enhancement of gasoline production in a novel hydrogen-permselective membrane reactor in...

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 4 ( 2 0 0 9 ) 3 9 6 5 – 3 9 7 6

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Enhancement of gasoline production in a novel hydrogen-permselective membrane reactor in Fischer–Tropschsynthesis of GTL technology

A.A. Forghani, H. Elekaei, M.R. Rahimpour*

School of Chemical and Petroleum Engineering, Department of Chemical Engineering, Shiraz University, P.O. Box 71345, Shiraz, Iran

a r t i c l e i n f o

Article history:

Received 28 December 2008

Received in revised form

6 February 2009

Accepted 7 February 2009

Available online 25 March 2009

Keywords:

Fischer–Tropsch synthesis

Pd–Ag membrane

Membrane reactor

Heterogeneous model

GTL technology

* Corresponding author. Tel.: þ987112303071E-mail address: [email protected] (

0360-3199/$ – see front matter ª 2009 Interndoi:10.1016/j.ijhydene.2009.02.038

a b s t r a c t

In this work a novel reactor configuration with hydrogen-permselective membrane is

proposed for Fischer–Tropsch synthesis. In this configuration the synthesis gas is fed to the

tube side and flows in co-current mode with reacting gas mixture that enters in the shell

side of the reactor. In this way, the synthesis gas is heated by heat of reaction which is

produced in the reaction side. Hydrogen can penetrate from the feed synthesis gas side

into the reaction side as a result of a hydrogen partial pressure difference. The outlet

synthesis gas from tube side is recycled to shells and the chemical reaction is initiated in

catalytic bed. Therefore, the reacting gas in shell side is cooled simultaneously with

passing gas in tube and saturated water in outer shell. In this study, the results of novel

membrane reactor (MR) are compared with a conventional Fischer–Tropsch synthesis

reactor (CR) at identical process conditions in terms of temperature, gasoline and CO2

yields, H2 and CO conversion as well as selectivity.

This novel membrane Fischer–Tropsch reactor improves the selectivity of hydrogenation

with hydrogen passing through membrane and increases production of high octane

gasoline from synthesis gas on bifunctional Fe-HZSM5 catalyst. The model was checked

against conventional Fischer–Tropsch synthesis reactor (CR) in pilot plant of Research

Institute of Petroleum Industry. Simulation results show 4.45% enhancement in the yield of

gasoline production, 6.16% decrease in the undesired product formations, and a favorable

temperature profile along the membrane Fischer–Tropsch reactor in comparison with

conventional reactor.

ª 2009 International Association for Hydrogen Energy. Published by Elsevier Ltd. All rights

reserved.

1. Introduction operating conditions [1]. There has been an increased atten-

The conversion of synthesis gas (COþH2 mixtures) into

liquids, and more specifically clean fuels and chemical feed-

stock via Fischer–Tropsch synthesis (FTS), is currently of

increasing interest. This catalytic synthesis leads to a wide

variety of products such as gasoline and diesel, whose abun-

dance depends on the catalysts employed, as well as on

; fax: þ987116287294.M.R. Rahimpour).ational Association for H

tion in the development of GTL (gas to liquid) technology in

the last couple of years. The main incentives for this conver-

sion are the increased availability of natural gas in remote

locations for which no nearby markets exist, environmental

pressure to minimize the flaring of associated gas, the

growing demand for middle distillate transportation fuels

(gasoil and kerosene) especially in the Asia–Pacific regions,

ydrogen Energy. Published by Elsevier Ltd. All rights reserved.

mailto:[email protected]

http://www.elsevier.com/locate/he

http://www.sciencedirect.com


and improvements in the cost effectiveness of GTL tech-

nology, resulting from the development of more active cata-

lyst and improved reactor design. Fischer–Tropsch reactor

being the heart of gas to liquid conversion processes has great

significance in the economics of the overall plant [2].

Due to the high demand on gasoline in the world and its

higher price relative to that of diesel, production of gasoline

from the Fischer–Tropsch process becomes more favorable.

Fuels produced from the Fischer–Tropsch synthesis have high

quality such as a very low aromaticity and absence of sulfur

that help in diminishing environmental impact. However, due

to the limitation of Schulz–Flory distribution (use for predict-

ing the polymer molecular weight distribution) [3], the yield of

the hydrocarbons with in the range of those presented in

gasoline is low, at the same time the octane number of FT

gasoline is lower than that of the gasoline obtained from crude

oil processing, since the FT gasoline mainly consists of

n-paraffin. To promote the yield and quality of the gasoline

from Fischer–Tropsch synthesis, bifunctional catalysts have

received extensive attention in the recent years [4,5].

F–T synthesis is either low temperature F–T (LTFT) process

or high temperature F–T (HTFT) process depending on the

product required. High temperature process is mainly used for

the production of gasoline and linear olefins while low

temperature process is applied for the production of waxy

material [2]. There is no liquid phase around the catalyst

particles in HTFT process and this is the main advantage of

HTFT process relative to LTFT process. M. Ahmadi Marvast

et al. considered a water-cooled fixed bed F–T reactor with

length of 12 m [6]. In their work, investigation of the effects of

reactor temperature and gasoline production was carried out,

and all the reactants’ conversion and components’ production

performed in a high temperature F–T (HTFT) process.

In the present work, the application of shell and tube Pd–Ag

membrane reactor in Fischer–Tropsch synthesis is assessed.

By this reactor configuration the reacting and synthesis gases

are in contact with each other in shell and tube schemes,

increasing production rate. The presence of a permselective

membrane in a reacting system can improve the selectivity of

hydrogenation [7]. A maximum value of the hydrogen flow is

reached for membranes composed of Pd alloyed with approx-

imately 23 wt % silver. Similar to Pd/Ag, other alloys, e.g. Pd/Y or

Pd/Ce show high hydrogen permeability and good mechanical

Fig. 1 – Schematic diagram of a conventional fixed bed

Fischer–Tropsch reactor.

stability [8]. Palladium-based membranes have been used for

decades in hydrogen extraction because of their high perme-

ability and good surface properties and because palladium, is

100% selective for hydrogen transport [9]. Key requirements for

the successful development of palladium-based membranes

are low costs as well as permselectivity combined with good

mechanical, thermal and long-term stability [8]. A Pd–Ag alloy

permeator tube has been operated from ambient temperature

up to 400 �C with hydrogen pressure of 100–150 kPa with

a resulting maximum elongation of about 1% [10].

These properties would make palladium-based membranes

such as Pd–Ag membrane very attractive. In this way, diffusion

of hydrogen through the Pd–Ag film from the shell side to

reaction space makes it possible to promote the conversion of

carbon monoxide.

There are several researches on application of palladium

membrane in reactors. Saracco et al. studied the potentials

and problems of high temperature membrane reactors [11].

Lin et al. found that membrane reactors for methanol steam

reforming at 300–400 �C can be a practical route for hydrogen

production [12]. Coronas and Santamaria discussed trade-off

between selectivity and permeability, porous membranes aim

for higher values of the latter, which often means sacrificing

selectivity. A notable exception is zeolite membranes with

a promising future of applications to catalytic reactors [13].

Rahimpour and Ghader [14] investigated the Pd–Ag membrane

reactors’ performance for methanol synthesis. They consid-

ered steady-state homogeneous model for methanol reactor.

Gallucci and Basilo studied co-current and counter-current

configurations of membrane reactor for methanol steam

reforming [15]. Rahimpour and Lotfinejad presented

a dynamic model for studying Pd–Ag membrane in a dual-type

reactor for methanol production [16], showing that methanol

production can be increased in membrane dual-type reactor.

Nair and Harold recently carried out an analysis of conven-

tional Pd and Pd/Ag membranes [17]

In this work, modeling of the conventional F–T synthesis

reactor (CR) in pilot plant of Research Institute of Petroleum

Industry (RIPI) was performed and then a new configuration to

improve the gasoline production in the single stage F–T

synthesis reactor by applying selective permeation of hydrogen

Table 1 – FTS pilot plant characteristics [19].

Parameter Value

Tube dimension [mm] Ø38.1� 3� 120,00

Molar ratio of H2/CO in feed 0.96

Feed temperature [K] 569

Reactor pressure [kPa] 1700

Cooling temperature [K] 566.2

Catalyst sizes [mm] Ø2.51� 5.2

Catalyst density [kg m�3] 1290

Bulk density [kg m�3] 730

Number of tubes 1

Tube length [m] 12

GHSV [h�1] 235

Bed voidage 0.488

Feed molar flow rate [g mol s�1] 0.0335

0 0.2 0.4 0.6 0.8 1 1.2 1.4 1.6 1.8 2 2.2 2.41

2

3

4

5

6

7

8

H2/CO ratio

gaso

lin

e yield

(g

r/g

r feed

)

0 2 4 6 8 10 120

0.10.20.30.40.50.60.70.80.9

1

length (m)

H2/C

O ratio

0 1 2 3 4 5 6 7 8x 10-3

0

0.5

1

1.5

2

2.5

hydrogen injection

(mol/s per tube)

H2/C

O ratio

0 1 2 3 4 5 6 7 8x 10-3

6.2

6.4

6.6

6.8

7

7.2

7.4

7.6

7.8

hydrogen injection

(mole/s per tube)

gaso

lin

e yield

(g

r/g

r feed

)

a b

c d

Fig. 2 – (a) Effect of the inlet H2/CO ratio on the C5D yields, (b) H2/CO ratio profile vs. reactor length, (c) effect of H2 addition on

the H2/CO ratio and (d) effect of H2 addition on C5D yield.

i n t e r n a t i o n a l j o u r n a l o f h y d r o g e n e n e r g y 3 4 ( 2 0 0 9 ) 3 9 6 5 – 3 9 7 6 3967

from synthesis gas and adding it to reaction side is suggested.

In this way, diffusion of hydrogen through the Pd–Ag film from

the tube side to reaction side makes it possible to provide

suitable level of hydrogen along the reactor increasing gasoline

production. The integration of membranes in a catalytic reactor

allows to either dose one of the reactants in a controlled

manner in order to achieve optimal axial concentration profiles

corresponding with higher product yields and higher product

selectivity [18].

In this work, Pd–Ag membrane was placed theoretically in

conventional F–T reactor of Research Institute of Petroleum

Industry (RIPI). Numerical simulation was utilized to compare

results of novel membrane reactor with conventional reactor

at same process conditions such as pressure, temperature,

and feed composition.

2. Process description

As will be described in the following sections, the Fischer–

Tropsch synthesis has been investigated in a conventional

reactor (CR) and membrane reactor (MR). Both reactors are

packed with bifunctional Fe-HZSM5 catalyst (metal part: 100

Fe/5.4 Cu/7K2O/21SiO2, acidic part: SiO2/Al2O3¼ 28).

2.1. Conventional Fischer–Tropsch synthesis reactor

In industrial fixed bed Fischer–Tropsch reactors, multi tubular

reactors cooled by pressurized boiling water are often used.

Fig. 1 shows a schematic of the conventional reactor (CR) for

Fischer–Tropsch synthesis. Table 1 presents the characteris-

tics of the conventional fixed bed reactor developed by

Research Institute of Petroleum Industry [19].

It is very important to adjust a H2/CO ratio for Fischer–

Tropsch synthesis [20]. Feed H2/CO ratios of about 1 reduce the

synthesis gas conversion and the CH4 selectivity, while the C5þ

selectivity and olefin/paraffin ratio for C2–C4 is increased that is

more suitable for high temperature F–T (HTFT) process. Fig. 2(a)

shows the effect of the inlet H2/CO ratio on the production yield

of gasoline along the conventional reactor. As can be seen, C5þ

production is maximal at the H2/CO ratio of 0.8. In order to reach

the highest production rate, the H2/CO ratio should be equal or

close to the optimum H2/CO ratio along the reactor. Fig. 2(b)

displays H2/CO ratio variations along the conventional fixed bed

reactor for Fisher–Tropsch synthesis. Regarding this figure, the

H2/CO ratio of the reacting gas decreases along the reactor up to

values much lower than the optimum. In this case study, the

optimum composition of the reactants is achieved by injecting

hydrogen into the H2-poor reacting gas. Fig. 2(c) presents the

effect of hydrogen addition on H2/CO usage ratio. The contin-

uousadditionofhydrogenleadsto an increase inhydrogenmole

percent in the reacting gas, resulting in a higher H2/CO ratio in

H2-poor reacting gas. Fig. 2(d) implies that by injecting

0.002 mol s�1 of hydrogen to the H2-poor gas (per each tube), an

optimum processing gas is obtained, resulting in maximum C5þ

production. Hence, hydrogen injection into the system is

required. In the membrane concept, hydrogen is withdrawn

Fig. 3 – Schematic diagram of a membrane Fischer–Tropsch reactor.


from the H2-rich synthesis gas stream by a membrane unit and

permeated to the reaction side in order to adjust the H2/CO

usage ratio.

2.2. Membrane Fischer–Tropsch reactor

Fig. 3 shows the schematic diagram of a membrane Fischer–

Tropsch reactor in co-current configuration. This system

consists of two concentric pipes like tube–shell system. The

tube wall in this system is hydrogen-permselective and

hydrogen partial pressure gradient between the shell and

tube permits diffusion of hydrogen through the Pd based

membrane layer. Therefore, the mass and heat transfer

processes simultaneously occur between both sides result-

ing in higher heat transfer and controlling of H2/CO ratio.

This simulation study is based on a Pd–Ag layer thickness of

15 mm. In this configuration, the fresh synthesis gas is fed to

the tube side (permeation side) and is preheated by the

generated heat in reaction side (shell side). Pure hydrogen

permeates to reaction side in order to control and maintain

the suitable hydrogen gradient in the whole length of the

reactor [21]. Then, the heated synthesis gas is routed from

recycle stream through shell in a co-current mode with feed

synthesis gas and the chemical reaction is initiated by the

catalyst. The reacting gas is also cooled with cooling satu-

rated water which flows around it. In fact, the heat of

reaction is transferred to both cooling water and fresh

synthesis gas. After leaving the shell, the product containing

hydrocarbons goes to hydro cracking unit. Catalyst charac-

teristics and specifications of membrane F–T reactor are

listed in Table 2.

3. Mathematical model

A one-dimensional heterogeneous model comprising a set of

heat and mass transfer equations and the kinetics of the main

reactions is applied in this work to simulate the membrane

Fischer–Tropsch synthesis reactor. The membrane reactor

model is based on the following assumptions.

- Steady-state conditions.

- Plug flow is considered in reactor and tube side.

- Reaction rates developed by M. M. Montazer Rahmati et al.

[22] were used to obtain the main reactions rates.

- Axial dispersion of heat is neglected.

In this simple model it is assumed that gradients of

temperature and concentrations between the solid phase and

gas phase can be ignored and the equations for the two phases

can be combined [23]. The general fluid-phase balance typi-

cally accounts for accumulation, convection, and reaction. In

the current work, axial dispersion of heat is neglected and the

heat loss by coolant is considered.

3.1. Reaction network

The Fischer–Tropsch components include H2, CO, CO2, H2O,

CH4, C2H6, C3H8, n-C4H10, i-C4H10 and C5þ. The following reac-

tions are considered as leading Fischer–Tropsch reactions [22]:

COþ 3H2!R1

CH4 þH2O

2COþ 4H2!R2

C2H4 þ 2H2O

Table 2 – Catalyst and specifications of membranesystem.

Parameter Value

Catalyst density [kg m�3] 1290

Catalyst equivalent diameter [m] 3.83� 10�3

Molar ratio of H2/CO in feed 0.96

Flow rate per tube [g mol s�1] 0.0335

Feed temperature [K] 565

Shell side pressure [bar] 18

Tube side pressure [bar] 35

Cooling temperature [K] 555

Bulk density [kg m�3] 730

Tube length [m] 12

Inner radius of Pd–Ag layer [mm] 19.05

Outer radius of Pd–Ag layer [mm] 19.065

Reactor radius [mm] 27

Number of tubes 1

Bed voidage 0.488

Catalyst thermal conductivity [kJ m�1 s�1 K�1] 0.00625

Fig. 4 – An element of length Dz, in MR.

Table 4 – Comparison between model results with pilotplant data for fresh catalyst.


2COþ 5H2!R3

C2H6 þ 2H2O

3COþ 7H2!R4

C3H8 þ 3H2O

4COþ 9H2!R5

n-C4H10 þ 4H2O

4COþ 9H2!R6

i-C4H10 þ 4H2O

6:05COþ 12:23H2!R7

C6:05H12:36ðC5þÞ þ 6:05H2O

COþH2O 4R8

CO2 þH2 ðWater gas shift ðWGSÞ reactionÞ

The reaction rate equation is as follows and the kinetic

parameters are given in Table 3:

Ri ¼ 0:278ki$exp

��Ei

RT

�Pm

COPnH2

hmol kg�1

cat s�1i

(1)

3.2. Proposed model for F–T synthesis in pilotplant of RIPI

AnF–Tsynthesispilotplantwasdesignedandconstructedbythe

RIPI and National Iranian Oil Company (RIPI–NIOC). This reactor

(CR) has been modeled regarding the following assumptions: (a)

one-dimensional plug flow; (b) axial dispersion of heat is negli-

gible compared to convection; and (c) gases are ideal.

Table 3 – Kinetic parameter data [22].

Reaction No. m n k E

1 �1.0889 1.5662 142583.8 83423.9

2 0.7622 0.0728 51.556 65018

3 �0.5645 1.3155 24.717 49782

4 0.4051 0.6635 0.4632 34885.5

5 0.4728 1.1389 0.00474 27728.9

6 0.8204 0.5026 0.00832 25730.1

7 0.5850 0.5982 0.02316 23564.3

8 0.5742 0.710 410.667 58826.3

The mass and energy equations for the bulk gas phase can

be written as follows:

�ft0

Ac:dyi

dzþ av$ct$kgi

�yis � yi

�¼ 0 i ¼ 1;2;.;N� 1 ð2Þ

�ft0

Ac$cpg$

dTdzþ av$hfðTs � TÞ þ pDi

Ac$UshellðTshell � TÞ ¼ 0 (3)

where, yi and T are the gas phase mole fraction and temper-

ature, respectively.

The boundary conditions for the bulk phase are expressed

by:

z ¼ 0; yi ¼ yi;in; T ¼ Tin (4)

The mass and energy balance equations for the catalyst

pellets can be formulated as follows:

kgi$av$ct

�yi � yis

�þ rB$h$ri ¼ 0 i ¼ 1; 2;.;N� 1 (5)

av$hfðT� TsÞ þ h$rB$X8

j¼1

rj

��DHfj

�¼ 0 (6)

Parameter Pilot plant Predicted Error %

XCO (%) 77.94 78.08 0.18

XH2 (%) 92.83 93.48 0.7

C5þ selectivity 42.55 51.9425 21.92

CO2 selectivity 339.07 315.12 �7.3

CH4 selectivity 44.15 44.23 0.1812

H2O selectivity 120.67 91.41 �24.3

C2H4 selectivity 3.95 3.47 �12.3

C2H6 selectivity 11.78 14.05 19.08

n-C4 selectivity 11.07 9.7789 �11.74

i-C4 selectivity 14.45 11.36 �21.43

C3H8 selectivity 9.33 6.374 �31.7

0 2 4 6 8 10 120

10

20

30

40

50

60

70

80

Length (m)

CO

C

on

versio

n (%

)

0 2 4 6 8 10 120

20

40

60

80

100

Length (m)

Hyd

ro

gen

C

on

versio

n (%

)

MRCR

MRCR

a b

Fig. 5 – (a) Hydrogen and (b) carbon monoxide conversion profiles along the membrane and conventional reactor.


Where, yis and Ts are the mole fractions on the catalyst surface

and solid-phase temperature, respectively.

3.3. Membrane fixed bed F–T reactor model

3.3.1. Reaction sideAs shown in Fig. 4, an element of length Dz, is considered. On

the basis of expressed assumptions, the mass and energy

balances for solid phase can be written as:

KgiavCt

�yis � yi

�¼ hrirB i ¼ 1; 2;.;N� 1 (7)

avhfðTs � TÞ ¼ rB

XN

i¼1

hri

��DHfi

�(8)

Also the mass and energy balances for gas phase can be

written as:

Fr

Ashell

dyi

dz¼ aH

As

� ffiffiffiffiffiffiPt

H

q�

ffiffiffiffiffiffiffiPsh

H

q �þ avCtKgi

�yis � yi

�(9)

Fr

ACpg

dTdz¼ pDro

AUshellðTshell � TÞ þ pDi

AUtubeðTtube � TÞ

c shell c

þ aH

As

� ffiffiffiffiffiffiPt

H

q�

ffiffiffiffiffiffiffiPsh

H

q �CphðTtube � TÞ þ avhfðTs � TÞ ð10Þ

where Fr is total molar flow rate of gas in reaction side for each

tube. T and yi are the temperature and concentration of

component i in the fluid phase and aH is hydrogen permeation

rate constant. rB is density of bed and PHsh and PH

t are partial

pressures of hydrogen in reaction and tube side. Ac is cross-

sectional area of reaction side. Dro and Di are inner diameter of

reaction side and outer diameter of tube side, respectively.

DHf,i is enthalpy of formation of component i. Moreover Eq. (9)

applies only for component of hydrogen. Several auxiliary

correlations for estimation of heat transfer coefficients are

presented in Appendix A.

The boundary conditions are as follows:

z ¼ 0; yi ¼ yi;in; T ¼ Tin (11)

where yi0 and T0 are the mole fraction of component i and

temperature of feed, respectively. Effect of permeated hydrogen

on energy balance is negligible because its amount is small.

3.3.2. Tube sideThe mass and energy balance equations is written only for

hydrogen in the tube side:

dFti

dz¼ aH

� ffiffiffiffiffiffiPs

H

p�

ffiffiffiffiffiffiPt

H

q �(12)

where i denotes H2.

Ft

AcCpg

dTtube

dz¼pDiUtube

AcðT� TtubeÞ þ

aHcpH

As� ffiffiffiffiffiffiPt

H

q�

ffiffiffiffiffiffiPs

H

p �ðT� TtubeÞ

(13)

The boundary conditions are as follows:

z ¼ 0; yi ¼ yif ; T ¼ Tf (14)

where yif and Tf are the mole fraction of component i and

temperature of product, respectively.

3.3.3. Hydrogen permeation through Pd–Ag membraneIn Eqs. (9), (10), (12) and (13), aH is hydrogen permeation rate

constant and is defined as [23]:

aH ¼2pLP

lnðRo=RiÞ(15)

where Ro and Ri stand for outer and inner radius of Pd–Ag

layer. The permeability of hydrogen through Pd–Ag layer as

a function of temperature is as follows [24,25]:

P ¼ P0expð�EP=RTÞ (16)

where the pre-exponential factor P0 above 200 �C is reported

as 6.33� 10�8(mol /m2 s Pa1/2) and activation energy Ep is

15.7 kJ mol�1 [25,26].

4. Numerical solution

The governing equations of model form a system of

coupled equations comprising algebraic, partial differential

0 2 4 6 8 10 120

0.2

0.4

0.6

0.8

1

1.2

1.4

1.6

Length (m)

Iso

-B

utan

e Y

ield

(g

r/g

r feed

*100)

0 2 4 6 8 10 120

1

2

3

4

5

6

7

8

Length (m)

Gaso

lin

e Y

ield

(g

r/g

r feed

*100)

MRCR

MRCR

MRCR

MRCR

MRCR

MRCR

0 2 4 6 8 10 120

1

2

3

4

5

6

Length (m)

Pro

pan

e Y

ield

(g

r/g

r feed

*100)

0 2 4 6 8 10 120

0.1

0.2

0.3

0.4

0.5

0.6

0.7

0.8

Length (m)

N-B

utan

e Y

ield

(g

r/g

r feed

*100)

0 2 4 6 8 10 120

0.5

1

1.5

2

2.5

3

3.5

4

Length (m)

Eth

an

Y

ield

(g

r/g

r feed

*100)

0 2 4 6 8 10 120

0.1

0.2

0.3

0.4

0.5

0.6

0.7

Length (m)

Eth

ylen

e Y

ield

(g

r/g

r feed

*100)

a b

c d

e f

Fig. 6 – (a) Gasoline, (b) propane, (c) n-butane, (d) iso-butane, (e) ethane, and (f) ethylene production yields.


and ordinary differential equations. After rewriting the

model equations at steady-state conditions, a set of

differential algebraic equations (DAEs) is obtained. To solve

this set of equations, backward finite difference approxi-

mation is applied here. Doing this, the DAEs change to

a non-linear algebraic set of equations. The reactor is then

divided into 30 separate sections and the Gauss–Newton

method is used to solve the non-linear algebraic equations

in each section. For solution of the model as a function of

time, the catalyst deactivation model is coupled with

model equations.

5. Simulation and discussion

5.1. Model validation

Model validation was carried out by comparison of CR model

results with the RIPI pilot plant data [19] under the design

specifications and input data. The characteristics of the pilot

plant are tabulated in Table 1. As can be understood from

Table 4, the estimated results are in good agreement with the

experimental data.

0 2 4 6 8 10 120

5

10

15

20

25

30

35

40

45

Length (m)

CO

2 Y

ield

(g

r/g

r feed

*100)

MRCR

MRCR

0 2 4 6 8 10 120

1

2

3

4

5

6

7

Length (m)

CH

4 yield

(g

r/g

r feed

*100)

a b

Fig. 7 – (a) Carbon dioxide, and (b) methane production yields.


5.2. Simulation results

The results of membrane F–T reactor (MR) model are depicted

in the following figures. Fig. 5(a) and (b) illustrates the

comparison of hydrogen and carbon monoxide conversion for

conventional F–T reactor (CR) and membrane F–T reactor

(MR).

As shown in Fig. 5, reactants’ conversion profiles of

membrane system are higher than conventional system

profiles due to adding hydrogen to the reaction side by the use

of Pd–Ag membrane tube, so diffusion of hydrogen promotes

the conversion of carbon monoxide.

Fig. 6(a)–(f) compares desired hydrocarbon products such

as gasoline, propane, normal-butane, iso-butane, ethane and

ethylene production yields, respectively. It is clearly seen that

the performance of membrane Fischer–Tropsch reactor

enhances the formation of the desired products, especially

C5þ; in fact hydrogen permeation in shell side improves

gasoline production, as demonstrated from Fig. 2. A compar-

ison of undesired products, along the membrane and

conventional reactor systems is shown in Fig. 7 (a) and (b). In

these figures, it is obvious that membrane reactor operates

properly against conventional reactor, owing to decrease the

production of carbon dioxide and methane as undesired

Fig. 8 – The comparison of reacting gas temperature

profiles along the two types of reactor systems.

products. Diffusion of hydrogen to reaction side affects water

gas shift reaction in favor of CO2 consumption, but on the

other hand, enhancement in water production is not appre-

ciable (water yield in CR is 14.03 (gr/gr feed� 100) while in MR

is 14.2 (gr/gr feed� 100)).

Fig. 8 presents the reacting gas temperature profile for two

reactor configurations. It is observed that the temperature

control of the MR is easier. There is a jumped temperature for

both systems at first 0.5 m of reactor. For simulation purposes,

the maximum temperature for the Fe-HZSM5 catalyst to

remain active is assumed to be 620 K [6]. As shown in Fig. 8, in

conventional reactor system, risk of temperature runaway

makes alert in closing to hotspot.

As discussed before, the heat of reaction, generated in shell

side is removed by cooling water and tube fresh synthesis gas.

The wall temperature increase occurred by heat transfer

between shell and tube, has a positive effect as well, although

too high temperatures cause the overtaking of membrane

technological threshold [27].

According to Eqs. (15) and (16) hydrogen permeation rate is

a function of feed gas temperature and hydrogen partial

pressure difference between shell and tubes. Fig. 9 shows the

0 2 4 6 8 10 121

1.5

2

2.5

3

3.5

4

length (m)

H2 p

artial p

ressu

re d

ifferen

ce (b

ar1/2)

Fig. 9 – Profiles of hydrogen partial pressure difference vs.

length of membrane reactor.

0 2 4 6 8 10 120

1

2

3

4

5

6

7x 10-6

length (m)

perm

eatio

n rate (m

ol/s)

Length (m)

Membrane thickness (micron)

0 20 40 60 80 05

10150

0.5

1

1.5

2x 10-5

H2 p

erm

eatio

n rate (m

ol/s)

a b

Fig. 10 – Hydrogen permeation rate profile (a) vs. length of membrane reactor and (b) in terms of reactor length and

membrane thickness.


hydrogen partial pressure difference along the membrane

reactor. Upward trend of this profile is arisen from entrance

co-current feeds in this reactor and owing to this, pressure

drop increases along the reactor and lead to enlargement of

hydrogen permeation rate as shown in Fig. 10(a). On the other

hand, behavior of permeation rate profile refers to feed

synthesis gas behavior; an increasing trend is identified for

the hydrogen permeation rate profile due to increase of feed

gas temperature up to 2 m of reactor length. Then enhance-

ment of permeation rate goes on with smaller slope as

a consequence of increase in H2 partial pressure difference

and decrease in tube temperature. Fig. 10(b) shows the effect

of membrane thickness on the H2 permeation rate profile.

Reducing thickness of permselective layer raises hydrogen

permeability and so gasoline production.

In this regard, the hydrogen mole fraction of tube side is

reduced along the reactor. Fig. 11 exhibits the lessening

circumstance of hydrogen concentration along the membrane

tube.

Fig. 12 illustrates the effect of feed synthesis gas temper-

ature on the hydrogen permeation rate. Since hydrogen

permeability follows the Arrhenius law according to Eq. (16),

0 2 4 6 8 10 120.4875

0.488

0.4885

0.489

0.4895

0.49

length (m)

hyd

ro

gen

m

ole fractio

n o

f tu

be

Fig. 11 – Tube hydrogen mole fraction along membrane

reactor.

increasing temperature promotes the hydrogen permeability.

Also, hydrogen permeation depends on the hydrogen partial

pressure square root difference between the reaction zone

and the permeation zone. Permeation rate increases when

tube temperature decreases, is caused by hydrogen partial

pressure square root difference augmentation.

Fig. 13 presents hydrogen permeation rate variations in

terms of total tube side pressure and reactor length. As shown

in this figure, increasing tube pressure enhances hydrogen

permeability along the reactor; especially slope of increasing

is sharp between tube pressures of 30–35 bar.

Moreover as shown in aforesaid figure, hydrogen perme-

ation increases along the reactor length.

Fig. 14 displays gasoline yield in terms of inlet temperature

of gas phase in shell side along the reactor. This investigation

shows that when the inlet reacting gas temperatures increase,

the gasoline production yield declines along the reactor with

a smooth slope.

Fig. 15 displays the effect of the H2/CO ratio on the gasoline

yield profile along the membrane reactor length. In this profile,

an optimum H2/CO ratio which maximizes the gasoline yield

production along the reactor is observed. Reactant ratios of

Fig. 12 – The hydrogen permeation rate vs. tube

temperature profile.

03

69

12

303540455055600

0.2

0.4

0.6

0.8

1

x 10-5

Length (m)

Tube pressure (bar)

pe

rm

ea

tio

n ra

te

(m

ol/s

)

Fig. 13 – Profile of H2 permeation rate vs. pressure of tube

and length of reactor.

05

1015

560

565

570

5750

2

4

6

8

length (m)Inlet temperature

of reacting gas (K)

gaso

lin

e yield

Fig. 14 – Profile of gasoline yield in terms of inlet reacting

gas temperature and length of reactor.

0 0.59 0.75 0.96 1.22 1.56 1.86

0

4

80

2

4

6

8

10

H2/CO ratio

length (m)

C5+

yield

(g

r/g

r feed

*100)

Fig. 15 – Three-dimensional profile of gasoline yield in

terms of H2/CO ratio and length of reactor.

C5+ i-C4 n-C4 C3H8 C2H6 C2H4 CH4 CO2 H2O0

50

100

150

200

250

300

350

components

Selectivity

MRCR

Fig. 16 – Selectivity comparison of the components for the

two types of reactor systems.


0.852 result in most C5þ production yield while in the present

work, this ratio was chosen 0.96. Increasing H2/CO ratio in

excess of unity is unfavorable for the increase in overall yield of

gasoline products.

A comparison of hydrocarbon products selectivity between

conventional reactor and membrane reactor systems is pre-

sented by Fig. 16. According to this comparison, CR system

shows unfavorable results relative to MR system. Membrane

reactor enhances the C5þ selectivity which is an object of the

proposed design and likely declines methane and carbon

dioxide (undesired products) selectivity. Therefore, the

membrane reactor is adherent with FT process requirements.

6. Conclusion

The simulations of a conventional fixed bed Fischer–Tropsch

reactor show that the optimum composition of the reactants is

achieved by injecting hydrogen into the H2-poor reacting gas.

Therefore, a membrane reactor concept for Fischer–Tropsch

synthesis is developed to increase the yield of desired hydro-

carbon products especially gasoline and decrease undesired

products such as CO2 and methane. The most trends have

been in the direction of applying palladium membranes into

conventional fixed bed Fischer–Tropsch reactor to utilize the

hydrogen-permselectivity and to control hydrogen dosing

capabilities of membrane. The mathematical model of CR is

validated against the RIPI pilot plant data and the MR model

results are compared with results of CR. Temperature,

component yields, reactants conversion, hydrogen partial

pressure difference and hydrogen permeation rate profiles are

investigated and the simulation results of this new reactor

configuration show 4.45% additional gasoline production yield

and 6.16% reduction in carbon dioxide formation while

increasing water production is not considerable. This feature

suggests that the concept of membrane Fischer–Tropsch

synthesis reactor system is an interesting candidate for

increasing the gasoline production from synthesis gas.


Acknowledgment

The authors would like to thank Iranian Oil and Gas Co. for

financial support. Also cooperation of Mr. Ali Khosravanipour

Mostafazadeh during preparation of this article is appreciated.

Appendix AAuxiliary correlations

The mass transfer coefficients between the gas phase and the

solid phase in fixed bed reactor have been taken from Cusler

[27].

kgi ¼ 1:17Re�0:42Sc�0:67i ug � 103 (A-1)

Re ¼ 2Rpug

m(A-2)

Sci ¼m

r$Dtm10�4

(A-3)

The diffusivity of each component in the gas mixture is

given by [28]:

Dim ¼1� yiP

isj

yi

Dij

(A-4)

Dij ¼10�7$TB=2

ffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffi1

Miþ 1

Mj

s

r�

nB=2ci þ n

B=2cj

�2 (A-5)

where, Dij is the binary diffusivity calculated using the Fuller–

Schetter–Giddins equation [18]. Mi and nci are the molecular

weight and critical volume of component i.

The overall heat transfer coefficient between the circu-

lating boiling water of the shell side and the bulk of the gas

phase in the tube side is given by the following correlation:

1Ushell

¼ 1hiþ

Ai$lnDoDi

2pLKwþ Ai

Ao$

1ho

(A-6)

Table A.1 – Molecular weight and critical volume of thecomponents.

Component Mi (g mol�1) vci (m3 mol�1)� 106

C5þ 85.084 370

i-C4H10 58.123 262.7

n-C4H10 58.123 255

C3H8 44.096 200

C2H6 30.07 145.5

C2H4 28.054 129.1

CO2 44.01 94.0

CO 28.01 18.0

H2O 18.02 56.0

H2 2.02 6.1

CH4 16.04 99.0

N2 28.01 18.5

where, hi is the convection heat transfer coefficient between

the gas phase and the reactor wall and is obtained by the

following correlation [29]:

hi

cprm$�cpm

k

�2=3

¼ 0:458eB

$

�rudp

m

��0:407

(A-7)

where, eB is the void fraction of the catalytic bed and dp is the

equivalent catalyst diameter and the other parameters are

related to bulk gas phase.

To calculate the heat transfer coefficient of boiling water in

the shell side at high pressure, Leva correlation is applied [29]:

ho ¼ 282:2 P4=3$DT2 0:7 < P < 14 MPa (A-8)

Appendix BNomenclature

Ac [m2] cross-section area of tube

Ai [m2] inner area of tube

As [m2] lateral area of tube

Ashell [m2] cross-section area of shell

av [m2 m�3] specific surface area of catalyst pellet

cpg [J mol�1 k�1] specific heat of the gas at constant

pressure

cpgt[J mol�1 k�1] specific heat of the tube gas at constant

pressure

cpH [J mol�1 k�1] specific heat of the hydrogen at

constant pressure

cps [J mol�1 k�1] specific heat of the catalyst at constant

pressure

ct [mol m�3] total concentration

Di [m] tube inside diameter

Dij [m2 s�1] binary diffusion coefficient of

component i in j

Dim [m2 s�1] diffusion coefficient of component i in

the mixture

Dro [m] reaction outside diameter

hf [W m�2 K�1] gas-catalyst heat transfer coefficient

hi [W m�2 K�1] heat transfer coefficient between fluid

phase and reactor wall

ho [W m�2 K�1] heat transfer coefficient between

coolant stream and reactor wall

ft0 [mol s�1] total molar rate in tube at entrance of

reactor

Fr [mol s�1] total molar rate for shell side

Ft [mol s�1] total molar rate for tube side

K [W m�1 K�1] conductivity of fluid phase

Kw [W m�1 K�1] thermal conductivity of reactor wall

kgi [m s�1] mass transfer coefficient between gas

and solid phase for component i

L [m] length of reactor

Mi [g mol�1] molecular weight of component i

N [–] number of components

P [bar] total pressure

Pa [bar] atmospheric pressure

PtH [bar] tube side pressure

PshH [bar] shell side pressure

P [mol m�1 s�1Pa�1/2] permeability of hydrogen through Pd–

Ag layer


P0 [mol m�1 s�1 Pa�1] pre-exponential factor of hydrogen

permeability

R [J mol�1 K�1] universal gas constant

Re [–] Reynolds number

Ri [m] inner radius of Pd–Ag layer

Ro [m] outer radius of Pd–Ag layer

ri [mol kg�1 s�1] reaction rate of component i

rbi [mol kg�1 s�1] reaction rate of component i in bubble

phase

Sci [–] Schmidt number of component i

T [K] bulk gas phase temperature

Ts [K] temperature of solid phase

Tsat [K] saturated temperature of boiling water

at operating pressure

Tshell [K] temperature of water stream

Ushell, Ut [W m�2 K�1] overall heat transfer coefficient

between coolant and process streams

ug [m s�1] linear velocity of gas phase

yi [mol mol�1] mole fraction of component i in the

fluid phase

yis [mol mol�1] mole fraction of component i in the

solid phase

z [m] axial reactor coordinate

Greek letters

aH; [mol m�1 s�1 Pa�0.5]hydrogen permeation rate constant

DHf,i [J mol�1] enthalpy of formation of component i

DH298 [J mol�1] enthalpy of reaction at 298� K

m [kg m�1 s�1] viscosity of fluid phase

nci [cm3 mol�1] critical volume of component i

r [kg m�3] density of fluid phase

rB [kg m�3] density of catalytic bed

rp [kg m�3] density of catalyst

h [�] catalyst effectiveness factor

Superscripts and subscripts

f feed conditions

s at catalyst surface

sh shell side

t tube side

r e f e r e n c e s

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