a 2d model for the cylinder methane steam reformer using

WHEC 16 / 13-16 June 2006 – Lyon France

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A 2D model for the cylinder methane steam reformer using electrically heated alumite

catalyst Qi ZHANG1, Ahmad IQBAL1, M. SAKURAI1, T. Kitajima1, H. TAKAHASHI2, M.NAKAYA3,

T.OOTANI3, H. KAMEYAMA1 1.Department of Chemical Engineering, Faculty of Engineering,

Tokyo University of Agriculture and Technology, Tokyo 184-0012, Japan 2.Ishikawajima-shibaura Machinery CO., LTD, Nagano 390-8717 Japan

3.Yokogawa Electrical Corporation, Tokyo 180-1750 Japan

ABSTRACT: A cylinder methane steam reformer applied for domestic fuel cell by using Electrically Heated Alumite Catalyst (EHAC) was developed. Because of the novel advantages of EHAC, i.e., high thermal conductivity, self-electrified character and special shape flexibility, the cylinder constructed methane reformer was expected to have a higher start-up time and better thermal performance than conversional fixed-bed reactors. For Ni/EHAC, an L-H kinetic model for methane steam reforming was proposed by using differential reactor. Adopting this kinetic model, a transient triple-phase lag model was established to investigate the heat and mass transfer phenomena inside. The simulation results compared favourably with experimental data of a cylinder constructed reactor. The results indicated that heating the reactor by electrifying the EHAC rather than a heater outside, the start-up time could be shortened from 25min into 5min. And the transverse temperature difference of the EHAC was no more than 0.2K which showed an excellent thermal performance of the structured catalyst. Meanwhile the parametric study of catalyst channel, velocity, reactor length and etc. was also investigated in this work.

KEYWORDS : Methane Steam Reforming, Transient Simulation, Electrically Heated Alumite Catalyst, Plate reactor

Introduction

Methane steam reforming is regarded as a major route for the industrial production of H2 rich gas for fuel cell system, as the natural gas containing 80-95% of methane. Many reviews on steam reforming of hydrogen discuss the conventional process, which is carried on supported Ni catalysts in multitubular reactors operated at the temperature and pressure respectively ranging from 773K(inlet) to 1073K(outlet) and 20 to 40bar and S/C in the feed 2-4.(Van Hook,1980; M.Zanfir,2003) However, this process suffers from several drawbacks, such as, poor heat transfer coefficient of the catalyst bed, diffusion limitations, great pressure drop and large volume size.

The public interest in energy issues increasing recently requires further efficiency in energy utilization. What’s more, downsizing of system followed by cost cut-down is also demanded. To meet such requirements, on a different viewpoint from the conventional fix-bed reactor, constructed wall type reactors are widely under development. A large number of studies have been conducted on the wall type reactor, since Katz (1959) took up the problems of gas diffusion and the reaction in a tube wall reactor. (Kameyama, 1995; el’nov et al. 2003;Zanfir and Gavriilidis, 2003; Fukuhara et al. 2004).

For the preparation of plate catalyst, though various methods were reported, such as sol-gel

coating, nano-particle-based coating, and chemical vapor deposition, thermal endurance of the catalysts were seldom mentioned. Basically at high temperatures, it is difficult to avoid coating layers shelling from the base material due to the difference of their thermal expansivity. To avoid such problem, a novel plate catalyst with high thermal endurance and conductivity by anodization technology was invented by Kameyama, 1996. A commercial Al plate was anodized to form alumina films on its two outer surfaces, and γ-alumina films with high surface areas can be obtained by transformation treatment. Since the alumina film is derived from the surface of the base material, a close uniformity between the alumina layer and its inside can effectively prevent the mismatch in thermal expansion. This kind of catalysts was used in various reaction systems, such as, hydrocarbon combustion (Kameyama et al.1995), de-NOx reaction (Guo et al. 2003) and VOC combustion (Wang et al.2005) and etc. However, the Al2O3/Al catalyst is limited to be used in high temperature reactions (>850K) because the base metal Al has a low melting point (<930K).

To solve the problem above, a commercial Al/CrAlloy/Al clad plate was applied in this work. By anodization, Al2O3/CrAlloy/Al2O3 could be obtained. The thermal endurance test at 1073K showed it has a high heat resistance and no Al2O3 layer was found to shell from the alloy layer after 40000 times repeat tests. (Zhang et al., 2004) On the other hand,


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this novel plate catalyst has good flexibility, which makes it possible to be made into serried or whirled shape for various catalytic reactors. Another interesting point is its electrically heating character. Because of the existence of alloy layer, the catalyst can be directly electrified, as shortening the start-up time from several hours to few minutes.

To design the plate wall reactor, many computational investigations on mass and heat transference were discussed by using one-dimensional or two-dimensional model in recent years. ( Tronconi and Groppi, 2001;Zanfir, 2003; Bel’nov et al.,2003) Since the numerical solution of the Navier–Stokes equations coupled with chemistry models is formidably complex and time-consuming because the chemistry contributes to the stiffness of the equation system, in most of those studies plug flow was assumed to simplify the model. However, few papers validate the plug flow model compared with the experimental data.

In this work, a cylinder reactor for methane steaming reforming system with the novel Ni-Al2O3/Aolly catalyst was proposed. Based on the specific structure of the catalyst, a two-dimension coaxial model was discussed to investigate the mass and heat transfer phenomena. The simulation results show a good agreement with the experiment results. Transient model compare the start-up time of two different heating conditions, i.e., electrically heating through the catalyst and heating through the heater around the reactor. The results indict that heating through the catalyst should be a sufficient strategy to shorten the start-up time. On the basis of the model, some knowledge was collected on the effects of catalyst thermal conductivity, channel height, reaction temperature and the length of catalyst bed. 1. Experimental 1.1 Catalyst preparation

A commercial Al/CrAlloy/Al clad plate (aluminum clad), was degreased in aqueous solutions of sodium hydroxide and nitric acid, and then rinsed with deionized water. Both sides of the clad plate were anodized in 4 wt% oxalic acid solution for 5 hours with an electric current density of 60A·m-2 and 303K to transform the

Fig.1 Cross section of Al2O3/alloy plate.

(1) alloy layer (2)anodized Al2O3 layer

Calcination

Aluminum clad plate

Pretreatment

Anodization

Hydration treatment

Impregnation

20% H2 reduction

Ni/Al2O3/alloy catalyst

Pore widening treatment

Calcination

NaOH 3 min, HNO3 1 min

298K 2h

773K 3h in air 923K 6h

4 wt% oxalic acid solution60A∙m-2 303K 5h

4 wt% oxalic acid solution 303K 2h

353K 1h 623K 1h in air Calcination

Aluminum clad plate

Pretreatment

Anodization

Hydration treatment

Impregnation

20% H2 reduction

Ni/Al2O3/alloy catalyst

Pore widening treatment

Calcination

NaOH 3 min, HNO3 1 min

298K 2h

773K 3h in air 923K 6h

4 wt% oxalic acid solution60A∙m-2 303K 5h

4 wt% oxalic acid solution 303K 2h

353K 1h 623K 1h in air

Fig.2 Flowchart of catalyst preparation aluminum to alumina on the surfaces. Subsequently, a pore widening treatment was carried out in the oxalic acid solution for 2 hours and then the plate was rinsed with deionized water, naturally dried and calcined at 623 K in the air. Thirdly, a hydration treatment at 353 K was performed in deionized water for 2 hours to adjust the pore size distribution. Finally, γ-alumina films were obtained by calcination at 773K for 3 hours (Kameyama, 1995). Both the γ-alumina films and the alloy layer inside were used as a whole support. The cross section of Al2O3/alloy was shown in Figure 1. The thickness of Al2O3 layer is 40µm on each side and alloy layer 80µm in the middle.

Ni was selected as the active metal for the methane steam reforming reactions. The catalyst support was impregnated in the required amounts of (CH3COO)2Ni·4H2O aqueous solutions for 2 hours at 298 K. The resulting plate was naturally dried for 4 hours and then calcinated at 773 K for 3 hours. The catalyst was used in the coaxial cylinder reactor. A 6-hour 20% H2 reduction would be carried out at 923 K before the catalyst was used. The flowchart of the catalyst preparation is presented in Figure 2. 1.2 Reactor and activity test

As a basic research and to validate the model, a cylinder reactor using Ni-Al2O3/alloy catalyst was developed (Figure 3(a)). The catalyst is whirled to form 6 channels with 1mm height in the reactor as shown in Figure 3(b). Figure 3(c) shows a schematic drawing of the reactor structure. Catalyst bed is set at the middle of the heater, with 1cm height pre-heating part, and to avoid the effect of the outlet surroundings, heater is 1cm longer than the catalyst bed at the end. Thermal insulation material is set in the center of the reactor where 3 J-type thermocouples are set at 0.015, 0.03, 0.045 m along the axial direction. Heater’s temperature is controlled at 1073 K. Outlet gas concentrations are detected by FID and TCD of GC-14B (Shimadzu Co. Ltd). The detailed characters of the reactor are shown in Table 1, which is also taken as the basic case for the simulation. 2. Modeling

1 2


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2.1 Kinetics Since the reactor simulations concern the reaction items, rate equations for steam reforming and water-gas shift reactions have to be combined in the calculations. The main chemical reactions involved in the process are: CH4+H2O→CO+3H2, ∆H0

298 = 206.2 kJ/mol CO+H2O→CO2+H2, ∆H0

298 = -41.2 kJ/mol The reaction kinetics for methane steam reforming was derived from our experimental results. The kinetic characteristics for EHAC were obtained using a differential reactor, avoiding concentration and temperature gradients by adopting the kinetic mechanism cited by Numaguchi and Kikuchi (1988). Assuming both methane and steam were absorbed on the catalyst with dissociation and the surface reactions producing CO and CO2 were the rate controlling steps, the rate equations were reported as follows:

12

14

224)/( 1

31

1 nOH

mCH

HCOOHCH

ppKppppk

w−

= (1)

22

24

222)/( 22

2 nOH

mCH

HCOOHCO

ppKppppk

w−

= (2)

where, m1=0.6, m2=0.55, n1=0, n2=0. The equilibrium constant, Km (ｍ=1, 2) was estimated from thermodynamic data. The corresponding kinetic rate parameters are summarized in Table 2.

(a) (b)

(c)

Fig. 3 the scheme of the reactor( for a clarity, the scheme is not drawn in real scale) (1)catalyst bed (2) pre-heating part (3)heater (4) thermal insulation material (5) thermocouples

2.2Assumption and descriptions The cross-section scheme of the catalyst bed

is shown in Figure 4(a). The boundary conditions, however, is very complex if considering such structure for a two-dimensional model. Since the screwed circle number is many enough, they are regarded as coaxial circles approximately in this model (Figure 4(b)). To diminish the difference between the catalyst area in the model and the real one, a shape factor, ε, defined as modelcat / AA=ε was applied. Figure 4(c) shows the scheme of 3 domains, i.e., gas phase, catalyst layers and alloy layer, which is called as a triple-phase lag model. The mass transfer, heat transfer and momentum transport were considered in the three domains respectively. To describe the model, several assumptions are made as follows: ① The whirled catalyst layers are regarded as

coaxial structure, and the catalyst bed is consist of 6 concentric layers;

② The whole system is considered as a steady state for reactor operation

③ The gases are assumed as ideal gasses and suitable for ideal gas law;

④ The chemical reactions would take place only on the catalyst layer

⑤ Since in the experimental result, no more than10Pa pressure drop is found along the reactor, in the simulation, the pressure drop is neglected in all runs.

The mathematic partial differential equations are given in Table3, where catalyst surfaces, alloy surfaces are regarded as the interior boundaries in this model. 2.3 Numerical simulation

The PDEs (partial differential equations) as shown in Table 3 were solved by using FEMLAB, a PDE solver tool from COMSOL Inc., which uses the finite element method. We used quadratic triangular elements and a non-uniform mesh, with small elements located at the gas-solid interface and solid phase. Since the Navier–Stokes model can take as long as a few hours to converge to a solution, depending on how the initial value begins. In this work, firstly the time dependent solver was used to get a friendly initial value and then switched to nonlinear solver to solve the problems. The results indicate that it took just no more than 10mins to reach a satisfying converge. Table 1 the characters of the coaxial reactor and

base case of the simulation Gas phase

H2O/CH4 3

Pressure 1.01 MPa Inlet temperature 473 K Heater

temperature 1073 K

Flow rate 0.02 mol/min~0.3mol /minCatalyst

Metal content Density

Ni/Al2O3/alloy 3 gNi/m2

cat (9 wt% ) Al2O3 :1970 kg/m3

Alloy: 3900 kg/ m3

1ｍｍ

1

Gas

1


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Thickness

Thermal- Conductivity

Al2O3 layer:40 µm Alloy layer 80 µm Al2O3 layer: 1.039 W·m-

1·K-1 Alloy layer:11.3 W·m-1·K-1

Reactor Total Length Catalyst bed

height Preheating part Radius Channel height

0.06 m 0.04 m 0.01 m 0.015 m 0.001 m

Table 2 Arrhenius parameter values for reforming

and water-gas shift reactions (Numaguchi & Kikuchi,1988)

Constant A E[kJ mol-1]

k1 [molPa-0.404kgcat

-1s-1] 1.14E9 163

k2 [molPa-1kgcat

-1s-1] 18.7 66

K1[Pa-2] exp(-22430/T+16.068) K2[-] exp(4400/T-4.036)

(a) (b)

(c) Fig. 4 Schematic diagram of the coaxial triple-phase lag model 3. Results and discussions 3.1 Verification of the heater heating model

The validity of the simulation in a steady-sate was firstly confirmed. For the plate catalyst, F/S is applied to describe the space velocity, where F is flow rate and S is the total surface area of the catalyst. Figure 5 shows a comparison of model results with experimental measurements. The calculations of are in well agreement with the experimental data for outlet methane fraction, H2 fraction and methane conversion, which demonstrates that the reaction kinetic model proposed by this work. Figure 5 (c) describes the temperature distribution at z=0.015, 0.03, 0.045m in

Table 3 Model equations Model equations Gas phase Material balance

0)1()( 2

2

=∂∂

+∂∂

−∂∂

+∂∂

⋅+∂∂

ry

ry

rD

ry

zy

uty ii

iriii

Heat balance

0)1( 2

2

=∂∂

+∂∂

−∂∂

+∂∂

rT

rT

rzTuCp

tTCp

gg

er

g

g

g

g λρρ

Navier-stock equations

)1())(( 2

2

2

2

ru

rru

zug

ru

zuu

tu

∂∂

+∂∂

+∂∂

+=∂∂

+∂∂

⋅+∂∂ υρρρ

Catalyst layer no slip conditions, ui,j=0 Material balance

0)()1(2

1,2

2

=−∂∂

+∂∂

+∂∂ ∑

=jgcatjji

iiieff

i Mwvry

ry

rD

ty

ερρ，

Heat balance

0)()1(2

12

2

2

2

=−∂∂

+∂∂

+∂∂

+∂

∂ ∑=j

catjj

catcatcat

cat

alloy

wHzT

rT

rT

rtTCp ερλρ

Alloy layer For heater heating model Lap lace equations

0)()](1[ =∂

∂∂∂

+∂∂

∂∂

+∂

∂z

Tz

rTrrrt

T alloyalloy

alloy

For electrically heating through the catalyst model With electrically heat-up through the catalyst

ele

alloyalloycatj

alloy

qr

rTrrl

Tlt

TCp =∂

∂∂∂

+∂

∂∂∂

+∂

∂ )])(1()([λρ

No mass transfer 0, =∀∀ iyzr

Interior boundaries surfacecatalyst, =∀ rz

ryD

ryDr

cati

ieffcat

gi

ig ∂∂

−=∂∂

,ρρ

)(2

1∑=

+∂

∂−=

∂∂

jcat

catjj

cat

cat

g

er wHrT

rT ρδλλ

no slip 0=u alloyface, =∀ rz

0=∂

∂=

∂∂

ry

ry alloy

icati

rT

rT cat

cat

g

er ∂∂

−=∂∂ λλ

Boundaries Inlet conditions r∀ 0,ii yy = ｔ＝0

0TT = ),0(),( 0,zzr uuuu == Right boundary (near the center) z∀ ,t

Mass and heat insulation 0=

∂∂

rT alloy

0=∂∂ry i 0=u

Left boundary (near the heater) For heater heating model

Gas phase

Alloy layerInner Catalyst layer

Outer catalyst layer


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z∀ ,t

)( TTarT

ww

g

er −⋅=∂∂

− λ 　0=∂∂ry i 0=u

For electrically heating through the catalyst model

0=∂

∂rT g

0=∂∂ry i 0=u

0

0.2

0.4

0.6

0.8

0 2 4 6 8 10F/S[mol/min/m2]

Mol

e Fr

actio

n [-

]

0

0.2

0.4

0.6

0.8

1

0 2 4 6 8 10F/S [mol/min/m2]

Met

hane

con

vers

ion

[-]

900

920

940

960

980

1000

0 2 4 6 8 10F/S[mol/min/m2]

Tem

pera

ture

in th

ere

acto

r[K

]

Fig. 5 Validity of the simulation algorithm and

kinetic rate used. Experimental data: Outlet mole fraction :(▲) methane; (◇)H2. (●) outlet methane conversion. Temperature :(○)at z=0.015m;(△)at z=0.03m; (□)at z=0.045m; ( )Calculated line varied by F/S.

the center of the reactor compared with the data measured by the thermocouples. It can be found that at lowest flow rate, the model under-predicts the temperature, which can be explained that axial heat transfer of two layer catalyst was just depended on the gas phase in the model, and the heat transfer in the whole catalyst was neglected. The actual reactor should have more effective heat transfer than the model especially when the convective heat transfer is not very distinctive. When the flow rate increased, it can be seen that the effect of the catalyst’s self heat transfer becomes smaller. However, at higher flow rates, the calculated value over-predicts the temperatures, likely due to an under-assessment of heat loss in

the model. In all cases, the differences between the calculated temperature and experimental values are no more than 20K which does not affect the reaction performance distinctly as shown in the Figure 5 (a) and (b). By the comparisons above, the coaxial model and

kinetics of Numaguchi is proved to be reasonable to simulate the whirled catalyst structure.

3.2 Thermal behavior of the reactor The average transverse temperature difference

of the catalyst, ∆Tcat, along the catalyst bed is shown in Figure 6(a), which is defined as

6

)T(T∆T

6

1mm

catsideinner

catsideoutlet

cat∑=

−= (6)

The high thermal conductivity of the clad catalyst makes it possible to have an efficient heat transfer at ∆Tcat is less than 0.3K. Further, Figure 6(b) shows the reactor gas phase temperature difference varied with different flow rate. ∆Tg defined as the difference between the average temperature of gas phase in the first channel (nearest the heater) and 6th channel (nearest to the center),

channel sixth thechannel1st theTT∆T g −= (7) ∆Tg is

found to be no more than 40K. In contrast to 250K difference in temperature of outside wall and mean gas phase temperature inside the tube usually observed in conversional reformer (Zanfir, 2003),

0

0.050.1

0.15

0.2

0.25

0.3

0 0.01 0.02 0.03 0.04Catalyst bed height [m]

Tra

nsve

rse

tem

pera

ture

diffe

renc

e of

the

cata

lyst

,Δ

Tca

t [K]

(a)

0

10

20

30

40

0 0.02 0.04 0.06Reactor length [m]

Tra

nsve

rse

tem

pera

ture

diffe

renc

e of

rea

ctor

,∆T

g [K]

(b) Fig.6 Transverse temperature difference along the

reactor or catalyst bed length (a) ∆Tcat ; (b)

∆Tg varied with different flow

rate:( )0.02mol/min ( )0.2mol/min ( )0..3mol/min

(a)

(b)

(c)


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the plate type reactor shows an excellent heat transfer performance. Meanwhile, it can be seen that ∆Tg is increasing in the pre-heating part, reaching the maximum at the inlet of the catalyst bed. However, ∆Tg

shows a declining profile along the catalyst bed. This can be explained by the fact that the conductivity of gas in the pre-heating part is lower than that of gas coupled with catalyst walls.

550

650

750

850

950

0 5 10 15 20 25 30

Time [min]

Out

let T

empe

ratu

re [K

]

(a)

0

0.05

0.1

0.15

0.2

0.25

0.3

0 5 10 15 20 25 30Time (min)

Out

let C

H4

mol

e fr

actio

n

(b)

Fig.7 the transient profile by two different heating methods with flow rate 0.1mol/min 3.3 Start-up performance of the reformer

The start-up behavior of a methane steam reformer is a crucial issue for the design of the domestic fuel cell. However, start-up times of the reformers are reported to be more than 1 hour. In this contribution, start-up acceleration strategies for the reformer are discussed for a given apparatus design and mass an acceleration of the start-up, it is only possible to increase the heat input into the system. In this work, we discussed two different heating methods, i.e., heating by the heater and electrically heating through the catalyst. Figure7 shows the comparison of the start-up performance in case of the two methods. It can be seen that the start-up time is no more than 5 mins by electrically heating through the catalyst while it takes more than 25 mins by heater heating. What is more, from Figure 11 (b), it is found that the outlet methane mole fraction shows a shower value by electrically heating through the catalyst which suggests that direct heat-up to the catalyst makes a better reaction performance than the heat-up from outside, which can be explained by the different temperature distribution in the reactor. As Figure 8(a) shows that the temperatures are greatly higher at the locations near the heater while the temperature

shows rather even distributions at different radius locations when by electrical heat-up through EHAC (Figure 8 (b)). Since it is obvious that even temperature distribution can improve the reaction performance, the inside heat-up is preferred in the catalytic reactor.

0 0.15 0.3 0.45 0.6 0.75 1

Dimensionless reactor radius Fig.8 Temperature distribution in the reactor bed by

different heat-up methods (a) heat-up by heater (b) electrical heat-up through the catalyst

3.3 Reaction performance improvement Channel height is an important parameter to affect the performance of the reactor. Figure 9 shows the profile of methane conversion as a function of channel height by keeping the product F/S and catalyst bed length 0.04 m. Methane conversion shows a decreasing profile with channel height increased. To explain this result, mass transfer effect is discussed firstly. Herein, Fourier

number, described asd

sFoττ

= , was adopted to

investigate the mass transfer phenomena, where,

sτ is the local space time:

)centerchannel,(zvzL

s−

=τ , (8)

and dτ is defined as the time necessary for the reactant module of center to reach the catalyst wall by diffusion.

)centerchannel,(4

channel

zDA

CHd =τ (9)

where, Achannel means cross sectional area of the channel (Cussler, 1997; Zanfir,2003). It is obvious that when Fo>1, the reactant molecules from the channel center would have enough time to reach the wall before they exit the reactor. Otherwise, the

850 900

910

920

940

950

960970

Catalyst

4 3 2 1 0

900

920

940

950

970 960

(a)

(b)


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molecules would not have enough chance to have a reaction before they leave the reactor. Comparing Fo along the catalyst bed for channel height 1mm, 3mm and 5mm, the results are shown in Figure 10.

00.20.40.60.8

1

1 3 5 7 9Channel height [mm]

Met

hane

con

vers

ion

[-]

Fig.9 Outlet methane conversion as a function of

channel height at consistent F/S

-3

0

3

0.00 0.01 0.02 0.03 0.04Length of catalyst bed [m]

ln(F

0) [-

]

Fig.10 logarithm of the Fourier number as a

function of catalyst bed for different channel heights at constant F/S. ( )channel height 1mm; ( ) channel height 3mm; ( )channel height 5mm

The smaller channel height is, the shortest the catalyst bed portion where ln(Fo)<0 can be seen. Furthermore, when channel height reaches 5mm, all the ln(Fo) is found to be less than 0 along the catalyst bed which demonstrates a bad transverse mass transfer causing a lower methane conversion The heat transfer characteristics were also studied to identify the reasons for the decreased performance as a result of the increasing the channel height. Figure 11 presents the overall heat transfer coefficient calculated from the following equation

avreactionoutreactor,,inreactor,, )()( TUAqTCTCG outpinp ∆=−− (10) where, qreaction means heat assumed by the reactions, which is calculated as :

24 2211 COjjCHjjreaction FwHxFwHq ==== += (11) and av)( T∆ is calculated by logarithmic mean temperature difference:

)/ln(/)()( 2121av TTTTT ∆∆∆−∆=∆ (12) where, ∆T1 means the temperature difference between gas and heater at the inlet of the reactor and ∆T2 means the temperature difference between gas and heater at the outlet of the reactor.

0

4

8

12

16

20

1 3 5 7 9Channel height [mm]

Ove

rall

coef

ficic

ient

of

heat

tran

sfer

, U [W

/m2 /K

]

Fig.11 Calculated overall coefficient of heat transfer

of the coaxial reactor with different cannel heights at consistent F/S.

0

0.2

0.4

0.6

0.8

1

0.5 1 1.5 2 2.5 3 3.5Preheater length [cm]

Met

hane

con

vers

ion

[-]

Fig.12 effects of pre-heater length and catalyst bed

length at different flow rate. (△): F=0.02mol/min (○): F=0.1mol/min

The figure indicates that the heat transfer coefficient, U, is decreasing with the channel height increased. It is also noted that when the channel height reaches 10mm, the coefficient becomes 1/2 of that of 1mm channel height. It can be considered as another reason for the lower methane conversion with larger channel height. From investigations above, it can be found out that smaller channel height can achieve a good efficient heat and mass transfer in the reactor.1mm should be the optimized channel height in this case.

On the other hand, it should be noted that even the channel height reaches 1 mm, the methane conversion is not very satisfying. Two aspects were herein taken into consideration, i.e., the reaction temperature and amount of the catalyst. Figure 12 discusses the effects of pre-heater and catalyst bed length with flow rate at 0.02 mol/min and 0.1 mol/min to investigate the influence of reaction temperature and catalyst amount From the results , it can be seen that when the pre-heater height increased from 1cm to 2 cm , methane conversion has a significantly improvement. However, when the catalyst bed length Lcat, is 4 cm, the conversions have a little change by increasing the pre-heater length, and just 70% conversion can be obtained with flow rate at 0.1 mol/min. while when the catalyst length is doubled and pre-heater length is set as 3 cm, methane conversion can be achieved near 95% even if the low rate changed from 0.02

Lcat=4 cm

Lcat=8 cm


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mol/min to 0.1 mol/min. Meanwhile, effect of inlet temperature Tin was also investigated. When Tin is increased from 473 K to 675 K in condition of Lpreheater and Lcat remain 1 cm and 4 cm , methane conversion just has increased by 3~5% for all flow rates. These results indicate that increasing the pre-heating length and catalyst bed should be effective ways to improve the reaction performance for the reactor. Conclusion

A transient two-dimensional coaxial model is developed for the methane reforming reactor by using Ni-Al2O3/alloy clad catalyst prepared by anodization technology. The simulation results show a good agreement with the experimental data, which proves it reasonable to use coaxial model to simply the whirled catalyst. Furthermore, the following knowledge was collected on the basis of simulation results. 1) In the case of the heater heating, the transverse catalyst temperature, ∆Tcat , is less than 0.3K, and ∆Tg is less than 40K, which shows a high thermal conductivity of the clad catalyst 2) The transient results indict that, electrically heating through the catalyst, a novel direct heat-up to the catalytic reactor system, shorten the start-up time dramatically and contribute a even temperature distribution in the reactor. 3) For the coaxial reactor, smaller channel height is, better efficient heat and mass transfer can be achieved in the reactor. 4) Adjusting the pre-heating height and catalyst bed length are also efficient ways to increase the reactor’s performance.

Nomenclature A =cross section area [m2]

wa =thermal conductivity of the wall [W·m-1·K-1]

Cp = heat capacity [J·Kg-1· K-1]

effD = effective diffusion coefficient [m2·s-1]

Dr =gas diffusion coefficient [m2·s-1] F = flow rate [mol/min] Fo =Fourier number [-] G =mass velocity [Kg·s-1] H = heat of reaction [kJ·mol-1] L =reactor length [m] M =average molecular weight [kg·mol-1] P =Pressure [Pa]

r =radial coordinate in the reactor [m]

R =gas constant [J·mol-1·K-1] Re =Reynolds number [-] R.E. =relative error of the model [-] T =Temperature [K] Tw =temperature of the heater [K] u =velocity [m·s-1]

U =overall heat transfer coefficient [W·m-2·K-1]

v =stoichiometric coefficient [-]

w =reaction rate [mol·kgcat-1·s-

1]

x =methane conversion [-] y =mole fraction [-]

z =axial coordinate in the reactor [m]

<Greek letter> erλ =gas thermal conductivity [W·m-1·K-1] ρ =density, [kg·m-3] sτ =local space time [s] dτ =diffusion space time [s]

catδ =thickness of catalyst layer [m] υ =kinematical viscosity [kg·m-1s-1] ε =shape factor [-] <Subscript and superscript> alloy =alloy layer cat =catalyst or catalyst bed gas, g =gas phase

i =component of the reactions j =reaction system(j=1, steam reforming ;

j=2, water gas shift) Appendix Heat capacity: bTaCp += +cT2

Thermal conductivity: 75.0

00 )(TT

er λλ =

Diffusion coefficient: 75.0)0

(0,, TT

irDirD =

The Dr,i is calculated for a binary mixture between i component species and H2(for reforming ) and N2(for combustion). The effective diffusion number:

11197,

−

+−⋅=

riD)iMTRp(ieffD

τε

Idea gas law )

,

,(∑=jiMjiy

RTP

jρ

Reference Cussler, E. L.; Diffusion: Mass Transfer in Fluid

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a 2d model for the cylinder methane steam reformer using

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