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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 4

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3D simulation of hydrogen production by ammoniadecomposition in a catalytic membrane reactor

Andrea Di Carlo a, Alessandro Dell’Era b, Zaccaria Del Prete c,*aDepartment of Chemistry, Chemical Engineering and Materials, University of L’Aquila, 67100 Campo di Pile (AQ), ItalybDepartment of Mechanics and Energetics, University "Guglielmo Marconi”, 00193 Rome, ItalycDepartment of Mechanical and Aerospace Engineering, SAPIENZA University of Rome, via Eudossiana 18, 00184 Rome, Italy

a r t i c l e i n f o

Article history:

Received 14 January 2011

Received in revised form

3 June 2011

Accepted 6 June 2011

Available online 16 July 2011

Keywords:

Hydrogen storage

Ammonia cracking

Membrane reactor

CFD model

* Corresponding author. Tel.: þ39 06 4458555E-mail address: zaccaria.delprete@unirom

0360-3199/$ e see front matter Copyright ªdoi:10.1016/j.ijhydene.2011.06.029

a b s t r a c t

Ammonia decomposition in an integrated Catalytic Membrane Reactor for hydrogen

production was studied by numerical simulation. The process is based on anhydrous NH3

thermal dissociation inside a small size reactor (30 cm3), filled by a Ni/Al2O3 catalyst. The

reaction is promoted by the presence of seven Pd coated tubular membranes about 203 mm

long, with an outer diameter of 1.98 mm, which shift the NH3 decomposition towards the

products by removing hydrogen from the reaction area. The system fluid-dynamics was

implemented into a 2D and 3D geometrical model. Ammonia cracking reaction over the Ni/

Al2O3 catalyst was simulated using the TemkinePyzhev equation.

Introductory 2D simulations were first carried out for a hypothetic system without

membranes. Because of reactor axial symmetry, different operative pressures, tempera-

tures and input flows were evaluated. These introductory results showed an excellent

ammonia conversion at 550 �C and 0.2 MPa for an input flow of 1.1 mg/s, with a residual

NH3 of only a few ppm. 3D simulations were then carried out for the system with

membranes. Hydrogen adsorption throughout the membranes has been modeled using the

Sievert’s law for the dissociative hydrogen flux. Several runs have been carried out at 1 MPa

changing the temperature between 500 �C and 600 �C to point out the conditions for which

the permeated hydrogen flux is the highest. With temperatures higher than 550 �C we

obtained an almost complete ammonia conversion already before the membrane area. The

working temperature of 550 �C resulted to be the most suitable for the reactor geometry.

A good matching between membrane permeation and ammonia decomposition was

obtained for an NH3 input flow rate of 2.8 mg/s. Ammonia reaction shift due to the pres-

ence of H2 permeable membranes in the reactor significantly fostered the dissociation: for

the 550 �C case we obtained a conversion rate improvement of almost 18%.

Copyright ª 2011, Hydrogen Energy Publications, LLC. Published by Elsevier Ltd. All rights

reserved.

1. Introduction together with high safety levels, have to be overcome in order

Hydrogen production and storage are still themain challenges

for the realization of a sustainable hydrogen economy. The

difficulties in reaching a satisfactory fuel gravimetric density,

9; fax: þ39 06 4881759.a1.it (Z. Del Prete).2011, Hydrogen Energy P

to achieve extensive hydrogen utilization. H2 gas-phase

storage presents well-known problems related to the

unavoidable need of high pressure vessel and yet results in

low volumetric density (30 kg/m3 at 70 MPa), whereas solid

ublications, LLC. Published by Elsevier Ltd. All rights reserved.

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 411816

storage inside metal and nanofabricated hydrides still suffers

the lack of a material matching all suitable requirements of

energetic density, temperature/pressure work condition and

adsorption/desorption kinetic. Liquid cryogenic H2 has

a better volumetric density (70 kg/m3 at - 253 �C) however,

around 30% of its heating value is required for liquefaction

and, to minimize hydrogen boil-off, it needs heavily insulated

and expensive tanks [1].

A different storagemethod, which is recently gainingmore

and more interest, is the storage in non-cryogenic hydrogen-

rich liquids, which involves the production on-demand of

hydrogen by thermal cracking or chemical decomposition.

One of the first liquid which seemed suitable for hydrogen

storing purposes has been methanol; having a good lower

heating value (19.9 MJ/kg) and a hydrogen volumetric content

of 87 kg/m3 (12.5% on a mass basis). But it has also several

drawbacks, as the need to conduct the reaction with water

(reducing the effective energy density carried by methanol)

and the presence of carbon-oxides in the final products [2].

Taking into consideration features as reforming temperature

and efficiency, ammonia appears today a more interesting

hydrogen carrying fuel than methanol for many applications

[3]. In the last century, it has been widely used as fertilizer,

cleaning agent, and explosive. With the growing need of

a densehydrogencarrying fuel, anhydrousammonia (NH3) has

received more and more attention as a candidate source for

atomic hydrogen storage, offering for that purpose optimal

physical properties. In fact, liquid anhydrous NH3 has

a hydrogen volumetric content of 120 kg/m3 (17.6% on a mass

basis) i.e. 1.7 times the amount of cryogenic liquidhydrogenon

a volume basis and an energy density of 17.8 MJ/kg. It does not

contain carbon atoms that cause the presence of pollutant

carbon-oxides between the decomposition products, and it

does not need additional water to keep the decomposition

reaction going. On the bad side, ammonia is a toxic compound

and, even if it can be detected by the human nose at very low

concentration (10e20 ppm), it can be lethal if inhaled at high

concentration (higher than 2000 ppm). Nonetheless, it is

produced and distributed world-wide in millions of tons per

year, so procedures for safe handling have beendeveloped and

are well established in every country. No significant ammonia

supply problems are known with regard to geographic distri-

bution, nor with regard to storage and transport, which can be

done in liquid phase at ambient temperature and low pressure

(<2 MPa) or at atmospheric pressure and moderate low

temperature (�30 �C) [4] [5]. Despite ammonia synthesis is an

energy demanding process, its production cost is yet lower

than 1 $/kg, while hydrogen production cost varies from 3.5 to

5.5 $/kg (depending on delivery phase: gaseous or liquid).

Ammonia decomposition process has been studied a long

time for many reasons, e.g. for removal of toxic ammonia

traces from coal gasification streams, or as reaction associated

with the process of metal nitriding [6e8]. Ammonia decom-

position can be regarded as the reverse process of the

synthesis reaction, occurring in industry at approximately

500 �C and 250 atm (DH ¼ �92.4 kJ/mol). Thermal cracking of

ammonia into hydrogen and nitrogen proceeds according to

the simple endothermic reaction:

2NH353H2 þN2 (1)

(DH¼ 66.5 kJ/mol) [9]. Ammonia cracking can be operated at

temperatures between 700 and 900 �C and unlike ammonia

synthesis, low pressure is preferred [5]. Usually, the use of an

appropriate catalyst can significantly reduce reaction

temperatures and can improve conversion rates. Theoretical

studies observed that for ammonia, the optimal synthesis

catalyst is not necessarily the optimal decomposition catalyst

[10]; nonetheless, synthesis catalysts are often used for the

decomposition process. Typical catalysts used both for

ammonia synthesis and cracking include iron oxide, molyb-

denum, ruthenium, and nickel [5]. Ru catalysts perform better

than Fe catalysts, particularly at low temperature and close to

the thermodynamic equilibrium, but are considerably more

expensive than Fe ones and have a shorter catalytic lifetime

[11]. More common and less expensive materials, that employ

for example Ni-alumina compounds in the catalysts, are

suggested bymany authors [12e14], claiming good conversion

results in a wide temperature and pressure range (e.g.

9e36 atm and 400e600 �C) [15]. Several studies about

ammonia decomposition over Ni-alumina catalyst were

carried out in the last decade, also because a significant

development in membrane science stimulated increasing

research efforts on catalytic membrane reactors [15e17].

Numerical studies that simulate the process in a catalytic

reactor over Ni-Al2O3, reported results as good as the

complete decomposition of ammonia already at a tempera-

ture of 400 �C and a pressure of 5 atm [17]. It has to be high-

lighted though, that many of the results quoted here

originated from purification processes in gasification plants

and not from studies aiming at hydrogen storage and

production with ammonia.

There are interesting examples in the literature which

focused on miniaturized reactors for portable devices with

a volume of only a few mm3 [18,19], or military applications

with a power range up to 50e60 Watt [4]; no catalytic

membrane reactors (CMR) was yet employed in these studies.

As the matter of fact, the utilization of membranes has the

double advantage of enhancing the ammonia decomposition

by immediately removing hydrogen from the reaction area,

shifting therefore the decomposition reaction towards the

products, and (especially with palladium membranes)

providing hydrogen with high purity, which is ideal when

feeding PEM fuel cells; e.g. to ensure a long durability, PEM fuel

cells require residual ammonia to be reduced below ppm

levels, since exposure of the acidic PEM electrolyte to

ammonia causes severe and irreversible loss of the fuel cell

performances [5]. Nowadays few companies can provide

commercial CMR. The REB Research� reactor simulated in our

study uses palladium-coated metal sandwich membranes,

which should insure the hydrogen permeating through the

membranes is pure at 99.9999%, regardless of any hydrogen

back-pressure change caused by a variable fuel cell load

demand [20]. This type of reactor has been realized to generate

high quality hydrogen by means of the simultaneous steam

reforming of a liquid fuel and the purification of the reformed

gas, obtained by permeating only hydrogen through the

selective membranes.

The work of Chein et al. [17] is the first example where a 2D

computational fluid dynamic (CFD) modeling was applied to

the catalytic decomposition of NH3 based on a simple

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 4 11817

chemical reaction model. A cylindrical reactor packed with

NiePt/Al2O3 catalyst particles was used as the physical

domain. This catalyst was chosen because the well-

established reaction model could be based on experimental

data. The NH3 decompositionwas predicted both theoretically

and numerically. Their results showed how adopting

a chemical reaction model similar to that used for methanol

reforming with a CuO/ZnO/Al2O3 catalyst, the numerical

model predicts quite satisfactory the experimental results.

Nonetheless, no examples employing CMR to reform

ammonia have been found in the literature. Therefore, the

aim of this work is to evaluate the behaviour of a micro CMR

for ammonia decomposition by using a CFD model. Kinetic

and mass transfer effects of the cracking reaction were

considered in the model. A 2D model was first employed to

evaluate the conversion efficiency of the micro-reactor

working without membranes. Then, a 3D simulation was

carried out to evaluate the gain in conversion efficiency

brought by the Pd/Ag membranes. Using a 3D simulation for

the CMR was necessary because of the internal asymmetry

introduced by the presence of the membranes inside the

reactor.

2. Reactor configuration

A schematic illustration of the membrane micro-reactor,

simulated with mesh points is shown in Fig. 1.

The reactor is basically a 25 cm long inconel tube with an

innerdiameterof12mmandathicknessof2mm.Seven tubular

membranesmade of palladium coating high permeability alloy

are placed at the bottom side of the reactor; each candle has

a total length of 15 cm and an outer diameter of 2 mm. There-

fore, the first 10 cm at the top of the reactor are not engaged by

the membranes. The catalyst particles are packed inside the

whole reactor. Preheated gaseous NH3 is fed into the reactor at

Fig. 1 e Schematic representa

a volumetric flow rate Q through the top inlet (see Fig. 1). This

reactor configuration has been designed to let the NH3 strongly

dissociate already along the first 10 cm, before running into the

Pd-membranes. Hydrogen is then selectively permeated

through the membranes due to the H2 pressure difference

among the outer and inner side of the Pd-membranes and sent

to thehydrogencircuit,while the retentategas, composedofN2,

un-permeated H2 and residual unconverted NH3, outgoes the

reactor from a different bottom outlet (see Fig. 1). The Ni/Al2O3

catalyst we considered is a commercial nickel-based catalyst,

ground to a mean particle diameter of 500 mm.

3. Model

The model we adopted is based on the fundamental concept

of interpenetrating continua for multiphase mixtures.

According to this theory, different phases can be present at

the same time in the same computational volume. Such an

idea is made possible by the introduction of a new dependent

variable, the void fraction 3i of each phase i. In this study the

phases we considered were the gas phase and the solid phase

(catalyst). The fundamental equations of mass, momentum,

chemical species and energy conservation are then solved for

the gas phase, while the solid phase is seen as a packed bed.

3.1. Dissociation model

To compute the kinetic expression of the catalyzed reactions,

several processes that act against the reactions, and are

therefore a resistance for it, have to be considered. For a single

porous particle they are:

C Gas film layer resistance: reagents and products diffuse

respectively from the gas towards the external surface

or from external surface towards the gas of the catalytic

particle.

tion of the micro-reactor.

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 411818

C Pores diffusion resistance: the inside of the particle

contains most of the catalytic surface, therefore the

reactions occur prevalently inside the particle. So the

reagents have to diffuse inside the particle, while the

products have to diffuse from the inside to the outer

surface of the catalyst.

C Superficial phenomena resistance: the reagents moving

inside the catalyst are adsorbed by the solid surfaces,

then they react following their kinetic mechanism and

the products are desorbed to the gas phase.

These three distinct phenomena are schematized in Fig. 2.

As shown in the figure, the gas film layer resistance could be

considered in serieswith the parallel resistance resulting from

pore diffusion resistance and superficial phenomena

resistance.

To take into account the different resistance mechanisms,

the chemical species evolving in the reaction have to be

considered in the gas phase (g) and in the solid-catalyst phase

(c). The following chemical species balance could be set:

DDt

�3grgymðgÞ

�¼ V$

�3gDmðgÞVymðgÞ

�þ rghmðgÞSc

�ymðcÞ � ymðgÞ

�(2)

where m ¼ NH3, H2, N2.

The terms yg and yc are referred to the mass fraction of

species m in the gas phase ( g) and in the solid-catalyst phase

(c) respectively, Dm( g) is the effective diffusion coefficient of

the species in bulk, calculated using the Wilke equation [21]:

DmðgÞ ¼ 1Pnsm

ynðgÞDmnðgÞ

(3)

The binary diffusion coefficient Dmn( g) has been calculated

using the Chapman-Enskog [22] equation:

DmnðgÞ ¼ 0:0018583

ffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiffiT3

�1

MmðgÞþ 1MnðgÞ

�s1

ps2mnUmn

(4)

where the expressions to calculate the required parameters

can be found in Bird [22]. Sc is the catalyst surface per unit

volume:

Sc ¼ 6dc

(5)

Fig. 2 e Section of a catalytic particle containing an ideal

pore.

The presence of the membrane will be discussed later.

Finally hm( g) is the mass transfer coefficient deduced using

RanzeMarshall correlation:

hmðgÞdcDmðgÞ

¼ 2þ 0:6ðScÞ1=3�Rep

�1=2(6)

A simpler model to consider the parallel effect brought by

the pore diffusion resistance and the superficial phenomena

resistance is based on the calculation of a so-called “effec-

tiveness factor”. To define the problem, the kinetic mecha-

nism for ammonia cracking reaction (eq. (1)) has to be set. For

this scope the formulation proposed by Dittmeyer et al. [16]

was used:

R ¼ k

"�p2NH3

p3H2

�b

�pN2

K2eq

�p3H2

p2NH3

�1�b#

(7)

where

k ¼ k0e�E=RT (8)

The necessary coefficient to define the kinetic Arrhenius k,

the equilibrium constant Keq can be find in [15]. The reaction

rate of the species in the solid phase rm(c) is defined using the

reactions stoichiometry:

rNH3ðcÞ ¼ �2hR

rH2ðcÞ ¼ 3hR

rN2ðcÞ ¼ hR

(9)

To close the reaction it was imposed that the amount of the

consumed/produced species in the solid phase were equal to

the species transferred from the gas phases to the solid phases

and vice versa:

rghgSc

�ymðcÞ � ymðgÞ

�¼ 3crcrmðcÞMM (10)

The momentum balance for the gas phase is given by the

NaviereStokes equation, modified to include an interphase

momentum transfer term:

v�rg3gng

�vt

þ V$�rg3gngng

�¼ V$

�sg�þ rg3gg� 3gVp� bgcng (11)

where g is referred to the gas phase, c to the catalyst solid

phases, p is the thermodynamic pressure, b the interface

momentum transfer coefficient, and sg the gas phase viscous

stress tensor. In particular b is deduced by Ergun equation:

bgc ¼ 15032cmg

3cd2c

þ 1:753crg

ngdc

(12)

The conservation of energy for the three phases is for gas-g,

catalyst-c:

vðri3iHiÞvt

þ V$ðri3iHiniÞ ¼ �3ivpvt

þ si : V$ni

þVkiVTi þXisj

hhij

�Tj � Ti

�� bij

�nj � ni

�2i

þXk

rkðcÞDHk (13)

where H is the specific enthalpy of the different phases, k the

thermal conductivity, h the convection coefficient between

Table 1 e Physical parameters used in the simulations.

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 4 11819

phases, T the temperature of the phases and r is the reaction

rate that occurs in the different solid phase.

dp(m) 500e-6

3p 0.5

s 4

rav(m) 1e-6

rg(kg/m3) EOS

Fig. 3 e Effectiveness Factor for different pressures and

temperatures.

3.2. Effectiveness factor calculation

In order to solve the transport equations of the process, the

effectiveness factor (h) must be assumed. Due to literature

lack about this term for ammonia cracking, it has to be

calculated in parallel with transport equations by solving the

particle equation of the catalyst. The effectiveness factor is

formally determined by solving the parabolic equation

derived from the species balance in the porous catalyst

particle; this further transport equation should be added to

the equation systems in each point of the computational

domain.

The system of equations that must be solved for particles

is:

3pvrgym

vt¼ 1

r2v

vr

�r2Dp;mrg

vym

vr

�þ rm (14)

Convective terms were assumed to be insignificant inside

the catalyst.

The effective diffusion coefficient in the particle (Dp,m) is

calculated with Eq. (15).

Dp;m ¼ Dp0;m3ps

(15)

Dp0,m was calculated using Bosanquet equation as reported

in Hayes [23]:

Dp0;m ¼ 11

Dmnþ 1Dkm

(16)

Dmn was calculated with Eq. (4) while Dkm is the Knudsen

diffusion, calculated as:

Dkm ¼ 23rav

ffiffiffiffiffiffiffiffiffiffiffi8RTpMm

s(17)

Because of the high importance the effectiveness factor

calculation has, a Crank-Nicolson method was used to solve

the systemof equations. In fact, the Crank-Nicolsonmethod is

unconditionally stable with a second order accuracy (in time

and space). A dedicated FORTRAN function was implemented

to solve Eq. (14).

Then the results were used to calculate h, necessary for the

definition of the reaction rates for the different chemical

species (Eq. (9)).

The effectiveness factor was finally calculated as:

h ¼3ZRc0

rm�ym;T

�r2dr

R3c rm

�y5mðcÞ;T

� (18)

where Rc is the radius of the catalyst particle. The Trapezoidal

rule was used to solve the numerator of Eq. (18).

In order to have a definition of reaction rates (Eq. (9)) as

accurate as possible, Eqns. (14) and (18) should be solved at

each iteration of Eqns. (2), (10), (11) and (13), using the

composition obtained at the catalyst surface ( ym(c)) as

boundary condition for Eq. (14). This procedure would

increase enormously the computational time, especially in

the case of 3D simulations. Therefore a particle equation was

solved separately for different gas composition and process

conditions. The results at steady state were then interpolated

and inserted in the code for the reactor simulation. Simulation

parameters are shown in Table 1:

Effectiveness factors obtained by solving the particle

equations are reported in Fig. 3. As depicted in the figure, an

increase of pressure makes the slope of the effectiveness

factor steeper. This effect is due to the hydrogen production

itself; as a matter of fact, the H2 partial pressure increase

reduces the conversion kinetic rate (see Eq. (7)). Under such

conditions the kinetic rate becomes the limiting step and the

effectiveness factor increased to 1. An opposite trend was

observed when increasing the reaction temperatures. A

temperature increase improves the reaction rates for

a broader hydrogen partial pressure range. An effectiveness

factor equal to 1 was reached quite later.

As already observed, kinetic of the reaction remains a key

problem for ammonia cracking because hydrogen production

reduces it. The use of a selectivemembrane could enhance the

reaction speed removing the hydrogen from the products as

soon as it is obtained, limiting therefore its partial pressure

and increasing significantly the reaction rate (see Eq. (7)).

3.3. Hydrogen permeation rate

The composite membranes considered in this study are made

of a very thin layer of palladium alloy. They are deposited as

a continuous layer on the outer surface of a thermostable

support. Details of themembranes geometry and composition

are given in Buxbaum [24]. The hydrogen permeation rate

Fig. 4 e Comparison between numerical simulation results

at 500 �C and Chellappa’s experimental results [25].

Fig. 5 e NH3 conversion at 450 �C (a) flow velocity: 2 cm/s;

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 411820

through the compositemembraneswas extrapolated from the

results published in Buxbaum [24]. We chose the permeation

rate to obey a linear power pressure law because a better fit

was obtained with this law:

QH2 ¼ 0:223exp

��1903T

�ðPH2u � PH2dÞ

�mol=m2s

�(19)

where T is the absolute temperature in kelvin. PH2u and PH2d

the hydrogen partial pressure (atm) upstream and down-

stream the membrane.

(b) flow velocity: 10 cm/s.

4. Results and discussion

4.1. 2D simulation without Pd-membrane

The predictions of NH3 decomposition efficiency as a function

of catalyst loading to NH3 molar flow rate ratio (W/F) for

various W under a reactor wall temperature of 500 �C are

compared in Fig. 4 with the experimental results reported by

Chellappa et al. [25]. The inlet pressure is 1 bar and the inlet

temperature is 450 �C.Even if further experimental validations would be, of

course, advisable, Fig. 4 shows our numerical simulations in

good agreement with Chellappa’s experimental results.

The first simulations were performed for a system with no

membranes, mainly with the aim of establishing a set of oper-

ative conditions to use with themembrane reactor. Because of

the axial symmetry of the reactor geometry, 2D simulations

were carried out. The following test conditionswere set: reactor

wall temperature at 450, 500 and 550 �C; operative pressure at

0.2, 0.5 and 1MPa;NH3 superficial velocity inside the reactor at 2

Table 2 e Operative conditions adopted for 2Dsimulations.

Mass flow (mg/s)

NH3 inlet flow(Tin ¼ 450 �C)

p ¼ 0.2 MPa p ¼ 0.5 MPa p ¼ 1 MPa

us ¼ 2 cm/s 1.3 3.25 6.5

us ¼ 10 cm/s 6.5 16.25 32.5 Fig. 6 e NH3 conversion at 500 �C (a) flow velocity: 2 cm/s;

(b) flow velocity: 10 cm/s.

Fig. 7 e NH3 conversion at 550 �C (a) flow velocity: 2 cm/s;

(b) flow velocity: 10 cm/s.

Fig. 8 e a) NH3 conversion; b) H2 molar fraction on

normalized reactor axis at different temperature with

us [ 1 cm/s.

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and 10 cm/s. The dependence of the NH3 input mass flow from

pressure and temperature is reported for the two superficial

velocities we considered in our 2D simulations in Table 2.

As already mentioned, ammonia decomposition is favored

by high temperatures and low pressure. Changing the super-

ficial velocity (NH3 input flow) instead affects the residence

time of the reactant gases: for example, a fast flow inside the

reactor reduces the gas residence time and the conversion

efficiency.

Fig. 5a) and b) show the ammonia conversion obtained at

450 �C as a function of the normalized reactor axes L/L0. These

results are quite poor for all the pressure values we consid-

ered: 0.2, 0.5 and 1 Mpa, and for both flow velocities of 2 cm/s

and 10 cm/s. Nonetheless, the negative influence of a high

pressure on the reaction kinetics is clearly evidenced; as

expected, the dissociation is favored by lower pressures.

Table 3 e Summary of 2D numerical simulation results.

450 �C

0.2 MPa 0.5 MPa 1 MPa 0.2 M

2 cm/s xH2 0.523 0.392 0.298 0.7

NH3 g/Nm3 227 360 456 45

10 cm/s xH2 0.385 0.266 0.195 0.6

NH3 g/Nm3 368 489 561 150

Even if they are not yet satisfactory, Fig. 6a) and b) show the

ammonia conversion improvements when the wall tempera-

ture is increased to 500 �C. When the temperature is risen at

about 550 �C and the flow velocity is limited at 2 cm/s, the

ammonia conversion appears to be extremely good, as it can

be seen particularly on Fig. 7a). For an operative pressure of

0.2 MPa, the NH3 dissociation appears to approach thermo-

dynamic equilibrium (value ¼ 1).

Therefore, a very high ammonia conversion rate (near the

thermodynamic equilibrium) emerged only for the lower

pressure and flow conditions and at relatively higher

temperatures (550 �C).

500 �C 550 �C

Pa 0.5 MPa 1 MPa 0.2 MPa 0.5 MPa 1 MPa

03 0.605 0.506 0.747 0.726 0.675

145 245 1.68 22 73

0.465 0.364 0.724 0.645 0.551

286 389 24 104 199

Fig. 9 e a) NH3 conversion; b) H2 molar fraction on

normalized reactor axis at different temperature with

us [ 2 cm/s.

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 411822

In Table 3we summarized the hydrogenmolar fraction and

the residual un-dissociated ammonia (in g/Nm3) still present

in the flow outgoing the 25 cm long “no membrane” reactor,

for all the conditions that were tested.

Only with the lowest pressure and flow values the

hydrogen molar fraction is near 0.75, value which indicates

the complete NH3 dissociation, while residual ammonia

outgoing the reactor at 550 �C, 0.2 MPa and 2 cm/s is 1.68 g/

Nm3, value that corresponds to 1300 ppm, which is less than

the lethal dose of 2000 ppm.

4.2. 3D simulation with Pd-membrane

The micro-reactor provided with tubular Pd-membrane rep-

resented in Fig. 1 was finally studied by numerical simulation.

The principal aim of these calculations was to evaluate how

much hydrogen permeation and purification through the Pd-

membranes could improve the ammonia decomposition

process, also at lower reaction temperatures (500 �C) and for

different NH3 input flow rates.

Because of the reactor asymmetric geometry (see Fig. 1),

a 3D simulation was unavoidable. In order to have a high H2

permeation through themembranes, simulationswere carried

out only at 1 MPa. Reactions at lower pressure (�0.5 MPa) were

not considered because those total pressures would not have

beenenough tohaveavailable the sufficientH2partial pressure

forhydrogenpermeationthroughthemembranes.Theprocess

was simulated varying the temperature from 500� to 550� and600 �C with a flow superficial velocity us equal to 1 cm/s

(3.25 mg/s) and 2 cm/s (6.5 mg/s).

The NH3 conversion rate and the H2 molar fraction

obtained simulating the process at three different tempera-

tures with a flow superficial velocity us ¼ 1 cm/s are shown in

Fig. 8.

Apparently, the best resultswere obtained for 600 �C, in this

case no appreciable un-dissociated ammonia was observed in

the retentate gas. However, it can be noticed that residual

ammonia in the output retentate gas was negligible already at

550 �C. The effect of H2 separation membranes can be seen

looking at Fig. 8. A change in all curve slopes can be observed at

thepointwhere themembranes start to separate thehydrogen

just produced. The presence of the membrane has a relevant

effect for the 550 �Ccase.At this temperature it clearly emerges

how the membranes enhance the NH3 dissociation, over the

chemical equilibrium, till the complete conversion. In the case

of 600 �C,most of the reaction took place in a very small area of

the reactor, immediately after the gas entrance and an almost

complete ammonia conversion was already obtained before

the membrane area. That was probably due to the high

temperaturewhich improves both equilibriumconversion and

kinetic rate. Nonetheless, even if at 600 �C the effect appears

not crucial, it is still visible how the membranes enhance the

conversion reaction. A shorter reactor or a higher input gas

flow would be advisable to have a well-balanced working

reactor and to improve H2 production. The case of 550 �C was

probably the most suitable for our reactor geometry: this

temperature is considerably lower than the ones generally

employed for ammonia synthesis process; hydrogen perme-

ation through the membranes played a role to improve

ammonia decomposition exploiting the whole reactor length

in an efficientway. An excellentmatching betweenmembrane

permeation and ammonia decompositionwas obtained in this

case. A 99.93% ammonia conversion was reached while the

equilibrium conversion at 550 �C is 99.92%.

In order to have the most efficient use of the membranes it

would be preferable to keep the hydrogen fraction of the

outgoing retentate gas around 0.1. This goal was met in Fig. 8

for both temperatures of 600 �C and 550 �C and for an input

flow rate of 1 cm/s. On the other hand, to design a cost

effective process the highest possible gas throughput

compatible with the device volume should be achieved.

Therefore, a further investigation was carried out increasing

the reactor feeding rate, considering an ammonia superficial

velocity of 2 cm/s. The results obtained for an ammonia

feeding flow rate of 2 cm/s are illustrated in Fig. 9.

As expected, increasing the superficial velocity decreases

the reactor performance in terms of ammonia dissociation.

For this set of simulation better results were obtained for

a process temperature equal to 600 �C. At 2 cm/s the hydrogen

molar fraction in the retentate gas varied between 0.29 for

500 �C and 0.18 for 600 �C, therefore the hydrogen amount not

permeated through the membranes and outgoing the reactor

with the retentate gas appears to be high.

These results are probably due to the short “residence

time” the input gas is allowed to spend inside the reactor,

Fig. 11 e Hydrogen flow permeated through the

membranes for different ammonia temperatures and input

flows.

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 4 11823

which is also the time the gas has to sweep the membranes.

Moreover, for the 500 �C case, the permeation rate is reduced

also because of the Arrhenius dependence of the permeation

rate with the temperature. This can explain the small differ-

ence of residual hydrogen observable in the retentate gas

between 500 �C and 600 �C.Regarding the role of the separation membranes on the

conversion rate, they bring their highest contribution at 550 �Cfor an ammonia feeding rate of 1 cm/s. Hydrogen permeation

through the membranes shifted the ammonia reaction,

helping significantly the conversion. Ammonia conversion in

the outgoing gas was 99.93% %. A comparison of ammonia

concentration in the retentate gas for both reactors, with and

without membranes, for a process temperature of 550 �C and

a flow rate of 1 cm/s, can be appreciated in Fig. 10.Without the

membranes nearly 82% of the input ammonia was dissoci-

ated, while with the membranes the conversion rate was

99.93%. This outcome corresponds to an improvement of

almost 18%.

Finally, the presence of the membranes resulted in an

enhancement of the dissociation process in all cases. As

expected the best results were obtained at higher tempera-

tures, because high conversion rates increase the partial

pressure of hydrogen and thus the H2 permeation through the

membrane.

In Fig. 11 we reported the hydrogen flow permeated

through the membranes for different ammonia temperatures

and input flows. A slight decrease of the curve slopes can be

observedwhen the process temperature increases. This result

is clearer for a flow velocity of 2 cm/s than for 1 cm/s. By

doubling the input ammonia flow, a doubling of the perme-

ated hydrogenwould be expected; this was almost obtained at

550 �C, but was completely obtained only for a temperature

equal to 600 �C. Due to a higher flow, reduction in ammonia

residence time, which implies lower conversion and lower

permeation, is compensated only with higher temperatures.

There are several studies reported in the literature about

Pd-membrane reactor employed to purify the hydrogen gas

derived from ammonia dissociation [6,15,26]. However,

different approaches and different operative conditions were

used: in all the cited works a sweep gas was used to remove

Fig. 10 e Comparison of NH3 conversion in the retentate

gas for the case of 550 �C, us [ 1 cm/s with and without

membranes.

the produced hydrogen from the permeation side; co-current

or counter-current sweep gas demonstrated to affect the

simulation results. In our model hydrogen is not swept by any

gas, it flows away from the membranes due to the pressure

difference downstream the membranes: hydrogen is the only

permeated gas and thus partial pressure is assumed to be

1 bar. Operative pressures in [6,15,26] were much higher (over

30 bar) than the one proposed in this work (10 bar as

a maximum). Obviously, chemical equilibrium, kinetic and

permeation rates are affected by the operative pressure thus

results should be compared with caution. Finally, the

permeation law used in those simulations differed from that

used in this work.

For all these reasons the authors believe that a direct

comparison of the results would be of little meaning while

a thorough experimental validation of each model would be

more appropriate. The experimental validation of ourmodel is

on the way and will be the subject of a next communication.

5. Conclusion

Ammonia cracking inside a commercial, small size catalytic

membrane reactor has been investigated by numerical simu-

lation. The CMR was considered filled with a fixed bed of

catalyst particles. The extra- and intra-particle diffusions

were coupled to kinetic mechanisms of reaction. The mass

transfer coefficient for the extra-particle diffusion was

calculated by literature correlation while the intra-particle

diffusion was considered by the use of an effectiveness

factor. To calculate the effectiveness factor a particle diffusion

model was solved. The outcomes of this model were coupled

to the transport equations of the gas phase.

The results showed that to get a satisfactory NH3 decom-

position with a no-ruthenium loaded catalyst, one must

operate at least at a temperature of 550 �C. Increasing the

reactor pressure seems to make the dissociation conditions

worse but, a pressure considerably higher than 1 bar is needed

at the reformate side to dispose of a sufficient hydrogen

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 6 ( 2 0 1 1 ) 1 1 8 1 5e1 1 8 2 411824

partial pressure and get a good permeation through the

membranes. These conditions are reached around 10 bar.

Hydrogen permeation through the membranes then shifts

the ammonia reaction, helping significantly the conversion:

an overall improvement in ammonia dissociation of almost

18% was obtained at P ¼ 10 bar and T ¼ 550 �C with an

ammonia input flow velocity of 1 cm/s. Moreover, even if with

different effectiveness, the presence of the membranes

contributed significantly to the dissociation process for all

cases. For an input flow velocity of 1 cm/s at both tempera-

tures of 550 �C and 600 �C we obtained a hydrogen molar

fraction at the reactor output of 0.1. But when the input flow

velocity was doubled to 2 cm/s the output hydrogen fraction

varied from 0.18 for the 600 �C case to 0.29 for the 500 �C case.

As expected, the best results were obtained at higher

temperatures; however, at this input flow rate, the hydrogen

amount outgoing the reactor together with the retentate gas

and, therefore, lost appears to be high. Residual hydrogen in

the retentate gas might be exploited through combustion to

recover heat for the process, however further investigations

would be needed to evaluate this conjecture. Despite balance

of plant optimization issues are very important to design an

efficient real system, a heat recovery strategy is beyond the

scope of this research.

Finally, as the optimal working conditions for the 30 cm3

CMR we simulated here, we can indicate a temperature of

550 �C, a pressure of 10 bar, and an NH3 input flow velocity of

1 cm/s, which corresponds to an NH3 mass flow of 2.82 mg/s.

An experimental investigation to verify the results obtained

here by numerical simulations will constitute the subject of

a future communication.

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