Nonoxidative and Oxidative Propane Dehydrogenation over Bimetallic

Article pubs.acs.org/IECR

Nonoxidative and Oxidative Propane Dehydrogenation over Bimetallic Mo−Ni/Al2O3 Catalyst Arman Siahvashi, Dean Chesterfield, and Adesoji A. Adesina* Reactor Engineering & Technology Group, School of Chemical Engineering, The University of New South Wales, Sydney, Australia 2052 ABSTRACT: Propane nonoxidative and oxidative dehydrogenation reactions were investigated over bimetallic alumina supported Mo−Ni catalyst in a quartz fixed-bed reactor at 0.1 MPa within the temperature range of 773−973 K. Time-on-stream analysis showed that Mo−Ni catalyst is stable at 4 h reaction time during ODH, while it significantly deactivated during NODH due to heavy carbon deposition confirmed by TOC. Hydrogen production was considerable during both reactions. C3H6 production rate is significantly higher during NODH at T > 823 K. An increase of the propane partial pressure resulted in a decrease in propane conversion to 83.2%, and selectivity to C3H6 significantly enhanced from 7.1 to 16.3%. An increase in oxygen partial pressure increased propane conversion from 74 to 95%, while the selectivity to C3H6 was considerably diminished at 923 K.

1. INTRODUCTION Today’s world trend attests the growing demand of alkenes, such as ethylene and propylene. Generally, alkenes are produced via steam cracking, fluid-catalytic-cracking (FCC), and catalytic dehydrogenation. However, the first two methods are believed to be insufficient to satisfy the expanding alkenes market; therefore, catalytic dehydrogenation has been found as an alternative route for the production of alkenes and it is of paramount importance from the industrial standpoint due to its high selectivity to the corresponding short-chain alkenes.1,2 Propane dehydrogenation is an endothermic process, implying that high temperatures ranging from 723 to 973 K are needed to achieve a high yield of propylene (C3H6) at 1 atm. During propane dehydrogenation, the following reactions take place.3−5 C3H8 ↔ C3H6 + H 2

(1)

C3H8 → CH4 + C2H4

(2)

C2H4 + H 2 → C2H6

(3)

C3H8 + H 2 → CH4 + C2H6

(4)

(5)

C3H8 + 5O2 → 3CO2 + 4H 2O

(6)

C3H8 + 3.5O2 → 3CO + 4H 2O

(7)

© 2013 American Chemical Society

(8)

C3H6 + 3O2 → 3CO + 3H 2O

(9)

CO + 0.5O2 → CO2

(10)

C3H8 → CH4 + C2H4

(11)

where reaction 5 is propane oxidative dehydrogenation and reactions 6 and 7 are the oxidation of C3H8 to CO2 and CO respectively. Both reactions 8 and 9 correspond to the oxidation of propylene to CO2 and CO. Finally, the last two reactions are ascribed to the oxidation of CO and propane cracking, respectively. Reactions 5−11 reveal the dehydrogenation of alkane along with water production. In principle, the removal of hydrogen from the reaction system increases alkene yields and is carried out by adding an oxidant (e.g., molecular oxygen) to produce H2O instead of H2. This enables the ODH reaction to be thermodynamically favorable at any level of alkane conversion and alkene yield.5 However, comparatively little attention has been paid to the formation of H2 during ODH. Indeed, at higher temperatures (above 773 K), where the total oxygen conversion is reached, H2 production cannot be neglected. Several papers have confirmed the formation of H2 in the ODH process.9−15 It was suggested that the main pathways for hydrogen formation during ODH are nonoxidative dehydrogenation, coking, and free radical reaction.9,11 Thermodynamic calculations associated with the main ODH and NODH reactions are reported in Table 1. It is evident that NODH reaction is feasible at temperatures above 923 K, while ΔG of the ODH reaction is negative at all temperatures. The main objective of the present study is to compare nonoxidative and oxidative propane dehydrogenation in the temperature range 773−973 K and to investigate the

where reaction 1 is the main dehydrogenation reaction accompanied by the probable side reactions (reactions 2−4) in which H2, C3H6, CH4, C2H4, and C2H6 are produced. A major issue with propane dehydrogenation is the equilibrium limitation and as a result, oxidative dehydrogenation (ODH) is utilized to circumvent this problem. Unlike the nonoxidative dehydrogenation of propane, oxidative dehydrogenation is exothermic and irreversible.6,7 The following set of reactions was proposed by Barsan et al.8 for oxidative dehydrogenation of propane to propylene: C3H8 + 0.5O2 → C3H6 + H 2O

C3H6 + 4.5O2 → 3CO2 + 3H 2O

Received: Revised: Accepted: Published: 4017

September 5, 2012 January 7, 2013 February 20, 2013 February 20, 2013 dx.doi.org/10.1021/ie302392h | Ind. Eng. Chem. Res. 2013, 52, 4017−4026

Industrial & Engineering Chemistry Research

Article

was obtained using the same instrument via NH3- and CO2temperature-programmed desorption (TPD). C3H8-TPD analysis was also performed to identify potential adsorption site(s) during reaction. The relevant probe gases for all TPD runs were 10% NH3/N2, 10% CO2/N2, and C3H8, respectively. The desorbed NH3, CO2, and C3H8 concentrations were monitored by a thermal conductivity detector during heating in helium flow to 973 K at various heating rates of 10, 15, 20, and 30 K min−1. Before each procedure in the Autochem 2910, prepared catalyst samples underwent pretreatment in 45 mL min−1 H2 for 2 h at 973 K. Powder X-ray diffraction (XRD) patterns of the calcined catalysts were obtained on an X’Pert Pro Multipurpose X-ray Diffraction (MPD) system using Cu Kα radiation (λ = 0.154 nm) operated at 40 mA and 45 kV. The diffractograms were analyzed using X’Pert Score Plus software. Thermogravimetric runs were conducted on a ThermoCahn TherMax 200 TGA unit to generate the weight change profiles of the catalyst. Approximately 100 mg of the uncalcined specimen was initially loaded into the sample boat. A temperature-programmed calcination (TPC) experiment for fresh catalyst was performed using high purity air at 55 mL min−1 and ramped to 973 at 5 K min −1. Temperatureprogrammed reduction (TPR) was conducted in 55 mL min−1 of 50% H2/Ar mixture using the same temperature-program cycle. TPO-TPR and TPR-TPO runs were also carried out with the same instrument to identify the oxide phases present in spent catalysts. Carbon content in the used catalyst was also determined using a Shimadzu Solid Sample Module SSM5000A coupled to a Total Organic Carbon (TOC) Analyzer 5000A. 2.3. Catalytic Studies. The reaction runs were conducted in a stainless-steel fixed-bed reactor with 15 mm i.d. (length = 300 mm) quartz tube fitted with a temperature controller. Figure 1 illustrates the essential features of this setup. The runs were carried out in the range 773−973 K at a constant total pressure of 0.1 MPa. Feed containing C3H8 (NODH) and O2 (ODH) (diluted in Ar) at different ratios was passed downward through the reactor bed packed with 0.5 g of catalyst. The outlet gases were analyzed by a TCD-equipped Shimadzu GC8A gas chromatograph serviced by an Alltech Hayesep DB 100/

Table 1. Thermodynamic Calculations for Reactions 1 and 5 As a Function of T(K) reaction

ΔH298° (kJ mol−1)

ΔG(T) (kJ mol−1)

C3H8 ↔ C3H6 + H2 C3H8 + 0.5O2 ↔ C3H6 + H2O

124.87 −117.43

129.3−0.14T −116.9−0.086T

performance of alumina-supported bimetallic 5% Mo−10% Ni catalyst under these reactions.

2. EXPERIMENTAL SECTION 2.1. Catalyst Preparation. Alumina-supported catalyst with a composition of 5 wt % Mo−10 wt % Ni was prepared using the coimpregnation method. The γ-alumina support (Saint-Gobain Norpro) was first crushed and sieved to 140− 425 μm before thermal treatment at 1073 K for 6 h at 20 K min−1. The alumina support was then impregnated with quantitative amounts of an aqueous solution of (NH4)6Mo7O24·4H2O and Ni(NO3)2 (Sigma−Aldrich, Australia), added simultaneously as precursors of Mo and Ni, respectively. Slurry impregnation was carried out in a Mettler−Toledo T90 autotitrator system at a controlled temperature of 303 K. The slurry was stirred continuously for 3 h at 2250 rpm followed by oven drying for 24 h at 393 K to remove excess water. The resulting dried product was then calcined at 1073 K for 5 h in air at a heating rate of 5 K min−1. Calcination at 50 K above the maximum reaction temperature employed ensures that subsequent changes (if any) in textural properties (surface area or pore structure) would not be due to thermally induced stresses under reaction conditions.16 The calcined catalyst was crushed and sieved to 140−250 μm before in situ H2-activation and reaction in a fixed-bed reactor. 2.2. Catalyst Characterization. Various characterization techniques were used to secure the innate properties of the prepared catalyst. Quantachrome Autosorb-1 unit was used for multipoint BET surface area, pore volume, and pore size measurements from N2-physisorption at 77 K. H2 chemisorption at 383 K to determine the metal dispersion, metal surface area, and active particle size was carried out on a Micromeritics AutoChem 2910 unit. Acidic and basic character of the catalyst

Figure 1. Schematic diagram of the experimental setup. 4018

dx.doi.org/10.1021/ie302392h | Ind. Eng. Chem. Res. 2013, 52, 4017−4026


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120 column. All gas flow rates were adjusted using Brooks 5850E mass flow controllers. Prior to the reaction, in situ reduction of calcined catalyst was performed in 45 mL min−1 of H2 with a heating rate of 10 K min−1 and held at 973 K for 2 h. The reactor was then cooled down under a blanket of Ar at ambient temperature to achieve the desired reaction temperature. In order to minimize the transport disguised kinetics during the data analysis, gas hourly space velocity (GHSV) of 12 000 mL gcat−1 h−1 and a catalyst particle size range of 140− 250 μm were employed. Moreover, the product line was fitted with drierite (CaSO4) to remove any moisture before product analysis on the GC.

3. RESULTS AND DISCUSSION Figure 2 portrays the derivative weight change profile of the catalyst during temperature-programmed calcination using air. Figure 3. X-ray diffraction profile of calcined Mo−Ni catalyst.

Table 2. Physicochemical Properties of Alumina Support and Reduced Mo−Ni Catalyst calcined γ-Al2O3 BET surface area (m2 g−1) pore volume (mL g−1) pore diameter (nm) metal dispersion (%) metal surface area (m2 gcat−1) active particle size (nm)

167.2 ± 1.3 0.83 ± 0.004 18.67 ± 0.3

Mo−Ni/Al2O3 106.7 0.56 21.04 0.61 0.52 170.3

± ± ± ± ± ±

1.1 0.005 0.2 0.013 0.018 2.4

dissolution of alumina (hence, reduction in BET area to 106.7 m2 g−1) and pore blockage by the metal oxide particles. The relatively low percentage metal dispersion (0.61%), metal surface area (0.52 m2 g−1), and active particle size (170.3 nm) were consistent with the high metal loading (5% Mo−10% Ni) used. Table 3 summarizes the results obtained from NH3- and CO2-TPD. Data from the different heating rates (10−30 K

Figure 2. Weight change profile of Mo−Ni catalyst during temperature-programmed calcination.

It is evident that full decomposition of the precursors is achieved well below 700 K. It has been suggested that the intermetallic oxide phase of NiMoO4 is the reason for the major peak at 475 K followed by the surface adsorbed nitro species (MoO3 and NiO with a NO2 ligand at the metal oxides)17,18 which are formed by the calcination of ammonium molybdate and then adsorbed on the nickel surface as shoulders at 545 and 590 K.19 Even so, precipitation reaction between the hydroxides of Ni and Mo during impregnation may have given rise to a complex bimetallic hydroxide which subsequently decomposed to the NiMoO4 phase upon calcination. The formation of metal oxide phases during calcination is confirmed by the X-ray diffraction of the calcined Mo−Ni catalyst and alumina support depicted in Figure 3. The patterns attest the existence of multiple metal oxide phases such as NiMoO4 (2θ = 31.3°) and NiAl2O4 (2θ = 37.1°, 45.4°, 59.8°, and 66.3°). A small peak corresponding to NiO was registered at 2θ = 56.1°, while no reflection was detected for MoO3 either due to low Mo loading (5 wt %) used or the formation of XRD-amorphous fine MoO3 particles. The data in Table 2 indicate that the alumina support has higher BET surface area (167.2 m2 g−1) than the Mo−Ni catalyst. The difference in textural properties between the pure alumina and the catalyst suggests that the addition of Mo and Ni in the bimetallic system may be due to acid attack and

Table 3. Acid and Basic Site Characteristics of the Support and Reduced Catalyst Obtained from NH3 and CO2-TPD peak number ΔHd,NH3(kJ mol−1) acid site concentration (μmol m−2) ΔHd,CO2(kJ mol−1) basic site concentration (μmol m−2) acidic/basic site ratio

I II I II total I II I II total

calcined γ-Al2O3 46.64 ± 0.5 2.21 ± 0.06 2.21 44.31 69.88 0.19 0.41 0.6 3.68

± ± ± ± ± ± ±

0.06 2.1 1.7 0.004 0.006 0.01 0.08

Mo−Ni 37.26 58.89 2.88 1.03 3.91 48.59

± ± ± ± ± ±

0.7 1.8 0.05 0.02 0.07 2.9

0.47 ± 0.003 0.47 ± 0.003 8.31 ± 0.06

min−1) were employed to estimate the heat of desorption (ΔHd) using eq 12.20,21 ⎛ −ΔH ⎞ ⎛ −ΔH A ⎞ ln β d⎟ d sat ⎟ ⎜⎜ ⎜⎜ = + ln ⎟ ⎟ R T R C Tp2 ⎝ g p ⎠ ⎝ ⎠ g 4019

(12)



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where β is the heating rate (K min−1), Tp is the peak temperature for species desorption, Rg is the gas constant, A sat is the adsorbed saturated amount, and C is a desorption rate constant. A plot of ln β/Tp2 versus 1/Tp enables the calculation of −ΔHd . In addition, the acid/basic site concentration was calculated using

On the other hand, CO2-TPD (cf. Figure 4b) revealed two distinct peaks, suggesting the presence of two basic sites on the catalyst. The first peak positioned at 400−430 K is indicative of a weak Lewis basic site with a CO2 heat of desorption of 44.31 kJ mol−1, while the second peak at 700−800 K is attributed to a strong Lewis site.22 Compared to basic sites, the acid sites are stronger (ΔHd = 37.3 and 58.9 kJ mol−1), and the total acid site concentration of the catalyst (3.91 μmol m−2) is higher than the total basic site concentration (0.47 μmol m−2), indicating an innate acidic catalyst. This is also confirmed by the acid/ basic site ratio of 8.31. The C3H8-TPD profile is illustrated in Figure 5, where a major peak is shown in the temperature range 600−640 K and

yF

ζ=

i dt ∫ 22414

w

(13)

where yi = mole fraction of desorbed gas component, i; F = total outlet volumetric flow rate (mol min−1); and w = weight of sample (g). Figure 4a,b displays the NH3-TPD and CO2-TPD profiles for calcined Al2O3. Alumina support showed one peak for NH3TPD (cf. Figure 4a) with heat of desorption 46.6 kJ mol−1. The low temperature peaks at 520−573 K are representative of the existence of weak Lewis acid centers22 since values of ΔHd,NH3 greater than 125 kJ mol−1 are typical of Bro̷ nsted acid sites.23

Figure 5. C3H8-TPD profile of reduced Mo−Ni catalyst.

a second peak between 680 K and 710 K. This suggests the existence of two types of sites for C3H8 adsorption. Given that the Mo−Ni catalyst has two types of acid sites, it is likely that these centers were also responsible for C3H8 adsorption. Table 4 reports the propane heat of desorption and site concentration for reduced Mo−Ni catalyst, indicating that propane may adsorb on both acid and metal sites. Table 4. C3H8-TPD Results of Reduced Mo−Ni Catalyst attributes −ΔHd,C3H8 (kJ mol−1) desorbed C3H8 (μmol m−2)

peak number I II I II

Mo−Ni 97.4 93.2 1.8 2.4

± ± ± ±

1.7 2.1 0.092 0.11

3.1. Reaction Runs. The transient rate profiles during nonoxidative and oxidative propane dehydrogenation for H2, CH4, and C3H6 formation rates as well as propane conversion over bimetallic Mo−Ni/Al2O3 catalyst are illustrated in Figure 6a−d. The experiment was carried out at 923 K with PC3H8 = 15.19 kPa for NODH and PC3H8 = 15.19 kPa and PO2 = 10.13 kPa for ODH in 4 h. Time-on-stream (TOS) analysis shows that more H2 was formed during ODH reaction compared to NODH. The H2 formation rate increased in the first hour of the ODH reaction, followed by a gentle decline. However, H2 production during NODH considerably decreased due to catalyst deactivation via carbon deposition. Basically, the

Figure 4. (a) NH3-TPD and (b) CO2-TPD of calcined alumina support. 4020



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Figure 6. Time dependent analysis of (a) H2, (b) CH4, and (c) C3H6 formation rates and (d) C3H8 conversion during NODH (PC3H8 = 15.19 kPa) and ODH (PC3H8 = 15.19 kPa and PO2 = 10.13) at 923 K.

Figure 7. (a) CO and CO2 formation rate and (b) H2/CO ratio as a function of time during ODH at 923 K.

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Figure 8. Production rate of (a) H2, (b) CH4, (c) C3H6, and (d) C3H8 consumption rate at different temperatures (PC3H8 = 15.19 kPa for NODH and PC3H8 = 15.19 kPa and PO2 = 10.13 kPa for ODH).

Figure 9. Temperature-dependent analysis of (a) CO and CO2 formation rate and (b) H2/CO ratio during ODH (PC3H8 = 15.19 kPa and PO2 = 10.13 kPa).

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Figure 10. Effect of reactant partial pressure on C3H8 conversion and product selectivity- H2, CO, CO2, CH4, and C3H6 for (a) propane (fixed PO2 = 15.19 kPa) and (b) oxygen (fixed PC3H8 = 20.26 kPa) partial pressure at 923 K.

Table 5. Parameter Estimates for NODH and ODH at 923 K, Using Nonlinear Least Squares Method k × 106 at 923 K mol gcat−1 s−1 kPa−(α + β) reaction species C3H8 H2 CO CO2 CH4 C3H6

α

β

NODH

ODH

NODH

ODH

1.12 ± 0.017 1.98 ± 0.015

1.55 ± 0.011 5.46 ± 0.042 6.57 ± 0.071 2.12 ± 2.1 × 10−3 5.1 ± 0.063 1.42 × 10−3 ± 4.2 × 10−5

0.98 ± 0.031 0.69 ± 4.7 × 10−3

0.88 ± 2.7 × 10−3 0.237 ± 4.1 × 10−3 −0.21 ± 2.8 × 10−3 −0.53 ± 4.1 × 10−3 1.13 ± 0.048 1.93 ± 0.022

0.93 ± 1.7 × 10−3 0.14 ± 1.3 × 10−3

0.93 ± 5.5 × 10−3 0.96 ± 3.3 × 10−3

existence of carbon nanofibers along with a large number of whisker-like carbon filaments is expected in this regime as reported by Kepinski and Borowiecki.24 In contrast, the C3H6 production rate was lower during ODH (cf. Figure 6c), suggesting that C3H6 and oxygen were consumed via the C3H6 + 3O2 → 3CO2 + 3H2 reaction, which is also confirmed by a higher H2 formation rate observed in Figure 7a. Interestingly, propane conversion remained constant at almost 95% during ODH at 923 K, while it sharply decreased during NODH, indicating that oxygen during ODH facilitates the gasification of carbon deposited on the Mo−Ni surface leading to better stability compared to NODH. Figure 8a−d shows the effect of temperature on the H2, CH4, C3H6, and C3H8 reaction rates. Hydrogen production rate (cf. Figure 8a) increased with temperature during both ODH and NODH, while ODH has revealed higher H2 formation rate. In contrast, Figure 8c shows that C3H6 production rate is significantly higher during NODH at T > 823 K. The propane consumption rate linearly increases with temperature during NODH, while ODH showed weaker dependence on temperature. Figure 9a,b displays similar plots for CO and CO2 formation rates and the H2/CO product ratio. There was significant production of H2 during ODH, which may be due to a cracking phenomenon on one hand and a rise to a complete O2 conversion on the other hand. During the reaction, H2 formation may be due to

NODH

ODH 0.057 ± 2.1 × 10−4 0.584 ± 1.8 × 10−3 0.91 ± 6.1 × 10−3 0.51 ± 7.2 × 10−3 −0.02 ± 7.3 × 10−4 −1.036 ± 0.024

Cx H1 − x + (x /2)O2 → xCO + [(1 − x)/2]H 2

From these reactions, it is apparent that 1/3 ≤ x ≤ 1 is consistent with the presence of CH2 and CH surface species rather than CH3 during ODH and NODH as also confirmed by Fourier transform infrared (FTIR) studies.17,18 Sinev et al.9 mentioned that nonoxidative dehydrogenation and coke formation mechanisms were responsible for H2 generation as the most predominant pathways in ODH. Furthermore, they suggested that free radicals in catalytic oxidation of light alkanes are the other probable cause of H2 formation.11 Ballarini et al.25 also suggested that the mechanism was the combination of both ODH and NODH, in which the former was the prevalent reaction on the fully oxidized catalyst for a short reduction period and the latter was the more dominant operating mechanism on the reduced catalyst for the longer reduction period. More interestingly, it was reported10 that, based on a decrease in CO selectivity, the additional route for the formation of H2 was the water gas shift (WGS) reaction: CO + H 2O ↔ CO2 + H 2

o ΔH298 = − 41 kJ mol−1

(14)

Figure 10a,b shows the influence of propane and oxygen partial pressures on C3H8 conversion, H2, and C3H6 selectivity at 923 K. An increase of the propane partial pressure (cf. Figure 10a) resulted in a decrease in propane conversion from 91.34 to 83.2%. H2 selectivity showed a slight increase, while that of C3H6 rose significantly from 7.1 to 16.3%. An increase in the oxygen partial pressure (cf. Figure 10b) increased propane conversion from 74 to 95%, while the selectivity to C3H6

C3H8 → (2/x)Cx H1 − x + CH4 + [(3x − 1)/x]H 2

with additional H2 during ODH arising from the oxidation of the surface CxH1−x species via 4023



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Table 6. Arrhenius Treatment for NODH and ODH Ea (kJ mol−1)

pre-exponential factor A × 104 NODH C3H8 H2 CO CO2 CH4 C3H6

2.76 ± 0.051 2.48 ± 0.034

0.64 ± 8.2 × 10−3 30.6 ± 0.71

ODH

NODH −4

0.032 ± 8.3 × 10 0.152 ± 7.7 × 10−3 0.521 ± 6.4 × 10−3 0.078 ± 6.9 × 10−4 3.81 × 10−4 ± 5.2 × 10−6 0.011 ± 2.5 × 10−4

ODH

42.37 ± 0.81 34.83 ± 0.76

25.68 29.57 28.12 2.17 −21.08 33.39

31.88 ± 0.93 77.57 ± 1.02

± ± ± ± ± ±

0.57 0.61 0.76 0.08 0.54 0.71

dropped considerably. These observations are consistent with the study of Barasan et al.8 3.2. Empirical Modeling. The C3H8 dehydrogenation rate along with the formation rates of H2, CH4, and C3H6 corresponding to NODH and H2, CO, CO2, CH4, and C3H6 for ODH are properly fitted to a power-law model: α ( ±ri)NODH = kppropane

i = H 2 , CH4 , C3H6 , C3H8

(15)

α β ( ±ri)NODH = kppropane poxygen

i = H 2 , CO, CO2 , CH4 , C3H6 , C3H8

(16)

in which k is the rate constant, ppropane and poxygen are the partial pressure of reactants in the NODH and ODH reaction, while α and β are the corresponding orders of reactions. Table 5 reports the parameter estimates obtained from the regression of the rate data using the nonlinear least-squares method via (POLYMATH 6.1) software. Both NODH and ODH rates for CH4 and C3H8 exhibited nearly first order dependency on propane partial pressure while the C3H6 rate was essentially second order in propane during ODH with the H2 formation rate showing weaker positive dependency on the same reactant. Interestingly, both CO and CO2 rates were inhibited by propane partial pressure during ODH but were not formed during NODH. However, all reacting species experienced positive (≤1) order reliance on oxygen partial pressure with the exception of C3H6 during ODH. A similar inhibition by oxygen was also reported by Meunier et al.26 Activation energy, Ea (kJ mol−1), and the frequency factor, A (mol gcat−1 s−1 kPa−1), were calculated using the Arrhenius treatment at different temperatures from 773 to 973 K . The results are provided in Table 6. 3.3. Characterization of Spent Catalyst. Under our experimental conditions, carbon deposition was inevitable. To investigate the amount of carbon deposited on the catalyst surface during NODH and ODH, the total organic carbon (TOC) measurement was carried out. Figure 11 gives information about the TOC (%) as a function of propane partial pressure during NODH and ODH (PO2 = 15.19 kPa). As seen, for both NODH and ODH, carbon deposition increases with the propane partial pressure. However, the amount of coke formed during ODH is smaller than that of NODH. The existence of residual carbon on ODH catalysts suggests the high surface capacitance for carbon storage due to rapid propane dehydrogenation and/or that the oxygen partial pressure higher than 15.2 kPa was necessary for complete carbon removal under the reaction conditions. The TOC behavior may be adequately expressed by the sigmoidal logistic equation:

Figure 11. Total organic carbon deposited on Mo−Ni/Al2O3 catalyst during NODH and ODH (PO2 = 15.19 kPa) at 923 K.

a

TOC (%) = 1+

PC3H8 b

( )

(17)

c

where PC3H8 is the partial pressure of propane (kPa), a is the carbon deposition capacitance of the bimetallic Mo−Ni catalyst, b is the critical propane partial pressure, and c is the anticoking factor. Table 7 lists the results obtained from the nonlinear regression. Table 7. Parameter Estimates of Equation 17 reaction

a

b

c

R2 value

NODH ODH

26.25 24.5

−2.71 −2.8

22.72 21.25

0.975 0.961

To explore the nature of carbon deposited on the catalyst during NODH and ODH, TPO-TPR and TPR-TPO experiments were carried out. Figure 12 show the catalyst weight change after 4 h of NODH and ODH reactions at 923 K . It is manifested from Figure 12a that a total deposited carbon burnoff occurred during the first TPO regime, characterized by a sharp drop in the catalyst ̀ weight. However, during ODH, the catalyst experienced a smaller weight change in agreement with the lower coke formation confirmed by TOC analysis. Subsequently, in the TPR step, metal oxide reduction caused a marginal drop in the catalyst weight. On the other hand, the TPR-TPO cycle (cf. Figure 12b) shows that during NODH little or negligible carbon was removed by H2 during TPR, while a slight weight drop was observed during ODH. This was followed by the TPO scheme 4024



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the presence of at least two types of carbon pools, known as Cα (reactive with both O2 and H2) and Cβ (reactive only with O2).

■

AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

■ ■

ACKNOWLEDGMENTS The authors are grateful to the Australian Research Council for financial support.

■

NOMENCLATURE β = heating rate Tp = peak temperature for species desorption Rg = gas constant A sat = adsorbed saturated amount C = desorption rate constant −ΔHd = heat of desorption yi = mole fraction of adsorbed effluent gas F = total outlet volumetric flow rate w = weight of sample k = rate constant ±ri = rate of reaction a = carbon deposition capacitance b = critical propane partial pressure c = anticoking factor REFERENCES

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Figure 12. Carbon gasification deposited on the catalyst using (a) TPO-TPR and (b) TPR-TPO schemes.

in which air was used as a reactant gas to achieve a complete carbon removal. This behavior is indicative of the existence of at least two reactive carbon species, Cα and Cβ. The former, atomic carbon, is reactive with both H2 and O2 (mostly during ODH), while the latter, dehydropolymerized carbon reacts with only O2.

4. CONCLUSIONS Bimetallic Mo−Ni catalyst over an alumina support was prepared and employed in propane nonoxidative (NODH) and oxidative (ODH) dehydrogenation processes. The BET surface area of the catalyst was measured as 106.7 m2 g−1. NH3 and CO2-TPD analysis confirmed the presence of different acidic and basic sites on the catalyst surface. Under our experimental conditions, H2 production rate was higher during ODH while the production rate of C3H6 was significantly lower compared to NODH. Propane conversion remained constant during ODH over 4 h, suggesting that Mo−Ni has good stability. The H2 and C3H6 formation rate increased with temperature during both processes, while H2 formation rate was significantly higher during ODH. Propane conversion slightly decreased with C3H8 partial pressure, while it linearly increased with O2 partial pressure. The behavior observed during NODH and ODH was adequately fitted to a power-law model. TOC analysis confirmed the existence of carbon during NODH and ODH. TPO-TPR and TPR-TPO cycles verified 4025



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Nonoxidative and Oxidative Propane Dehydrogenation over Bimetallic

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