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Article
Oxidative Dehydrogenation of n‑Butane to C4 Olefins Using Lattice
Oxygen of VOx/Ce-meso-Al2O3 under Gas-Phase Oxygen-Free
Conditions
Muhammad Y. Khan, Sagir Adamu, Rahima A. Lucky, Shaikh A. Razzak, and Mohammad M. Hossain*
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ABSTRACT: High-performance, fluidizable VOx/Ce-meso-Al2O3 catalysts were prepared by an excessive solvent approach. The
prepared catalysts were characterized using various physicochemical techniques in order to secure desired properties. XRD, Raman,
and FTIR analyses indicated the presence of amorphous VOx phases on Ce-meso-Al2O3. Nitrogen adsorption isotherm analysis
confirmed a mesoporous framework with a high surface area of the catalysts. H2-TPR reduction showed an active and stable behavior
of the catalysts in repeated reduction and oxidation cycles. The NH3-TPD and NH3 desorption kinetics analysis revealed that the
synthesized catalysts have moderate acidities and low activation energies of NH3 desorption, suggesting weak metal−support
interactions. The VOx/Ce-meso-Al2O3 catalysts were evaluated for n-butane oxidative dehydrogenation (BODH) using a fluidized
CREC riser simulator under gas-phase oxygen-free conditions. The reaction time and reaction temperature were varied between 5
and 25 s and 450−575 °C, respectively. It was found that BODH with 0.2 wt % cerium-doped VOx/meso-Al2O3 catalysts gives the
highest selectivity (62.3%) of C4 olefins with a conversion of 10.6% at 450 °C and 5 s. Furthermore, the fluidizable VOx/Ce-mesoAl2O3 catalyst showed a stable performance over repeated feed injections followed by catalyst regeneration cycles for BODH.
catalysts stand as good candidates for BODH.9,11−16 However,
there are still critical issues related to vanadium catalysts that
need to be addressed. For example, the availability of the
desired amorphous vanadium oxide species is one of the key
aspects for achieving high selectivity of desired C4 olefins.11−16
On the other hand, the support type, vanadium oxide−support
interaction, amount of vanadium oxide loading, and method of
catalyst synthesis play important roles in achieving highly
selective catalysts. It has been shown that high vanadium
loadings may yield a nonselective crystalline V2O5 phase for
BODH.17 As a result, limited VOx loadings (5−10 wt %) are
recommended.18 Similarly, it is reported that the doping of NiMoO4 catalysts with highly basic elements of the IA or IIA
groups tends to increase olefin selectivity.19,20 In this respect, it
is speculated that the metal−support interactions and the
acid−basic nature of the support are key factors affecting C4
olefin selectivity.21 To clarify these matters, nickel oxides and
vanadium oxides have been impregnated on the following
supports: Al2O3, SiO2, MCM-41, MgO, ZrO2, USY, Ti-HMS,
and NaY.12,22−27 Among these materials, magnesium oxide
stands as a good support for VOx-based catalysts.28 However,
given the need of improving mechanical strength properties for
fluidizable catalysts, γ-Al2O3 has been considered as a
1. INTRODUCTION
C4 olefins, which include C4 alkenes (1-butene, iso-butene,
and cis-butene) and 1,3-butadiene, are important feedstocks
for the petrochemical industries. These olefins are processed in
the production of variety of polymers and chemicals such as
styrene rubber, nitrile butadiene rubber, polybutadiene,
polyamides, alkylates, and maleic anhydride.1−3 With the
growing world population and improvements in the quality of
human life, demands for the aforementioned products/
chemicals (so does the C4 olefins) have increased significantly,
which are projected to grow further in the coming years.4
Conventionally, C4 olefins are obtained from various sources,
including (i) dehydrogenation of light alkanes, (ii) fluid
catalytic cracking of heavy oils, and (iii) hydrocracking of
hydrocarbon feedstocks.5 Despite their technical maturity, the
conventional processes experience several disadvantages,6−8
such as (i) high-temperature reaction conditions due to the
highly endothermic nature of involved reactions, (ii)
limitations by thermodynamic constraints and low product
selectivity, (iii) severe catalyst deactivation by coke formation,
and (iv) energy-intensive product separations. In this regard,
C4 olefin production via n-butane oxidative dehydrogenation
(BODH) has been considered to be very promising, given the
possibility of addressing the issues related to the conventional
processes.9−11 Nevertheless, the critical challenge for BODH is
to overcome the low yields of the desired C4 olefin products
by developing more selective BODH catalysts.
Keeping the above into consideration, there are studies in
the open literature investigating the various aspects of BODH
catalysts. Among the studied catalysts, vanadium oxide-based
© 2020 American Chemical Society
Received: January 22, 2020
Revised: April 24, 2020
Published: May 1, 2020
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preferable support for BODH.29−31 Nevertheless, alumina has
a strong acidic character, which leads to low V 2 O 5
dispersions.16 Thus, modification of the γ-Al2O3 support by
doping it with group IIA alkaline elements, such as Mg, Ca, Sr,
and Ba, has been proposed.12 The effect of the addition of
transition metals (Ni, Fe, and Co) on the BiO-γ-Al2O3 support
was also investigated as a catalyst modifier.29−31 It was shown
that the ternary system (including all-metal species Ni, Fe, and
Co) impregnation on the BiO-γ-Al2O3 support improved the
selectivity of butadiene reasonably. An increase in catalyst
activity was linked to the improved basicity of the alumina
support.29 One should, however, notice that this type of γAl2O3 doping remains intrinsically limited given the lack of
catalyst stability.
Furthermore, the selectivity to the desired C4 olefin
products can also be improved by selecting adequate reactor
operating conditions. In this regard, BODH under gas-phase
oxygen-free conditions could be an interesting alternative to
enhance C4 olefin selectivity by suppressing COx formation,
which is similar to the ODH of propane (to propylene) and
ethane (to ethylene), as demonstrated by the present research
group.8,33 This process involves twin interconnected fluidized
bed reactors, (i) a ODH reactor and (ii) catalyst regenerator,
and solid catalyst particles can be circulated between the
fluidized beds. In this approach, the feed n-butane will be
dehydrogenated by reacting with the lattice oxygen of the solid
metal oxide catalyst. The oxygen-depleted catalyst will be sent
to the catalyst regenerator to be reoxidized in contact with air
and recycled back to the ODH reactor.
Given all the abovementioned facts, a new Ce-doped
fluidizable VOx/meso-Ce-Al2O3 catalyst for BODH under
gas-phase oxygen-free conditions has been investigated. The
proposed catalyst is specially tailored for its dual roles: (i)
catalyze the BODH reaction and (ii) act as an oxygen carrier
for the BODH reaction. To accomplish this, cerium is used for
alumina support modification, which upon oxidation, gives a
CeO2 phase. Cerium oxide has outstanding properties to limit
support acidity. Also, ceria tends to increase catalyst thermal
stability and reduces BODH catalyst deactivation. In this
respect, Solsona et al.32 and Khan et al.33 reported improved
catalyst activity of NiO and VOx, respectively, due to the
incorporation of cerium. It is thus proposed that the vanadium
oxide impregnation on cerium-modified alumina in the present
study be prepared using the excessive solvent impregnation
method.34 It is also considered to characterize the developed
BODH catalyst using the following state-of-the-art characterization techniques: XRD, nitrogen adsorption isotherms, TGA,
TPR/TPO, NH3-TPD, FTIR, and Raman as well as NH3
desorption kinetics. Finally, the superior BODH catalyst
performance for n-butane conversion and C4 selectivity
experimental runs were developed in a specially designed
fluidized CREC riser simulator under gas-phase oxygen-free
conditions. This reactor has the special feature of mimicking
the operating conditions of circulating fluidized beds, possible
reactor candidates for this BODH process.
Article
were used without any further purification. Both meso-Al2O3 (MAs)
and Ce-meso-Al2O3 (Ce-MAs) catalyst supports were synthesized
using the facile free method,35,36 which has also been adopted in our
previous work.33,37,38 In the MAs preparation, a 1.0 molar solution of
ammonium carbonate was used to partially hydrolyze 0.1 molar
solution of aluminum nitrate nonahydrate under continuous stirring.
The hydrolysis of the aluminum nitrate solution was continued until a
monolithic gel formation was obtained, which was dried at 30 °C for
24 h. Afterward, the dried gel was heated at a rate of 1 °C/min to 150
and 200 °C for 12 h. Subsequently, MAs crystals were calcined at 300
°C at the same heating rate of 1 °C/min for 12 h. Cerium-doped
alumina supports (Ce-MAs) were synthesized using the same
procedure. However, before the hydrolysis step, an appropriate
amount of cerium nitrate solution (yielding 0.2, 1.0, 3.0, and 5.0 wt %
Ce in MAs) was added to the aluminum salt solution.
The excessive solvent impregnation technique was employed to
impregnate 5 wt % vanadium oxide on all synthesized supports. In a
typical impregnation of VOx, an appropriate weight of vanadium
acetylacetonate salt was dissolved in the excessive volume of toluene
under vigorous stirring. The prepared supports (MAs and Ce-MAs)
were added to this vanadium salt solution and were allowed to mix
under continuous magnetic stirring. After 6 h, the catalyst was
separated from toluene in a centrifuge rotating at 4000 rpm. The
separated catalyst was then heated to 30 °C for 24 h. The dried
sample was then heated (0.5 °C/min) in the furnace at 140 °C for 6
h. The powder catalysts were then reduced in the fluidized conditions
using a gas mixture of H2 and Ar in a vertically mounted Thermocraft
furnace. Subsequently, the reduced catalysts were calcined in the same
furnace at 750 °C for 10 h at a ramping rate of 0.5 °C/min. Table 1
Table 1. Catalyst Designation
VOx/Ce-meso-Al2O3 catalyst composition
no.
wt % V
wt % Ce
wt % meso-Al2O3
nomenclature
1
2
3
4
5
6
0.0
5.0
5.0
5.0
5.0
5.0
0.0
0.0
0.2
1.0
3.0
5.0
100
95
94.8
94
92
90
MAs
0.0 Ce
0.2 Ce
1.0 Ce
3.0 Ce
5.0 Ce
presents the wt % values of different species present in the six different
catalysts prepared in the present study. The wt % values shown in
Table 1 are determined as per the synthesis method. As we can see, all
the cerium-doped catalysts have the same amount of vanadium (5 wt
%) over them; hence, they can be named such that they indicate the
amount of Ce present in them. For the sake of convenience, we have
used the vocabulary of Table 1 throughout this document. For
example, the 0.2 Ce acronym represents the 5.0 wt % VOx/0.2 wt %
Ce-γ-Al2O3 catalyst. The particle size of all the synthesized catalysts
was in a range of 15−110 μm and a Sauter mean diameter of 95 μm.
The sizes of the synthesized catalysts are within the range of group A
of Geldart’s powder classification, which is similar to commercial fluid
catalytic cracking catalysts.
2.2. Catalyst Characterization. The catalysts’ BET (Brunauer−
Emmett−Teller) surface areas and pore properties were measured on
Quantachrome ASIQwin equipment. In the typical experimental
procedure, approximately 150 mg of a catalyst sample was loaded into
a U-tube sample holder, and it was preheated under N2 for 3 h at 350
°C. In the adsorption analysis, the bath temperature was maintained
at 77 K using liquid N2. The specific surface area was determined
using the BET method, and the pore volume and diameter were
derived from the BET isotherm using the Barret−Joyner−Halanda
(BJH) procedure.
X-ray diffraction (XRD) measurements were conducted on a MiniFlex II Rigaku. The equipment is provided with nickel-filtered Cu Kα
radiation with a wavelength =0.15406 nm, 15 mA, and 30 kV. The
2. EXPERIMENTAL SECTION
2.1. Catalyst Synthesis. Cerium nitrate hexahydrate (Ce(NO3)3·
6H2O), aluminum nitrate nonahydrate (Al(NO3)3·9H2O), vanadium
acetylacetonate, toluene, and ammonium carbonate were used in the
synthesis of VOx/Ce-meso-Al2O3 catalysts. The aforementioned
chemicals were purchased from Sigma-Aldrich except for ammonium
carbonate, which was obtained from Fisher Limited. These chemicals
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Figure 1. (a) Pictorial view of the main parts of the riser simulator. (b) Schematic diagram of the CREC riser simulator experimental setup.39
°C min−1. The thermal conductivity detector (TCD) was employed
to analyze the exit gas composition. Similarly, NH3-TPD measurements help to determine the total acidity of the powdered samples. In
all TPD analyses, 100 mg of the calcined catalyst was loaded into the
sample holder and was pretreated at 300 °C using He gas for 3 h (30
cm3 min−1). In the subsequent step, the catalyst was heated to 100 °C
in a gas mixture of 5% NH3/He at a flow rate of 50 cm3 min−1 for 1 h.
The gaseous NH3 molecules were flushed out of the catalyst bed by
purging He for an hour at 50 cm3 min−1. NH3 desorption was studied
up to 750 °C with a heating rate of 10 °C min−1 under a He flow rate
of 50 cm3 min−1. NH3 in the effluent stream was monitored using a
TCD detector.
Raman spectrums of the synthesized catalysts were determined on
a Horiba Raman spectrometer (iHR 320). The apparatus is equipped
with a CCD detector that helps to minimize the elastic beam
scattering. In a typical experimental procedure, the following settings
angle of the incident beams was set between 10 and 90° and a step
size of 0.02 with a scanning rate of 3°/min.
Thermogravimetric analyses (TGA) were performed on a TA SDTQ600. TGA measurements help in determining the thermal stability
of the prepared catalysts. In a typical experimental procedure, 10 mg
of the sample was heated at a rate of 10 °C min−1 to 950 °C with a N2
purging flow rate of 100 cm3 min−1.
The H2-TPR (temperature-programmed reduction) and NH3-TPD
(temperature-programmed desorption) experiments were conducted
on a Micromeritics (AutoChem-II) apparatus. Along with the redox
properties, TPR/TPO analyses also help in evaluating the catalyst
stability. In each experimental run, an equal amount of the catalyst
sample (100 mg) was placed in the quartz U-tube. Before the
reduction step, the samples were pretreated in argon (99.9%) at 300
°C for 3 h (50 cm3 min−1). A gas mixture of 5% H2/Ar (50 cm3
min−1) was used for the reduction at 850 °C at a ramping rate of 10
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were used: excitation λ = 532 nm and laser intensity = 50% at 50−
2500 spectrums. Before the collection of spectrums, each catalyst was
dehydrated for 1 h at 500 °C.
2.3. Reactor System. The n-butane ODH (BODH) reactions
were carried in the fluidized bed CREC riser simulator using only the
lattice oxygen of the catalyst. This mini-batch reactor has a 53 cm3
reaction volume and is specially designed to investigate the ODH
catalyst’s activity and reaction kinetics in a fluidized condition (riser/
downer).39 The reactor is shown in Figure 1, illustrating that the
reactor has two major sections: (a) the upper part and (b) the lower
part. The complete reactor assembly includes the following principal
parts: grids, impellers, heaters, retaining rings, thermocouples, and
cooling jackets. The reactor is connected to the vacuum chamber via
an automatic four-port valve that directs the reaction products to be
collected to the vacuum chamber. An online GC−MS was connected
to the vacuum chamber through a six-port valve to analyze these
reaction effluents.22,40
In all the BODH experiments, 1.0 g of the synthesized catalyst was
loaded into the reactor. Prior to each run, the reactor and vacuum
chamber were purged with Ar gas for 15 min, and a leak test was
conducted. Then, the leakproof reactor’s temperature was increased in
a stepwise manner to a desired reaction temperature value under
continuous flushing with Ar gas to eliminate hot spots and the
atmospheric gas-phase oxygen interface. Upon reaching the temperature set point, the reactor was degassed to 1 atm pressure using a
vacuum pump. Finally, the reaction vessel was disconnected from the
vacuum chamber with the help of an isolation valve. In the following
step, the vacuum chamber pressure was further decreased to 0.26 atm
using a vacuum pump. Before the BODH reactions, the catalyst bed
was fluidized using an impeller rotating at a speed of 5000 rpm. All
BODH experiments were performed in the fluidized bed condition
with no gas-phase oxygenthis is to make sure that only the lattice
oxygen of the catalysts is used. It is possibly important to mention
here that the intense mixing of the small catalyst particles (15−110
μm) and gaseous n-butane (feed) eliminates any possibility of mass
transfer limitations. The BODH reactions were studied in a
temperature range of 450−600 °C, and the resident time was varied
within 5−25 s. The reactor temperature was gradually raised from the
ambient temperature to the desired temperature value under a
continuous flow of Ar. After each experimental run, the spent catalyst
was oxidized with zero air at 575 °C for 10 min, and upon completion
of the regeneration step, the supply of air was stopped. Subsequently,
the reactor was cleaned with argon gas for 15 min. The pure n-butane
(99.97%) is used as a feed in all BODH experiments. All experiments
were performed at a pressure = 1 atm, impeller speed = 4000 rpm
(used for fluidization), and feed-to-catalyst ratio = 1 mL of n-butane/
1.0 g of catalyst. One milliliter or the equivalent of 2.48 mg of pure nbutane was then fed into the reaction chamber via a preloaded airtight syringe.
Upon the completion of the reaction, a valve automatically
connected the reactor with the vacuum chamber. Thus, this allowed
the reaction effluents to be analyzed by an online GC (Agilent
7890A) equipped with FID and TCD detectors. After each BODH
reaction, the used catalyst was reoxidized at 575 °C for 10 min using
zero air. The products of n-butane ODH (BODH) reaction were
identified using an online GC. Further computation with the GC
results enables us to calculate conversion (X) and selectivity (Si) with
the help of the following formulas:
i moles of feed reacted yz
zz × 100
conversion, X (%) = jjj
k moles of n − butane fed {
Article
desorption isotherms of the bare calcined meso-Al2O3 support
and 0.2 Ce and 1.0 Ce catalysts are displayed in Figure 2. The
Figure 2. N2 adsorption isotherms of bare MAs, 0.2 Ce, and 1.0 Ce
samples.
complete nitrogen adsorption isotherm analysis, namely, BET
surface area, BJH pore size, and pore volume, are given in
Table 2 with a cross-correlation coefficient value of >0.99 and
Table 2. BET Surface Area, Pore Volume, and Pore Size of
the Bare Support
catalyst
SBETa (m2/g)
VBJHb (cm3/g)
rBJHc (nm)
MAs
0.0 Ce
0.2 Ce
1.0 Ce
3.0 Ce
5.0 Ce
296
231
152
142
99
88
0.8
0.6
0.4
0.4
0.3
0.2
10.2
9.3
8.4
5.6
4.2
3.6
a
SBET, surface area of the catalyst bVBJH, pore volume crBJH, average
pore width
95% confidence interval. All the analyzed samples showed a
type IV isotherm, assigned to the mesoporous materials that
typically form a monolayer, which is followed by the
multilayers, finally resulting in the capillary condensation that
starts at a relative pressure of ∼0.6. The surface area of the
calcined meso-Al2O3 catalyst support was found to be 296.5
m2g−1, which is within the range of surface area as reported in
the literature35,36 and consistent to the value that the authors
found in previous works.33,34,37 As can be observed in Table 2,
the surface area of the bare MAs decreases from 296.5 to 88.4
m2g−1 for the 5.0 Ce catalyst. Therefore, with an increase in
cerium wt % and an overlayer of vanadium in the catalyst
matrix, the BET surface area, BJH pore width, and BJH pore
volume of all the samples decrease. Thus, more highly doped
catalysts will have a smaller BET surface area. This adverse
effect is due to the breakdown and blockage of the MAs pores,
which are due to the presence of dopant (cerium) and
impregnated overlayer of VOx. The average pore sizes of the
MAs (meso-Al2O3 support) and all other catalysts are in a
range of 2−50 nm, which according to IUPAC nomenclature,
is classified as mesoporous.41 The BET analysis has confirmed
the formation of the high surface area of mesoporous catalysts.
(1)
i moles of product i yz
zzz × 100
selectivity of product i, Si (%) = jjjj
k moles of feed reacted {
(2)
3. RESULTS AND DISCUSSION
3.1. Catalyst Characterizations. 3.1.1. Nitrogen Adsorption Isotherm Analysis. As a reference, the adsorption/
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Article
This could be related to the excellent performance of the
catalysts in ODH reactions.
3.1.2. X-ray Diffraction. The XRD patterns of the prepared
support and catalysts are shown in Figure 3. The XRD patterns
Figure 4. TPR profiles of the freshly prepared catalysts with varying
amounts of dopant wt %: (a) 0.2 Ce, (b) 1.0 Ce, (c) 3.0 Ce, and (d)
5.0 Ce samples.
Figure 3. XRD patterns of MAs, 0.0 Ce, 0.2 Ce, 1.0 Ce, 3.0 Ce, and
5.0 Ce samples.
of CeO2 in γ-Al2O3 also inhibits the transformation of the γAl2O3 phase to α-Al2O3 phase.46
The H2-TPR profile for the calcined 0.0 Ce catalyst (0.0 wt
% Ce, 5.0 wt % V, and 95.0 wt % γ-Al2O3) showed only one
asymmetric peak extending from 270 to 475 °C. The reduction
observed in the preceding temperature range is assigned to the
monomeric and polymeric reduction of the amorphous VOx
with the oxidation states of V+5 and V+4, respectively.8 A
similar observation of VOx reduction in a temperature range of
280−500 °C was also reported by Reddy and Varma,47
Martinezhuerta et al.,48 and our previous articles.33,47,48 The
cerium-doped 0.2 Ce, 1.0 Ce, 3.0 Ce, and 5.0 Ce catalysts also
displayed low-temperature H2-uptake peaks. However, due to
the presence of both VOx and CeO2 phases (as seen in XRD),
the TPR peak at low temperatures is assigned to the reduction
of both VOx and CeO2 phases. The contribution of CeO2 in
the consumption of H2 is because of the surface reduction of
the CeO2 phase having an oxidation state of Ce+4, as also
observed by Kaŝpar et al.49 Furthermore, the temperature of
the maximal H2-TPR consumption (Tmax1) has increased with
the amount of cerium (dopant) in the support. The increase in
reduction temperature is believed to be due to the formation of
stable monomeric surface species.50 The TGA analysis also
exhibited a similar trend where higher-cerium-containing (3.0
Ce and 5.0 Ce) catalysts showed more excellent thermal
stability.51 Furthermore, in XRD characterizations, 3.0 Ce and
5.0 Ce catalysts showed intense diffraction peaks for the CeO2
phase.
The reduction at high temperatures with the maximum at
800 °C is solely attributed to the bulk reduction of the CeO2
phase, which is only noticeable in 3.0 Ce and 5.0 Ce catalysts,
as presented in Figure 4. The catalyst reduction at high
temperatures shows a strong metal−support interaction.51,52
This high-temperature reduction peak is absent in both 0.2 Ce
and 1.0 Ce catalysts due to a very low Ce concentration
(CeO2) in these catalysts.33 However, as the loading of cerium
increases, so does the formation of the crystalline CeO2 phase,
as previously confirmed by XRD analysis. The increase in
CeO2 formation results in the high-temperature H2 consumption peak.
of the bare MAs exhibit peaks at 2θ = 37.0°, 45.5°, and 67.0°,
corresponding to mesoporous γ-Al2O3 (JCPDS 10-0425).35
The modification of MAs with different cerium wt % values
leads to a change of intensity and shape of XRD peaks of γAl2O3, except for the 0.2 Ce catalyst. As evident in Figure 3, no
diffraction lines for the crystalline CeO2 phase were noticed for
the 0.2 Ce catalyst (0.2 wt % cerium-doped catalyst),
suggesting the amorphous CeO2 phase formation at a low
loading of cerium. With increasing the cerium loading beyond
0.2 wt %, the intensity of the cerium oxide diffraction lines also
increased. According to JCPDS 34-0394, the diffraction lines at
2θ = 28°, 33°, 47.2°, 56.1°, 58.5°, 76.5°, and 79.0° are the
characteristic spectrum of the cubic crystal structure of the
CeO2 phase.42 At a higher loading of cerium (>3.0 wt %) in
the catalyst, the diffraction peaks corresponding to the ceria
phase become well resolved and sharper. The presence of the
intense peaks of CeO2 at higher loadings, i.e., 1.0 Ce, 3.0 Ce,
and 5.0 Ce, decreases the intensity of the main lines of the γAl2O3 support at 2θ = 37.0° and 45.5°.43 Meanwhile, no
diffraction lines were observed for crystalline VOx (besides
MAs, every catalyst is impregnated by 5 wt % vanadium; Table
1), suggesting either crystalline VOx particles are very small
(size of ≤3 nm) and highly dispersed or there is amorphous
VOx phase formation.8,44,45 This will be further discussed in
section 3.1.6 of this article (Raman spectroscopy). Finally, no
diffraction peaks were observed for AlV3O9, which means that
the prepared catalysts are free of solid solution formation
between VOx and MAs, or we can say that there is no solid
reaction between the VOx and MAs.
3.1.3. Reduction of Catalysts by H2-TPR. The H2-TPR
profiles of the calcined samples with different loadings of
cerium (0, 1, 2, 3, and 5 wt %) are depicted in Figure 4.
According to the TPR peaks of the catalysts in Figure 4, 0.0
Ce, 0.2 Ce, and 1.0 Ce catalysts showed only one broad lowtemperature H2 reduction peak in a temperature range of 270−
475 °C, while 3.0 Ce and 5.0 Ce catalysts also showed a hightemperature H2 uptake peak centering at 800 °C. The presence
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The detailed parameters of the maximum reduction
temperature and total H2 consumption are illustrated in
Table 3. The low-temperature H2 consumption peaks were
Table 3. H2 Consumption of Calcined Catalysts with
Different wt % Values of Cerium (0.0, 0.2, 1.0, 3.0, and 5.0)
low-temperature
samples reduction (Tmax1, °C)
0.0
0.2
1.0
3.0
5.0
Ce
Ce
Ce
Ce
Ce
424.0
430.0
456.5
463.6
473.8
high-temperature
reduction (Tmax2, °C)
total H2
consumption
(cm3 g−1)
796.5
797.0
39.2
35.8
44.3
45.0
48.37
located between 270 and 475 °C, while the higher-temperature
reduction peaks were centered at ∼802 °C. The total H2
consumption for all the synthesized catalysts is the summation
of reductions by VOx and CeO2 phases except for the 0.0 Ce
catalyst. The complete H2 oxidation of 0.0 Ce, 0.2 Ce, and 1.0
Ce is almost comparable due to the presence of the same
loading of VOx (5 wt %) on each cerium-modified support.
However, the highly doped catalysts (3.0 Ce and 5.0 Ce) also
result in the secondary high-temperature reduction peak, which
contributes to more H2 uptake. As previously discussed, H2
consumption at this temperature is due to the bulk reduction
of the CeO2 phase, and from a BODH perspective, the high
percentage of CeO2 is generally unfavorable as it is reported to
promote the complete combustion reactions (CO and CO2).53
The total H2 consumption values given in Table 3 are the
average of the five repeated reduction cycles. All the catalysts
showed a higher H2 consumption during the first cycle of the
reduction, and this is because of a decrease in nitrate species
on the freshly prepared catalysts that tends to increase the H2
consumption. However, with the completion of the first cycle
of reduction, the surface nitrate species are eliminated; thus,
the H2 uptake becomes constant in the rest of the TPR cycles.
Thus, one can predict that 5.0 Ce will be more reactive (will
give the highest HC conversion), but because of the presence
of a high percentage of easily reducible CeO2 species, it may
not be selective for olefin production.
3.1.4. NH3-TPD Analysis. It is generally believed that olefin
selectivity in the ODH reaction mainly depends on the overall
acidity of the catalysts.54,55 NH3-TPD was thus employed to
calculate the total acidity and metal−support interactions. NH3
is widely used for the determination of the aforementioned
properties because of its smaller size and a wide range of
temperature applications.56
The effect of cerium doping on the overall acidity of the
VOx/γ-Al2O3 catalysts can be viewed via NH3-TPD profiles,
displayed in Figure 5. The typical values of maximal desorption
temperatures and total acidity of all samples are summarized in
Table 4. The acidic sites are classified into three main
categories (weak, medium, and strong) based on the maximal
desorption temperature.57,58 All the tested samples showed an
initial broad low-temperature desorption peak at around 195
°C, which is attributed to NH3 desorption from weak acid
sites, whereas a high-temperature desorption peak at ∼544 °C
indicates the NH3 adsorbed on acidic sites with strong acidic
strength.
The synthesized γ-Al2O3 support showed a total acidity of
11.4 cm3 g−1, while the acidity of the unmodified VOx/γ-Al2O3
(0.0% Ce) catalyst was found to be 13.1 cm3 g−1, which is due
Figure 5. NH3-TPD profiles of 0.2 Ce, 1.0 Ce, 3.0 Ce, and 5.0 Ce.
Table 4. NH3-TPD Analyses of MAs and 0.0 Ce, 0.2 Ce, 1.0
Ce, 3.0 Ce, and 5.0 Ce
total acidity
catalysts (cm3 NH3/g cat)
γ-Al2O3
0.0 Ce
0.2 Ce
1.0 Ce
3.0 Ce
5.0 Ce
11.4
13.1
7.6
8.1
10.1
10.7
maximum desorption
temperature for weak
acid sites (°C)
maximum desorption
temperature for
strong acid sites (°C)
200
196
200
196
193
192
395
550
547
543
542
to the acidic nature VOx (Table 4). However, the total acidity
of cerium-modified catalysts (0.2 Ce, 1.0 Ce, 3.0 Ce, and 5.0
Ce) was found to be lower than those of both bare γ-Al2O3 and
VOx/γ-Al2O3. A similar trend of increasing acidity of the
alumina support upon impregnation of VOx has previously
been reported by Al-Ghamdi and de Lasa8 and Datka et al..59
The total acidity of the 0.2 Ce catalyst is 7.64 cm3 g−1; in
comparison with the 0.0 Ce catalyst, the presence of cerium in
the catalyst has significantly reduced the overall acidity of the
VOx/Ce-γAl2O3 catalyst. However, a further increase in the
loading of cerium to 1.0 wt % slightly increases the total
acidity, but still the acidity is lower than those of the bare
support (γ-Al2O3) and cerium-free VOx/γAl2O3 catalyst (0.0
Ce). The slight increase in acidity upon addition of cerium is
assigned to the synergetic effect of both CeO2 and VO2 phases,
promoting more strong acidic sites in comparison to weak
acidic sites, which is reflected by the significant shift of the
high-temperature peak (also the area under the curve). This is
in agreement with our previous work33 and was also observed
by Wu et al.60 and Lee et al.,61 suggesting that a higher loading
of cerium creates new strong acidic sites on the support’s
surface and increases the overall acidity.
3.1.5. NH3-TPD Kinetics. The activation energy (Edes) and
the pre-exponential factor (kdesO) of NH3 desorption were
determined by fitting the NH3-TPD data into the kinetics
model. The knowledge of Edes and kdesO is essential to
understand the metal−support interaction and ease of product
(C4 olefin) desorption from the catalyst pores, respectively.
The estimated kinetics parameters are applicable under the
following assumptions: the single activation energy of
desorption (Arrhenius equation) and Edes are independent of
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abundant amount of lattice oxygen,45 making them potentially
more active for the reactions (higher conversion in BODH).
However, at a much higher loading of cerium, the catalysts will
become less selective toward BODH reactions due to more
loosely bound lattice oxygen and higher acidity.66 On the other
hand, the catalysts having less cerium wt % (0.2 Ce and 1.0
Ce) require greater energy to desorb NH3, showing relatively
stronger metal−support interactions than the one having a
higher cerium wt % (3.0 Ce and 5.0 Ce). These make them, in
turn, less reactive but highly selective for the BODH reaction
(higher olefin selectivity). Sedor et al. have also demonstrated
a similar trend for the Ni/Al2O3 catalysts, which is the fact that
a catalyst having a small amount of metal oxide results in
higher activation energy, which corresponds to a strong metal−
support interaction.40
It must be pointed out that these TPD kinetics modeling
results agree with the TPR characterization, which showed
higher reactivity for the catalysts with more cerium wt %,
resulting in a higher H2 uptake (see Table 3). XRD
characterization also agrees with the NH3-TPD results by
presenting sharp peaks of CeO2 for 3.0 Ce and 5.0 Ce catalysts
(see Figure 3). According to Trovarelli, the higher concentration of CeO2 wt % in the catalyst, the greater the availability
of the lattice oxygen.46
3.1.6. Raman Spectroscopy. Figure 6 reports the Raman
spectrum of the synthesized catalyst in the range of 100−1200
the surface coverage of adsorbates; there is a uniform
distribution of NH3 concentration over the catalyst bed;
NH3 desorption from the catalyst surface is governed by the
second-order rate process; the desorption process is irreversible (no readsorption); molecular transport and convection
transport resistances are negligible; and the catalyst bed
temperature increases linearly with time.60,62 NH3-TPD
experiments were performed by considering the aforementioned assumptions. The desorption constant (Kd) is given by
Arrhenius’ eq 3. By performing an ammonia balance across the
catalyst bed and after mathematical simplifications, the rate of
desorption (rdes) is given by eq 4
i −E y
Kd = kdesO expjjj des zzz
k RT {
(3)
ij −E ij 1
i dθ y
1 yzyz
n
rdes = −Vmjjjj des zzzz = kdesO θdes
expjjjj des jjj − zzzzzzz
j
Tn z{{
k dT {
k R kT
and
(4)
The change in volume of desorbed NH3 with respect to
temperature can be obtained as63,64
2
ij E ij 1
kdesO ijj
Vdes yzz
ij dVdes yz
1 yzyz
jj
zz =
jjj1 −
zzz expjjjj− des jjjj − zzzzzzzz
β k
Vm {
Tn {{
k dT {
k R kT
Article
(5)
where θdes, Vm, Tn, n, and β represent the catalyst surface
coverage, monolayer volume, centering temperature, order of
desorption process, and heating rate, respectively. Equation 5
was fitted to experimental data by a NonlinearModelFit builtin function in Mathematica. All NH3-TPD data were obtained
using a catalyst mass = 0.1 g and β = 10 °C min−1. The
desorption parameters given in Table 5 were estimated with a
correlation constant value (R2) > 0.99, confidence interval =
95%, degree of freedom = 3769, and large negative Akaike
Information Criterion (AIC) value.
Table 5. Estimated Kinetics Parameters at a Heating Rate of
β = 10 °C min−1
samples
0.0
0.2
1.0
3.0
5.0
Ce
Ce
Ce
Ce
Ce
Edes (kJ/mol)
R2
AIC
Vdes (mL/g cat)
±
±
±
±
±
0.99
0.99
0.99
0.99
0.99
−46,361
−45,379
−41,149
−44,664
−38,882
13.1
7.6
8.1
10.1
10.7
28.0
19.7
14.3
11.6
11.3
0.1
0.1
0.1
0.1
0.1
Figure 6. Raman spectrums of MAs, 0.0 Ce, 0.2 Ce, 3.0 Ce, and 5.0
Ce catalysts.
cm−1. The absence of Raman peaks for MAs is because of the
ionic characteristic of the Al−O bond.67 The VOx spectrum
not only is a function of the vanadium loading but also
depends on the type and concentration of the surface species.
In the 0.2 Ce catalyst, the only weak band is detected at 1130
cm−1. This spectrum of VOx at 1130 cm−1 is almost invisible in
the 0.0 Ce catalyst. The transmittance in the 1130 cm−1 region
is assigned to the widely spread surface species of
monovanadate with an isolated tetrahedral geometry. The
3.0 Ce and 5.0 Ce catalysts also exhibit a similar Raman band
at 1130 cm−1 but with a much higher intensity of the line. The
5.0 Ce catalyst also shows the second Raman band in 950 cm−1
due to the V−O−V bond, suggesting crystalline V2O5 phase
formation. However, no XRD lines were detected for V2O5
crystals, possibly due to their tiny sizes. Thus, the 5.0 Ce
catalyst showed the presence of both polyvanadate and
monovanadate surface species.17 Raman spectroscopic analysis
In Table 5, lower desorption energies for all cerium-doped
catalysts (0.2 Ce, 1.0 Ce, 3.0 Ce, and 5.0 Ce) are noticed as
compared with the cerium-free catalyst (0.0 Ce). Furthermore,
the calculated desorption energies are also lower than those
reported in the literature by Bakare et al. for VOx/γ-Al2O3
catalysts.45 The lower desorption energies mean that less
energy is required to desorb the absorbed NH3 molecules from
the catalyst’s pores, suggesting weaker metal−support interactions. The estimated desorption energies in the present study
lie between the range reported in the literature.45,64,65 In
comparison to the 0.0 Ce catalyst, cerium-doped catalysts
generally will be more selective for the BODH because it will
easily desorb the olefins. The catalysts with more cerium wt %
(3.0 Ce and 5.0 Ce) relatively require the least activation
energy for desorption, which in turn corresponds to the
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Article
is in complete harmony with XRD analysis (see Figure 3) and
H2-TPR experiments (see Table 3). The increase in cerium
loading in the catalysts reduces the dispersion of the VOx that
promotes the formation of the VOx crystalline phase. The
Raman spectrum for the CeO2 phase appeared at 460 cm−1
wavenumbers, and the intensity of the CeO2 line increases with
cerium wt % in the catalyst; a good resolve peak can be easily
seen in the spectra of 3.0 Ce and 5.0 Ce catalysts. The band at
460 cm−1 corresponds to the CeO2 phase with a fluorite
structure, which was also confirmed by the XRD characterization (see Figure 3).68
3.2. ODH in the CREC Riser Simulator. Prior to the
actual BODH experiments, different thermal runs were
conducted to determine the possible contribution of pyrolysis
to the total yield of desired products. These blank runs were
performed on the empty reactor at 10 s residence time and
within the temperature range of 450−600 °C. Figure 7
Figure 8. Effect of cerium loading on n-butane conversion, C4 olefin
selectivity (including all the isomers, iso-butylene, 1-butene, 2-butene,
etc.), and COx selectivity at 500 °C (standard deviation within ±2.5%
of replicate experiments (w = 1.0 g, feed = 1 mL, t = 10 s)).
the 0.2 Ce catalyst (50.2%). The lower C4 olefin selectivity
over the 0.0 Ce (5.0 wt % VOx/γ-Al2O3) catalyst is mainly due
to its higher acidity, which promotes cracking reactions. On
the contrary, the 0.2 Ce catalyst had previously shown the least
acidic nature (see Table 4), so it minimizes the cracking
reactions and, in return, promotes C4 olefin selectivity. That is
why a higher C4 olefin yield is observed over the 0.2 Ce
catalyst (7.7%) as compared with the 0.0 Ce catalyst (6.2%).
Among the Ce-promoted VOx/Ce-γ-Al2O3 catalysts, the
highest C4 olefin selectivity of 50.2% with 15.4% conversion is
obtained by the 0.2 Ce catalyst (0.2 wt % cerium-doped VOx/
γ-Al2O3). With a further increase in cerium content, there is
both an augmented n-butane conversion and COx selectivity.
One can also notice a trend toward lower C4 olefin selectivity
at a higher amount of cerium-containing catalysts. For
example, the 5.0 Ce catalyst showed the lowest C4 olefin
selectivity of 30.1% with the highest n-butane conversion of
22%.
Figure 8 results can be supported using both TPD and TPR
characterization data. TPR demonstrates that higher CeO2
leads to more reducible species on the surface of the catalyst
and higher lattice oxygen availability (as shown in Table 3 with
the higher H2 uptake). The higher H2 consumption means
more reactivity. Also, TPD provides lower desorption energies,
suggesting less metal−support interactions, making them more
reactive for n-butane conversion. That is why a higher
conversion is observed on 5.0 Ce and 3.0 Ce catalysts as
compared to 0.2 Ce and 1.0 Ce catalysts. Raman and XRD
analyses revealed a greater concentration of the CeO2 phase on
5.0 Ce and 3.0 Ce catalysts. The contact of these catalysts with
n-butane results in the formation of more COx because of the
easily reducible CeO 2 phase, which favors complete
combustion reactions. Similarly, trends of lower alkane
selectivity with higher loadings of cerium have also been
reported by Martin et al.,19 Maldonado-Hodar et al.,20 and Xu
et al.12 The higher-cerium-containing catalysts have revealed
higher acidity (see TPD Table 4), which favors the cracking
reactions and gives <C4 hydrocarbons (HC). Once <C4 HC
have formed, these degraded fractions react in an unselective
manner with the lattice oxygen and result in the formation of
Figure 7. n-Butane conversion vs temperature for thermal blank runs
(standard deviation within ±2.5% of replicate experiments (w = 1.0 g,
feed = 1 mL, t = 10 s)).
illustrates a direct correlation between n-butane conversion
and reaction temperature, provided that the reaction is
performed anaerobically. The lower conversion values
(<3.5%) depicted in Figure 7 suggest that pyrolysis of nbutane would have a negligible effect on BODH reactions. In
particular, also for a selected temperature of 450 °C and
reaction time of 10 s, it was noticed that the thermal n-butane
conversion was limited to 1.2%. Therefore, based on the
thermal blank runs, n-butane conversion in the BODH
reactions will solely be attributed the VOx/Ce-MAs catalysts.
Figure 8 reports the effect of cerium loading on VOx/Ce-γAl2O3 performance under anaerobic conditions at 500 °C and
10 s. It can be observed that the 0.0 Ce (VOx/γ-Al2O3) catalyst
gives a 17.4% n-butane conversion with a 35.5% C4 olefin
selectivity (including all the isomers, iso-butylene, 1-butene, 2butene, etc.) and 42% COx selectivity. Coke was not
considered in selectivity calculation because of its minimal
concentration. The availability of oxygen in the catalyst helps
to minimize coke formation.45 It appears that the 0.0 Ce
catalyst shows a slightly higher n-butane conversion (2%) than
the 0.2 Ce catalyst, which is due to a higher BET surface area
of this catalyst (230.8 m2 g−1). However, C4 olefin selectivity
of the 0.0 Ce catalyst is significantly lower (35.5%) than that of
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more COx.12,19,20 The indirect formation of COx on 5.0 Ce
and 3.0 Ce catalysts also contributes to higher COx selectivity.
On the other hand, the catalysts containing lesser cerium
(0.2 Ce and 1.0 Ce catalysts) have strong metal−support
interactions, as shown by the small uptake of H2 and higher
desorption energy (see TPD kinetics Table 5), suggesting
lower conversion due to more compactness for lattice
oxygen.8,9,17 However, higher C4 olefin production on 0.2
Ce and 1.0 Ce catalysts is again due to lower acidity (less
formation of <C4 HC) and more selective surface oxide
species.66 Based on the highest C4 olefin yield/selectivity
observed on the 0.2 Ce catalyst, as revealed in Figure 8, only
the 0.2 Ce catalyst will be investigated in the subsequent part
of the present study.
The typical product distribution for the BODH over the two
extreme conditions of the temperature (450 and 600 °C) on
the 0.2 Ce catalyst is shown in Figure 9. The BODH products
Article
Owens, Owen et al., Patel and Andersen, and Armendariz et
al.16,66,69,70 Therefore, the elevated reactor’s temperature
makes both n-butane and C4 olefins more vulnerable to
form cracking products (<C4 hydrocarbons), which upon
complete oxidation, result in COx.71
The performance of the 0.2 Ce catalyst is further assessed by
studying the effect of residence time on the selectivity of C4
olefins, keeping constant reaction temperatures (450 °C and
550 °C), as shown by Figure 10. A similar trend is noted, as
Figure 10. C4 olefin selectivity and n-butane conversion vs reaction
time at various reaction temperatures (deviation within ±2.5% of
replicate experiments; w = 1.0 g, feed = 1 mL).
previously represented by the effect of temperature on product
selectivity (see Figure 9). Although C4 olefins are the main
reaction product, their selectivity decreased significantly with
an increase in contact time. At a reaction time of 5 s and 450
°C, the maximum selectivity of 62.4% is obtained for C4
olefins; whereas, the 25 s and 550 °C reaction conditions give
the lowest C4 olefin selectivity of 28.1%. We can conclude that
long residence times favor conversion of the feed; that is, the
longer contact time of the hydrocarbons with the lattice
oxygen of the catalyst, the higher the concentration of
byproducts (COx and <C4 hydrocarbons). A similar trend is
also reported by Volpe et al. and Cavani et al. that at long
residence times, there is more possibility of direct and indirect
complete combustion of n-butane and C4 olefins, respectively.22,72 It is obvious that the selectivity of COx follows the
reverse trend (higher concentration COx at higher residence
time), as compared with the selectivity of C4 olefins. It was
found in the present study that elevated reaction temperatures
and longer residence contact times favor the COx formation.
As mentioned before, this is because of an increase in the rate
of (1) direct complete oxidation of n-butane, (2) combustion
of formed C4 olefins to COx, and (3) indirect formation of
COx by complete oxidation of cracked products (<C4 HC)
with the unselective lattice oxygen of the catalyst.
Figure 11 illustrates four successive BODH reactions
performed at the same operating conditions (residence time
= 5 s, temperature = 450 °C, and catalyst = 0.2 Ce). As in
BODH reactions, only the lattice oxygen of the catalyst is
consumed; therefore, after each experimental run, the spent
catalyst was regenerated (oxidized in air at 575 °C for 10 min),
which is used for subsequent BODH runs. These experiments
are important as they demonstrate the reproducibility of results
and, hence, the consistent nature of catalysts. Previously, the
Figure 9. Typical product distribution of BODH reaction at 450 and
600 °C over 0.2 Ce (standard deviation within ±2.5% of replicate
experiments (w = 1.0 g, feed = 1 mL, t = 10 s)).
include methane, ethane, ethylene, propane, propylene, COx,
and C4 olefins, while it is clear that C4 olefins and COx (CO +
CO2) are the main reaction products. Figure 9 illustrates the
negative effect of reaction temperature on the selectivity of C4
olefins. This is a typical trend in BODH reactions, as stated by
Madeira and Portela that low temperatures result in higher
olefin selectivity.6 We have also noticed a similar behavior in
which a low temperature corresponds to a lower feed
conversion but results in higher selectivity of C4 olefins.
Therefore, we can conclude, as said in previous literature, that
a low reaction temperature favors olefin formation over the 0.2
Ce catalyst. However, at high reaction temperatures, exactly an
opposite trend was observed. An elevated reaction temperature
supports the feed conversion at the expense of olefin
selectivity. As seen in Figure 9, the higher n-butane conversion
at 600 °C contributes more toward COx and cracking product
(<C4 hydrocarbons) formation, indicating that higher reaction
temperatures make catalysts more selective for unwanted
products (COx and <C4 hydrocarbons).9 The reason for
higher conversion is because of the more activation of C−H
bonds of n-butane and C4 olefins, which results in a higher
concentration of COx and degradation products in the effluent
stream. This behavior has also been reported by Kung and
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520 °C displaying a 56% C4 olefin selectivity at a 5% n-butane
conversion in a fixed-bed reactor under a gas-phase oxygen-free
environment.27 Comparable selectivity results were achieved
by Lemonidou et al.,73Rubio et al.,75 Dejoz et al. using
magnesium-based catalysts.76
It is however argued that in order to achieve a stable BODH
catalytic process, its implementation in circulating fluidized
beds requires a limitation on the catalyst flow. Table 5 shows
that the VOx/Ce-γAl2O3 (0.2 Ce) catalyst of the present study
achieves at 450 °C a 62.4% C4 olefin selectivity at a 10.7% nbutane conversion. The promising performance is accomplished under gas-phase oxygen-free conditions and low coke
formation (less than 1.3 wt %). Thus, this BODH catalyst
displays great stability under repeated n-butane injections. This
commendable activity of BODH at 450 °C, confirmed using
TPR analysis, can be assigned to the high catalyst specific
surface area and high density of active sites. As a result, this
BODH catalyst appears to be very suitable for continuous
circulating fluidized operation at 450 °C, with minimal catalyst
regeneration in between BODH cycles.
Figure 11. C4 olefin selectivity and n-butane conversion for
successive BODH with interstage catalyst reoxidation over 0.2 Ce
(standard deviation within ±2.5% of replicate experiments (w = 1.0 g,
feed = 1 mL, T = 450 °C, t = 5 s)).
4. CONCLUSIONS
The effect of cerium doping on fluidizable VOx/γ-Al2O3
catalysts was investigated both using surface science characterization and the gas-phase oxygen-free BODH in a CREC riser
simulator batch reactor. The following are the main findings:
(i) BET analysis revealed a mesoporous framework with a
high specific surface area. An inverse relationship
between the specific surface area and cerium content
was also noticed.
(ii) XRD and Raman analyses show no peaks for the
crystalline V2O5 phase, confirming the formation of a
highly dispersed amorphous phase. For cerium, a
crystalline CeO2 phase was confirmed in the 3.0 and
5.0 Ce-doped catalysts.
(iii) TPR/TPO characterization displayed low- and hightemperature H2 reduction peaks for all the prepared
catalysts, while the cyclic TPR/TPO and TGA showed a
stable catalytic performance and thermal stability for the
synthesized catalysts, respectively.
(iv) The coexistence of both strong and weak acidic sites has
been observed during NH3-TPD characterization. The
NH3-TPD kinetics studies have shown a relatively weak
metal−support interaction for the synthesized catalysts
with activation energy of ammonia desorption following
the subsequent trend: 5.0 Ce < 3.0 Ce < 1.0 Ce < 0.2 Ce
< 0.0 Ce.
(v) Among the synthesized catalysts, the 0.2 Ce catalyst
showed a maximum selectivity of 62.4% (C4 olefin) at
450 °C and 5 s. Furthermore, the best catalytic
performance was observed at low reaction temperatures
and shorter residence times.
(vi) The cerium-modified VOx/γ-Al2O3 displayed very low
coke levels at 450 °C and 5 s reaction time, making it
very adequate for its applications in continuous
circulating fluidized bed operation.
repeated TPR/TPO characterizations of the catalysts have also
indicated the stable performance of the catalysts (see Table 3).
Table 6 reports the comparison of the performance of the
BODH catalyst of the current work with others reported in the
Table 6. Comparison of BODH Catalyst Performance
catalyst
VOx/CeγAl2O3
VOx/MCM41
V2O5/MgOAl2O3
V2O5/MgOZrO2
VOx/USY
VOx/USY
Xa
(%)
Tb
(°C)
tc
Sd
(%)
reactor
system
10.7
450
5s
62.4
47.4
550
1h
57
fluidized
bed
fixed bed
Wang et al.25
30.3
600
64.3
fixed bed
Xu et al.12
32.9
500
6h
43.1
fixed bed
Lee et al.74
5
520
4
56
fixed bed
Garcia et al.27
68
fixed bed
Volpe et al.22
Lemonidou et
al.73
Rubio et al.75
Dejoz et al.76
min
8.2
520
4
min
V2O5/MgO
29.5
520
54
fixed bed
V2O5/MgO
MoO3V2O5/
MgO
31.8
24.2
500
550
55.8
69.5
fixed bed
fixed bed
Article
reference
this work
a
X, n-butane conversion bT, reaction temperature ct, reaction time dS,
selectivity of C4 olefins
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a fixed-bed reactor.25 Similarly, Xu et al. have revealed 30.3% nbutane conversion with 64.3% C4 olefin selectivity using a 5%V2O5/MgO-Al2O3 catalyst at 600 °C in a fixed-bed reactor.12
In a separate study by Volpe et al., VOx as an active metal was
impregnated on the surface of USY, α-Al2O3, NaY, and γ-Al2O3
catalyst supports. Among these catalysts, VOx/USY showed
the best catalytic activity of 68% C4 alkene selectivity with a
8.2% n-butane conversion at 520 °C. In their analysis, gasphase O2-free conditions in a fixed-bed reactor were used.22
Furthermore, Garcia et al. also reported a VOx/USY catalyst at
■
AUTHOR INFORMATION
Corresponding Author
Mohammad M. Hossain − Department of Chemical
Engineering and Center of Research Excellence in
Nanotechnology, King Fahd University of Petroleum and
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Hossain, M. M. Kinetics of oxidative cracking of n-hexane to olefins
Minerals, Dhahran 31261, Saudi Arabia; orcid.org/00000002-7780-5910; Phone: +966-13-860-1478;
Email: mhossain@kfupm.edu.sa; Fax: +966-13-860-4234
Authors
Muhammad Y. Khan − Department of Chemical Engineering,
King Fahd University of Petroleum and Minerals, Dhahran
31261, Saudi Arabia
Sagir Adamu − Department of Chemical Engineering, King Fahd
University of Petroleum and Minerals, Dhahran 31261, Saudi
Arabia
Rahima A. Lucky − Department of Chemical and Biochemical
Engineering, University of Western Ontario, London, Ontario
N6A 3K7, Canada
Shaikh A. Razzak − Department of Chemical Engineering, King
Fahd University of Petroleum and Minerals, Dhahran 31261,
Saudi Arabia
Complete contact information is available at:
https://pubs.acs.org/10.1021/acs.energyfuels.0c00220
Notes
The authors declare no competing financial interest.
■
ACKNOWLEDGMENTS
The author(s) would like to acknowledge the support provided
by the Deanship of Scientific Research (DSR) at King Fahd
University of Petroleum & Minerals (KFUPM) for funding this
work through project no. IN161022.
■
Article
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