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Chemical Engineering Science 58 (2003) 867875

www.elsevier.com/locate/ces

Heterogeneous modeling for xed-bed FischerTropsch synthesis:Reactor model and its applications

Yi-Ning Wang, Yuan-Yuan Xu, Yong-Wang Li, Yu-Long Zhao, Bi-Jiang Zhang

State Key Laboratory of Coal Conversion, Institute of Coal Chemistry, Chinese Academy of Sciences, P.O. Box 165, Taiyuan 030001,

Peoples Republic of China

Abstract

A comprehensive one-dimensional heterogeneous reactor model is developed to simulate the performance of xed-bed FischerTropschreactors for hydrocarbon production. The detailed mechanistic kinetics is combined into the reactor model along with considering the fact

that the catalyst pores are lled with liquid wax under realistic conditions. The equilibrium between the gases in the bulk and the wax in

the catalyst pores is correlated by using a modied SRK equation of state (MSRK EOS). The model is solved by using Gear method to

integrate the reactor model with the embedded pellet model discretized by orthogonal collocation on nite elements. The validity of the

reactor model is tested against the measured data from dierent-scale demonstration processes. Satisfactory agreements between model

predictions and experiment results are obtained. Detailed numerical simulations are performed to investigate the eect of major process

parameters on the reaction behavior of xed-bed FTS systems with recycle operation.

? 2003 Elsevier Science Ltd. All rights reserved.

Keywords: FischerTropsch synthesis; Fixed bed reactor; Detailed kinetics; Pellet model; Heterogeneous reactor model; Recycle operation

1. Introduction

FischerTropsch synthesis (FTS), in which syngas is con-

verted into a wide product spectrum consisting of a complex

multi-component mixture of predominantly linear hydrocar-

bons and oxygenates, is a promising option for environmen-

tally sound production of transportation fuels and chemical

feedstocks from coal or natural gas. The various types of

reactors (including xed bed, uidized-bed, ebulliating-bed,

and slurry phase) have been considered in the history of FTS

process development, characterizing with the most suitable

particle size of the catalyst used. The xed-bed Fischer

Tropsch process, being one of the most competing reactor

technologies, occupies a special position in FTS industrial

practices, as persuasively exemplied by the large-scale

commercial operations of Sasol (Dry, 1996) and Shell

(Sie, 1998).

It is well known that the economic viability of syngas

conversion is determined by capital costs and average prod-

uct price. In this respect, the manufacture of syngas is by

far the most capital-intensive part of a gas conversion plant

Corresponding author. Tel.: +86-351-4130-337; fax: +886-351-

4050-320.

E-mail address: [email protected] (Y.-W. Li).

(Dry, 1996, 1999). Therefore, the FischerTropsch synthe-

sis step should aim at utilizing syngas as eciently as pos-

sible, and selectivity considerations are then extremely im-

portant in the design of the FischerTropsch synthesis sec-

tion (Geerlings et al., 1999). To achieve an optimum per-

formance for the complete process, the catalyst and the

reactor should be comprehensively optimized. Evidently,

due to the highly complexity of FTS reaction system, the

proper establishment of heterogeneous reactor model, from

which selectivity information as well as heat-transfer infor-

mation can be derived in a quantitative fashion, is of critical

importance.Typical industrial FTS processes with xed-bed reactors

normally produce complex mixtures consisting of hydrocar-

bons ranging from methane to wax. For the reasons of re-

ducing pressure drop and facilitating heat removal, catalyst

particles of a few millimeters in sizes are generally needed

to be used in xed-bed reactors, contributing to the existence

of intraparticle pore-diusion limitations (Sie, 1998). As a

result of diusion limitation and capillary condensation, the

catalyst pores are often lled with a stagnant phase formed

by the heavy waxy products (Anderson, Seligman, Schultz,

Kelly, & Elliott, 1952; Post, vant Hoog, Minderhoud, &

Sie, 1989; Zimmerman, Rossin, & Bukur, 1989; Madon &

0009-2509/03/$ - see front matter? 2003 Elsevier Science Ltd. All rights reserved.

doi:10.1016/S0009-2509(02)00618-8
mailto:[email protected]:[email protected]

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868 Y.-N. Wang et al./ Chemical Engineering Science 58 (2003) 867 875

Iglesia, 1994; Sie & Krishna, 1999). Therefore, the compre-

hensive modeling of the FischerTropsch xed-bed reactors

is closely related to the quantitative description of the in-

teractions between complex chemical kinetics and special

transport phenomena involved.

A few mathematical modeling studies on Fischer

Tropsch xed-bed reactors have been reported in literature(Atwood & Bennett, 1979; Bub, Baerns, Bussemeier, &

Frohning, 1980; Jess, Popp, & Hedden, 1999). Atwood and

Bennett (1979) have proposed a one-dimensional, hetero-

geneous plug ow model to investigate parameter eects

on large-scale commercial reactors. Later, Bub et al. (1980)

developed a two-dimensional, pseudo-homogeneous, plug

ow model, which was used to predict product distribution.

Recently, Jess et al. (1999)proposed a pseudo-homogeneous

two-dimensional model to conduct the conceptual design

of a xed-bed reactor for converting nitrogen-rich syn-

gas. However, all the investigations were based on the

traditional lumped kinetics, in which detailed selectivity

information is obscured. Moreover, the intraparticle diu-

sion eect has not been considered in the reactor models

by Bub et al. (1980) and by Jess et al. (1999), due to

the introduction of pseudo-homogeneous treatment. De-

spite the intraparticle diusion eect was not ignored by

Atwood and Bennett (1979), the oversimplied analytical

expression of eectiveness factor, which is derived from

the assumption that the kinetics is rst-order in CO, makes

the fundamental analysis of intraparticle diusion-reaction

behavior impossible.

With a view to realizing a more comprehensive descrip-

tion of FTS xed-bed reactor, we have systematically devel-

oped the detailed kinetics model (Wang, 2001; Wang, Xu,Xiang, Li, & Zhang, 2001), the generalized gaswax equi-

librium correlation (Wang, Li, Bai, Zhao, & Zhang, 1999),

and the diusion-reaction model for wax-lled catalyst pel-

let (Wang et al., 2001). As a continuing part of these pre-

liminary eorts, a comprehensive one-dimensional hetero-

geneous reactor model for xed-bed FischerTropsch syn-

thesis is constructed in this contribution. The validation of

the proposed model is performed by using experimental data

from the FTS processes at dierent scales. Detailed numeri-

cal simulations are carried out so as to reach a better under-

standing of the reaction behavior of xed-bed FTS systems

with recycle operation.

2. Development of reactor model

2.1. Model assumptions

In view of the operation characteristics of the xed-bed

FTS reactor system in industrial cases, the following as-

sumptions are introduced: (1) The pores of the catalyst

pellets in the reactor are assumed to be lled with waxy

products (Post et al., 1989; Sie & Krishna, 1999; Wang et

al., 2001), (2) the real state of the bulk containing a very

small amount of heavy hydrocarbons (wax) could be in a

saturated gas state with wax or in a dispersed colloid state

of wax fog, (3) at the interface between the catalyst pel-

let and the bulk gas phase, gasliquid equilibrium can be

assumed to be approached if the eect of the gas lm on

mass transfer can be neglected (Froment & Bischo, 1990),

(4) interparticle heat conduction and heat transfer by radia-tion are assumed to be less important than convective heat

transport.

2.2. Mathematical model

A one-dimensional and heterogeneous model is chosen in

this work to simulate a tubular xed-bed FischerTropsch

reactor. The mass and energy equations for the bulk gas

phase can be written as follows:

d(usci)

dz=

3

R3

P

p(1 B)

RP0

NR

j=1

ij Rjr2 dr

(i = 1; NC); (1)

usgCP;mdTg

dz=

3

R3Pp(1 B)

RP0

NRj=1

(Hj) Rjr2 dr

+ 4U

dt(Tw Tg): (2)

For the pressure drop, the following equation is used:

dP

dz=fp

gu2s

dp: (3)

The mass and energy equations for the catalyst pellets can

be formulated as follows (Wang et al., 2001):

De; i1

r2d

dr

r2

dcs;i

dr

=P

NRj=1

ij Rj

(i = 1;NPG); (4)

Ke1

r2d

dr

r2

dTs

dr

=P

NRj=1

(Hj) Rj: (5)

The initial conditions for the bulk phase and the boundary

conditions for the catalyst pellet are given as

z= 0: ci = ci;0; P=Pin; Tg = Tin; (6)

r= 0:dcs; i

dr= 0;

dTs

dr= 0; (7)

r=Rp: yi =LiVi

xi;dTs

dr=

hf

Ke(Ts Tg); (8)

where the mole fractions of component i in the gas phase

and in the liquid phase at pore mouth are yi = ci=ci and

xi = cs;i =cs;i , respectively.

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Y.-N. Wang et al./ Chemical Engineering Science 58 (2003) 867 875 869

Table 1

Comparison between measured and predicted values for the dierent-scale FTS processes

Experiment scale: Single-tube test Pilot-plant test Industrial-demonstration

Experiment date 1991 1997 1989 1994Experimental locus Taiyuan Taiyuan Daixian JinchengTube dimension 38 3 7000 38 3 7000 32 3 4000 38 3 7000

Number of tubes 1 1 114 1327Catalyst sizes (mm) 2:5 510 2:5 510 2:5 510 2:5 510Catalyst density (kg=m3) 1100 1100 1100 1100Reactor pressure (Mpa) 2.5 2.55 2.5 2.5Cooling temperature (K) 528.2 518.2 523.2 518.2GHSV (h1) 500 447 425 324Recycle ratio (Rcyc) 3.0 3.07 3.5 3.32Experimental results:XCO (%) 74.0 73.79 70.0 74.0

C+5 =g=Nm3 (CO + H2) 80 100 89.10 85.75 87102

Predicted results:XCO (%) 72.12 74.02 71.78 73.01

C+5 =g=Nm3 (CO + H2) 85.11 89.34 80.32 101.56

The detailed mechanistic kinetics (Wang, 2001;

Wang et al., 2001) is incorporated into the reactor model.

The FischerTropsch components included are CO, H2,

CO2, H2O;N2 and hydrocarbons (n-parans and n-olens)

up to C50. The equilibrium between the gases in the bulk

and the wax in the catalyst pores is correlated by using a

modied SRK equation of state (Wang et al., 1999). The

physico-chemical parameters required by the pellet model

are similar to those used in the pellet modeling ( Wang et al.,

2001). The bed-side heat transfer coecient (i) and the

cooling-media-side heat transfer coecient (c), are taken

into account in the calculation of the overall heat transfercoecient (U). And i is calculated by Levas correlation

(Leva, 1949) and c by VDI-Warmeatlass correlation as

cited in literature (Hartig & Keil, 1993). The friction fac-

tor for pressure drop, fp, is calculated by using Hicks

correlation (Hicks, 1970). The viscosity () and the heat

conductivity (g) for gas bulk phase is calculated by using

Chungs correlation (Reid, Prausnitz, & Poling, 1987). The

supercial gas velocity (us) along the reactor axis is calcu-

lated by means of the total mass ux and the local density

of gas mixture (g), and the latter is evaluated by PR EOS

(Peng & Robinson, 1976).

2.3. Numerical method

The transfer and reaction equations for catalyst pellets

are discretized by orthogonal collocation on nite elements

(Finlayson, 1980). The resulting nonlinear equations are

solved by using HYBRID1 method, which is based on a

modied Powell algorithm (More, Garbow, & Hillstrom,

1980). The bulk equations of initial value type are integrated

by using Gear method, coupling with the embedded pellet

model. The numerical integration of the reaction rates and

the heat of the FTS reactions (in Eqs. (1)(2)), is performed

by Gauss quadrature formula.

3. Simulation and discussion

3.1. Model validation: predictions vs. experiments

In the course of engineering scale-up, FischerTropsch

demonstration units at dierent scales have been estab-

lished by Institute of Coal Chemistry, Chinese Academy

of Sciences (ICC-CAS) (Zhang, 1993; Li, Zhou, Li, &

Hu, 2000). The typical experimental data are collected and

summarized in Table 1. To simulate the FTS processes, our

reactor simulator is combined into Design II (WinSim Inc.,

1998), a commercial process simulation software. Duringthe process simulations, the detailed ux information con-

cerning the FTS reactor unit (i.e. the 1st stage reactor) is

derived from our reactor simulator and then transferred to

Design II by means of its internal INLINE FORTRAN.

In addition, a stoichiometric model (Li, 1989), which is

constructed on the basis of experimental data, is coupled

into the process to describe the 2nd stage reactor (i.e. the

reformer with ZSM-5 zeolite). For comparison, dierent

process simulation diagrams (as typically illustrated in Fig.

1) are conceptually constructed according to the realistic

experimental cases.

Table 1 shows the predicted values and the experimental

results for the single-tube test, the pilot- plant test, and the

industrial-demonstration test. It can be seen that the simu-

lated results of CO conversion are in good agreement with

the experimental ones. Moreover, the index of C+5 yield is

also satisfactorily recovered.

Detailed comparison is further conducted for predict-

ing our latest single-tube experiments with recycle oper-

ation. Judging from the comparison of various indexes

including syngas conversion, C+5 yield, usage ratio, and

o-gas compositions, it is evident that the simulation

results are close to the measured values, as shown in

Table 2.

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Fig. 1. Process scheme for xed-bed FischerTropsch synthesis. X: heat exchanger; M: mixer; F: asher; C: compressor; R01: FTS reactor; R02: reformer

(with ZSM-5 zeolite).

Table 2Detailed comparison between measured and predicted values for the industrial single-tube tests in 1997

Run No. RUN-9701B-78 RUN-9701B-37

Calculated Measured Calculated Measured

XCO (%) 74.02 73.79 76.56 79.37XH2 (%) 52.08 55.49 54.97 64.28H2 : CO Usage ratio 1.33 1.42 1.42 1.61O-gas composition (%):H2 49.13 48.57 47.68 43.96N2 7.24 8.00 11.50 12.06CO 14.10 15.15 12.51 12.81CH4 5.81 3.17 6.45 3.61CO2 22.73 22.4 20.84 24.00

C+5 =g=Nm3 (CO + H2) 89.34 89.10 90.04 92.90Inlet temp. (1st stage)/K 518.2 518.2Max. temp. (1st stage)/K 525.0 519.2523.2 525.3 519.2525.2Outlet temp. (1st stage)/K 522.4 523.0

3.2. Detailed simulations

Despite gas recycle operation has usually been applied in

industrial xed-bed FTS processes (Sie & Krishna, 1999),

there is still a lack of quantitative understanding on this sub-

ject, due to the complexity nature of FTS system. To ll the

gap, the xed-bed FTS system with recycling is considered

to be the focus of this section. The conceptual owsheet dia-

gram, which is simplied according to the FischerTropsch

processes in ICC-CAS (Zhang, 1993; Li et. al., 2000), is

presented in Fig. 2. In our simulation, a component split-

ter is used to approximate the ZSM-5 reformer reactor, in

which the C2C4 olens from the FTS reactor are further re-

formed into the desired C+5 products. Based on this approx-

imation, a FORTRAN program, in which an additional ex-

ternal iteration program is imposed on our reactor simulator,

is developed to realize detailed simulation of this process. A

comprehensive set of numerical results is demonstrated in

Table 3 and Fig. 3.

3.2.1. Tube diameter eects

The eect of tube diameter on the axial temperature prole

in the reactor under industrial operation conditions is shown

in Fig. 3(A). With increasing the tube diameter, the temper-

ature of hotspot will ascend, and its position will move to

the rear part of the reactor, and the temperature at the re-

actor exit keeps a continuously increasing trend. This phe-

nomenon is not dicult to explain. In reality, the increase

of tube diameter will increase the burden of heat removal

of unit heat-transfer area, making the heat removal become

more and more dicult and thus leading to the elevation of

the hotspot. From the viewpoint of the allowable reaction

temperature (the upper limit is assumed to be 543:2 K in

our cases), the diameter of the tube for loading the Fe-Cu-K

industrial catalyst is required to be less than 60 mm ID.

The eect of tube diameter on the conversions, the se-

lectivities, and the C+5 yield is shown in Table 3. From the

C-atom-based selectivity, it can be seen that, the variation

of tube diameter has a slight inuence on the selectivity of

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Fig. 2. Conceptual process diagram for the FischerTropsch process in ICC-CAS.

CO2 but a great inuence on the selectivities of CH4 and

C+5 products. With the increase of tube diameter, the SCH4obviously increases while the SC+5 decreases seriously. Ev-

idently, this kind of trend is directly related to the increase

of the average bed reaction temperature resulting from the

increase of tube diameter. At the same time, the increase

of average reaction temperature will promote the conver-

sion of the syngas, making the CO conversion and even

the overall CO + H2 conversion keep increasing trends

with increasing the tube diameter. Generally, the increase

of tube diameter increases the XCO+H2 ; however, it plays a

negative role in enhancing the yield eciency of C+5 prod-

ucts (unit: C+5 g=NM3 CO + H2 converted). Therefore, the

overall yield of C

+

5 products (unit: C

+

5 g=NM

3

CO + H2in fresh syngas) will nally decrease with the increase of

tube diameter. This indicates that increasing tube diame-

ter is unfavorable for enhancing the overall yield of C +5products.

3.2.2. Recycle ratio eects

The recycle operation of the syngas generally means the

increase of supercial gas velocity in reactor by keeping

the space velocity of the fresh syngas unchanged. Fig. 3(B)

demonstrates the eect of recycle ratio on the axial tempera-

ture prole in the reactor. When the recycle operation is not

imposed, namely Rcyc = 0:0, the temperature in the regionclose to the reactor inlet rapidly increases and the thermal

stability of the reactor is then destroyed. However, when

proper recycle ratios are adopted, it can be seen that the

hotspot in the reactor descends evidently, indicating that the

recycle operation is of importance to maintain the thermal

stability in the xed-bed reactor. With the further increase

of recycle ratio, the temperature change along the axial di-

rection is further reduced, and most part of the reactor bed

is under an isothermal operation, which creates a favorable

condition for the thermal stability of the reactor.

The eect of recycle ratio on the conversions, the selec-

tivities, and the C+5 yield is shown in Table 3. It can be seen

that the increase of recycle ratio suppresses the formation

of the undesired products, CO2 and CH4, and thus enhancesthe formation of the desired products, C+5 . The increased

selectivity to the C+5 products leads to an increase in the

yield eciency of C+5 products, as indicated in Table 3.

On the other hand, although increasing recycle ratio results

in the increase of supercial gas velocity, the decrease of

the average reaction temperature in the reactor, and the

decrease of the CO conversion, the overall conversion of

syngas, XCO+H2 , keeps an increasing trend with the increase

of recycle ratio, due to the increase of H 2=CO usage ratio.

Therefore, judging from the C+5 index, we can see that with

increasing the recycle ratio the overall yield of C+5 prod-

ucts will be accordingly enhanced. These results suggestthat the recycle operation contributes favorably to the in-

crease of the yield of desired products. However, it should

be added that from economic standpoint the power con-

sumption is also a decisive factor for the nal selection of

recycle ratio.

3.2.3. Cooling temperature eects

The eect of cooling temperature on the axial tempera-

ture prole in the reactor is shown in Fig. 3(C). It can be

seen that, when the cooling temperature is lower than the

inlet gas temperature, there exists no hotspot in the reactor.

With increasing the cooling temperature, the FTS reactiontakes place with rapid rates in the region close to the reactor

inlet, and a large amount of reaction heat is generated. As a

consequence, the heat removal gradually becomes relatively

dicult compared with the heat generation, leading to the

appearance of hotspot in the reactor. However, the eect of

cooling temperature makes a little change on the position of

hotspot (if any). The appreciable changes are that with the

increase of cooling temperature the hotspot position moves

slightly toward the rear region of the reactor. Although the

reactor bulk temperature obviously increases with increas-

ing the cooling temperature, the rates of heat generation are

suciently commensurate with the rates of heat removal for

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0.0 0.2 0.4 0.6 0.8 1.0520

525

530

535

540

545

550

555

560

dt=32 mm dt=40 mm dt=50 mmdt=60 mm dt=70 mm dt=80 mm

T,

K

Dimensionless reactor length, z/L

0.0 0.2 0.4 0.6 0.8 1.0520

525

530

535

540

545

550

555

560

Rcyc= 0.0Rcyc= 2.0Rcyc= 3.0Rcyc= 4.0Rcyc= 5.0

T,

K

Dimensionless reactorlength, z/L

0.0 0.2 0.4 0.6 0.8 1.0505

510

515

520

525

530

535

540

545

550

Tw=503 K Tw=513 KTw=523 K Tw=533 K

T,

K


0.0 0.2 0.4 0.6 0.8 1.0522

524

526

528

530

532

P=15 barP=20 barP=25 barP=30 bar

T,

K


(A)

(C) (D)

(B)

Fig. 3. Eect of process parameters on the axial temperature prole in the xed-bed reactor. Base conditions: dt= 32 mm, L= 7:0 m, GHSV= 500 h1,

Rcyc = 3:0, Tw = Tg;0 = 523 K, Pin = 25 bar. Fresh syngas composition (%): CO: 30.59; H2: 57.75; CO2: 7.0; N2: 4.08; CH4: 0.58.

most part of the reactor bed, making the reactor system keep

a relatively constant temperature.

The eect of cooling temperature on the conversions, the

selectivities, and the C+5 yield is shown in Table 3. Due to

the increase of cooling temperature, the reactor bulk tem-

perature is evidently increased and the overall consumption

rates of the reactants are also increased, being favorable for

the enhancement of syngas conversion. By analyzing the

C-atom-based selectivity, we can know that the increase of

cooling temperature brings an obviously negative eect on

the formation of desired products (C+5 products). Although

the increase of cooling temperature increases the syngas con-

version, XCO+H2 , it markedly lowers the yield eciency of

C+5 products. From the yield indexes in Table 3, the overall

yield of C+5 products will hence decrease remarkably with

increasing cooling temperature. Therefore, the selection of

lower cooling temperatures is recommended, on the premier

of keeping a satised syngas conversion. Furthermore, for

the complicated FTS system, it is not suitable to pursuit

higher syngas conversions by enhancing the cooling tem-

perature, unless some cases like catalyst deactivation occur

and then enhancing cooling temperature can be considered

so as to keep the conversion to a certain level.

3.2.4. Pressure eects

The eect of reaction pressure on the axial temperature

prole in the reactor is shown in Fig. 3(D). The increase

of reaction pressure corresponds to the increase of reactant

concentrations. With increasing pressure, the temperature at

hotspot increases and the hotspot position moves slightly

toward the inlet of the reactor. Moreover, it can be seen that

the temperature at the reactor exit also increases with the

increase of pressure. Evidently, due to the pressure increase,

the bulk temperature in the reactor generally increases and

the conversion rates of the reactants accordingly increase

(see Table 3). This means that the increase of pressure has

a signicant eect on the enhancement of CO conversion.

According to the temperature eect on the product selec-

tivity, it can be expected that, the enhancement of reactor

bulk temperature resulting from the pressure increase, will

contribute unfavorably to the formation of heavy hydrocar-

bon. By comparing the eects of cooling temperature and

reaction pressure (Table 3), it shows that the variation trends

of syngas conversion and the yield eciency of C+5 prod-

ucts with the pressure are very similar to those with cooling

temperature. Nonetheless, the increase of pressure has a rel-

atively weak eect on the reduction of the yield eciency of

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C+5 products, contributing to the nal dierence in the over-

all yield of C+5 products between the two kinds of eects.

That is, the overall yield of C+5 products will increase with

increasing reaction pressure while decrease with increasing

cooling temperature.

4. Conclusions

A one-dimensional heterogeneous model is developed to

simulate the xed-bed FischerTropsch reactor. The detailed

kinetics is embedded into the reactor model along with con-

sidering the fact that the catalyst pores are lled with liquid

wax under realistic conditions. The equilibrium between the

gases in the bulk and the wax in the catalyst pores is cor-

related by using a modied SRK EOS. The reactor model

is tested against the measured data from the dierent-scale

demonstration processes, and satisfactory agreements are

found between model predictions and experiment results.

The numerical investigation reveals that the proposed modelcan allow us to gain a quantitative insight into the compli-

cated xed-bed FTS system.

The simulation of a xed-bed FTS reactor with recycle

operation indicates that, the increase of tube diameter is un-

favorable to the increase of the overall yield of C+5 products

and the allowable diameter of the reactor tubes is less than

60 mm ID in our cases. The recycle operation not only con-

tributes favorably to maintaining the thermal stability of the

reactor system but also to enhancing the yield of desired

products. Nonetheless, the power consumption should often

be taken into account in the selection of recycle ratio. The

increase of cooling temperature can increase syngas conver-

sion, however it suppresses remarkably the overall yield of

C+5 products. Therefore the selection of lower cooling tem-

peratures is recommended, on the premier of keeping a sat-

ised syngas conversion. The increase of reaction pressure

has a signicant eect on the increase of CO conversion and

at the same time it favors the enhancement of the overall

yield of C+5 products.

Notation

ci bulk gas concentration of component i, mol=m3

ci;0 inlet gas concentration of component i, mol=m

3

cs;i liquid concentration of component i, mol=m3

CP;m residual heat capacity of jth reaction, J/kg K

dP pellet diameter, m

dt tube diameter, m

De; i eective diusivity coecient of key

component i, m2=s

fp friction factor for pressure drop

GHSV gas hourly space velocity of fresh syngas, h1

hf heat transfer coecient of external lm,

J=m2s K

Hj reaction heat of jth reaction, J/mol

Ke eective conductivity coecient, J=m s K

L reactor length, m

NC total number of components involved

NPG number of key components involved

NR total number of reactions involved

P reaction pressure, bar

Pin inlet reaction pressure, bar

r pellet dimension, mRcyc recycle ratio (recycle gas/fresh syngas)

Rj rate of jth reaction, mol=g s

RP pellet radius, m

Tg bulk temperature of gas phase, K

Tg;0 inlet temperature of gas phase, K

Ts pellet temperature, K

Tw cooling temperature, K

us supercial gas velocity, m/s

U overall heat transfer coecient, J=m2 s K

xi molar fraction of component i in liquid wax

yi molar fraction of component i in bulk gas

phasez reactor dimension, m

Greek letters

i the bed-side heat transfer coecient,

J=m2 s K

c the cooling-media-side heat transfer coe-

cient, J=m2 s K

ij stoichiometric coecient of component i in

jth reaction

B bed voidage, dimensionless

Li fugacity coecient of component i in liquid

wax

Vi fugacity coecient of component i in gas

phaseg the heat conductivity for gas bulk phase,

J=m s K

the viscosity for gas bulk phase, N s=m2

g bulk gas density, kg=m3

p catalyst pellet density, kg=m3

Acknowledgements

The authors gratefully acknowledge nancial support

from Chinese Academy of Sciences. The support of Alexan-

der von Humboldt foundation in Germany is also acknowl-edged by one of the authors (Y.-W. Li).

References

Anderson, R. B., Seligman, B., Schultz, J. F., Kelly, R., & Elliott, M.

A. (1952). Fischer Tropsch synthesis: Some important variables of

the synthesis on iron catalysts. Industrial and Engineering Chemistry,

44, 391.

Atwood, H. E., & Bennett, C. O. (1979). Kinetics of the FischerTropsch

reaction over Iron. Industrial and Engineering Chemistry Process

Design and Development, 18, 163.

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Bub, G., Baerns, M., Bussemeier, B., & Frohning, C. (1980). Prediction

of the performance of catalytic xed bed reactors for FischerTropsch

synthesis. Chemical Engineering Science, 33, 348.

Dry, M. E. (1996). Practical and theoretical aspects of the catalytic

FischerTropsch processes. Applied Catalysis A, 138, 319.

Dry, M. E. (1999). FischerTropsch reactions and the environment.

Applied Catalysis A, 189, 185190.

Design II Chem. Tran. (1998). Reference Guide, Ver8.1, WinSim Inc.Finlayson, B. A. (1980). Nonlinear Analysis in Chemical Engineering.

New York: McGraw-Hill.

Froment, G. F., & Bischo, K. B. (1990). Chemical reactor analysis and

design (2nd ed.). New York: Wiley.

Geerlings, J. J. C., Wilson, J. H., Kramer, G. J., Kuipers, H. P. C. E.,

Hock, A., & Huisman, H. M. (1999). FischerTropsch Technology:

From active site to commercial process. Applied Catalysis A, 186,

2740.

Hartig, F., & Keil, F. J. (1993). Large-scale spherical xed bed reactors:

Modeling and optimization. Industrial and Engineering Chemistry

Research, 32(3), 424437.

Hicks, R. E. (1970). Pressure drop in packed beds of spheres. Industrial

and Engineering Chemistry Fundamentals, 9, 500.

Jess, A., Popp, R., & Hedden, K. (1999). FischerTropsch synthesis

with nitrogen-rich syngas: Fundamentals and reactor design aspects.Applied Catalysis A, 186, 321342.

Leva, M. (1949). Fluid ow through packed tubes. Chemical Engineering,

56(5), 115.

Li, Y. W. (1989). The development of process simulation software

FTPSS: Analysis of FT synthesis (pp. 35 37). M.S. dissertation,

Institute of Coal Chemistry, Chinese Academy of Sciences, Taiyuan,

China.

Li, G. H., Zhou, J. L., Li, Y., & Hu, J. W. (2000). A 3000-hour scale-up

test of an Fe-Cu-K catalysts for F-T synthesis in xed-bed reactor.

Journal of Fuel Chemistry and Technology, 28(1), 8991 (China).

Madon, R. J., & Iglesia, E. (1994). Hydrogen and CO Intrapellet

diusion eects in ruthenium-catalyzed hydrocarbon synthesis. Journal

of Catalysis, 149, 428437.

More, J., Garbow, B., & Hillstrom, R. (1980). User Guide for

MINPACK-1; Argonne National Labs, ANL-80-74, Argonne, IL.

Peng, D. -Y., & Robinson, D. B. (1976). A new two-constant equation

of state. Industrial and Engineering Chemistry Fundamentals, 15,

5964.Post, M. F. M., vant Hoog, A. C., Minderhoud, J. K., & Sie, S. T.

(1989). Diusion limitations in FischerTropsch catalysts. A.I.Ch.E.

Journal, 35, 11071114.

Reid, R. C., Prausnitz, J. M., & Poling, B. E. (1987). The Properties of

gases and liquids (4th ed.). New York: McGraw-Hill.

Sie, S. T. (1998). Process development and scaleup: IV. Case history

of the development of FischerTropsch synthesis process. Reviews in

Chemical Engineering, 14, 109137.

Sie, S. T., & Krishna, R. (1999). Fundamentals and selection of advanced

FischerTropsch reactors. Applied Catalysis A, 186, 5570.

Wang, Y. N. (2001). Modelization and simulation of xed-bed Fischer

Tropsch synthesis: Kinetics, pellet, and reactor (pp. 2039). Ph.D.

dissertation, Institute of Coal Chemistry, Chinese Academy of Sciences,

Taiyuan, China.

Wang, Y. N., Li, Y. W., Bai, L., Zhao, Y. L., & Zhang, B. J. (1999).Correlation for gas-liquid equilibrium prediction in FischerTropsch

synthesis. Fuel, 78, 911917.

Wang, Y. N., Xu, Y. Y., Xiang, H. W., Li, Y. W., & Zhang, B. J. (2001).

Modeling of catalyst pellets for FischerTropsch synthesis. Industrial

and Engineering Chemistry Research, 40, 43244335.

Zhang, B. J. (1993). Liquid fuels synthesized from coal (pp. 326374).

Shanxi: Shanxi Press of Science and Technology (China).

Zimmerman, W. H., Rossin, J. A., & Bukur, D. B. (1989). Eect of

particle size on the activity of fused iron FischerTropsch catalyst.

Industrial and Engineering Chemistry Research, 28, 406413.

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