas-shift

(WGS) rate plays an important role in predicting syngas

composition, and the reaction rate of the WGS, adopted from

Jones and Lindstedt [27], is found to be too fast because the

rate was obtained with the presence of a catalyst. Considering

that no catalyst is added in a typical gasifier, the water shift

reaction rate is purposely slowed down to make the syngas

composition consistent with syngas composition in the actual

production of a commercial entrained-flow gasifier with coal

slurry fed from the bottom. In a recent paper, when Lu and

Wang [28] calibrated the WGS reaction rate using the

Japanese air-blown, dry-feed CRIEPI gasifier, they

concluded that all of the originally published WGS reaction

rates were too fast. Therefore, by adding a backward WGS

reaction rate, the reaction speed doesn’t slow too much and

results in the same gas composition and temperature at the

gasifier exit as the case without adding the backward WGS

reaction rate. Moreover, they also clarified that the

calibrated WGS reaction rates obtained under air-blown and

dry-fed operating conditions may not be applicable to

slurry-fed or oxygen-blown gasifiers because the higher water

vapor concentration in slurry-fed gasifiers and higher

operating temperatures in oxygen-blown gasifiers may affect

the global WGS rate [28]. Chyou et al. [29] studied the effect

of 2nd stage design and coal type on the performance of

E-Gasifier. Results showed that dual-injector design can

minimize the circulation zone and improve the gasification

performance. Moreover, the tangential injection design saves

more energy as compared to the opposing-jets design. Lignite

in gasification leads to lower performance than Illinois #6

coal. However, the cheaper price of lignite makes it more

competitive in the condition of unit price of syngas HHV, and

thus the lower gasification performance is acceptable.

2. NUMERICAL SCHEME

The CFD solver used in this study is the commercial

CFD code ANSYS FLUENT V.12.0. FLUENT is a

finite-volume-based CFD solver written in C language and

has the ability to solve fluid flow, heat transfer, and chemical

reactions in complex geometries, and it supports both

structured and unstructured meshes. The second-order

discretization scheme is applied. The SIMPLE algorithm is

used in this study as the algorithm for pressure-velocity

coupling.

The 3-D, steady-state Navier-Stokes equations are

solved in an Eulerian-Lagrangian frame of reference. All the

coal particles are treated as a discrete, secondary phase

dispersed in the continuous phase via the Discrete Phase

Model (DPM) with the stochastic tracking to consider the

turbulent dispersion effect. The particles are tracked for every

50th

continuous phase flow iterations. The P1 radiation model

is used, and the gravitational force is considered in the

modeling. The standard k-ε model is used to capture the

turbulence flow. The species transport equations are solved

through the Finite-Rate/Eddy-Dissipation Model. In this

model, both the finite rate and the eddy-dissipation rate are

used and compared, and the slower rate is selected to compute

the continuous phase reactions.

2.1 Governing Equations

The equation for conservation of mass is:

mS )(

(1)

The source term Sm is the additional mass including the

vaporization of water droplets and the devolatilization of coal

particles added to the continuous phase from the dispersed

second phase flow.

The momentum conservation equation is:

Fgp

)()( (2)

where p is the static pressure, is the stress tensor, and

g

and F

are the gravitational body force and external

body forces (e.g. that arise from interaction with the dispersed

phase), respectively.

The stress tensor is given by

IT

3

2 (3)

where µ is the molecular viscosity, I is the unit tensor, and the

second term on the RHS is the effect of volume dilation.

The energy conservation equation is:

4

ph

ii

i

i

Sx

TK

xhu

x

(4)

where the second term on the RHS is the source term for

particle-gas heat transfer, evaporation energy (latent heat), the

radiation energy, and reaction heat.

2.2 Turbulence Model

The turbulence kinetic energy, k, and its rate of

dissipation, ε, are obtained from the following standard k-

transport equations:

k

jk

t

j

i

i

Gx

k

xku

x)( (5)

and

kCG

kC

xxu

xk

j

t

j

i

i

2

21)(

(6)

In these equations, Gk represents the generation of

turbulence kinetic energy due to the mean velocity gradients.

σk and σε are the turbulent Prandtl numbers for k and ε. The

turbulent (or eddy) viscosity, μt, is computed by combining k

and ε as follows:

ε

kCρμ μt

2

= (7)

the empirical model constants are C1ε = 1.44, C2ε = 1.92, Cμ =

0.09, σk = 1.0, σε= 1.3 [30].

2.3 Discrete Phase

In the Lagrangian approach, each particle is tracked.

The force balance on the discrete phase particles can be

written as

p

pxpD

p guuF

dt

du

)()(

(8)

where FD(u-up) is the drag force per unit particle mass:

24

182

eD

pp

D

RC

dF

(9)

uudRe

pp

(10)

where u is the fluid phase velocity, up is the particle velocity, μ

is the molecular viscosity of the fluid, ρ is the fluid density, ρp

is the particle density, and dp is the diameter of the particle

which is treated as a spherical shape. Re is the relative

Reynolds number calculated from the slip velocity between

the fluid phase velocity and the particle velocity.

The track of the particle is evaluated by the

instantaneous velocity, which is related to the mean velocity

and the fluctuating velocity as:

'* uuu

(11)

where the fluctuating velocity in Eq. (11) is defined by the

stochastic discrete random walk model.

The relationships for calculating the burning rate of the

char particles are presented and discussed in detail by Smith

[31]. The reaction rate R (kg/s) of particle surface species

depletion as a particle undergoing an endothermic reaction in

the gas phase is given as:

YRAR p (12)

N

nkinD

RpRR

(13)

where Ap is the particle surface area (m2), Y is the mass

fraction of surface species in the particle, η is the effectiveness

factor (dimensionless), R is the rate of particle surface species

reaction per unit area (kg/m2-s), pn is the bulk concentration of

the gas phase species (kg/m3), D is the diffusion rate

coefficient for the reaction, Rkin is the kinetic rate of reaction

(units vary), and N is the apparent order of reaction (In this

case, N is 0.5). The kinetic rate of reaction r is defined as:

)/( RTEn

kin eATR

(14)

When the coal slurry particle is injected into the

gasifier, the water is assumed to be atomized to small droplets

and undergoes evaporation. The rate of vaporization is

controlled by concentration difference between the surface

and gas streams, and the corresponding mass change rate of

the droplet can be given by:

CCkd=dt

dmsc

p 2 (15)

where kc is the mass transfer coefficient and Cs is the

concentration of the vapor at the particle’s surface, which is

evaluated by assuming that the flow over the surface is

saturated. C∞ is the vapor concentration of the bulk flow,

obtained by solving the transport equations. The value of kc

can be calculated from empirical correlations by [32, 33]:

330506002 ..d

cd ScRe..=

D

dkSh (16)

where Sh is the Sherwood number, Sc is the Schmidt number

(defined as ν/D), D is the diffusion coefficient of vapor in the

bulk flow. Red is the Reynolds number, defined as uν/D, u is

the slip velocity between the droplet and the gas and D is the

droplet diameter.

5

When the droplet temperature reaches the boiling point,

the following equation can be used to evaluate its evaporation

rate [34]:

pfgpdd ChTTC

dd=

dt

dm//1lnRe46.00.2 5.02

(17)

The energy equation for the particle accounts for

convection, radiation, devolatilization, and surface reactions,

is given as:

reacp

hfgp

pRppPpp

pp

Hdt

dmfh

dt

dm

TATThAdt

dTcm

44

(18)

where mp

and Tp

are the particle mass and particle

temperature, respectively. θR is the radiating temperature,

defined as:

4/14/ GR

(19)

and G is the incident radiation, which is related to the

radiative intensity as:

4

0IdG

(20)

The radiation in the gasifier is described by the P1

model [35]. The convection coefficient (Nusselt number, Nu)

between the particle and the gaseous flow field is evaluated by

the empirical relation proposed by Ranz and Marshall [32] as:

3/12/1 PrRe6.00.2 d

p

k

hdNu

(21)

2.4 Devolatilization

The devolatilization reaction of the coal particle is

described by the two-competing-rates model proposed by

Kobayashi [36] as:

)/(

111 pRTE

eAR

(22)

)/(

222 pRTE

eAR

(23)

These two kinetic rates are used as a weight function by

an expression for devolatilization as:

dtdtRRRRmmf

tm tt

ap

v

021

02211

0,0,

)(exp)()1(

)(

(24)

where α1 and α2 are yield factors, fω is the mass fraction of

moisture, mp is the mass of particle, and ma is the mass of ash.

The value of the constants are A1 = 2×105, A2 = 1.3×10

7, E1 =

1.046×108 J/kg mol, and E2 = 1.67×10

8 J/kg mol.

As the volatiles diffuse out of the coal particle, the

heterogeneous reactions on the particle surface occur. Those

particle surface reactions are described by the implicit

relations of Smith et al. [31] as:

r,kkrpr,k RYAR (25)

rN

r

rknrkinrk

D

RpRR

,0

,,,

(26)

RTErnrrkin

reTAR/,

,

(27)

2.5 Chemical Reaction

The conservation equation of the species is in the

following general form:

iiii SRJY

)(

(28)

where Ri is the net rate of production of species i by chemical

reaction and Si is the rate of mass creation by addition from

the dispersed phase sources. iJ

is the diffusion flux of

species i, which arises due to the gradients in concentration

and temperature.

The source term for species i due to all reactions is

computed as the sum of the Arrhenius reaction sources over

the NR reactions that the species participate in:

RN

r

riiwi RMR

1

,,ˆ

(29)

where Mw,i is the molecular weight of species i and riR ,ˆ

is

the Arrhenius molar rate of production/consumption of species

i in reaction r.

The Finite-rate/Eddy-dissipation-rate model computes

both Arrhenius rates and turbulent mixing rates, and the

smaller one is chosen for the homogeneous reactions. Only

finite rates are used for the heterogeneous reactions.

2.6 Condensation and Re-vaporization of Tar

As mentioned above, the second-stage temperature of

E-STR reactor will not be sufficiently high enough to crack

the tar, and a catalytic material is needed for tar cracking.

Without including the actual catalytic kinetics, this study

intends to implement sub-models of liquid tar (condensation)

and its re-vaporization into the existing gasification simulation

model, specifically it considers low-temperature operating

conditions and paves the road for future incorporation of

catalytic kinetics.

In the tar condensing sub-model, a threshold

condensation temperature is specified. When the local

temperature is lower than the designated condensation

temperature (assumed as 780 K), the gaseous volatiles will

condense to liquid form (liquid tar) and release the energy of

latent heat. On the other hand, when the temperature of liquid

tar is higher than the specified re-vaporization temperature

6

(assumed as 810 K), it will crack into three species (CO, H2,

C2H2) and absorb energy for cracking reaction.

The tar condensation and re-vaporization sub-models

are described by the UDF (User-Defined Function), and

interpreted into the E-STR model during the iteration.

2.6 Finite-Rate/Eddy-Dissipation-Rate Model

2.6.1 Arrhenius rate:

)RT/Eexp(TAk

]C[k)(R̂

r

n

rr,f

N

jjr,fr,ir,ir,i

r,j

1

(30)

where riri ,, ,

are the reactant and product stoichiometric

coefficients, respectively, r,j

are the rate exponents for

reactant j, in reaction r, n is the temperature exponent of

reaction r, Er is the activation energy, R is the universal gas

constant, Ar is the pre-exponential factor, and Cj is the molar

concentration of species j.

2.6.2 Eddy-Dissipation Rate:

Nj jr,j

P Pir,i

)P(r,i

Rr,R

Rir,i

)R(r,i

)P(r,i

)R(r,ir,i

M

Y

kABMR

)M

Y(

kAMR

)R,Rmin(R

(31)

where YR and YP are the mass fractions of reactant and

produce species, respectively. A is the Magnussen constant for

reactants (4.0), B is the Magnussen constant for products

(0.5), M is the molecular weight, and the R and P subscripts

are the reactants and products.

The global gasification reaction rates used in this study

are listed in Table 1 [22].

Table 1 Global reaction rate constants in this study

Reactions Ar Er (J/kmol)

C(s) + 1/2O2 → CO 0.052 6.1E+07

C(s) + CO2 → 2CO 0.0732 1.125E+08

C(s) + H2O → CO + H2 0.0782 1.15E+08

CO + 1/2O2 → CO2 2.2E+12 1.67E+08

H2 + 1/2O2 → H2O 6.8E+15 1.68E+08

CH2.761O0.264 → 0.479H2 + 0.264CO

+ 0.356CH4 + 0.19C2H2 Eddy-dissipation rate only

C2H2 + O2 → 2CO + H2

CH4 + 1/2O2 → CO + 2H2

CO + H2O → CO2 + H2 2.75E+02 8.38E+07

3. MODEL DESCRIPTION

3.1 Computational Model

Figure 3(a) showed the schematic diagram of a

previously developed E-Gas model, where its height (H) is 12

m, and length (L) is 8 m. 80% of coal/slurry and all oxidant

are fed on both side of 1st stage, while the remaining 20% of

coal/slurry is fed at 2nd

stage. In the E-STR model, the length

is shortened from 8 m to 2 m by literally removing the

horizontal cylindrical section, as shown in Figure 3(b). The

fuel feeding configuration is also modified by feeding all

coal/slurry at the 2nd

stage and all oxidant at the 1st stage. The

unreacted char leaves the reactor and recycles back to the 1st

stage fed port. A reasonable recycling rate will be deduced in

the following section.

(a) (b)

Figure 3 Schematic of the computational model of (a) E-Gas,

(b) E-STR

The grid sensitivity study was conducted with three

grids (94,683, 186,200, and 379,573). The difference of the

species between the cases of 94,683 and 186,200 is about

3.5%, and it reduces to 1.9% at the case of 379,573. The

difference of outlet temperature between the two cases is less

than 1%. To save computational time, the grid containing

186,200 tetrahedral meshes is used.

3.2 Boundary Condition

The Illinois #6 coal is used in this study, and its

properties are shown in Table 2 and Table 3, respectively.

Table 2 Proximate analysis of feedstock coal

Proximate analysis

(wt. %)

Illinois #6 Coal

wet

Moisture 11.12

Ash 9.7

Volatiles 34.99

Fixed Carbon 44.19

sum 100

HHV (MJ/kg) 27.1

L

H

D2

D1

Geometry

•D1 = 2m

•D2 = 1.6m

•H = 12m

•L = 8m

•V = 42.24m3

Coal

slurry

Oxidant

Coal

slurry

Oxidant

Syngas

Coal slurry

7

Table 3 Ultimate analysis of feedstock coal

Ultimate analysis

(wt. %)

Illinois #6 Coal

wet

Moisture 11.12

C 63.75

H 4.5

N 1.25

Cl 0.29

O 6.88

S 2.51

Ash 9.7

sum 100

To simplify the computational model, the masses of N,

Cl and S shown in the proximate analysis are lumped as ‘N2’

and treated as the gas phase fed together with the oxidant in

the model. The treatment of the ash transformation (i.e.

transformation to slag) is neglected; instead, SiO2 is used as a

gas species to represent ash in the model. For a real gasifier,

there is a hopper at the bottom of the gasifier to discharge the

slag. However, because of the assumption stated above, in this

study, both the physical behavior of slag formation and the

hopper design at the bottom are excluded in the simulation.

The total slurry mass flow rate is 39.7 kg/s, and the total

oxidant mass flow rate is 22.9 kg/s. Two operating pressures

(28 and 70 atm) are considered. Refractory brick is used in the

gasifier, and the wall is assumed to be adiabatic (i.e. perfect

insulation). The mass ratio of O2/Coal (DAF) is 0.92, and the

coal slurry concentration (defined as mass of coal (MF)/mass

of coal slurry) is 0.67. The detailed inlet and boundary

conditions and the corresponding parameter ratios of the base

case are summarized in Error! Not a valid bookmark

self-reference.. In the base case, no unreacted char is

recycled, and the operating pressure 28 atm is considered.

Table 4 Feedstock and operating conditions of the base case

Flow rate (tons/day)

- Coal (as received):

- Slurry water:

- Coal slurry:

2545

831

3376

First Stage Flow rate (tons/day)

- Oxidant:

1947

Second Stage Flow rate (tons/day)

- Coal Slurry Feed:

3376

Operating Condition:

- Pressure (atm):

- Inlet Oxidant Temp. (K):

- Inlet Feedstock Temp. (K):

28/70

411

450

Operating Parameters:

- O2/Coal (DAF):

- Coal (MF)/Coal Slurry:

- O2/C:

- H2O/C:

0.92

0.67

1.14

0.69

4. RESULTS AND DISCUSSION

4.1 Base Case

The results of E-Gas and preliminarily modified E-STR

model (decrease length, change fed configuration without char

recycling) are compared in this section. The operating

pressure of the base E-STR case is still the same as E-Gas at

28 atm. The purpose of making this base case comparison is to

find out the difference between the two by only changing the

geometry of the gasifier. The result, shown in Table 4,

indicates that similar syngas compositions and outlet

temperature are observed between E-Gas and E-STR without

char recycling. This implies that the lengths of both gasifiers

are sufficiently long to achieve almost equilibrium status at

the exit of the gasification process. When the E-STR is

operated at higher pressure (70 atm), the difference between

these two cases becomes obvious.

Table 4 Exit syngas temperature and compositions of base

E-STR case and E-Gas

Parameters E-Gas E-STR base

Mole

fraction

(%)

CO 40.1 44.1

H2 18.4 16.5

CO2 9.7 5.9

H2O 21.4 24.4

Exit temperature (K) 1891 1894

The comparison of temperature contours in Fig. 4 shows

that E-STR without char recycling and operating at 28 atm

results in higher interior temperatures near the 2nd

stage

injector than that of E-Gas. The fuel feeding configuration is

different; the fuel is separately fed in E-Gas, so its

temperature is lower; while all the fuel is fed at the 2nd

stage

in E-STR, so its temperature is higher.

(a) (b)

Figure 4 Comparison of Mid-plane temperature distributions

for (a) E-Gas model and (b) E-STR model

8

Comparison of particle resident times are shown in

Figure 5.1 Longer particle resident time is observed in

E-STR due to concentrated injection at 2nd

stage while it is

separately injected in E-Gas. Therefore, the particles in the

E-STR reactor require a longer time to be fully consumed.

Besides, concentrated injection generates stronger momentum,

racing coal particles impinging on the opposite wall from the

injector. In the case of E-Gas, however, particles with 20%

mass flow rate are push upward by the flow coming from 1st

stage without reaching the opposite wall.

(a)

(b)

Figure 5 Comparison of particle resident time (seconds) for

(a) E-Gas model and (b) E-STR model

In the present results, no unreacted char is seen leaving

the base E-STR reactor. Since in a real operation, a certain

percentage of char is unreacted at the gasifier exit

modification should be considered to reduce the char reaction

rate to match the operation data.

4.2 Recycling Rate and Char Reaction Rate for E-STR

Gasifier

In this section, the amount of recycled char is verified

for the E-STR gasifier. The recycling rate obtained from the

study of Watanabe and Otaka [12] in the CCP project is about

60%, while it is about 7.6% from the Wabash plant in the

NETL’s report [37]. We took the above two char-recycling

rates as the maximum and the minimum extent for this study.

Since this range is wide, two additional cases, 15% and 30%,

are also investigated and compared as shown in Table 5.

To approximately match the unreacted char at the

gasifier exit to the corresponding amount specified at the

recycle char inlet, the char reaction rates for 60% and 30%

recycle char cases are reduced (as listed in Table 5).

For the case of 60% char recycle, the extremely high

temperature was observed due to the interaction between the

recycled char and 95% oxygen. Such a high temperature will

make the reactor unable to sustain and should be avoided in

design. Besides, the CGE (Cold Gas Efficiency) was far

below the desired level (~80%). For the case of 30% char

recycle, its CGE is much better than 60%. The local hot spot,

however, is still observed. A reasonable temperature and

suitable CGE are achieved in the case of 15%. The last case

(7.6%) results in the highest CGE among compared cases and

uniform interior temperature at the expense of higher outlet

temperature (~ 2000 K), making it inconsistent with the

concept of the E-STR system. The overall gasification

performance can also be evaluated by comparing the HHV

values. Again, lower recycled char rates produce higher HHV

values as depicted in Table 5.

It can be concluded from the above discussion that the

case of the 15% char recycling rate led to reasonable CGE,

outlet temperature, and the situation of the hot spot is also

alleviated. Therefore, the condition of 15% recycling rate will

be employed in the following investigation.

Table 5 Comparison results under different char recycling rate

Char Recycle 60% 30% 15% 7.6%

CGE 42.6% 74.9% 84.1% 87.4%

Exit Temperature 1099 1300 1653 1981

HHV(MJ/kg) 5.86 8.58 9.05 9.00

Adjustment of

char reaction rates 1x10

-4 1x10

-4 5x10

-5 5x10

-5

Source [12] Adjust Adjust [37]

4.3 Condensation & Re-vaporization Sub-models

As explained earlier, the sub-models of condensation

and re-vaporization of volatiles are developed and

implemented through the user-defined functions.

Condensation:

Volatiles (gaseous) → Volatiles (liquid) + Heat

Re-Vaporization:

Volatiles (liquid) + Heat → CO + H2 + C2H2

9

The volatiles are first released according to the

two-competing-rates devolatilization model proposed by

Kobayashi [34] as shown in equations (22-24). Typically, the

devolatilization and the subsequent volatiles thermal cracking

occur and complete under high temperature near the

combustion region in the gasifier. In a low-temperature

operation condition, such as in the E-STR, the heavy volatiles

or tar vapors have not been thermally cracked to lighter gases.

Therefore, they would condense if the local temperature

becomes lower than the corresponding condensation

temperature of some species due to the endothermic

gasification process. Although the condensation temperature

varies among different compositions of volatiles and tars, for

simplicity, one single condensation temperature 780K is

assigned. The condensation is assumed to occur

instantaneously as the local temperature is lower than 780K

and thermal energy due to phase change will release. On the

other hand, a single re-evaporation temperature (810K) is also

assigned, and the re-evaporation is also assumed

instantaneously. The 30K difference between the

condensation and re-evaporation temperatures is arbitrarily

chosen for two reasons: (a) to avoid spatially sharp changes

from condensation to re-evaporation in very short distance

and (b) to simulate the resistance and finite time needed to

condense and re-evaporate. The second reason is further

explained below.

Assume the actual saturation temperature of the

volatiles is 795K – this means that the condensation and

evaporation temperatures are the same at 795K. Thus, the

assigned condensation (780 K) and re-evaporation (810 K)

temperatures are 15K lower and higher than 795 K

respectively. This implies that the volatiles need to be

sub-cooled 15K to effectively condense as well as superheated

15 K to re-evaporate. Typically, condensation will not occur in

the middle of no-where, rather a solid surface is required for

condensation to occur and grab on. A solid surface can be the

gasifier wall, duct/pipe surface, or a tiny solid (i.e. a nucleate)

surface. Since there are abundant particles in a gasifier

including char dust, ashes, and soot to serve as numerous

nucleate sites for condensation to occur, it is reasonable to

assume that condensation can occur at any computational cell

once the temperature of that cell drops below the assigned

condensation temperature. Since the rates of condensation and

re-evaporation are not known and finding their rates in not

within the scope of this study, ± 15 K is arbitrarily chosen

to consider the resistance involved in vapor diffusion and

convective process during the condensation and

re-evaporation (including the subsequent thermal cracking

process) processes. The accuracy of the assigned condensation

and re-evaporation temperatures is not the focused of this

study; rather, the development and implementation of these

sub-models are the goals of this study.

The volatiles condensation sub-model is first developed

and implemented in E-STR without including the

re-evaporation model. To investigate if the condensation is

adequately modeled, the result of condensation-only is

compared in Table 7 with the base case and not including

condensation and recycling of the unreacted char.

Table 7 Effect of condensation and re-vaporization sub-model

Base Condensation

only

Both

sub-model

Mole

fraction

(%)

CO 36.91 35.42 38.8

H2 24.48 21.46 26.7

CO2 10.07 10.43 8.11

H2O 20.61 19.98 18.33

CH4 2.27 1.29 2.45

Volatiles(G) 1.05x10-6 3.68x10-6 4.25x10-7

Volatiles(L) － 6.47 0

Unreacted Char (%) 20 11 18

Temperature (K) 1463 1598 1505

The comparison shows that incorporating the volatiles

condensation sub-model results in a formation of about 6.47%

condensed liquid volatiles and an increase of about 135 K in

exit temperature, due to the release of latent heat. The above

comparison indicates that the proposed volatiles condensation

sub-model successfully captures the phenomenon of the

volatiles condensation.

The result of the subsequent implementation of the

re-evaporation sub-model is also shown in Table 7. The

liquid volatiles from previous condensation are reduced to

zero due to re-vaporization in the base E-STR when no

unreacted char is recycled, and the gasifier's temperature is

high. The exit temperature also drops to the level of the base

case due to energy supply for the evaporation latent heat.

When both the condensation and re-vaporization sub-models

are included, the result returns to being similar to that of the

base case.These two sub-models are considered to be

successfully developed and implemented. Both sub-models

can also be improved to incorporate more complex physics or

be calibrated and fine-tuned to match experimental data, when

the data become available in the future.

4.4 Parametric Studies

To investigate other potential operating conditions, a

parametric study is performed by varying the fuel distribution

between the first and second stages, the operating pressure,

and the amount of oxygen supply (O2/Coal ratio).

4.4.1 Fuel Re-Distribution between the first and second

stages

In this case, the fuel distribution is changed by shifting

20% coal-slurry to the first stage (originally 100% at the

second stage) and feeding recycled char at the second stage

(originally at the first stage). The entire oxygen is still injected

at the first stage like the base case. Two alternate fuel

redistribution cases are studied – one with the condensation

and re-evaporation sub-models, the other without. Results of

syngas composition and exit temperature are compared in

Table 8.

10

Table 8 Effect of fuel-redistribution on syngas composition

and exit temperature with and without condensation and

re-evaporation sub-models

Base

2-stage fuel

with

Sub-models

2-stage fuel

without

sub-models

Mole

fraction

(%)

CO 36.91 36.42 34.41

H2 24.48 24.16 22.63

CO2 10.07 11.56 12.74

H2O 20.61 18.58 20.24

CH4 2.27 3.04 3.65

Volatiles(G) 1.05x10-6 4.47x10-7 4.76x10-7

Volatiles(L) － 0 -

Unreacted Char (%) 20 15 12

Exit Temperature (K) 1463 1615 1558

HHV(MJ/kg) 8.15 7.91 7.33

As shown in Table 8, the effect of fuel re-distribution on

syngas composition is noticeable but not large. The exit

temperature increases for the fuel re-redistribution cases. To

further examine the temperature distributions, the interior

temperature distributions along the height of reactor are

compared in Fig. 6.

Figure 6 Effect of fuel re-distribution on temperature

distributions with and without condensation and

re-evaporation sub-models

The variation of the exit temperature and the interior

temperature distributions are quite different. In the base case,

the average temperature in the 1st stage (with recycled char) is

about 2200 K. In the case of fuel re-distribution, there is a low

temperature region in the 1st stage since the recycled char is

injected at the 2nd

stage. With 80% coal-slurry and recycled

char, the temperature at the 2nd

stage is higher than that of

base case. Therefore, using fuel re-distribution leads to lower

first stage temperature and higher second stage temperature

but similar exit temperature. The temperature distribution of

the base case is more preferable because the wall temperature

is low, and the refractory can last longer.

4.4.2 Operating pressure

The operating pressure for E-STR reactor is 70 atm.

Previous calculations were performed under 28 atm. In this

section, the operating pressure is raised to 70 atm. The result

is compared with the base case in Table 9.

Table 9 Effect of operating pressure and oxygen amount on

syngas composition and exit temperature

Base 70 atm

100% O2

28 atm

50% O2

Mole

fraction

(%)

CO 36.91 40.81 21.89

H2 24.48 29.21 19.97

CO2 10.07 10.82 9.15

H2O 20.61 10.52 34.13

CH4 2.27 2.84 5.35

Volatiles(G) 1.05x10-6 1.3x10-9 6.1x10-5

Volatiles(L) － 0 0

Unreacted Char (%) 20 0 82

Temperature (K) 1463 1532 1100

HHV(MJ/kg) 8.15 9.38 5.73

Under the higher operating pressure at 70 atm, there is

no unreacted char at the exit of the gasifier, and the mole

fraction of CO increases about 4 percentage points. The

temperature at the second stage increases from 3,400K (in Fig.

6) to about 4,000K in Fig. 7 and the exit temperature from

1,463 K to 1,532K, indicating the effect of significantly

increased pressure. Please note that the adiabatic temperature

of solid carbon can reach about 4,667K under atmospheric

condition [38].

4.4.3 O2/Coal ratio

To further reduce temperature to match the value (750oC

or 1023 K) shown in Fig. 2 for high-pressure E-STR operating

condition, the amount of oxygen is reduced by decreasing

the oxygen/coal ratio from 0.92 in the base case to one half

the value (0.46). It is understood that a significant reduction of

oxygen will lead to poor carbon conversion rate and

unsatisfactory gasificaiton process. However, as an

exploration study, it is interesting to see how a significantly

reduced O2/Coal ratio would affect the gasification process

under a high operating pressure condition. The results are

compared in Table 9. It is shown that a reduction of 50%

oxygen supply effectively reduces temperature to 1100 K —

close to the level of design value at 1023 K. However, the

unreacted char increases to 80%, and the CO yield decreases

to 20%. The interior temperature distribution is compared

with base case in Fig. 7.

The case with 50% O2 leads to a higher temperature at

the 1st stage due to less mass introduced. After the 2

nd stage

injector, lower oxygen supply results in reduced combustion

and lower temperature all the way to the gasifier exit, as

shown in Fig. 7. The distribution of CO and H2 are compared

in Fig. 8. For the case of 50% O2, the mole fractions of CO

and H2 increase faster in the beginning, but remain constant at

11

20% in the rest of the gasifier after 5 m. In the base case, both

CO and H2 continuously increase to the position about 8 m.

Figure 7 Comparison of temperature distributions under two

pressures and two oxygen supply conditions

Figure 8 Comparison of CO, H2 and O2 distributions with half

oxygen-fed condition at 28 atm

Among the conducted parametric study, the reduction of

oxygen supply is the most effective way to lower the gasifier

temperature. When the interior temperature is low enough, the

effect of liquid volatiles/tar would be significant. The molar

ratio of H2/CO is 0.663, 0.716, and 0.912 for the base case, the

case of 70 atm, and the case of reduced oxygen, respectively.

The results indicate that high pressure and low oxygen supply

leads to a higher, more favorable H2/CO ratio for synthetic

natural gas (SNG) production.

However, the amount of condensed liquid volatiles/tar

seems lower under current simulation without using catalysts.

This may be due to the significantly reduced oxygen supply or

the incorrect values being assigned to the

condensation/re-vaporization temperatures. Further study will

examine more closely these two parameters.

5. CONCLUSION

The objective of this study is to modify the existing

CFD model, consider the low-temperature and high pressure

features of the newly introduced E-STR gasifier, aiming for

producing higher H2/CO ratio syngas for SNG production.

Due to its low-temperature, volatiles condensation and

re-evaporation sub-models are developed and implemented.

Due to its high operating pressure condition, different fuel and

oxidant feeding distributions are compared. The results are

summarized below.

Similar syngas compositions and outlet temperatures are

observed between E-Gas and E-STR without char recycling.

Higher interior temperature is obtained in E-STR because all

coal slurry is fed at the 2nd stage rather than being

separately injected as in E-Gas. This also leads to longer

particle residence time at the 2nd stage for E-STR.

Incorporating the volatiles condensation-only

sub-model results in a formation of about 6.47 % liquid

volatiles and an increase of outlet temperature about 135 K,

due to the release of latent heat. When the re-vaporization

model is incorporated together with the condensation model,

under the test condition, the liquid volatiles evaporate with an

amount depending on the setpoint of evaporation temperature.

It is thus concluded that the volatiles condensation and

re-evaporation sub-modles have been successfully developed

and implemented. These sub-models will be useful under the

condition that the gasifier temperature is intentionally kept

low, like in the E-STR gasifer.

An attempt was made to reduce the reactor

temperature by changing the fuel feeding mass distribution

between the first and second stage as well as changing the

oxygen/coal ratio. By introducing 20% coal-slurry at the 1st

stage and feeding the recycled char at the 2nd

stage, the

temperature is reduced near the 1st injector but is increased at

the 2nd

stage.

When the operation pressure increases from 28 atm

to 70 atm, the temperature increases at the 2nd

stage. The mole

fractions of both H2 and CO in the exit syngas increase, but

the outlet temperature maintains about the same.

Under the studied condition, the present model can't

predict reduced temperature inside the E-STR gasifier as

claimed by the manufacturer [3]. By reducing the oxygen/coal

ratio from 0.92 to 0.46 (i.e., with a 50% reduction), the

temperature inside the gasifier shows an increase near the 1st

stage but reduces in the rest of the gasifier with a reduction of

exit temperature from 1463 K to 1100 K.

The overall result indicates that operations with high

pressure at 70 atm and low stoichiometric ratio lead to

favorable higher H2/CO ratios suitable for producing

synthetic natural gas (SNG).

ACKNOWLEDGEMENT

This research is supported by a grant from INER

(Institute of Nuclear Energy Research) program “Clean

12

Carbon as Sustainable Energy (CaSE)” under the framework

of National Energy Program (NEP) at National Science

Council of Taiwan, ROC.

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