1-s2.0-S0360319910017696-main

qud

e

tion

and minimum Gibbs energy model for reactions. The reforming and

a typical case were compared with those from the model of Ye et al. [15] which assumed

a plug flow for bed hydrodynamics. The model predictions show satisfactory agreement

rogen

an intermediate in a number of chemical and metallurgical

reformingor partial oxidationofnatural gas inparallel fixedbed

reactors within huge top-fired or side-fired furnaces, coupled

with Pressure Swing Adsorption (PSA) for hydrogenpurification

[2]. Although this technology has beenwidely used for decades,

it still suffers from several disadvantages such as low catalyst

CH4 H2O CO 3H2, DH298o 206 kJ/mol (R1)

CO H2O CO2 H2, DH298o 41 kJ/mol (R3)

The first two reactions are strongly endothermic, and both

lead to significant increase in molar flow rates as the reaction

* Corresponding author. Tel./fax: 86 20 22236985.

Avai lab le at www.sc iencedi rect .com

w.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9E-mail address: [email protected] (D. Xie).processes, for example in the production of ammonia and

methanol, upgrading of heavy hydrocarbon feed stocks, iron

ore reduction and food processing [1].

Most of the worlds hydrogen supply is generated by steam

CH4 2H2O CO2 4H2, DH298o 165 kJ/mol (R2)

Wateregas shift (WGS):worlds growing concerns about energy supply, security, air

pollution, and greenhouse gas emissions. Hydrogen offers

long-term potential for an energy system that produces near-

zero emissions based on each countrys domestically avail-

able resources. It is also amajor industrial commodity, used as

Fluidized Bed Membrane Reactors (FBMR) have been proposed

and developed by several research groups [1,4e7].

The principal reactions involved in catalytic steam

methane reforming are [8]:Steam methane reforming (SMR):Available online 20 September 2010

Keywords:

Hydrogen

Fluidized bed

Reforming

Membrane

Modeling

1. Introduction

An energy economy based on hyd0360-3199/$ e see front matter 2010 Profedoi:10.1016/j.ijhydene.2010.08.130with experimental data in the literatures. The influences of reactor pressure, temperature,

steam to carbon ratio, and permeate side hydrogen partial pressure on solid carbon, NHxand NOx formation were studied using the model.

2010 Professor T. Nejat Veziroglu. Published by Elsevier Ltd. All rights reserved.

could alleviate the

effectiveness, low heat transfer rates, large temperature gradi-

entswithin the bed and thermodynamic equilibriumconstraint

on chemical reaction [3e6]. To overcome these problems,25 August 2010

Accepted 28 August 2010separation processes were coupled by the mass balance. The model assumed a continu-

ously stirred tank reactor for the fluidized bed hydrodynamics. The model predictions forReceived in revised form membrane separationArticle history:

Received 29 May 2010

An equilibrium model of steam methane reforming coupled with in-situ membrane sepa-

ration for hydrogen production was developed. The model employed Sieverts Law forReaction/separation coupled emethane reforming in fluidize

Donglai Xie*, Weiyan Qiao, Ziliang Wang, W

MOE Key Laboratory of Enhanced Heat Transfer & Energy Conserva

a r t i c l e i n f o a b s t r a c t

journa l homepage : wwssor T. Nejat Veziroglu. Pilibrium modeling of steambed membrane reactors

ixing Wang, Hao Yu, Feng Peng

, South China University of Technology, Guangzhou 510640, China

e lsev ie r . com/ loca te /heublished by Elsevier Ltd. All rights reserved.

It did not consider the oxidant addition to the bed, nor solid

carbon formation as product. The coupling of the chemical

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 en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9 11799proceeds. Equilibrium conversions of both reforming reac-

tions benefit from high temperatures and low pressures,

whereas the wateregas shift reaction (R3), being exothermic

and having no change in the number of molars, benefits

thermodynamically from lower temperatures and is inde-

pendent of pressure.

To sustain the above endothermic reactions (R1) and (R2),

oxygen or air can be introduced into the system with the

following oxidation reactions taking place [9]:

CH4 0.5O2 CO 2H2, DH298o 36 kJ/mol (R4)

CH4 1.5O2 CO 2H2O, DH298o 607 kJ/mol (R5)

CH4 2O2 CO2 2H2O, DH298o 802 kJ/mol (R6)

Side reactions, like the catalytic dry reforming of methane,

methane cracking, etc. could also present accompanying the

above principle reactions, including [9e11]:

CH4 CO2 2CO 2H2, DH298o 247 kJ/mol (R7)

CH4 C 2H2, DH298o 75 kJ/mol (R8)

C H2O CO H2, DH298o 131 kJ/mol (R9)

C 0.5O2 CO, DH298o 111 kJ/mol (R10)

C CO2 2CO, DH298o 172 kJ/mol (R11)

Modeling of FBMR presents interesting challenges because

of the coupling of selective diffusion through the permeable

membrane with chemical reactions and mass transfer on the

reactor side [12]. It is useful to investigate the effects of key

operating parameters on the reactor performance, and also

the performance of the reactor systemcan be explored beyond

the range of parameters that can be studied experimentally

due to limitations imposed by economic and safety

considerations.

Two types of methods can be tried for the modeling of

chemical reactions, the kinetic approach and thermodynamic

approach. Many kinetic models have been proposed to simu-

late the steam methane reforming process for hydrogen

production, whose features have been summarized well by

Xie [13]. Predictions from these models were claimed by their

authors to be in good agreement with experimental data.

However, Kinetic models are limited to small numbers of

reactions and species with clearly defined mechanism. For

complex systems, the reaction mechanisms often require

extensive study. Also these models were solved by FORTRAN,

Matlab or other computer programs, which are not easily

accessible to design engineers in industry. For operating

conditions of interest (i.e., temperature 550e900 C andabsolute pressure 0.5e3.0 MPa), the steam reforming and

wateregas shift reactions, or the oxidation reactions in the

presence of catalyst are fast enough that the production of

hydrogen closely approaches the equilibrium values [14]. As

a result, thermodynamic equilibrium analysis provides

a simple and direct basis for practical applications. To theknowledge of the authors, two thermodynamic equilibrium

models have been developed to simulate the hydrogenreaction and membrane separation processes was achieved

by a sequential modular approach. The FBMRwas divided into

several successive steam methane sub-reformers and

membrane sub-separators. The process is represented by

(m 1) sub-reformers and m sub-separators. The reactor off-gases from the ith sub-reformer are fed to the ith sub-sepa-

rator, the non-permeated gases from the ith sub-separator are

fed to the (i 1)th sub-reformer, and the permeated hydrogenfrom the ith sub-separator accumulates in the (i 1)th sub-separator. Hence the coupling of the reaction and separation

process is a pseudo-coupling.

In practice, if autothermal operation is preferred, it is

better to use air, rather than pure oxygen as oxidant.When air

is introduced, the following side reactions involving nitrogen

could take place:

N2 3H2 2NH3 DH298o 98 kJ/mol (R12)

N2 2H2 N2H4 DH298o 95 kJ/mol (R13)

N2 2O2 2NO2 DH298o 68 kJ/mol (R14)

N2 O2 2NO DH298o 181 kJ/mol (R15)

N2 2O2 N2O4 DH298o 10 kJ/mol (R16)

N2O2 2N2O DH298o 163 kJ/mol (R17)

The possibility of the formation of NOx and NHx with the

presence of nitrogen needs to be studied, which has not been

considered previously.

In this work, a thermodynamic model is developed to

model the fluidized bed membrane reactor for hydrogen

production. In the model, the hydrogen separation process by

the membrane is coupled with the steam methane reforming

process with or without air addition for heat supply. The

formation of solid carbon, NOx and NHx was also considered.

2. Model development

2.1. Primary assumptions

A fluidized bed membrane reactor for pure hydrogen produc-

tion by steam methane reforming is shown schematically in

Fig. 1. Preheated high-temperature (usually 500e800 C) andhigh-pressure (usually 1e3 MPa) natural gas and steam areproduction process in fluidized bed membrane reactors in

literature. The model of Grace et al. [14] considered a Contin-

uously Stirred Tank Reactor (CSTR) for the bed hydrody-

namics, and pure oxygen for autothermal operation. Solid

carbon formationwas studied using themodel. The amount of

hydrogen separated by membrane was used as an input for

themodel. Hence the reaction process and separation process

was not coupled. The model of Ye et al. [15] considered the

FBMR as a Plug Flow Reactor (PFR) for reactor hydrodynamics.premixed and fed into the reactor. The reactor contains Nickel

based catalyst and palladium (or its alloy) membranemodules

in tubular or planner shapes. Usually palladiumesilver or

palladiumecopper membrane modules are widely employed

due to their relatively high permeability and longevity

comparing to pure palladium modules. For these membrane

modules the reactor temperature should be controlled below

700 C to avoid damaging the membranes. For autothermaloperations, air is introduced into the reactor to produce the

heat for sustaining the endothermic reforming reactions.

Otherwise, external heatneeds tobe supplied to the reactor. To

reduce thepermeate sidehydrogenpartial pressure, sweepgas

(usually steam) can be introduced into thepermeate side of the

membrane modules. To simplify the simulation of steam

methane reforming and hydrogen separation processes in the

FBMR, the domain sketched by the dashed box in Fig. 1 is

considered for model development.

To represent the characteristics of fluidized bed reactors,

the model assumes:

1. The system is at steady state and reaches thermodynamic

equilibrium.

2. The reacting gas and catalyst are perfectly mixed, i.e., the

reactor is a continuously stirred tank reactor.

3. Uniform temperature within the fluidized bed.

4. Pressure gradients are ignored both within the reactor bed

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 911800and within the membrane.

5. Heat losses are negligible.

6. Ideal gas laws are applicable to all reaction gases.

7. Tomake themodelmore general, consider CH4, H2O, O2 and

N2 in the feed stream. The reactor of gas constitutes of 15

species, including CO, CO2, H2O, CH4, H2, O2, N2, C2H2, C2H4,

NO, NO2, N2O, NH3, N2H4 and solid carbon.

8. Only hydrogen can penetrate through the membrane i.e.,

only hydrogen and sweep gas are considered in the

permeate side.

Methane & steam Sweep gas

Membranes module

Air

Hydrogen + Sweep gas

Domain considered in the model

Non-permeate product gasFig. 1 e Schematic of a fluidized bed membrane reactor.2 2 4 2 2 2 2 2 2 4 2

N2O, NH3, N2H4 and solid carbon.

The elements of C, O, N, and H conserve prior to and after

the reactions. Hence for the elements of C, O and N,

X159. Hydrogen permeation through the membrane follows Sie-

verts Law [16], i.e.:

QH h k CepPMRH PMMH

e

EPRT

(1)

The hydrogen flux follows the Sieverts law when the

hydrogen pressure exponentM is equal to 0.5, which is usually

valid for thick Pd films [17]. Deviations from the Sieverts law

(M> 0.5) were reported for very thinmembranes [18,19]. Based

on a hydrogen permeation model, Ward and Dao [20] showed

that at temperatures above 400 C, M was equal to 0.5 for

membranes thicker than 1 mm. Usually to use Sieverts Law

correctly with exponent 0.5, the thickness of membrane

should be higher than 10 mm [17]. The co-existence of H2O, CO,

CO2 or CH4 has been reported to have a negative influence on

the membrane separation performance [21]. This negative

effect is considered in the permeation efficiency factor e h,

together with some other factors influencing the membrane

performance, like the existence of a membrane substrate. In

practice, h should be determined experimentally. h is reported

in literature to be from 0.39 to nearly 1.0 [3,12,22].

10. Although solid carbon is considered in the current model,

the purpose is to study the operating regime that the solid

carbon may appear and which should be avoided in prac-

tical applications. Simulation results in the following part

of this study show that solid carbon is formed under some

extreme conditions. For fluidized beds, carbon formation

is not concerned under regular operation conditions due to

the high bed temperature uniformity [15]. Hence the

influence of carbon accumulation on membrane and

catalysts on their performances is not considered in the

current model.

2.2. Governing equations

Themost commonly used function to identify the equilibrium

state is Gibbs free energy, which is a suitable parameter to

calculate the equilibrium compositions of the reaction

system. The equilibrium composition brings in temperature

dependence, without requiring detailed information

regarding the specific reaction or catalyst performance.

As the system reaches thermodynamic equilibrium, the

total Gibbs free energy of the system should reach minimum

under the operation temperature. Hence

vGtvni

0; i 1e15 (2)

i 1e15 represents the fifteen species in the system, in theorder of CO, CO , H O, CH , H , O , N , C H , C H , NO, NO ,i1niaik Ak; k 1;2; 3 (3)

namics on reactor performances is minimized. When

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 en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9 118012.3. Solution of the governing equations

Using Gibbs free energy minimization approach to predict the

product composition of a reaction system that reaches ther-

modynamic equilibrium is a well-established method. The

mathematical aspects of the method are well documented

[24e28]. It has been widely used in recent years for process

optimization in energy sectors, for example, gasification

[29e32] and hydrogen production from reforming related

processes [32e37]. In the current study, the model defined by

control equations of (1)e(8) was solved with two approaches.

A) Solved in Matlab with NewtoneRaphson algorithm. With

somemodifications on themass balance of hydrogen element

considering its permeation through membranes, the compu-

tational program developed in the current study is following

the work of Li et al. [30], where detailed algorithm information

is described. B) Solved inMicrosoft Excel using its solver tool.

In Excel, the minimum Gibbs energy problem becomes an

optimization problem. In the programming solver tool, the

Gibbs energy of the reactor product as a function of the system

pressure, temperature and molar flow rate of product is set as

an objective to reach minimum value. The variables are the

molar flow rates of product gases and solid carbon. The

constraints are those mass balance equations (3) and (4).

Thermodynamic properties of the species involved were

tabled in a separate work sheet and can be accessed by the

solver tool. By employing the solver tool, Excel can automat-

ically adjust the product molar flow rates to make the Gibbsk 1, 2, 3 represents the three elements of C, O and N,respectively.

For element of H, consider part of the hydrogen in the

reacting gas permeates through the membrane:

X15i1

niai;H 2QH AH (4)

The total Gibbs free energy of the system is linked to the

composition of the system by

Gt X14i1

niGi nCGc (5)

Gi is related to the system pressure and temperature by

Gi G0i RT lnyiPP0

(6)

where

yi niX14i1

ni

(7)

For solid carbon, its Gibbs free energy is related to the system

pressure by

GC VCP P0 (8)Where VC is the molar volume of solid carbon and equals to

4.58 106 m3/mol [23].energy of the system reaching minimum. In both approaches,

the thermodynamic data of Gi0 for gases were obtained frommembrane surface area is large enough, the equilibrium of the

reactions became the controlling step, and the influence ofJANAF Thermo Chemical Tables [38]. Both approaches can

give identical result.

2.4. Comparisons with PFR model of Ye et al. [15]

Strong gas back mixing exists in bubbling fluidized beds.

However, gases are still flowing upwards in the bed. Either PFR

or CSTR hydrodynamic models, or combination of these two

models, has been employed in the reformer models. If no

membrane modules were employed for in-situ hydrogen

separation, the thermodynamic models with either PFR or

CSTR hydrodynamic sub-models should give identical

reformer performance predictions.Whenmembranemodules

were employed, the driving force for hydrogen permeation

through membrane is related to the reactor hydrodynamic.

Hence the reactor performance depends on the selection of

hydrodynamic sub-models.

Consider a fluidized membrane reactor that has feed gas

streams of CH4 and steam. In-situ hydrogen separation

modules with membrane thickness of 10 mm and effective

surface area varying from 0 to 1.8 m2 are installed inside the

reactor. The corresponding Cep varies from 0 to 180 Km. The

base operation condition is P 2.0 MPa, T 650 C,FCH4 226 mol/h, FS 0, SC 3.0, OC 0 mol/h, P 2.0 MPa,PM 0.1MPa. Themodel of Ye et al. [15] was used to predict thereactor performance under the conditions of plug flow, and it

was comparedwith the predictions from the currentmodel for

the conditions assumingCSTR for reactor hydrodynamics. The

reactor performance was denoted by two parameters: CH4conversion and H2 yield, and they are defined as:

XCH4 nCO nCO2

FCH4(9)

YH2 QHFCH4

(10)

The predicted reactor performances from bothmodels were

illustrated in Fig. 2. When very few membrane surface area

was installed, i.e., Cep below 15 Km for the operation condi-

tions specified, the predicted methane conversion and

hydrogen yield from both models are very close, with the

predictions from PFR model slightly higher than those pre-

dicted fromCSTRmodel. Under such conditions, the hydrogen

yield increases almost linearly with increasing the membrane

permeation capacity, indicating that the separation is the

controlling step for hydrogen production. When Cep varies

from 15 to 60 Km, the predicted methane conversion and H2yield from PFRmodel are obviously higher than those from the

CSTR model. When membrane permeation capacity is high

enough, i.e., Cep higher than 60 Km for the condition studied,

the reactor performances from both models became close

again. This is because that the plug flow reactor can provide

a higher driving force for hydrogen permeation through

membranes than a CSTR. However, if very few membrane

surface areas are employed, the effect of reactor hydrody-membrane permeation capacity on reactor performance is

also suppressed.

3. Comparison of model predictions withexperimental data

0.20e0.28 mm) palladium membrane tubes, each has outside

diameter of 4.7 mm were installed in the reactor. Table 1

compares experimental data of methane conversion,

hydrogen production rate from the membrane, and product

gas composition with the predictions from the current model,

as well as the predictions from the PFR model of Ye et al. [15].

The permeation efficiency h is taken as 0.39, as suggested by

Adris et al. [12]. It can be seen from the table that except for

some points (for example CH4 concentration at 640 C and CO2concentration at 542 C), the predictions are in satisfactoryagreement with the experimental data in general. The

discrepancies could be caused by several factors, for example,

the error on selection of the membrane permeation efficiency

factor, errors on reactor temperature, pressure and gas

composition samplings, chemical reactions in the reactor

freeboard that is not considered in the current model, etc. It is

also worth mentioning that the predictions from both models

are almost identical. This is consistent with the previous

statement that when Cep is low, the predictions from both

CSTR and PFR models are very close.

Roy [39] tested a fluidized bedmembrane reactor with high

flux membrane tubes. The tubes had a substrate thickness of

76 mmand palladium coating of 5.2 mmon both the outside and

inside surfaces. The outer diameter of the tubes was

0 20 40 60 80 100 120 140 160 1800.4

0.5

0.6

0.7

0.8

0.9

1.0

1.1

1.2

0.0

0.5

1.0

1.5

2.0

2.5

3.0

3.5

4.0

Hyd

roge

n y

ield

(mo

l/mo

l CH 4

)

CH4

conver

sion ( -

)

Membrane permeation capacity (Km)

CH4 conversion CSTR CH4 conversion PFR H2 yield CSTR H2 yield PFR

Fig. 2 e Influence of Cep on methane conversion and

hydrogen yield. (FCH4 [ 226.41 mol/h, FS [ 0; SC [ 3.0,

OC [ 0, T [ 650 C, P [ 2.0 MPa, PM [ 0.1 MPa).

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 911802The FBMR process has been intensively studied by researchers

in recent years. Some experimental data are available from

the literature for model validations. Experimental results by

Adris et al. [12], Roy [39], andMahecha-Botero et al. [5] are used

to validate the current model performances.

Adris et al. [12] reported experimental results from a pilot-

scale fluidized bed membrane reactor for reaction tempera-

tures from 447 to 640 C and steam to carbon molar ratio (SC)of 2.4. The reactor has a diameter of 97 mm and length of

1.143 m. Twelve thin-walled (nominal wall thickness ofTable 1 e Comparison of equilibrium model predictions with eSC[ 2.4, OC[ 0, Fs[ 80 mol/h, Cep[ 0.4 Km, P[ 0.98 MPa, P

Bed temperature (C) 447

Methane conversion () Experimental data 0.12CSTR prediction 0.11

PFR prediction 0.11

H2 production (mol/h) Experimental data 1.70

CSTR prediction 1.73

PFR prediction 1.77

Product gas composition

(vol%, dry basis)

CH4 Experimental data 61.7


PFR prediction 62.7

CO Experimental data 0.10


PFR prediction 0.30

CO2 Experimental data 9.5

CSTR prediction 7.8

PFR prediction 7.6

H2 Experimental data 28.7


PFR prediction 29.43.175 mm. The following parameters were evaluated by

experimentation: the permeability pre-exponential factor k

was 7.85 109 mol (m s Pa0.72); activation energy Ep was11.5 kJ/mol; pressure exponent M was 0.72. Hence these

numbers were employed in the current model for this case.

Nine high flux membrane tubes were installed in a 100 mm

diameter reactor, where oxygenwas fed to provide the heat. In

the fluidized bed reactor, the coverage of membrane outside

surface by catalyst dust, exposure of membrane tubes to two

different phases (bubble phase and dense phase) and the gas

xperimental data of Adris et al. [12]. (FCH4 [ 74.2 mol/h,

M[ 0.4 MPa, h[ 0.39).

494 542 594 640

0.17 0.24 0.33 0.43

0.16 0.22 0.32 0.41

0.16 0.22 0.34 0.42

2.50 3.57 4.81 6.23

2.51 3.50 4.77 6.23

2.55 3.61 4.95 6.30

49.6 37.6 27.3 19.5

51.9 41.8 31.8 23.9

52.3 42.3 32.1 24.0

0.40 1.20 3.20 5.60

0.52 1.25 2.73 4.80

0.50 1.24 2.60 4.70

11.5 13.0 13.5 13.3

9.6 11.2 12.1 12.0

9.5 10.8 11.7 11.7

38.5 48.2 56.0 61.6

38.0 45.8 53.4 59.337.6 45.9 55.3 59.5

taL9

ers

Experimental Predih 0.9

0.9

0.9

0.9

0.9

0.9

0.8

0.7

0.7

0.8

0.8

0.7

0.8

0.8

0.9

0.8

0.8

0.8

0.8

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 en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9 11803phasemass transfer resistancemay affect themembrane tube

permeability [39]. They used permeation efficiency h to

account for the influence of above factors on the membrane

permeability. However, the actual value of h was not dis-

closed. In the current model, values of 0.3 and 0.1 were tried.

Table 2 gives results of model predictions in comparison with

experimental results. It can be seen that the predictions from

h 0.1 was closer to the experimental results.

0.68 650 4.1 0.44 0.76

0.68 650 4.1 0.45 0.77

0.68 650 4.1 0.5 0.79

0.68 650 4.1 0.56 0.81

0.68 650 4.1 0.62 0.82

0.68 650 4.1 0.35 0.74

0.68 650 3.1 0.35 0.66

0.68 650 2.4 0.35 0.59

0.99 600 3.1 0.40 0.52

0.99 625 3.1 0.40 0.56

0.99 650 3.1 0.40 0.6

0.68 576 4.1 0.45 0.62

0.68 600 4.1 0.45 0.67

0.68 625 4.1 0.45 0.73

0.68 649 4.1 0.45 0.77

0.68 600 4.1 0.45 0.68

0.78 600 4.1 0.45 0.67

0.88 600 4.1 0.45 0.63

0.99 600 4.1 0.45 0.62Table 2 e Comparison of model predictions with experimenCep[ 8.16 Km, P [ 0.68 MPa, PM[ 0.14 MPa, k[ 7.85 3 10

Bed P (MPa) Bed T (C) SC OC CH4 convMahecha-Botero et al. [5] tested a pilot-scale fluidized bed

membrane reactor for the production of hydrogen. The

reactor was operated under steam methane reforming and

autothermal reforming conditions, without membranes and

with membranes of different total areas. Heat was added

either externally or via direct air addition. The reactor is

contained in a stainless steel vessel of height 2 m and rect-

angular cross-sectional area of 48.4 cm2. The vessel can hold

up to six double-sided membrane modules with a nominal

permeation area of 300 cm2. The membrane modules con-

tained Pd/Ag foil of thickness 25 mm sealed onto a porous

metal backing with a barrier layer to prevent interdiffusion.

The authors suggested in another paper [3] that the

membrane efficiencies were ranging from w0.6 to 0.9. Here

the efficiency of 0.6 is applied in the current model. Table 3

compares the experimental results with the model predic-

tions. It can be seen that themodel slightly over-estimated the

natural gas conversion, the hydrogen concentration in the

reactor off gas and hydrogen production rate.

4. Effects of operation conditions on carbonformation

Coke formation is a serious operational problem in conven-

tional steam reforming. The presence of the solid carbon candeactivate the catalyst and, hence degrading the performance

of the reforming system [14]. The operating regime that the

solid carbon may appear should be avoided in practical

applications.

4.1. Influence of steam to carbon ratio and membranepermeation capacity

l data of Roy [39]. (FH2O [ 138.33 mol/h, FS[ 45 mol/h,mol/(m.s Pa0.72), Ep [ 11.5 KJ/mol, M[ 0.72).

ion () Hydrogen production (mol/h)cted0.3

Predictedh 0.1

Experimental Predictedh 0.3

Predictedh 0.1

2 0.87 18.6 43.0 21.7

3 0.86 18.6 42.6 21.6

4 0.88 18.5 41.4 21.2

5 0.90 17.6 39.8 20.6

6 0.92 17.1 38.1 19.9

1 0.85 19.1 45.1 22.5

5 0.77 20.5 50.9 24.1

7 0.70 21.4 54.8 25.1

4 0.63 19.0 40.2 23.5

0 0.69 21.0 45.2 26.6

6 0.75 23.2 49.9 29.7

8 0.70 14.6 32.3 15.8

4 0.76 15.9 36.1 17.8

9 0.82 17.5 39.7 19.8

3 0.87 18.6 42.5 21.5

4 0.76 14.8 36.1 17.8

4 0.86 16.0 39.2 23.5

3 0.85 16.3 42.1 25.2

5 0.85 17 44.9 27.06Fig. 3 shows the influence of steam to carbon ratio and

membrane permeation capacity on carbon formation under

the operation conditions of FCH4 1000 mol/h, FS 0, OC 0.2,P 2.0MPa, T 650 C, PM 0.1MPa. It can be seen that carbonformation is suppressed by higher steam to carbon ratio. The

in-situ removal of hydrogen through membranes enhanced

the carbon formation. The coking boundary moves toward

higher SC values as the reactor Cep is raised. For example, for

the operation conditions investigated, when Cep increased

from 1.0 Km to 10.0 Km, the coking boundary increased from

an SC value of 0.8e1.0. With higher membrane permeation

capacities per unit feed of methane and steam, more

hydrogen is removed from the reactor, and reaction (R8) is

pushed forward to produce more solid carbon.

4.2. Influence of temperature and membrane permeationcapacity

Fig. 4 shows the influence of temperature and membrane

permeation capacity on carbon formation for the operation

conditions of FCH4 1000 mol/h, FS 0, OC 0.2, P 2.0 MPa,T 650 C, PM 0.1 MPa. The range of the temperatureinvestigated is between 450 and 850 C. It should be pointedout that when palladium alloy membrane is installed, the

reactor temperature is usually maintained below 700 C forthe protection of membrane. It can be seen that temperature

is a strong factor for the formation of solid carbon. The

the carbon formation. The coking boundary moves toward

Table

3eCompariso

nofmodelpredictionswithexperimentaldata

ofMahech

a-Botero

etal.[5].(Fs[

0,SC[

3.0,Cep[

0.4

Km,P[

0.98MPa,k[

1.373

10L3mol/

(m$h$Pa0.5),Ep[

20.5

KJ/mol,M[

0.5,h[

0.6).

Expt

Membrane

area(m

2)

OC

Naturalgas

feed

(mol/h)

Reactor

pressure

(MPa)

Permeate

pressure

(MPa)

Methaneconversion()

H2molarfractionin

ROG(dry)(%

)Perm

eate

H2flow

(mol/h)

Expt

Prediction

Expt

Prediction

Expt

Prediction

10

040

0.75

N/A

0.29

0.30

43.5

53.7

N/A

20

040

1.00

N/A

0.22

0.27

33.3

51.0

N/A

30

0.35

40

1.00

N/A

0.40

0.45

22.2

31.4

N/A

40.09

040

0.75

0.10

0.25a

0.39

40.1

49.2

13.8

21.3

50.09

040

0.75

0.03

0.66

0.50

41.1

43.5

39.3

47.7

60.09

040

1.00

0.03

0.66

0.51

33.3

39.2

42.0

54.3

70.09

020

0.75

0.03

0.43a

0.66

29.4

35.6

41.5

40.6

80.09

020

1.00

0.03

0.67

0.69

20.6

30.8

41.1

45.3

90.18

0.35

40

1.00

0.10

0.46a

0.58

22.0

26.8

N/A

29.4

10

0.18

0.35

40

1.00

0.03

0.73

0.78

17.5

19.9

49.6

72.2

11

0.18

0.35

40

1.00

0.03

0.69

0.78

16.2

19.9

48.2

74.2

12

0.18

0.35

20

1.00

0.03

0.81

0.92

11.9

14.2

36.6

50.9

13

0.18

0.35

13.3

1.00

0.03

0.76a

0.96

8.7

11.2

32.6

37.5

14

0.18

013.3

1.00

0.03

0.73

0.95

8.7

15.5

39.7

47.6

aIn

bedgassamplingnotperform

ed.Themethaneconversionwascalculatedfrom

reactoroffgascompositionanalysis.

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1.0 1.10.00

0.05

0.10

0.15

0.20

0.25

0.30

0.35

0.40

Carb

on y

ield

(mol/

mol C

H 4)

Steam to carbon ratio (mol steam/ mol CH4)

Cep=0.4 KmCep=1.0 KmCep=5.0 KmCep=10.0 Km

Fig. 3 e Influence of steam to carbon ratio and membrane

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 911804lower temperatures as the reactor Cep is raised. For example,

for the operation conditions investigated, if Cep increased from

1.0 Km to 10 Km, the coking boundary decreased from 600 Cto 500 C. For the formation of solid carbon, reaction (R5) isendothermic and (R9eR11) are exothermic. With the temper-

ature increasing, reactions (R9eR11) are shifted to the right

direction while reaction (R8) went to the opposite direction.

The net effect of temperature on carbon formation behaves

then as shown in Fig. 4.

4.3. Influence of reactor pressure and membranepermeation capacityamount of carbon initially increased and then decreased with

the temperature increasing. Again, the membrane enhanced

permeation capacity on carbon yield. (FCH4 [ 1000 mol/h,

FS [ 0, OC [ 0.2, P [ 2.0 MPa, T [ 650 C, PM [ 0.1 MPa).Reactions (R8eR11) are related to solid carbon formation in the

reforming process. In these reactions, reaction (R8) should be

450 500 550 600 650 700 750 800 8500.00

0.05

0.10

0.15

0.20

0.25

0.30

Carb

on yi

eld

(mol/

mol C

H 4)


Temperature ( oC )

Fig. 4 e Influence of temperature and membrane


FS [ 0, SC [ 0.8, OC [ 0.2, P [ 2.0 MPa, PM [ 0.1 MPa).

0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.50.00

0.05

0.10

0.15

0.20

0.25

0.30

Carb

on y

ield

(mol/

mol C

H 4)

Cep=0.4 KmCep=1.0 KmCep=5.0 KmCep=10.0 KmCep=20.0 Km

Pressure (MPa)

Fig. 5 e Influence of pressure and membrane permeation

capacity on carbon yield. (FCH4 [ 1000 mol/h, FS [ 0,

OC [ 0.2, SC [ 1.0, T [ 650 C, PM [ 1.0 MPa).

400 450 500 550 600 650 700 750 800 8500

50

100

150

200

250

300

0.0000

0.0002

0.0004

0.0006

0.0008

0.0010

0.0012

N

H3

yiel

d (m

ol/m

ol C

H4)

Cep=0.4 Km Cep=1.0 KmCep=5.0 KmCep=10.0 Km

Temperature ( oC )

Mola

r fra

ctio

n o

f NH

3 (pp

m)


Fig. 7 e Influence of temperature and membrane

permeation capacity NH3 molar fraction and 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 en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9 11805the dominant process as it is the source of solid carbon. These

processes follow LeChateliers principle, so that the formation

of solid carbon benefit from low pressure. In an FBMR system,

a perm-selective membrane installed in the fluidized bed

reduces the adverse effect of high pressure by removing

product hydrogen. Ye et al. [15] concluded from their simu-

lation that over the range of conditions investigated, at low

permeation capacity, methane conversion decreases with

increasing reactor pressure. At medium permeation capacity,

the influence of pressure on methane conversion is almost

neutral. At high permeation capacity, methane conversion

increases with increasing reactor pressure. Here for the

formation of solid carbon, the same trends from the influence

of membrane permeation capacity and pressure were

observed as that of the methane conversion, as shown in

Fig. 5. At low permeation capacity (Cep 0.4e5.0 Km), carbon0.0 0.1 0.2 0.3 0.4 0.50.00

0.05

0.10

0.15

0.20

0.25

0.30

Carb

on y

ield

(mol/

mol C

H 4)

Cep=0.4 KmCep=1.0 KmCep=5.0 KmCep=10.0 KmCep=20.0 Km

Oxygen to carbon ratio (mol O2/ mol CH4)

Fig. 6 e Influence of oxygen to carbon ratio and membrane


FS [ 0, OC [ 0.2, SC [ 0.8, T [ 650 C, P [ 2 MPa,PM [ 0.1 MPa).formation decreases with increasing reactor pressure. At

medium permeation capacity (Cep 10.0 Km), the influence ofpressure on carbon formation is almost neutral. If a high

permeation capacity (Cep 20.0 Km) could be used, carbonformation would increase with increasing reactor pressure.

4.4. Influence of oxygen to carbon ratio and membranepermeation capacity

Obviously, oxygen input can reduce the amount of solid

carbon formation through reaction (R10). Again, the in-situ

removal of product hydrogen could enhance the carbon

formation. Fig. 6 shows the influence of OC and membrane

permeation capacity on carbon formation at the operation

condition of FCH4 1000 mol/h, FS 0, SC 0.8, T 650 C,P 2.0 MPa, PM 0.1 MPa. It can be seen that with increasing

(FCH4 [ 1000 mol/h, SC [ 3.0, OC [ 0.2, P [ 2 0.0 MPa,

PM [ 0.1 MPa).the oxygen to carbon ratio, the mole of carbon formed is

decreasing for all Cep values. The coking boundary moves

toward higher OC ratio as the reactor Cep is raised. For the

operation conditions investigated, when Cep increased from

0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.540

60

80

100

120

140

160

180

0.0000

0.0001

0.0002

0.0003

0.0004

0.0005

0.0006

0.0007

0.0008

0.0009

N

H3 yi

eld

(mol/

mol C

H4)

Mol

ar fr

actio

n o

f NH

3 (pp

m)


Pressure (MPa)


Fig. 8 e Influence of pressure and membrane permeation

capacity on NH3 molar fraction and yield. (FCH4 [ 1000 mol/

h, SC [ 3.0, OC [ 0.2, T [ 650 C, PM [ 0.1 MPa).

by-products was investigated for the conditions of

FCH4 1000 mol/h, FS 0, SC 0.5e5.0, OC 0.05e0.5,

membranes, the hydrogen concentration in the reactant gases

confirmed by the model simulation results as shown in Fig. 7.

0 1 2 3 4 5

90

100

110

120

130

140

0.0002

0.0003

0.0004

0.0005

0.0006

0.0007

0.0008

0.0009

0.0010



N

H3

yiel

d (m

ol/m

ol C

H

4)

Mola

r fra

ctio

n o

f NH

3 (pp

m)



permeation capacity on NH3 molar fraction and yield.

(FCH4 [ 1000 mol/h, FS [ 0, OC [ 0.2, T [ 650C,

P [ 2.0 MPa, PM [ 0.1 Mpa).

0.0 0.1 0.2 0.3 0.4 0.5

70

80

90

100

110

120

130

0.0002

0.0003

0.0004

0.0005

0.0006

0.0007

0.0008

0.0009

0.0010

0.0011

0.0012

N

H3

yiel

d (m

ol/m

ol C

H4

)

Mola

r fra

ctio

n o

f NH

3 (pp

m)

Oxygen to carbon ratio (mol O2/ mol CH4)


Cep=0.4 KmCep= 1.0 KmCep= 5.0 KmCep=10.0 Km


permeation capacity on NH3 molar fraction and yield.

(FCH4 [ 1000 mol/h, FS [ 0, SC [ 3.0, T [ 650C,

P [ 2.0 MPa, PM [ 0.1 MPa).

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 911806T 400e800 C, P 0.5e3.0 MPa, PM 0.1 MPa,Cep 0.4e10.0 Km. It was found that under such operation0.4 Km to 20 Km, the coking boundary for OC increased from

0.15 to 0.50.

5. Effects of operation conditions on NH3formation

When air is introduced into the reactor for autothermal

operation, the nitrogen in the air could involve with reactions

to generate NO, NO2, N2O, NO2, N2H4 and NH3 through reac-

tions (R12eR17), which are not desired. The formation of theseconditions, the concentration of NO, NO2, N2O, NO2, N2H4 in

0.0 0.5 1.0 1.5 2.0 2.5 3.0 3.5 4.0 4.5 5.0 5.50.08

0.10

0.12

0.14

0.16

0.18

0.20

0.22

0.24

0.26

0.28

0.30

Steam to carbon ratio (mol steam / mol CH4)

Mola

r fra

ctio

n (-)

Cep=0.4 Km Cep=0.4 KmCep=1.0 Km Cep=1.0 KmCep=5.0 Km Cep=5.0Km Cep=10.0 Km Cep=10.0 Km


permeation capacity on N2 and H2 molar fraction.

(FCH4 [ 1000 mol/h, FS [ 0, OC [ 0.2, T [ 650C,

P [ 2.0 MPa, PM [ 0.1 Mpa) (solid symbols: N2, empty

symbols: H2).is decreased. Hence the membrane helps to reduce the

formation of NH3 i.e., higher membrane permeation capacity

helps to reduce the NH3 concentration in the system. This isthe reactant gas is below 10 ppm. Hence the formation of

these species will not be discussed here, and only NH3formation is presented.

5.1. Influence of temperature and membrane permeationcapacity

Reaction (R12) is exothermic; therefore increasing operation

temperature can shift the equilibrium backward and suppress

NH3 formation. With the in-situ removal hydrogen with0.0 0.1 0.2 0.3 0.4 0.5

0.04

0.08

0.12

0.16

0.20

0.24

0.28

Mola

r fra

ctio

n of H

2 an

d N

2


Cep=0.4 Km Cep=0.4 KmCep=1.0 Km Cep=1.0 KmCep=5.0 Km Cep=5.0 KmCep=10.0 Km Cep=10.0 Km

Fig. 12 e Influence of oxygen to carbon ratio and

membrane permeation capacity on N2 and H2 molar

fractions. (FCH4 [ 1000 mol/h, FS [ 0, SC [ 3.0, T [ 650C,

P [ 2.0 MPa, PM [ 0.1 MPa) (solid symbols: N2, empty

symbols: H2).

NH3 concentration in the system.

3. NH3 concentration and yield increase with increasing the

operation pressure.

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 en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9 11807It can be seen that for the operation conditions studied, the

NH3 concentration in the product gas is between 50 and

220 ppm and the NH3 yield is around 3 105 mol/mol of CH4.

5.2. Influence of pressure and membrane permeationcapacity

Fig. 8 shows the influence of pressure and membrane

permeation capacity on the NH3 formation. It can be seen that

the NH3 concentration increases with increasing the opera-

tion pressure. This can be explained by LeChateliers principle

that NH3 formation is favored by high pressure as the product

moles are less than the reactants in reaction (R12). Again

membrane helps to reduce the formation of NH3.

5.3. Influence of steam to carbon ratio

When other operation conditions are fixed, the steam to

carbon ratio affects the NH3 formation by the following

processes:

A) Higher steam feed diluted the N2 concentration in the

reactant streams, which favors the NH3 formation;

B) At lower steam to carbon ratios, when it is increased, the

methane conversion is promoted, and lead to higher H2concentration in the reactant streams. If steam to carbon

ratio is further increased, the extra steam dilutes the

hydrogen in the stream.

Fig. 9 shows the influence of steam to carbon ratio on NH3molar fraction and yield for Cep 0.4, 1.0, 5.0, 10.0 Km. It can beseen that NH3 yield increases with increasing the steam to

carbon ratio. However, its molar concentration reached

maximum values around steam to carbon ratio of 1.0e1.5.

Fig. 10 shows the N2 and H2 concentration in the reactor.

When steam to carbon ratio increases, the N2 concentration is

decreasing, while the H2 concentration increase until SCreached approximately 2.5, and then decreases with further

increasing of the SC values.

5.4. Influence of oxygen to carbon ratio

When other operation conditions are fixed, the oxygen to

carbon ratio affects the NH3 formation by the following two

processes:

A) With higher oxygen to carbon ratio, the N2 concentration

in the reactant stream is increased, which favors the

formation of the NH3;

B) Higher air feed to the reactor consumed and diluted the H2concentration in the reactant streams which in return

decreased molar number of NH3 in the stream.

Fig. 11 shows the influence of oxygen to carbon ratio on

NH3 molar fraction and yield for Cep 0.4, 1.0, 5.0, 10.0 Km. Itcan be seen that NH3 yield increases with increasing the

oxygen to carbon ratio. However, its molar concentration

reached maximum values around oxygen to carbon ratio of0.35. Fig. 12 shows the N2 and H2 concentration variations in

the reactor with the steam to carbon ratio. It can be seen that4. NH3 yield increases with increasing steam to carbon ratio.

However, its concentration reached maximum values

around steam to carbon ratio of 1.0e1.5.

5. NH3 yield increases with increasing oxygen to carbon ratio.with increasing the oxygen to carbon ratio, the N2 concen-

tration increases and H2 concentration decreases. The overall

tendency of NH3 formation by the influence of oxygen to

carbon ratio then behaves as illustrated in Fig. 11.

6. Conclusions

A reaction/separation coupled thermodynamic equilibrium

model was developed to simulate hydrogen production with

fluidized bed membrane reactors from steam methane

reforming reactions with or without air addition. The model

assumes a CSTR for bed hydrodynamics. Themodel considers

CH4, H2O, O2 and N2 in the feed stream and the reactor off gas

constitutes of 15 species, including CO, CO2, H2O, CH4, H2, O2,

N2, C2H2, C2H4, and NO, NO2, N2O, NH3, N2H4 and solid carbon.

Model predictions on reactor performance show satisfactory

agreement with the experimental results of Adris et al. [12],

Roy [39], and Mahecha-Botero et al. [5].

Parametric studies of the carbon formation from the steam

methane reforming with the model show that:

1. Over the range of condition investigated, the in-situ removal

of hydrogen through membranes enhances the carbon

formation. Carbon formation is suppressed by higher steam

to carbon ratio.

2. Temperature is a strong factor for the formation of solid

carbon. Over the range of conditions studied, the amount of

carbon formed initially increased and then decreased with

the temperature increasing. Solid carbon yield decreases

with increasing steam to carbon ratio.

3. For the conditions studied, at low permeation capacity

(Cep 0.4e5.0 Km), carbon formation decreases withincreasing reactor pressure. At medium permeation

capacity (Cep 10.0 Km), the influence of pressure oncarbon formation is almost neutral. If a high permeation

capacity (Cep 20.0 Km) could be used, carbon wouldincrease with increasing reactor pressure.

4. O2 input can reduce the amount of solid carbon formation.

NH3 formation was also studied by the current model.

Parametric studies of the NH3 formation from the steam

methane reforming with the model show that over the range

of operations studied:

1. Increasing operation temperature can shift the equilibrium

backward and decrease NH3 formation.

2. The membrane helps to reduce the formation of NH3 i.e.,

higher membrane permeation capacity helps to reduce theHowever, its concentration reached maximum values

around steam to carbon ratio of 0.30e0.35.

(project # 2009ZZ0013) are gratefully acknowledged.

i n t e rn a t i o n a l j o u r n a l o f h y d r o g e n en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 911808r e f e r e n c e s

[1] Xie D, Grace JR, Lim CJ. Experimental study of gas and solidcirculation in an internally circulating fluidized bedNotation

aik number of atoms of the kth element present in each

molecule of species i in reacting system, e

Ak total molar flow rate of kth element in the feed,

mol h1

Cep membrane permeation capacity (membrane surface

area/thickness), Km

EP activation energy, J mol1

FCH4 molar flow rate of methane feed, mol h1

Fs molar flow rate of sweep gas, mol h1

Gt total Gibbs free energy, J mol1

Goi Gibbs free energy of species i at its standard state,

J mol1

Gi Gibbs free energy of pure species i at operating

conditions, J mol1

GC molar Gibbs free energy of solid carbon, J mol1

QH total molar flow rate of hydrogen extracted through

perm-selective membranes, mol h1

k pre-exponential factor, mol Km1 h1 MPaM

m number of sub-separators used in the model of Ye

et al. [15], e

ni molar flow rate of specie i in the product gas and

solid carbon, mol h1

M pressure exponent in Equation (1)

OC oxygen to air ratio, e

P total absolute pressure of the reactor, MPa

P0 standard state pressure, MPa

PRH absolute partial pressure of hydrogen in the reactor,

MPa

PMH absolute partial pressure of hydrogen in the

permeate side, MPa

PM absolute pressure in the permeate side, MPa

R universal gas constant, J mol1 K1

SC steam to carbon ratio, e

T temperature, CVC molar volume of solid carbon, m

3 mol1

XCH4 conversion of CH4, defined in Equation (9), e

YH2 H2 yield, defined in Equation (10), e

yi molar fraction of species i in the product gas, e

h permeation efficiency factor, eAcknowledgement

Financial support fromtheNationalHighTechnologyResearch

and Development Program of China (2009AA05Z102) and the

Fundamental Research Funds for the Central Universitiesmembrane reactor cold model. Chem Eng Sci 2009;64:2599e606.[2] Grace JR, Elnashaie SSEH, Lim CJ. Hydrogen production influidized beds with in-situ membranes. Int J Chem ReactorEng 2005;3:A41.

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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 en e r g y 3 5 ( 2 0 1 0 ) 1 1 7 9 8e1 1 8 0 9 11809

Reaction/separation coupled equilibrium modeling of steam methane reforming in fluidized bed membrane reactorsIntroductionModel developmentPrimary assumptionsGoverning equationsSolution of the governing equationsComparisons with PFR model of Ye et al. [15]

Comparison of model predictions with experimental dataEffects of operation conditions on carbon formationInfluence of steam to carbon ratio and membrane permeation capacityInfluence of temperature and membrane permeation capacityInfluence of reactor pressure and membrane permeation capacityInfluence of oxygen to carbon ratio and membrane permeation capacity

Effects of operation conditions on NH3 formationInfluence of temperature and membrane permeation capacityInfluence of pressure and membrane permeation capacityInfluence of steam to carbon ratioInfluence of oxygen to carbon ratio

ConclusionsAcknowledgementNotationReferences

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