MODELING A FIXED-BED REACTOR FOR THE ...

 

MODELING A FIXED-BED REACTOR FOR THE OXIDATIVE DEHYDROGENATION OF ETHANE ON A MULTIMETALLIC MIXED OXIDE CATALYST Gamaliel Che-Galiciaa, Roberto Quintana-Solórzanob, Jaime S. Valenteb, Richard S. Ruiz-Martíneza, Carlos O. Castillo-Araizaa* a

Departamento de Ingeniería de Procesos e Hidráulica, Universidad Autónoma MetropolitanaIztapalapa. México D. F. México. b Instituto Mexicano del Petróleo. México D.F. México.

Abstract This work reports on simulation results of the performance of an industrial scale wall-cooled tubular fixed-bed reactor for the oxidative dehydrogenation of ethane to ethylene over a highly active and selective multimetallic mixed oxide catalyst containing Mo, V, Nb and Te. The referred reactor is simulated by applying a 2-D pseudo-homogeneous model to which, effective heat and mass transport parameters from experiments in the absence of chemical reaction accounting for hydrodynamics on heat transfer with some detail were incorporated [Ind. Eng. Chem. Res. 2007, 46 (23), 7426-7435]. For addressing the chemical phenomena taking place over the said catalyst, a Langmuir-HinshelwoodHougen-Watson based kinetic model is developed. Kinetic model parameters are obtained by fitting available lab-scale experimental data. Simulations results indicate an ethane conversion of around 60 %, ethylene selectivity close to 90 % and a slight hot spot reaching 450 ºC, when the temperature of the coolant medium was 440 ºC. When increasing further the coolant temperature to 480 ºC, ethane conversion augments to 76 %, ethylene selectivity decreases to 80 % and the hot spot increases to 535 ºC. These results suggest that the wall-cooled fixed-bed reactor may be a profitable alternative to perform the oxidative dehydrogenation of ethane to ethylene at the industrial scale.

Introduction Ethylene is the major raw material of the petrochemical industry. This olefin is, for instance, the main component in the production of ethylene oxide, ethylene dichloride and polyethylene, among others important chemicals. Presently, ethylene is mainly produced via steam cracking of diverse hydrocarbon streams and, in a minor grade, via catalytic dehydrogenation of ethane and fluid-catalyticcracking, in the later as a side product [1]. These processes, however, exhibit several drawbacks related to energy requirements, catalyst deactivation by coke deposition, control of conversion and selectivity, products separation and thermodynamics [2,3]. By reason of these limitations and considering the increase in the worldwide demand of ethylene, the design of a new process to cope with these industrial requirements is mandatory. In this sense, catalytic oxidative dehydrogenation of ethane to ethylene (ODH-Et) appears as an attractive alternative to complement and even replace the current processes for ethylene production. Needless to say that ODH-Et which involves an exothermic reaction wherein ethane conversion is not limited by the thermodynamic equilibrium can be performed at reaction temperatures below 500 ºC leading to a considerable energy saving. Other advantages of ODH-Et with                                                                                                                 *  Corresponding author:

E-mail address: [email protected] (C.O. Castillo-Araiza)  

 

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  respect to existing commercial processes is that, in the former, the number of reaction products is limited, i.e. ethane, ethylene and carbon oxides, without coke deposition over the catalyst due to the presence of oxygen in the reaction mixture [4,5]. There are still some important issues to be tackled on by industry and academy in order to extend ODH-Et to the commercial level, two of the most important ones being the development of a highly active and selective catalyst, and the reactor configuration. In the context of the catalyst, several formulations have been evaluated, the MoVNbTeO multimetallic mixed oxide system corresponding to one of the most promising materials in view of its high efficiency [6-14]. In the process scenario, the most convenient reactor configuration is still under debate from a potential industrial point taking into consideration the high exothermicity of the main and, mostly, the secondary reactions associated to the ODH-Et. It is therefore clear that controlling the reaction temperature through the reactor is a key aspect to avoid a reactor run-away, protect the catalyst structure and maintain values of ethane conversion and ethylene selectivity within expected ranges. For this end, an efficient removal of the heat released by the chemical transformations occurring in the catalyst bed is a crucial matter. The aim of this work is to present and discuss simulation results after applying modeling to investigate the performance of an industrial wall-cooled fixed-bed reactor in the ODH-Et over a highly active-selective MoVNbTeO based catalyst. A two-dimensional (2-D) pseudo-homogeneous model, which couples kinetics with hydrodynamics, heat as well as mass transport phenomena, was developed from well-known reactor engineering basic principles. Kinetics was developed using lab-scale catalytic experiments [15], whereas reactor key transport parameters characterizing hydrodynamics, dispersive, conductive and interfacial heat as well as mass transport phenomena were obtained previously from a study in absence of chemical reaction [16].

Procedures Kinetic experiments The catalyst used in the experimental part consisted of a multimetallic mixed oxide catalyst containing Mo, V, Nb and Te with a nominal atomic ratio equal to Mo:V:Te:Nb=1:0.24:0.24:0.18 [15]. Experimental ODH-Et was carried out at atmospheric pressure in a 1.0×10-2 m internal diameter quartzmade fixed-bed reactor, which was operated isothermally. In all tests, 0.60 g of catalyst, which was previously sieved for an average particle size equal to 150 µm, was loaded into the reactor. The reaction temperature was controlled/measured with a thermocouple located at the middle of the catalyst bed. The reaction feed consisted of a mixture containing 9 % mol of ethane, 7 % mol of oxygen and 84 % mol of nitrogen. The inlet flow rate of all gases was measured using Brooks 5850i series thermal mass flow controllers. The composition of the reactor by pass and the reactor effluent was analyzed on-line in an Agilent 7890A gas chromatograph, which was equipped with two detectors, a flame ionization detector (FID) and a thermal conductivity detector (TCD), as well as an array of three columns, a 30 m × 0.53 mm × 40 µm HP Plot Q, a 30 m × 0.25 mm × 5 µm HP Plot Al2O3/S and a 30 m × 0.53 mm × 50 µm molecular sieve. Hydrocarbons were quantified in the FID whereas non hydrocarbons in the TCD. Reaction temperature and space-time were varied between 400-480 ºC and 23-70 gcat·h·molethane−1, respectively. Kinetic modeling Only a limited number of papers have been related to the kinetic analysis of the ODH-Et. Such models are, however, specific for each catalyst formulation [17-19]. The MoVNbTeO catalytic system is a novel material and, therefore, developing its own kinetics is necessary. As a first step to develop kinetics model, a general reaction network was proposed on the basis of experimental observations, vide Figure 1. It consists of a set of parallel and consecutive reactions, in which, ethylene is produced out of  

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  ethane (r1), whereas carbon oxides are generated from combustion of ethane (r2 and r3) and secondary combustion of ethylene (r4 and r5).

2C O2

O

O

2

1 2 O2

O

2

2H

2

3H

O

2

r3

C2H4

H 2O

2

O

2O

5

2

2

C2H6

r4 3O

r1

2

7

O

2

2

3H

2H

r2

2CO

r5

Figure 1. Global reaction network used for modeling the kinetics of the ODH-Et in accordance to the LHHW formalism. Considering the reaction network presented in Figure 1, a Langmuir-Hinshelwood-HougenWatson (LHHW) mechanism was proposed taking into consideration a set of assumptions: (i) there is a single type of active sites on the catalyst surface, (ii) competitive adsorption of reactants, ethane and oxygen, occurs, (iii) reaction products, i.e., ethylene, carbon oxides and water, are susceptible to be readsorbed over the active sites, (iii) adsorption of oxygen is dissociative, (iv) surface reactions are the rate-determining steps (RDS), (v) adsorption/desorption steps are quasi-equilibrated. Additionally, except for reaction 1, reaction orders for oxygen inlet partial pressure were included as adjustable parameters. The corresponding reaction rates in accordance to reaction network given in Figure 1 are expressed by these equations: r1 =

k'1 (pO )1 2 pC H 2

6

2

6

2

4

2

4

2

2

2

2

2

2

6

2

6

2

4

2

2

2

2

2

2

2

6

2

6

2

4

2

6

4

2

k'4 (pO ) n3 2 pC H 2

2

2

2

4

(1+ (K O pO ) + K C H pC H + K C H pC H + K H O p H O + K CO pCO + K CO pCO ) n3 +1 2

6

2

6

2

4

2

4

2

k'5 (pO ) n4 2 pC H 2

2

2

2

2

6

2

6

2

4

2

4

4

2

2

with: 3

2

(4)

2

(1+ (K O pO )1 2 + K C H pC H + K C H pC H + K H O p H O + K CO pCO + K CO pCO ) n4 +1 2

(3)

2

12

2

(2)

2

(1+ (K O pO )1 2 + K C H pC H + K C H pC H + K H O p H O + K CO pCO + K CO pCO ) n2 +1 2

(1)

2

6

4

k'3 (pO ) n2 2 pC H

2

 

2

(1+ (K O pO ) + K C H pC H + K C H pC H + K H O p H O + K CO pCO + K CO pCO ) n1+1

2

r5 =

2

12

2

r4 =

2

k'2 (pO ) n1 2 pC H 2

r3 =

6

(1+ (K O pO ) + K C H pC H + K C H pC H + K H O p H O + K CO pCO + K CO pCO )2 2

r2 =

2

12

2

(5)

 

k'1 = k1K1O2 K C H 2

2

(6)

6

k'2 = k 2 K On1 2 K C H 2

2

k'3 = k 3K On2 2 K C H 2

2

2

k'5 = k5K On4 2 K C H 2

2

(8)

6

k'4 = k 4 K On3 2 K C H 2

(7)

6

(9)

4

(10)

4

For the i-th reaction, ri denotes the specific reaction rate in mmol·g-1·h-1 and ki the corresponding rate coefficient in mmol·g-1·h-1 pn is the partial pressure of component n (Pa) and Kn is the adsorption equilibrium constant of component n (Pa-1). Specific reaction rates are then combined to give global reaction rates denoted for the n-th species by Rn for the different observed compounds: 5

R n = ∑ υ n,iri

(11)

i=1

where υn.i is the stoichiometric coefficient of the component n for the i-th reaction. Kinetic model parameters were estimated by minimizing a weighted objective function RSS(β), which includes the residual sum of squares of the molar flow rates of the different species: n resp

nexp

n=1

k=1

1 2 n RSS(β) = ∑ w n ∑ (Fk,n − Fˆk,n )2 ⎯⎯⎯⎯ → min

β ,β ,...,β

(12)

where β is optimal parameters vector, nexp is the number of independent experiments, nresp is the number of responses, Fk,n and Fk,n are the n-th experimental and predicted responses, respectively for the k-th observation, and wn is the weight factor assigned to the n-th response. The molar flow rate of component n was calculated by solving a system of ODEs, vide eq. (13), corresponding to the experimental reactor model equations operated in the integral regime. At the studied operating conditions, both intra and inter-particle transport limitations were negligible according to criteria reported in literature [20] and, hence, the lab-scale reactor was model as an isobaric, isothermal, one-dimensional pseudo-homogeneous continuous fixed-bed reactor:

dFn = Rn dW

(13)

with initial conditions: Fn = Fno , when W = 0

(14)

where Fn is the molar flow rate of component n (mmol·h-1) and W is the mass of the catalyst loaded in the reactor (g). The subroutine DVODE was used to solve the corresponding set of ODEs. The initial minimization of the objective function (eq. 12) in the model regression was carried out using the Rosenbrock method [21]. Then, the ODRPACK subroutine was used for fitting calculated values to the corresponding experimental data points. This set of subroutines can perform either weighted orthogonal  

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  distance regression or nonlinear least square problems for explicit and implicit models using multiresponse data with an implementation of the Levenberg-Marquardt method [22]. In order to estimate values of activation energies and pre-exponential factors as well as standard adsorption enthalpies and entropies, a non-isothermal parameter estimation, i.e., using the experimental data at the three different reaction temperatures (400, 440 and 480 °C), was performed. To overcome the correlation between parameters a reparameterization is required. Thus, the Arrhenius equation in the reparameterized form was as follows:

⎡ E ⎛ 1 1 ⎞⎤ k'i = exp ⎢ A 'i − A,i ⎜ − * ⎟ ⎥ R ⎝ T T ⎠ ⎥⎦ ⎢⎣

(15)

while the Van’t Hoff equation in the reparameterized form adopted the following form:

⎡ ΔS ΔH n ⎛ 1 1 ⎞ ⎤ K n = exp ⎢ n − ⎜⎝ T − T* ⎟⎠ ⎥ R R ⎣ ⎦

(16)

For the i-th reaction, A 'i is the natural logarithm of the pre-exponential factor (mmol·g-1·h-1·Pa(1+β) ), and EA,i is the activation energy (kJ·mol-1). T is the reaction temperature (K), T* is the averaged reaction temperature (K), ∆S°n is the standard entropy adsorption of component n (kJ·mol-1·K-1) and ∆H°n is the standard adsorption enthalpy of component n (kJ·mol-1). Industrial Reactor Model A 2-D pseudo-homogeneous wall-cooled fixed-bed reactor model that couples heat and mass transport phenomena to kinetics was proposed to describe the ODH-Et taking place over the MoVNbTeO based catalyst. The interfacial heat and mass transport resistances between the fluid and the catalyst particle were neglected considering the relatively high value of the inlet flow rates used during the operation of this type of industrial reactors [16]. Effective heat and mass transport parameters, i.e., keff, hw, Deffr, and Deffz, included in the reactor model were obtained from a previous study wherein the role of hydrodynamics on heat and mass transport phenomena in the absence of chemical reaction was assessed with some detail [16, 23-25]. The governing reactor model equations for mass and heat, on the basis of the assumptions given above, are the following: r ⎛ ∂ 2 C n 1 ∂C n ⎞ ∂C n ∂C n ∂2 Cn + vz = Deffr ⎜ 2 + + Deffz + ρb ∑ υ niri ∂t ∂z r ∂r ⎟⎠ ∂z 2 ⎝ ∂r i=1

ρf C pf

r ⎛ ∂ 2 T 1 ∂T ⎞ ∂T ∂T ∂2 T + v z ρf C pf = k effr ⎜ 2 + + k + ρ (−ΔH i )ri effz b∑ ∂t ∂z r ∂r ⎟⎠ ∂z 2 ⎝ ∂r i=1

(17)

(18)

The corresponding initial and boundary conditions are: t = 0;

 

C n = C n,ss

and

T = Tss

(19)

5

  z = 0;

v z C no = v z C n − Deffz

∂C n ∂z

v z ρf C pf To = v z ρf C pf T − k effz

(20)

∂T ∂z

(21)

z = L;

∂C n =0 ∂z

and

∂T =0 ∂z

(22)

r = 0;

∂C n =0 ∂r

and

∂T =0 ∂r

(23)

∂C n ∂T =0 k effr = h w (T − Tb ) and (24) ∂r ∂r where Cn is the molar concentration of component n in the gas phase (mol·m-3), Deffr is the radial mass dispersion coefficient (m2·h-1), Deffz is the axial mass dispersion coefficient (m2·h-1), ρb is the fixed-bed density (kg·m-3), ρf is the fluid density (kg·m-3), Cpf is the specific heat capacity of the fluid (J·kg-1·K-1), keffr is the radial effective thermal conductivity (W·m-1·K-1), keffz is the axial effective thermal conductivity (W·m-1·K-1) and hw is the wall heat transfer coefficient (W·m-2·K-1). The resulting reactor model was given in terms of a set of parabolic partial differential equations, which was solved numerically by the method of orthogonal collocation using 30 and 50 interior collocation points at the radial and axial coordinates, respectively, employing shifted Legendre polynomials [26]. The resulting set of ordinary differential equations was then solved by the Runge-Kutta-Fehlberg method [27]. The reactor was simulated by using operating conditions that may be relevant for industry. In particular, the particle Reynolds number was around 630 since transport parameters in absence of reaction were obtained at this flow condition, the inlet molar concentration of the reaction mixture was constant and equal to 9/7/84 for C2H6/O2/Inert. A parametric sensitive analysis was carried out varying the coolant temperature from 400 to 480 ºC and the inlet temperature of the reaction mixture from 100 to 200 ºC. Reactor dimensions were 10 m length (L) and 2.53 cm internal diameter (iD), while the catalyst packed in the bed consisted of spherical non-porous particles with a diameter (dp) of 8 mm. Thus, the ratio iD to dp was 3.16. r = Rt;

Results and Discussion Kinetics The estimated parameters and corresponding 95% probability confidence intervals are presented in Table 2. Qualitatively, values are in a good agreement with previous studies for the oxidative dehydrogenation of propane over a V-Ti-O catalyst [19, 28, 29]. From the information displayed in Table 2 it is first noticed that values of reaction orders associated to oxygen are lower than 1.0 indicating that reaction rates are less sensitive to oxygen inlet partial pressure compared to that of hydrocarbons (ethane and ethylene). On the basis of the value of activation energies, the main reaction leading to ethylene out of ethane is the one that demands the lowest value of activation energy to proceed, 70.0 kJ·mol-1. On the contrary, total oxidations producing CO from ethane (r3) and ethylene (r5) are the reactions with the largest activation energy demand, 120 and 121 kJ·mol-1, respectively. It is also important to mention that the magnitude of the activation energy indicates how sensitive the reaction rates are to temperature changes; more precisely, reactions with a relatively high activation energy are more significantly affected by temperature increases. On the basis of the previous reasoning,  

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  it is clear that the relative importance of oxidation reactions and, hence, COx products, augments with temperature. Also notice that even though activation energies are similar for total oxidations of ethane and ethylene, catalyst active sites favors the total oxidation of ethylene over the total oxidation of ethane as deduce from the values of the pre-exponential factor (vide Table 2); namely, A2 and A3 related to ethane oxidation are lower than A4 and A5 related to ethylene oxidation. Table 2. Kinetic parameters values and corresponding 95% probability confidence intervals of the LHHW model used to describe the ODH-Et over the MoVNbTeO catalyst. Parameter Value Parameter Value -8

70.0 ± 0.4

(1.17 ± 0.02) ×10

EA,2, kJ·mol

105.9 ± 1.0

A '3 , mmol·g-1·h-1·Pa-(1+β)a

(3.87 ± 0.05) ×10-9

EA,3, kJ·mol-1

120.2 ± 1.0

A '4 , mmol·g-1·h-1·Pa-(1+β)a

(6.05 ± 0.03) ×10-3

EA,4, kJ·mol-1

100.4 ± 2.7

A '5 , mmol·g-1·h-1·Pa-(1+β)a

(4.79 ± 1.71) ×10-4

EA,5, kJ·mol-1

121.3 ± 1.8

ΔS°O2, J·mol-1·K-1

(6.19 ± 0.06) ×101

ΔH°O2, kJ·mol-1

(5.77 ± 0.04) ×101

ΔS°C2H6, J·mol-1·K-1

(1.35 ± 0.01) ×101

ΔH°C2H6, kJ·mol-1

(4.75 ± 0.04) ×101

ΔS°C2H4, J·mol-1·K-1

(4.68 ± 0.03) ×101

ΔH°C2H4, kJ·mol-1

(4.69 ± 0.05) ×101

ΔS°CO2, J·mol-1·K-1

(4.36 ± 0.03) ×101

ΔH°CO2, kJ·mol-1

(9.79 ± 0.41) ×101

ΔH°CO, kJ·mol-1

N.S.b

N.S.b

ΔS°H2O, J·mol-1·K-1

(4.61 ± 0.03) ×100

ΔH°H2O, kJ·mol-1

(6.48 ± 0.05) ×101

n1

(4.91 ± 0.05) ×10-1

n3

(6.26 ± 0.09) ×10-1

n2

(1.45 ± 0.01) ×100

n4

(1.34 ± 0.02) ×10-1

β is the reaction order associated to the partial pressure of oxygen for reaction i. N.S. non-significant parameter, not reported.

Fi calculated [mmol/h]

a) 25

b) 25

+10 %

20

-10 %

15 10 5

Fi calculated [mmol/h]

b

-1

A '2 , mmol·g-1·h-1·Pa-(1+β)a

ΔS°CO, J·mol-1·K-1

a

EA,1, kJ·mol-1

5.53 ± 0.04

A '1 , mmol·g-1·h-1·Pa-(1+β)a

0

+10 %

20

-10 %

15 10 5 0

0

5

10

15

20

25

0

Fi observed [mmol/h]

5

10

15

20

25

Fi observed [mmol/h]

Figure 2. Parity plots comparing experimental with calculated for reactor outlet molar flow rates for: (a) ethane and (b) ethylene.  

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  Concerning the parameters of the Van’t Hoff’s equation, values of standard enthalpies and entropies of adsorption are physically consistent according to criteria given by Boudart et al. [30]. Namely, standard adsorption enthalpies are all negative while standard adsorption entropies are between 0.0 and the corresponding gas phase molecular standard entropy. The most negative value of the adsorption enthalpy holds for CO2 adsorption, -98 kJ·mol-1, while the least negative stands for the adsorption of ethylene, -47 kJ·mol-1. Regarding kinetic model adequacy, two parity plots, which contrast experimental versus calculated molar flow rates of ethane and ethylene, are included in Figure 2. Clearly, there is an adequate correspondence between the calculated and the observed molar flow rates, suggesting that the proposed kinetic model is capable to describe experimental data. The quality of parity plots for oxygen, water and carbon oxides is similar, which are not shown for brevities’ sake. Likewise the value of Freg (3224) is substantially higher than tabulated one is (2.79). Industrial Reactor Simulations Temperature and concentration predictions along the reactor at three different coolant temperatures (400, 440 and 480 °C) and two different inlet temperatures of the reaction mixture (100 ºC and 200 °C) are shown in Figures 3(a) - (d). 1.0

350 0.6

C2H6 O2 C2H4 CO2 CO

0.4

300

250

0.2

200 4

z [m]

6

8

1.0 0.8

d)

450

350 300

0.2

250

0.0

200 2

4

z [m]

6

8

300

0.2

250 200 2

4

z [m]

6

8

10

1.0

450 C2H6 O2 C2H4 CO2 CO

0.8 0.6

400 350

T [ºC]

400

0.4

350

0.4

0

500

0.6

0

0.6

400

0.0

550

C2H6 O2 C2H4 CO2 CO

450 C2H6 O2 C2H4 CO2 CO

0.8

10

T [ºC]

Xn [Dimensionless]

2

Xn [Dimensionless]

0

1.0

T [ºC]

Xn [Dimensionless]

0.8

0.0

c)

b)

400

T [ºC]

Xn [Dimensionless]

a)

0.4

300

0.2

250

0.0

10

200 0

2

4

z [m]

6

8

10

Figure 3. Dimensionless concentration and temperature profiles along the reactor axial position for three different simulation scenarios varying coolant temperature maintaining the inlet reaction mixture temperature at 200 °C:(a) 400 ºC, (b) 440 ºC, (c) 480 ºC; (d) contains an additional simulation performed at an inlet reaction mixture temperature of 100 ºC at a coolant temperature of 440 °C.  

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  Clearly, as the coolant temperature augments ethane conversion increases, ethylene selectivity decreases while the hot spot becomes more pronounced. More specifically, our simulation results show that when coolant temperature is equal to 400 ºC, ethane conversion is as high as 25 % and the selectivity to ethylene amounts to 95 % with the presence of a slight hot spot (ca. 405 °C) at about 1.3 m length of reactor. In an intermediate scenario, i.e., at a coolant temperature medium equal to 440 °C, ethane conversion increases to 59 %, ethylene selectivity decreases to 90 % and the hot spot is still slight with a temperature value of ca. 450 ºC. Augmenting further the temperature of the coolant fluid to 480 ºC increases ethane conversion to 75 %, decreases selectivity to ethylene to 80 % and results in a more evident hot spot at which temperature increases up to 530 ºC after about 1.0 m length of reactor. In an additional simulation result (vide Figure 3d), the inlet temperature of the reaction mixture was increased decrease from 200 to 100 °C while the temperature of the coolant medium was maintained at 440 °C. Evidently, such a decrease in the inlet temperature of the reaction mixture did not have an appreciably effect on the reactor performance as values of dimensionless concentrations and temperature remain practically unchanged. From all these results, it is learned that although using a relatively high coolant temperature yields a larger ethane conversion, it significantly increases the value of the hot spot which may cause a set of unwanted consequences in the process. Apart of decreasing the selectivity to ethylene which is the desired product, high values of temperature may alter the structure of the catalyst by sintering and, in a more complicated scenario, lead to an unsafe region of the reactor operation.

Conclusions This work reports simulations at operating conditions relevant for industry of the performance of an industrial wall-cooled fixed-bed reactor in the ODH-Et over a highly active-selective MoVNbTeO based catalyst. In order to address the chemical phenomena, a LHHW based kinetic model was developed and next coupled to a 2-D pseudo-homogeneous wall-cooled fixed bed reactor model. The kinetic model represents adequately experimental observations while corresponding kinetic parameters are physically meaningful. Simulation results indicate that a careful selection of the temperature of the coolant fluid is important for tuning the reactor performance. Though increasing the temperature of the coolant medium has a positive impact on ethane conversion, selectivity to ethylene decreases while an augment in the magnitude of the hot spot. Changing the inlet temperature of the reaction mixture, on the contrary, does not alter dramatically the performance of the reactor. On the basis of the results discussed in this document, the wall-cooled fixed-bed reactor represents a promising alternative for petrochemical industry to carry out the ODH-Et over selective catalyst such a MoVNbTeO based solid reaching values of ethane conversion of around 60 and an ethylene selective of around ca. 90% operating at a Rep=630, temperature of the coolant of 440 ºC and an inlet reaction mixture molar composition of 9/7/84 for C2H6/O2/Inert.

References 1. 2. 3.

 

Cavani, F., Ballarini N. and Cericola, A. (2007), “Oxidative dehydrogenation of ethane and propane: How far from commercial implementation?,” Catalysis Today, 127, pp. 113-131. Albonetti, S., Cavani, F., Trifirò, F. (1996), “Key Aspects of Catalyst Design for the Selective Oxidation of Paraffins,” Catalysis Reviews: Science and Engineering, 38, pp. 413-438. Blasco, T. and López Nieto, J.M. (1997), “”Oxidative dehydrogenation of short chain alkanes on supported vanadium oxide catalysts,” Applied Catalysis A: General, 157, pp.117-142. 9

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