Review pubs.acs.org/CR
Catalytic Dehydrogenation of Light Alkanes on Metals and Metal Oxides Jesper J. H. B. Sattler, Javier Ruiz-Martinez, Eduardo Santillan-Jimenez, and Bert M. Weckhuysen* Inorganic Chemistry and Catalysis, Debye Institute of Nanomaterials, Utrecht University, Universiteitsweg 99, 3584 CG Utrecht, The Netherlands 1. INTRODUCTION 1.1. Setting the Scene
Because of their extensive use as chemical building blocks, light olefins, such as propylene (propene) and ethylene (ethene), are among the most important types of compounds in the chemical industry. Light olefins are the feedstock employed in the production of a vast array of chemicals, including polymers (e.g., polyethylene and polypropylene), oxygenates (e.g., ethylene glycol, acetaldehyde, acetone, and propylene oxide), and important chemical intermediates (e.g., ethylbenzene and propionaldehyde).1−4 The demand of these building blocks has increased steadily over the last years, all major markets registering a positive growth in demand and production in recent years with the exception of the Eurozone, due to the effects of the recent economic downturn.5 In Figure 1, the uses as well as the recent and predicted future growth of the olefin supply and demand are shown.6 Steam cracking and fluid catalytic cracking (FCC) of naphtha, light diesel, and other oil byproducts are the most common methods for obtaining these light olefins. For example, in 2007 less than 3% of propylene was produced by on-purpose techniques, while FCC and steam cracking accounted for the bulk of the propylene produced.7 A number of factors including the high-energy demands of these processes, their low selectivity toward the production of particular olefins, dwindling petroleum reserves, and rising oil prices are driving the petrochemical industry to search for a more economical feedstock and more efficient conversion technologies. In recent years, hydraulic fracturing or “fracking” technologies have improved to the point where large volumes of shale gas can be extracted in a cost-effective manner. In fact, the United States already obtains one-quarter of their natural gas production from shale gas deposits, and this is expected to rise even further in the coming years due to the emphasis that is being given to the attainment of energy independence.8,9 This increased supply of natural gas has resulted in a drop in gas costs of about 75% relative to 2005 prices, making this shale gas very attractive both as an energy source and as feedstock for the production of transportation fuels or chemicals. This includes light olefins, which can be obtained by first converting natural gas into synthesis gas and by subsequently converting the latter either directly through the Fischer−Tropsch-to-olefins (FTO) process or indirectly through the methanol-to-olefins (MTO) route with an intermediate methanol synthesis step.10,11 As shown in Table 1, in addition to methane, shale gas deposits contain considerable amounts of natural gas liquids
CONTENTS 1. Introduction 1.1. Setting the Scene 1.2. Scope of This Review 2. Commercial Dehydrogenation Processes 2.1. Catofin Process 2.2. Oleflex Process 2.3. Other Patented Alkane Dehydrogenation Processes 3. Thermodynamics of Nonoxidative Alkane Dehydrogenation 4. Overview of Nonoxidative Dehydrogenation Catalyst Materials 4.1. Platinum-Based Catalysts 4.1.1. Nature of the Active Sites 4.1.2. Catalyst Deactivation 4.1.3. Role of the Support 4.1.4. Role of the Promoters 4.2. Chromium Oxide-Based Catalysts 4.2.1. Nature of the Active Sites 4.2.2. Catalyst Deactivation 4.2.3. Role of the Support 4.2.4. Role of the Promoters 4.3. Vanadium Oxide-Based Catalysts 4.4. Molybdenum Oxide-Based Catalysts 4.5. Gallium Oxide-Based Catalysts 4.6. Carbon-Based Catalysts 4.7. Other Formulations 5. Comparing Different Catalysts 5.1. In Terms of Reaction Mechanism and Deactivation Pathway 5.2. In Terms of Propane and Isobutane Dehydrogenation Performance 6. Conclusions and Emerging Areas of Research Author Information Corresponding Author Notes Biographies Acknowledgments References
10613 10613 10614 10615 10615 10615 10617 10618 10620 10620 10620 10621 10621 10622 10625 10626 10627 10629 10630 10630 10633 10634 10637 10637 10640 10640 10644 10646 10647 10647 10647 10647 10648 10648
Received: May 4, 2014 Published: August 27, 2014 © 2014 American Chemical Society
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Figure 1. Consumption of ethylene, propylene, and butylenes in 2004, 2010, and 2016 (predicted). In addition, the products made by these olefins are shown in the graph to the right. These (predicted) numbers originate from ref. 6.
worldwide, and many of these installations are already under construction.18−21 Notably, there is a similarly advantageous market situation for the dehydrogenation of butanes, because the same conditions apply. Increased demand, especially from the automotive industry, combined with a decrease in supply due to the high cost of naphtha cracking, has resulted in sharp increases in butadiene prices. This in turn has sparked investment in new butane dehydrogenation (BDH) installations in the United States, China, and Japan.22−25
Table 1. Gas Composition of Six Different Shale Gas Deposits in the United Statesa shale deposit Barnett Shale (avg of four wells) Marcellus Shale (avg of four wells) Fayetteville Shale (one well) New Albany Shale (avg of four wells, N2 not reported) Antrim Shale (avg of four wells) Haynesville Shale (one well)
methane (%)
ethane (%)
propane (%)
carbon oxides (%)
nitrogen (%)
86.8
6.7
2.0
1.7
2.9
85.2
11.3
2.9
0.4
0.3
97.3
1.0
0.0
1.0
0.7
89.9
1.1
1.1
7.9
62.0
4.2
1.1
3.8
29.0
95.0
0.1
0
4.8
0.1
1.2. Scope of This Review
Because dehydrogenation for the production of light olefins has become extremely relevant in recent times, we aim to provide the reader with a complete overview of the materials used to catalyze this reaction. First, we will introduce the subject by considering the industrial processes that have been patented and applied and the thermodynamics of the dehydrogenation reaction. Next, the Pt- and CrOx-based catalysts that are currently used in commercial processes will be examined in detail. Additionally, other metal oxides, such as GaOx, VOx, MoOx, and InOx, which have also proven to be promising dehydrogenation catalysts will be surveyed. For each catalytic material, relevant factors, such as the specific nature of the active sites, as well as the effect of support, promoters, and reaction feed on catalyst performance and lifetime, will be discussed. Finally, we will compare different catalysts both in terms of the reaction mechanism and deactivation pathways and in terms of catalytic performance. To the best of our knowledge, no attempt has been made to compare such a wide range of dehydrogenation catalysts in a single review article, although other reviews have been published in recent years dealing with the dehydrogenation of light paraffins. For example, Bhasin et al. wrote an extensive review on the industrial application of dehydrogenation technologies in 2001, which includes a comparative discussion of the state of the art of oxidative dehydrogenation technologies.26 Two more review articles by Vora and Bricker (UOP) were published in 2012. In the first review article, Vora gives a general overview of the dehydrogenation of light olefins and ethylbenzene on Pt-based catalysts, discussing important principles for catalyst and reactor design.27 In the second review paper, Bricker addresses catalyst design in more detail, emphasizing mass transfer and the use of selective hydrogen combustion to increase olefin yield.28 Other reviews focusing on a specific family of dehydrogenation catalysts have also been published. For example, Weckhuysen and Schoonheydt reviewed the dehydrogenation of alkanes on supported chromium oxide catalysts in 1999.29 Dehydrogenation of alkanes on vanadium and chromium oxides is also discussed by Jackson et al. in a book on metal oxide catalysis published in 2009.30 The chemistry of
a
Larger hydrocarbons, hydrogen, and oxygen are present in trace amounts and are not given.12
(NGL), such as ethane and propane, which are easily separated from the natural gas.13 Consequently, the recent boom in shale gas production has turned the United States into the lowest cost chemical producer outside the Middle East, as illustrated by the 126 investment projects totaling 84 billion dollars made up to September 2013 that have been tallied by the American Chemistry Council.14 The high availability of these relatively cheap natural gas liquids, specifically ethane, has resulted in a shift from oil-based naphtha to shale-based ethane for steam cracking installations in the United States, with ethane crackers being constructed and naphtha crackers being either dismantled or converted. Relative to naphtha cracking, steam cracking of ethane produces a negligible amount of olefins other than ethylene, which is why the supply of propylene has dropped and its price has risen sharply, creating opportunities for on-purpose catalytic technologies, such as the catalytic dehydrogenation of light paraffins into the corresponding olefins.15 Indeed, the profitability of propane dehydrogenation is determined to a large extent by the price difference between propane and propylene, which makes the current market conditions very favorable. An additional advantage of dehydrogenation technologies is that dehydrogenation is an on-purpose technique, which yields exclusively a particular olefin instead of a mixture of products. In fact, industrial dehydrogenation processes are currently optimized in such a way that they can produce olefins of polymer-quality purity. As of this writing, ca. 5 million tons of propylene is produced annually by propane dehydrogenation (PDH).16,17 However, this number is expected to increase significantly in the upcoming years as dozens of new PDH installations have been announced 10614
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vanadium oxide for oxidative dehydrogenation has been further discussed in a review by Albonetti et al.31 Another review by Cavani et al. discusses the oxidative dehydrogenation of ethane and propane on supported vanadium and molybdenum oxides. The authors conclude that, although the oxidative dehydrogenation of ethane has potential, the oxidative dehydrogenation of higher paraffins is still far from being commercially applied.32 The nonoxidative dehydrogenation of light paraffins on both supported and bulk transition metal oxides has also been discussed by Copéret in a review focused on C−H bond activation.33 Also of note are three reviews by Krylov, Wang, and Ansari, where the use of CO2 as a mild oxidant for the dehydrogenation of light alkanes is discussed.34−36 Finally, in a chapter by Caspary et al. in the most recent edition of the Handbook of Heterogeneous Catalysis, the dehydrogenation of alkanes is discussed broadly, albeit emphasis is placed on CrOx and Pt catalysts, on the kinetics and mechanism of the dehydrogenation reaction, and on reactor technology.37
2. COMMERCIAL DEHYDROGENATION PROCESSES To date, five main industrial processes for the dehydrogenation of light alkanes have been patented, two of which are currently being used to produce light olefins.38−41 As of April 2012, 5 million tons of propylene was being produced by the 14 propane dehydrogenation plants around the globe.16,17 The global demand of propylene was 103 million tons that year and is expected to increase by 4−5% on a yearly basis.42,43 Literally dozens of new installations have been planned or are already under construction, amounting to an additional 14 million tons of propylene to be produced annually by 2018.16,17,44 Additionally, six isobutane dehydrogenation installations are currently in operation, with three other plants planned.43 The majority of these new facilities will be built in China or the United States and will use either the Catofin (Lummus) or the Oleflex (UOP) technologies. A compilation of all existing and planned propane dehydrogenation installations is shown in Figure 2. Furthermore, in Table 2, a comparison is made between the two most important processes, Catofin and Oleflex. The STAR (UHDE) process will be applied for the first time in two installations in the Middle East. The other two processes, fluidized bed dehydrogenation (FBD) process (Snamprogetti and Yarsintez) and Linde-BASF PDH, have not yet been commercially applied.
Figure 2. World map showing an overview of propane dehydrogenation installations currently in operation (filled shapes) and installations that have been announced and are expected to start up before 2018 (unfilled shapes). The numbers represent the maximum production of each plant (×1000 tons per year). The locations of the facilities are approximations.16,17,44,45 For some of the newly announced PDH installations (□), the industrial process to be used remains to be announced.
reaction products.46 A schematic of a Catofin dehydrogenation installation is shown in Figure 3. The duration of the dehydrogenation step depends on the heat content of the catalyst bed, which decreases rapidly due to the endothermic nature of the reaction. Part of the heat required for the reaction is introduced to the reactors by preheating the reaction feed, additional heat being provided by adjacent reactors that are regenerating the coked catalysts. Moreover, fuel gas is added to the reactor during the regeneration step to generate supplementary heat,47 and inert material is added to the catalyst bed to increase its heat storage capacity.48 The catalyst remains in use for 2−3 years, and the progressive loss in activity with increasing time-on-stream is counteracted by gradual increases in temperature to afford a constant dehydrogenation activity throughout the entire catalyst life span.29,49
2.1. Catofin Process
2.2. Oleflex Process
The Catofin process, by CB&I Lummus, is based on the Houdry Catadiene process, which originally was exclusively used for the dehydrogenation of isobutane to isobutene. In turn, isobutene was employed for the production of methyl tertiary butyl ether (MTBE), a fuel additive used to raise the octane number of gasoline. For environmental reasons, the use of MTBE has declined in recent years, causing a shift in the use of Catofin installations to alternative purposes, such as propane dehydrogenation. A Catofin installation generally consists of 5− 8 parallel adiabatic fixed bed reactors containing a chromia− alumina catalyst. The reaction is run at temperatures of approximately 575 °C and pressures between 0.2 and 0.5 bar.26 Each reactor alternates between dehydrogenation, regeneration, and purge steps, each lasting a few minutes (15−30 min being needed for one full cycle). Each individual reactor is made to run continuously so there are always some units performing dehydrogenation reactions, while other reactors are being regenerated or purged, which results in a constant flow of
The Oleflex process by UOP uses a very different reactor design comprising fluidized bed reactors, a catalyst regeneration unit, and a product recovery section. A schematic representation is shown in Figure 4.50 The reaction is run using a Pt−Snbased catalyst at pressures between 1 and 3 bar and temperatures ranging from 525 to 705 °C. Three or four adiabatic radial flow reactors containing the catalyst are connected in series with preheaters in between. These gas flow preheaters represent the main source of heat of the reactor system. The catalyst flows through the system, the last reactor being connected to a continuous catalyst regeneration (CCR) unit that regenerates the catalyst by burning of any carbon deposits and redisperses the Pt on the support material by means of treating the catalyst with a chlorine−air mixture. The entire system is designed to operate continuously to obtain an uninterrupted stream of reaction products. The regenerated catalyst is then reintroduced to the first reactor, completing an 10615
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Table 2. Specifics of the Catofin and Oleflex Commercial Dehydrogenation Processes43,44,46,50−55 Catofin
Oleflex
catalyst formulation process license holder used for operating conditions reactor type
18−20 wt % CrOx on an alumina support, promoted by 1−2 wt % Na or K CB&I Lummus
CrOx/SiO2 > CrOx/MCM-41 ≫ CrOx/MgO ≈ CrOx/Si-2 (Si-2 is the abbreviation of silicalite-2 zeolite, which has been described as a silica analogue of the aluminosilicate zeolite ZSM-11194), a reverse order being found for the reduction temperature of the chromium species on the aforementioned catalyst supports. Cr−Mg−Al and Cr−Mg mixed oxides prepared using layered double hydroxide (LDH) precursors have been studied as catalysts for the nonoxidative dehydrogenation of ethane.195 Some of these catalysts were found to achieve ethane conversions of 27−30% and selectivity toward ethylene of 71−75%. The method used for introducing Cr into the LDH precursors was found to determine several properties of the resulting mixed oxides, such as surface area, catalytic performance, coking resistance, and thermal stability against sintering. Finally, the study of porous chromia-pillared α-zirconium phosphate materials has been particularly informative with regards to the influence that acidity has on catalytic behavior. In
4.3. Vanadium Oxide-Based Catalysts
Despite the excellent performance of Pt and CrOx catalysts in paraffin dehydrogenation, a number of issues including catalyst poisoning, the high cost of Pt, and environmental concerns associated with the use of Cr have spurred the search for alternatives. Vanadium oxides are known to be active for many hydrocarbon oxidation reactions, including dehydrogenation. Research on vanadium-based dehydrogenation catalysts started in the 1980s, although at that time the focus was on oxidative dehydrogenation using vanadium−magnesium mixed oxides.199,200 10630
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Figure 20. An artist’s impression of the surface of a CrOx/Al2O3 catalyst in which coordinatively unsaturated Cr3+ species catalyze the dehydrogenation reaction. Note that partially reduced CrOx species are formed on the surface of the catalyst, and that the Cr is partially dissolved in the alumina support. The (simplified) reaction mechanism is shown in the inset.
catalyst, as −OH groups serve as anchoring sites for the vanadium precursors used to prepare these catalysts. On the other hand, acid sites are detrimental for catalyst performance, because they facilitate cracking and induce coke formation. A schematic display of these vanadium species is shown in Figure 21.202 Several authors claim that two-dimensional polymeric V3+ or 4+ V species are the most active for dehydrogenation,203−206 the vanadium species bonded directly with the γ/θ-alumina support being the most active sites in PDH. Isolated vanadium species are also active, albeit they are susceptible to deactivation due to coke formation. On the contrary, large V2O5 crystallites are practically inactive in dehydrogenation.30,206 Prior to reaction, the vanadium in these catalysts is generally present as V5+. However, the reduction of the catalyst by the hydrocarbon feed results in the formation of V4+ and V3+.207−209 The reducibility of the vanadium species depends on their molecular structure, V−O−V and VO bonds being
The nature of the vanadium present on the catalyst depends to a large extent on the type of carrier, the support surface area, the metal loading, and the oxidation state of vanadium.201 Far below the theoretical monolayer coverage isolated vanadium species are present, at intermediate surface coverage polymeric species or sheets of vanadium oxide are favored, and above monolayer coverage V2O5 crystallites become dominant. Therefore, a combination of V−O-support, V−O−V, or V O moieties may be obtained, each having different catalytic activity. In addition, depending on the support, mixed oxides with Al, Zr, Ti, or Mg may be formed. The surface area-to-metal loading ratio determines which species will form. Indeed, Wu et al. reported that monomeric species are predominantly present below 1.2 V/nm2, polymeric vanadium species prevail between 1.2 and 4.4 V/nm2, and crystalline V2O5 species preponderate at even higher loadings. On the one hand, the presence of hydroxyl groups on the surface of the support is important to obtain a well-dispersed 10631
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Table 4. Summary of the Catalytic Data of the Cr-Based Dehydrogenation Catalysts Discussed catalyst (1) mesoporous 9 wt % CrOx/ Al2O3 (2) mixed oxide 40% Cr2O3/60% Al2O3 (3) 73 wt % Cr2O3-pillared α-ZrP
reaction temp (°C)
space velocity (h−1)
580 588 ± 2 550
(4) Ga−Cr mixed oxide on α-ZrP (20.1 wt % Cr, 9.3 wt % Ga) (5) 14−16 wt % CrOx/Al2O3
540
(6) 5 wt % CrOx/SBA-1
550
(7) 4 wt % CrOx/ZrO2
550
(8) (1 wt % Cr) Cr−Si−Zr Xerogel (9) Cr(2)Mg(12)Al(4)Ox (mixed metal oxide) (10) 5 wt % Cr, 10 wt % Ce/ SBA-15 (11) 20 wt % Cr, 1 wt % Na/ Al2O3
450
550
700 700 550
feed composition
conversion (%)a
selectivity (%)a
C3H8 = 5, He = 95
15.7−10.0
98
=
iC4H10 = 100
∼23−14
99.4
=
C3H8 = 7, He = 93
18.1−6.3
=
C3H8 = 7, He = 93
=
iC4H10 = 10, N2 = 90
WHSV = 1.2 WHSV = 0.3 WHSV = 0.86 WHSV = 3.2 GHSV = 3600 WHSV = 0.12
C3H8 = 6.67, CO2 = 33.3, He = 60 C3H8 = 2.5, N2 = 97.5 C3H8 = 1.9, N2 = 0.2, He = 97.1 C2H6 = 28.6, N2 = 71.4 C2H6 = 25, CO2 = 75 C3H8 = 10, N2 = 90
WHSV 3.3 WHSV 1.7 WHSV 2 WHSV 46.6
specific activity (s−1)b
kd (h−1)
catalyst lifec (h)
ref
0.52
1
197
6.87 × 10−4
0.48
1.25
154
87.1
1.23 × 10−5
0.24
5
162
27.6−4.4
80
5.05 × 10−5
0.42
5
196
16 (start first cycle), 8 (30th cycle) 37−26
98 ± 1
1.21 × 10−3
0.026
30
163
85−92
1.66 × 10−4
0.14
3.75
165
60.9−24.3
76.2−87.0
2.86 × 10−5
0.26
6
170
37.8−22.7 (35 and 400 m) 29.6
91.1−94.7
1.85 × 10−4
0.12
6.08
191
71
6.54 × 10−4
1
195
55−46
96−97
0.059
5
175
47−37
80−89
0.069
6
198
7.41 × 10−6
a First value is obtained at the start of the cycle, second at the end. bSpecific activity is defined as (mol olefin formed)/(mol Cr*t(s)). cCatalyst life = total time single cycle/experiment.
Figure 21. Various vanadium oxide species present on a catalyst support (S): isolated vanadium oxide species (a), dimers (b), two-dimensional chains of vanadium oxide species (c), and V2O5 crystallites (d). Reprinted from ref 201, copyright 2003, with permission from Elsevier.
Ogonowski et al. reported that the effect of CO2 depends strongly on the catalyst support. Relatively weak basic sites on carriers, such as those present on an activated carbon support, facilitate the reverse water gas shift reaction, which lifts the thermodynamic limitation of the reaction and increases alkane conversion. When these basic sites are too strong, as is the case on Al2O3 and ZnO2, CO2 irreversibly adsorbs and poisons these sites. This increases the acidity of the support, which in turn increases coke formation and lowers selectivity. If these basic sites are absent, as in the case of SiO2, no beneficial effects of CO2 are observed. Alumina is commonly used as a support for vanadium oxide catalysts, although Sokolov et al. concluded through testing a wide range of Si- and Al-based supports that using silica to dope the alumina support provides beneficial effects on catalyst stability and activity.214 Furthermore, as was shown by Harlin et al., the performance of supported vanadium catalysts can be improved through the addition of MgO, either as a promoter or as a component of a mixed oxide. According to these workers, the increase in dehydrogenation selectivity stems from the basic nature of MgO, resulting in reduced coke formation.208 An overview of the different VOx catalysts used for the dehydrogenation of light alkanes is shown in Table 5.
more easily reduced than V−O-support bonds. Harlin et al. studied the effects of vanadium oxide reduction with H2, CO, and CH4 prior to the dehydrogenation reaction. In all cases, the reduction was incomplete, and yielded a mix of vanadium oxide species, including V3+, V4+, and V5+ (see Figure 22). Higher initial catalytic activities were obtained in all cases, although the effect was strongest after reduction with CO. The authors suggested that when the catalyst was reduced with H2 and CH4, detrimental −OH groups were formed. V4+ was determined to be the most active species in butane dehydrogenation. Initially, deactivation is caused by the strong adsorption of reactants on the active sites and to a lesser extent to coke deposition. When the coke is removed through the calcination of the catalyst, the latter may deactivate further as vanadium oxide species sinter and form larger V2O5 crystallites or mixed oxides such as AlVO4, both of which are less active.206 V2O5 crystallites can revert to active polymeric vanadium species by treating the catalyst under an oxygen atmosphere at 600 °C.210 In several studies, CO2 is added to the feed as a mild oxidant to increase alkene yields.211−213 Different reasons for this beneficial effect of CO2 have been suggested, such as the removal of acid sites that catalyze unwanted side reactions, coke gasification, oxidation of overly reduced vanadium centers, and hydrogen removal by the reverse water−gas shift reaction. 10632
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600
As some of these catalysts require an activation period, the maximum (as opposed by the initial) conversion is considered to determine the deactivation rate. bFirst value is obtained at the start of the cycle, second at the end. cSpecific activity is defined as (mol olefin formed)/(mol Cr*t(s)). dCatalyst life = total time single cycle/experiment. eCatalyst is prereduced by CO prior to the dehydrogenation reaction. a
215 3 0.86 12−30 to 1- and 1−3 butene 70−15
212 0.72
i-C4H10 = 14.3, CO2 = 85.7 GHSV = 72 (at a pressure of C4H10 = 3, N2 = 97 0.5 bar) WHSV = 1.9 600
45.5−26.5
76.4−85.0
5.77 × 10−4
1.17
208 1.34; 3.12 9.32 × 10−4; 9.40 × 10−4 85−90 77−88 (to ibutene) 44−36; 54−35 WHSV = 5 580
(4) 4.9 wt % VOx−4.7 wt % Mg/Al2O3e (5) 5.4 wt % VOx−4.5 wt % Zr/Al2O3e (6) 4 wt % V−36 wt % MgOx/ Cact (7) 3.5 wt % VOx/Al2O3
62−38 WHSV = 5
i-C4H10 = 10, N2 = 90 i-C4H10 = 10, N2 = 90 580 (3) 5.2 wt % VOx/Al2O3c
32.9−8.5 (15−60 m) GHSV = 14 400 600
WHSV = 2.6, N2 = 90
(2) 3.5 wt % VOx/θ-Al2O3
nC4H10 = 100
0.25
207 0.085 11.5 9.46 × 10−4
205 1 2.22
202 1.33 1.04 11−3 nC4H10 = 10
55−76 (to all butene isomers) 56.1 (all butene isomers) 65−85 (to i-butene)
1.75 × 10
−4
520
conversion (%)b feed composition
(1) 2.2 wt % VOx/γ-Al2O3
Molybdenum oxides are also frequently used as catalysts in hydrocarbon conversion reactions, including dehydrogenation. To our knowledge, the first report of the dehydrogenation activity of a molybdenum oxide-based catalyst was of MoO3/ Al2O3 in the dehydrocyclization of n-heptane, which was published in 1946.216 Since that first report, the nonoxidative dehydrogenation of alkanes on MoOx has been studied scarcely, as emphasis has been given to the study of oxidative dehydrogenation on these materials. The chemistry of molybdenum oxide is comparable to that of vanadium oxide, as Mo can be present as MoOx monomers, polymers, or MoO3 crystallites, depending on the molybdenum loading, the type of support used, and catalyst preparation conditions.217 Harlin et al. noted that optimal activity was achieved with monolayer coverage of molybdenum oxide on an alumina support. In the fresh catalyst, molybdenum is present as Mo6+, but analogously to catalysts comprising vanadium and chromium oxide, MoO3 is reduced by the hydrocarbon feed during reaction to create the active species.218 These reduced sites are believed to be Mo4+ and Mo5+. Further reduction results in sites that are more active, but the latter catalyze cracking and deactivate fast due to coke deposition. In fact, at
space velocity (h−1)
4.4. Molybdenum Oxide-Based Catalysts
reaction temp (°C)
Figure 22. Vanadium 2p3/2 XPS spectra of a VOx/Al2O3 catalyst calcined (top) and reduced with CO (bottom). Reduction yields a mixture of V5+, V4+, and V3+. Reprinted from ref 207, copyright 2000, with permission from Elsevier.
catalyst
Table 5. Summary of the Catalytic Data of the V-Based Dehydrogenation Catalysts Discusseda
selectivity (%)b
specific activity (s−1)c
kd (h−1)
catalyst lifed (h)
ref
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loadings above monolayer coverage, the catalyst is more easily reduced, which results in a drop in catalyst activity.218,219 The same authors also point out that such a catalyst deactivates due to the formation of catalytically inactive Al 2 (MoO 4 ) 3 crystallites. To prevent the formation of this mixed oxide, they added magnesium oxide to obtain a molybdenum− magnesium mixed oxide that also shows dehydrogenation activity. It was found that an additional beneficial effect of magnesia addition is that the resulting mixed oxide has lower acidity, which increases alkene selectivity and reduces coke deposition. The reduction of the catalyst with hydrogen prior to reaction increases conversion, although selectivity decreases due to the formation of acid sites.220 Overall, MoO3 catalysts deactivate rapidly due to coking and generally afford a relatively low conversion and selectivity. The formation of molybdenum oxycarbides (MoOxCy) has been reported to occur during butane dehydrogenation on unsupported molybdenum oxide. These materials are also active in the dehydrogenation of paraffins, and a high activity close to the thermodynamic equilibrium is initially obtained, although over a period of a few hours the selectivity drops to zero and the reaction yields only methane, as shown in Figure 23. During this period, the oxycarbide becomes depleted from oxygen and a wide range of carbides active in hydrogenolysis are formed.222 To improve the stability of the oxide, steam can be added to the feed (which hampers the growth of these carbides) or the catalyst can be doped with vanadium (which increases the oxophilicity of the catalyst). X-ray diffractograms showing this transition are shown in Figure 24. However, the introduction of steam is not free of drawbacks, as it induces sintering. Hydrogen addition prevents coke formation and increases conversion, albeit it is believed that H2 competitively adsorbs on the active sites of the catalyst.221−223 Consistent with the fact that MoOxCy is a catalyst active in butane dehydrogenation, Harlin et al. showed that MoO3 deposited on SiC displays significant catalytic properties when hydrogen and steam are present in the feed.224 As compared to bulk MoO3 and MoO3 supported on SiO2 and Al2O3, higher selectivities and stability are obtained over MoO3/SiC, albeit at the cost of a lower conversion. This was explained by the increased reducibility of MoO3/SiC as compared to the MoO3/Al2O3 catalyst. A summary of the Mo-based catalysts used for the dehydrogenation of light paraffins is shown in Table 6. 4.5. Gallium Oxide-Based Catalysts
Figure 23. Yield of different carbon molecules during butane dehydrogenation on a MoOxCy catalyst at different steam partial pressures. When no steam is added, the selectivity to C4 alkenes drops rapidly and the methane yield reaches 100% after 5 h. The hydrogenolysis reaction is suppressed when steam is added to the feed. Reprinted from ref 221, copyright 1999, with permission from Elsevier.
The dehydrogenation activity of gallium oxide supported on ZSM-5, which was first reported in the late 1980s in studies dealing with conversion of propane to aromatics, has been the target of renewed attention in recent years.225−228 In the interim, both bulk and supported Ga2O3 have been used as dehydrogenation catalysts. The exact nature of the active sites in Ga-based dehydrogenation catalysts is still under debate. Akin to Al2O3, Ga2O3 has different polymorphs of which β-Ga2O3 is not only the most stable, but also the most active Ga species in PDH according to some authors.229 β-Ga2O3 has a monoclinic structure, in which gallium atoms are equally distributed between tetrahedral and octahedral configurations. This crystal structure has a high concentration of weak Lewis acid sites (coordinatively unsaturated tetrahedral ones on the surface of the material), which are deemed to be centers of dehydrogenation activity.230 Indeed, a heterolytic dissociation reaction mechanism has been
proposed where H− adsorbs on a Ga+ Lewis acid site and C3H7+ on a neighboring oxygen, although these species can also be adsorbed reversed.33,231 The assignment of Lewis acid sites as the active species was substantiated by Chen et al.,232 who, by relating the results of NH3-TPD measurements with activity data for different GaxAl10−xO15 mixed oxides, showed that a high concentration of Lewis acid groups in the form of coordinatively unsaturated tetrahedral Ga3+ cations is a prerequisite for dehydrogenation activity (see Figure 25). Xu et al. reported dehydrogenation activity on both Lewis and Brønsted acid sites on GaO3/H-ZSM-5. A zeolitic support displaying a low concentration of medium and strong acid sites and a relatively high concentration of weak acid sites was found to lead to the most active and stable catalyst.233,234 In addition, Michorczyk et al. concluded that acidic groups catalyze the 10634
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Figure 24. XRD diffractograms of the fresh Mo oxycarbide catalyst before dehydrogenation, but after pretreatment (top left), as well as after dehydrogenation reaction with no steam (top right), with 5 Torr of steam (bottom left), and with 18 Torr of steam (bottom right). In the absence of steam, the well-defined MoO2 and MoOxCy species disappear during reaction and are replaced by more amorphous MoOxCy species, which are enriched with carbon and presumably responsible for hydrogenolysis. The addition of steam prevents this transformation and with 18 Torr of steam, MoO2 and ordered MoOxCy species are still observed on the spent catalysts. Reprinted from ref 221, copyright 1999, with permission from Elsevier.
Table 6. Summary of the Catalytic Data of the Molybdenum Oxide-Based Dehydrogenation Catalysts Discusseda catalyst
reaction temp (°C)
space velocity (h−1)
feed composition
(1) 13.4 wt % MoO3/Al2O3
560
WHSV = 2
n-C4H10 = 10, N2 = 90
(2) 12.8 wt % MoO3−8.4 wt % Mg/Al2O3 (3) MoO3
550
WHSV = 2
(4) 14.2 wt % MoOx/SiC
560
WHSV = 2
n-C4H10 = 9.3, N2 = 83.9, H2O = 6.8 n-C4H10 = 10, N2 = 90, H2O = 3.1
conversion (%)b 21−18 (second min 25) 13−11
specific activity (s−1)c
kd (h−1)
catalyst lifed (h)
ref
50
7.21 × 10−5
2.78
0.17
218
82
7.66 × 10−5
1.12
selectivity (%)b
−5
42−8
55−84
2.96 × 10
0.28
7.5
221
22−17
38−71
5.41 × 10−5
0.32
1
224
a
As some of these catalysts require an activation period, the maximum (as opposed to the initial) conversion is considered to determine the deactivation rate. bFirst value is obtained at the start of the cycle, second at the end. cSpecific activity is defined as (mol olefin formed)/(mol Mo*t(s)). dCatalyst life = total time single cycle/experiment.
high energy barrier. Nevertheless, theoretical calculations by Pidko et al. indicate that these relatively stable [H−Ga−OH]+ sites are preferentially regenerated by water desorption resulting in a reduced active site.236 Murata and co-workers have observed an increase in activity after a prereduction step with hydrogen at 550 °C, while a simultaneous increase of Gaδ+H and Ga−OH sites was observed via NH3 TPD. These authors proposed that these sites also form during reaction and may be responsible for the dehydrogenation activity of the silica supported Ga2O3 catalyst.237,238 Contradictory information has been reported regarding the reducibility of Ga2O3. Rodriguez et al. reported that Ga2O3/ Al2O3 is not reduced by hydrogen at 700 °C, but the presence of Pd on the catalyst can assist the reduction of Ga3+ to Ga+ and
reaction after observing that promotion with a base, such as potassium, resulted in a decrease in activity.235 Several authors have performed DFT calculations in an effort to understand the dehydrogenation reaction mechanism on Ga2O3. Studies by Liu et al. regarding propane dehydrogenation on a perfect Ga2O3 (100) surface revealed that the C−H bond is activated by a surface oxygen that abstracts an hydrogen from the propane molecule, thus forming a propyl species, which coordinates with a Ga surface site and a hydroxyl group. A second hydroxyl group is formed via β-hydrogen elimination, but for these hydroxyl groups to recombine, either as H2O or as H2, a high energy barrier needs to be overcome. This barrier is much lower over a gallium hydride−hydroxide species, although the formation of the gallium hydride also involves a 10635
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with CrOx (vide supra), where a weak metal−support interaction results in an increased reducibility. Indeed, Meitzner et al. showed that Ga2O3/H-ZSM-5 could be reduced by hydrogen or hydrocarbons at dehydrogenation reaction conditions. Finally, regarding the reduction of Ga2O3, it is important to note that Ga2O is gaseous and metallic gallium is liquid at dehydrogenation reaction conditions.240 A number of studies have shown that the addition of CO2 to the reactant feed improves the olefin yield.241−245 Nakagawa et al. found that the addition of steam and CO2 decreased coke deposition on a Ga2O3/TiO2 catalyst by the Boudouart reaction, while also assisting ethylene desorption.241,242 However, other authors claimed that the increase in activity was caused by the reverse water−gas shift reaction (RWGS).243,244 Additionally, Michorczyk et al. noted that this effect could only be seen at relative low levels of coke deposition.235 It was suggested that CO2 could also poison the basic sites required for the dissociative adsorption of the alkane. Additionally, Hensen et al. studied the promoting effect of water with experiments and theory and observed an increase in the conversion of propane by cofeeding water.246 The authors attributed this promoting effect to the formation of partially hydrolyzed gallium species, more specifically, to the formation of binuclear hydroxyl-bridged Ga3+ cations. Different carriers have been used in gallium oxide-based dehydrogenation catalysts, the inclusion of acid and basic sites on the support being important to achieve a high conversion. Xu et al. reported that in the absence of an oxidizing agent, the conversion obtained using Ga2O3 supported on ZrO2, Al2O3, and TiO2 was high, medium, and low, respectively, while SiO2 and MgO-supported Ga2O3 were inactive in alkane dehydrogenation.233 The authors proposed the activity to be dependent on the presence of acid sites of medium strength on the catalyst surface. However, a different trend was observed when CO2 was present, as conversion was observed to be high on Ga2O3/ TiO2 and low on Ga2O3/ZrO2 and Ga2O3/Al2O3. This was attributed to the two contrary roles of CO2, which on the one
Figure 25. Concentration of Lewis acid sites versus propane conversion measured 15 min after the start of the PDH reaction on GaxAl10−xO15. A clear correlation between the conversion and the surface density of weak Lewis acid sites is observed. For these experiments, CO2 is used to improve conversion. Reprinted from ref 232, copyright 2008, with permission from Elsevier.
Ga0 at temperatures between 300 and 550 °C and form an alloy (in which Ga0 is immobilized). Indeed, it has been reported that the reduction of Ga2O3 can be promoted by the proximity of other, more easily reduced metals, such as Pt and Pd.131−133,238 For instance, on a Pt−Ga/H-beta catalyst, reduced gallium hydrides were observed after hydrogen reduction at 550 °C.239 It has also been proposed that the reducibility of these species depends on the interaction of gallium oxide with the catalyst support, which is also the case
Table 7. Summary of the Catalytic Data of Gallium Oxide-Based Dehydrogenation Catalysts Discusseda reaction temp (°C)
space velocity (h−1)
(1) 1.7 wt % Ga2O3/SiO2
550
(2) Ga2O3/MTS (mesoporous silica) ratio Ga/Si = 0.05 (3) β-Ga2O3
550
(4) Ga2O3
600
(5) 5 wt % Ga2O3/ZrO2
600
WHSV = 0.97 h−1 WHSV = 0.33 h−1 WHSV = 0.15 h−1 WHSV = 1.2 h−1 WHSV = 0.3 h−1
catalyst
500
(6) 5 wt % Ga2O3/TiO2 (7) Ga8Al2O15
(8) 5 wt % Ga2O3/H-ZSM-5 (w Si/Al = 97) (10) 0.001 wt % Pt, 1.2 wt % Ga/Al2O3
500
650 590
WHSV = 0.3 h−1
WHSV = 0.36 h−1 GHSV = 400 h−1
feed composition
selectivity (%)b
conversion (%)b
specific activity (s−1)c
kd (h−1)
catalyst lifed (h)
ref
C3H8 = 10, Ar = 90
23−20
86−80
4.97 × 10−4
0.036
5
249
i-C4H10 = 10, He = 90 C3H8 = 2.5, N2 = 97.5 C3H8 = 17, CO2 = 83 C3H8 = 2.5, N2 = 97.5 C3H8 = 2.5, CO2 = 5, N2 = 92.5 (a) C3H8 = 2.5, N2 = 97.5 (b) C3H8 = 2.5, CO2 = 5, N2 = 92.5 C2H6 = 3, CO2 = 15, Ar = 82 C3H8 = 100
46−40
58.3 to i-C4H8
5.52 × 10−5
0.092
2.67
230
33−12
95−95
6.95 × 10−7
0.21
6
229
31−3
95−82
3.55 × 10−5
0.67
4
235
39−6; 32−8
74
1.90 × 10−5
0.77
3
233
73
1.54 × 10−5
0.56
51.7−22.5
91.6−98.2
2.39 × 10−6
0.168
8
232
49.7−33.1e
91.7−98.0e
2.30 × 10−6
0.089
23.7−14.5
85.9−93.7
2.84 × 10−5
0.0086
70
244
39
86
0.25
247
a
As some of these catalysts require an activation period, the maximum (as opposed to the initial) conversion is considered to determine the deactivation rate. bFirst value is obtained at the start of the cycle, second at the end. cSpecific activity is defined as (mol olefin formed)/(mol Ga*t(s)). dCatalyst life = total time single cycle/experiment. eValues obtained at 15 min and 8 h. 10636
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activity. The authors hypothesized that the carbon species formed on the catalyst surface were in fact catalyzing the reaction. The highest activity was observed at 700 °C, a temperature at which the carbon deposits display a very ordered structure. In view of this, the authors decided to test their hypothesis by using carbon nanofibers (CNFs) as a catalyst under similar conditions. Tellingly, CNFs proved to be catalytically active. It was concluded that radicals formed on the surface of coke facilitate the dehydrogenation reaction, albeit at elevated temperatures the noncatalytic conversion of propane is also relatively high. However, elevated temperatures also lead to a drop in selectivity, as cracking becomes the dominant reaction. Interestingly, the latter was not the case in the work of McGregor et al., who reported a selectivity toward butenes of ∼80%. Furthermore, hybrid nanocarbon, comprising a diamond core and a graphitic shell, displayed activity for the dehydrogenation of propane, which was attributed to the presence of either ketone groups or structural defects in the graphitic material.251 In previous work, CNFs had been shown to be active in the oxidative dehydrogenation of n-butane, although in this case surface oxygen groups were believed to be the active sites.252 Furthermore, nitrogen doped carbon nanotubes (CNTs) showed some activity in the oxidative dehydrogenation of propane, although conversion and selectivity were relatively low.253 Liu et al. prepared carbon-based ordered mesoporous systems that showed activity in the dehydrogenation of propane in the absence of steam or oxidants.254 The active sites were proposed to be oxygen-containing groups (CO), which are formed during a pretreatment step involving nitric acid. The catalyst shows remarkable long-term stability, displaying conversion and selectivity values of 45% and 85%, respectively, after 100 h of time on stream at 600 °C. Experiments in which carbon nanotubes were tested under similar conditions resulted in low activity. Neylon et al. studied the activity of supported early transition metal (groups 5 and 6) nitrides and carbides in butane dehydrogenation,255,256 as the catalytic properties of these materials are believed to be similar to those of platinum. Interestingly, Mo, V, and W carbides and nitrides showed activity in both dehydrogenation and hydrogenolysis, displaying specific activities similar to those of a commercial Pt−Sn/Al2O3 catalyst. Niobium carbide and nitride did not show any activity. The group 6 compounds tested were found to be more active, albeit they primarily catalyzed hydrogenolysis. In Table 8 are shown the catalytic data of the carbon-based dehydrogenation catalysts as reported by McGregor et al. and Liu et al. In addition, silica supported metal sulfides, such as ZnS, FeS, CuS, CoS, MnS, NiS, and MoS, were tested for the dehydrogenation of isobutane, and were found to perform significantly better than their respective oxides. The materials show higher initial yields than commercial CrOx and Pt−Sn catalyst, but they deactivate over a few hours due to sulfur leaching.257
hand removes H2 from the surface of the catalyst by the reverse water−gas shift reaction, while on the other hand absorbs strongly on the basic sites of the support. Thus, the low conversion observed when zirconia or alumina were employed as the support can be ascribed to the relatively high amount of basic sites on these oxides. Notably, the structure of the support is also important, as a mesoporous Ga−Al mixed oxide exhibited an increased resistance to catalyst coking.245 Work on the promotion of GaOx-based dehydrogenation catalysts appears to be limited, although a patent by Iezzi et al. claims that a catalyst comprising 1.2 wt % Ga2O3,
Catalytic Dehydrogenation of Light Alkanes on Metals and Metal Oxides Jesper J. H. B. Sattler, Javier Ruiz-Martinez, Eduardo Santillan-Jimenez, and Bert M. Weckhuysen* Inorganic Chemistry and Catalysis, Debye Institute of Nanomaterials, Utrecht University, Universiteitsweg 99, 3584 CG Utrecht, The Netherlands 1. INTRODUCTION 1.1. Setting the Scene
Because of their extensive use as chemical building blocks, light olefins, such as propylene (propene) and ethylene (ethene), are among the most important types of compounds in the chemical industry. Light olefins are the feedstock employed in the production of a vast array of chemicals, including polymers (e.g., polyethylene and polypropylene), oxygenates (e.g., ethylene glycol, acetaldehyde, acetone, and propylene oxide), and important chemical intermediates (e.g., ethylbenzene and propionaldehyde).1−4 The demand of these building blocks has increased steadily over the last years, all major markets registering a positive growth in demand and production in recent years with the exception of the Eurozone, due to the effects of the recent economic downturn.5 In Figure 1, the uses as well as the recent and predicted future growth of the olefin supply and demand are shown.6 Steam cracking and fluid catalytic cracking (FCC) of naphtha, light diesel, and other oil byproducts are the most common methods for obtaining these light olefins. For example, in 2007 less than 3% of propylene was produced by on-purpose techniques, while FCC and steam cracking accounted for the bulk of the propylene produced.7 A number of factors including the high-energy demands of these processes, their low selectivity toward the production of particular olefins, dwindling petroleum reserves, and rising oil prices are driving the petrochemical industry to search for a more economical feedstock and more efficient conversion technologies. In recent years, hydraulic fracturing or “fracking” technologies have improved to the point where large volumes of shale gas can be extracted in a cost-effective manner. In fact, the United States already obtains one-quarter of their natural gas production from shale gas deposits, and this is expected to rise even further in the coming years due to the emphasis that is being given to the attainment of energy independence.8,9 This increased supply of natural gas has resulted in a drop in gas costs of about 75% relative to 2005 prices, making this shale gas very attractive both as an energy source and as feedstock for the production of transportation fuels or chemicals. This includes light olefins, which can be obtained by first converting natural gas into synthesis gas and by subsequently converting the latter either directly through the Fischer−Tropsch-to-olefins (FTO) process or indirectly through the methanol-to-olefins (MTO) route with an intermediate methanol synthesis step.10,11 As shown in Table 1, in addition to methane, shale gas deposits contain considerable amounts of natural gas liquids
CONTENTS 1. Introduction 1.1. Setting the Scene 1.2. Scope of This Review 2. Commercial Dehydrogenation Processes 2.1. Catofin Process 2.2. Oleflex Process 2.3. Other Patented Alkane Dehydrogenation Processes 3. Thermodynamics of Nonoxidative Alkane Dehydrogenation 4. Overview of Nonoxidative Dehydrogenation Catalyst Materials 4.1. Platinum-Based Catalysts 4.1.1. Nature of the Active Sites 4.1.2. Catalyst Deactivation 4.1.3. Role of the Support 4.1.4. Role of the Promoters 4.2. Chromium Oxide-Based Catalysts 4.2.1. Nature of the Active Sites 4.2.2. Catalyst Deactivation 4.2.3. Role of the Support 4.2.4. Role of the Promoters 4.3. Vanadium Oxide-Based Catalysts 4.4. Molybdenum Oxide-Based Catalysts 4.5. Gallium Oxide-Based Catalysts 4.6. Carbon-Based Catalysts 4.7. Other Formulations 5. Comparing Different Catalysts 5.1. In Terms of Reaction Mechanism and Deactivation Pathway 5.2. In Terms of Propane and Isobutane Dehydrogenation Performance 6. Conclusions and Emerging Areas of Research Author Information Corresponding Author Notes Biographies Acknowledgments References
10613 10613 10614 10615 10615 10615 10617 10618 10620 10620 10620 10621 10621 10622 10625 10626 10627 10629 10630 10630 10633 10634 10637 10637 10640 10640 10644 10646 10647 10647 10647 10647 10648 10648
Received: May 4, 2014 Published: August 27, 2014 © 2014 American Chemical Society
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Figure 1. Consumption of ethylene, propylene, and butylenes in 2004, 2010, and 2016 (predicted). In addition, the products made by these olefins are shown in the graph to the right. These (predicted) numbers originate from ref. 6.
worldwide, and many of these installations are already under construction.18−21 Notably, there is a similarly advantageous market situation for the dehydrogenation of butanes, because the same conditions apply. Increased demand, especially from the automotive industry, combined with a decrease in supply due to the high cost of naphtha cracking, has resulted in sharp increases in butadiene prices. This in turn has sparked investment in new butane dehydrogenation (BDH) installations in the United States, China, and Japan.22−25
Table 1. Gas Composition of Six Different Shale Gas Deposits in the United Statesa shale deposit Barnett Shale (avg of four wells) Marcellus Shale (avg of four wells) Fayetteville Shale (one well) New Albany Shale (avg of four wells, N2 not reported) Antrim Shale (avg of four wells) Haynesville Shale (one well)
methane (%)
ethane (%)
propane (%)
carbon oxides (%)
nitrogen (%)
86.8
6.7
2.0
1.7
2.9
85.2
11.3
2.9
0.4
0.3
97.3
1.0
0.0
1.0
0.7
89.9
1.1
1.1
7.9
62.0
4.2
1.1
3.8
29.0
95.0
0.1
0
4.8
0.1
1.2. Scope of This Review
Because dehydrogenation for the production of light olefins has become extremely relevant in recent times, we aim to provide the reader with a complete overview of the materials used to catalyze this reaction. First, we will introduce the subject by considering the industrial processes that have been patented and applied and the thermodynamics of the dehydrogenation reaction. Next, the Pt- and CrOx-based catalysts that are currently used in commercial processes will be examined in detail. Additionally, other metal oxides, such as GaOx, VOx, MoOx, and InOx, which have also proven to be promising dehydrogenation catalysts will be surveyed. For each catalytic material, relevant factors, such as the specific nature of the active sites, as well as the effect of support, promoters, and reaction feed on catalyst performance and lifetime, will be discussed. Finally, we will compare different catalysts both in terms of the reaction mechanism and deactivation pathways and in terms of catalytic performance. To the best of our knowledge, no attempt has been made to compare such a wide range of dehydrogenation catalysts in a single review article, although other reviews have been published in recent years dealing with the dehydrogenation of light paraffins. For example, Bhasin et al. wrote an extensive review on the industrial application of dehydrogenation technologies in 2001, which includes a comparative discussion of the state of the art of oxidative dehydrogenation technologies.26 Two more review articles by Vora and Bricker (UOP) were published in 2012. In the first review article, Vora gives a general overview of the dehydrogenation of light olefins and ethylbenzene on Pt-based catalysts, discussing important principles for catalyst and reactor design.27 In the second review paper, Bricker addresses catalyst design in more detail, emphasizing mass transfer and the use of selective hydrogen combustion to increase olefin yield.28 Other reviews focusing on a specific family of dehydrogenation catalysts have also been published. For example, Weckhuysen and Schoonheydt reviewed the dehydrogenation of alkanes on supported chromium oxide catalysts in 1999.29 Dehydrogenation of alkanes on vanadium and chromium oxides is also discussed by Jackson et al. in a book on metal oxide catalysis published in 2009.30 The chemistry of
a
Larger hydrocarbons, hydrogen, and oxygen are present in trace amounts and are not given.12
(NGL), such as ethane and propane, which are easily separated from the natural gas.13 Consequently, the recent boom in shale gas production has turned the United States into the lowest cost chemical producer outside the Middle East, as illustrated by the 126 investment projects totaling 84 billion dollars made up to September 2013 that have been tallied by the American Chemistry Council.14 The high availability of these relatively cheap natural gas liquids, specifically ethane, has resulted in a shift from oil-based naphtha to shale-based ethane for steam cracking installations in the United States, with ethane crackers being constructed and naphtha crackers being either dismantled or converted. Relative to naphtha cracking, steam cracking of ethane produces a negligible amount of olefins other than ethylene, which is why the supply of propylene has dropped and its price has risen sharply, creating opportunities for on-purpose catalytic technologies, such as the catalytic dehydrogenation of light paraffins into the corresponding olefins.15 Indeed, the profitability of propane dehydrogenation is determined to a large extent by the price difference between propane and propylene, which makes the current market conditions very favorable. An additional advantage of dehydrogenation technologies is that dehydrogenation is an on-purpose technique, which yields exclusively a particular olefin instead of a mixture of products. In fact, industrial dehydrogenation processes are currently optimized in such a way that they can produce olefins of polymer-quality purity. As of this writing, ca. 5 million tons of propylene is produced annually by propane dehydrogenation (PDH).16,17 However, this number is expected to increase significantly in the upcoming years as dozens of new PDH installations have been announced 10614
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vanadium oxide for oxidative dehydrogenation has been further discussed in a review by Albonetti et al.31 Another review by Cavani et al. discusses the oxidative dehydrogenation of ethane and propane on supported vanadium and molybdenum oxides. The authors conclude that, although the oxidative dehydrogenation of ethane has potential, the oxidative dehydrogenation of higher paraffins is still far from being commercially applied.32 The nonoxidative dehydrogenation of light paraffins on both supported and bulk transition metal oxides has also been discussed by Copéret in a review focused on C−H bond activation.33 Also of note are three reviews by Krylov, Wang, and Ansari, where the use of CO2 as a mild oxidant for the dehydrogenation of light alkanes is discussed.34−36 Finally, in a chapter by Caspary et al. in the most recent edition of the Handbook of Heterogeneous Catalysis, the dehydrogenation of alkanes is discussed broadly, albeit emphasis is placed on CrOx and Pt catalysts, on the kinetics and mechanism of the dehydrogenation reaction, and on reactor technology.37
2. COMMERCIAL DEHYDROGENATION PROCESSES To date, five main industrial processes for the dehydrogenation of light alkanes have been patented, two of which are currently being used to produce light olefins.38−41 As of April 2012, 5 million tons of propylene was being produced by the 14 propane dehydrogenation plants around the globe.16,17 The global demand of propylene was 103 million tons that year and is expected to increase by 4−5% on a yearly basis.42,43 Literally dozens of new installations have been planned or are already under construction, amounting to an additional 14 million tons of propylene to be produced annually by 2018.16,17,44 Additionally, six isobutane dehydrogenation installations are currently in operation, with three other plants planned.43 The majority of these new facilities will be built in China or the United States and will use either the Catofin (Lummus) or the Oleflex (UOP) technologies. A compilation of all existing and planned propane dehydrogenation installations is shown in Figure 2. Furthermore, in Table 2, a comparison is made between the two most important processes, Catofin and Oleflex. The STAR (UHDE) process will be applied for the first time in two installations in the Middle East. The other two processes, fluidized bed dehydrogenation (FBD) process (Snamprogetti and Yarsintez) and Linde-BASF PDH, have not yet been commercially applied.
Figure 2. World map showing an overview of propane dehydrogenation installations currently in operation (filled shapes) and installations that have been announced and are expected to start up before 2018 (unfilled shapes). The numbers represent the maximum production of each plant (×1000 tons per year). The locations of the facilities are approximations.16,17,44,45 For some of the newly announced PDH installations (□), the industrial process to be used remains to be announced.
reaction products.46 A schematic of a Catofin dehydrogenation installation is shown in Figure 3. The duration of the dehydrogenation step depends on the heat content of the catalyst bed, which decreases rapidly due to the endothermic nature of the reaction. Part of the heat required for the reaction is introduced to the reactors by preheating the reaction feed, additional heat being provided by adjacent reactors that are regenerating the coked catalysts. Moreover, fuel gas is added to the reactor during the regeneration step to generate supplementary heat,47 and inert material is added to the catalyst bed to increase its heat storage capacity.48 The catalyst remains in use for 2−3 years, and the progressive loss in activity with increasing time-on-stream is counteracted by gradual increases in temperature to afford a constant dehydrogenation activity throughout the entire catalyst life span.29,49
2.1. Catofin Process
2.2. Oleflex Process
The Catofin process, by CB&I Lummus, is based on the Houdry Catadiene process, which originally was exclusively used for the dehydrogenation of isobutane to isobutene. In turn, isobutene was employed for the production of methyl tertiary butyl ether (MTBE), a fuel additive used to raise the octane number of gasoline. For environmental reasons, the use of MTBE has declined in recent years, causing a shift in the use of Catofin installations to alternative purposes, such as propane dehydrogenation. A Catofin installation generally consists of 5− 8 parallel adiabatic fixed bed reactors containing a chromia− alumina catalyst. The reaction is run at temperatures of approximately 575 °C and pressures between 0.2 and 0.5 bar.26 Each reactor alternates between dehydrogenation, regeneration, and purge steps, each lasting a few minutes (15−30 min being needed for one full cycle). Each individual reactor is made to run continuously so there are always some units performing dehydrogenation reactions, while other reactors are being regenerated or purged, which results in a constant flow of
The Oleflex process by UOP uses a very different reactor design comprising fluidized bed reactors, a catalyst regeneration unit, and a product recovery section. A schematic representation is shown in Figure 4.50 The reaction is run using a Pt−Snbased catalyst at pressures between 1 and 3 bar and temperatures ranging from 525 to 705 °C. Three or four adiabatic radial flow reactors containing the catalyst are connected in series with preheaters in between. These gas flow preheaters represent the main source of heat of the reactor system. The catalyst flows through the system, the last reactor being connected to a continuous catalyst regeneration (CCR) unit that regenerates the catalyst by burning of any carbon deposits and redisperses the Pt on the support material by means of treating the catalyst with a chlorine−air mixture. The entire system is designed to operate continuously to obtain an uninterrupted stream of reaction products. The regenerated catalyst is then reintroduced to the first reactor, completing an 10615
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Table 2. Specifics of the Catofin and Oleflex Commercial Dehydrogenation Processes43,44,46,50−55 Catofin
Oleflex
catalyst formulation process license holder used for operating conditions reactor type
18−20 wt % CrOx on an alumina support, promoted by 1−2 wt % Na or K CB&I Lummus
CrOx/SiO2 > CrOx/MCM-41 ≫ CrOx/MgO ≈ CrOx/Si-2 (Si-2 is the abbreviation of silicalite-2 zeolite, which has been described as a silica analogue of the aluminosilicate zeolite ZSM-11194), a reverse order being found for the reduction temperature of the chromium species on the aforementioned catalyst supports. Cr−Mg−Al and Cr−Mg mixed oxides prepared using layered double hydroxide (LDH) precursors have been studied as catalysts for the nonoxidative dehydrogenation of ethane.195 Some of these catalysts were found to achieve ethane conversions of 27−30% and selectivity toward ethylene of 71−75%. The method used for introducing Cr into the LDH precursors was found to determine several properties of the resulting mixed oxides, such as surface area, catalytic performance, coking resistance, and thermal stability against sintering. Finally, the study of porous chromia-pillared α-zirconium phosphate materials has been particularly informative with regards to the influence that acidity has on catalytic behavior. In
4.3. Vanadium Oxide-Based Catalysts
Despite the excellent performance of Pt and CrOx catalysts in paraffin dehydrogenation, a number of issues including catalyst poisoning, the high cost of Pt, and environmental concerns associated with the use of Cr have spurred the search for alternatives. Vanadium oxides are known to be active for many hydrocarbon oxidation reactions, including dehydrogenation. Research on vanadium-based dehydrogenation catalysts started in the 1980s, although at that time the focus was on oxidative dehydrogenation using vanadium−magnesium mixed oxides.199,200 10630
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Figure 20. An artist’s impression of the surface of a CrOx/Al2O3 catalyst in which coordinatively unsaturated Cr3+ species catalyze the dehydrogenation reaction. Note that partially reduced CrOx species are formed on the surface of the catalyst, and that the Cr is partially dissolved in the alumina support. The (simplified) reaction mechanism is shown in the inset.
catalyst, as −OH groups serve as anchoring sites for the vanadium precursors used to prepare these catalysts. On the other hand, acid sites are detrimental for catalyst performance, because they facilitate cracking and induce coke formation. A schematic display of these vanadium species is shown in Figure 21.202 Several authors claim that two-dimensional polymeric V3+ or 4+ V species are the most active for dehydrogenation,203−206 the vanadium species bonded directly with the γ/θ-alumina support being the most active sites in PDH. Isolated vanadium species are also active, albeit they are susceptible to deactivation due to coke formation. On the contrary, large V2O5 crystallites are practically inactive in dehydrogenation.30,206 Prior to reaction, the vanadium in these catalysts is generally present as V5+. However, the reduction of the catalyst by the hydrocarbon feed results in the formation of V4+ and V3+.207−209 The reducibility of the vanadium species depends on their molecular structure, V−O−V and VO bonds being
The nature of the vanadium present on the catalyst depends to a large extent on the type of carrier, the support surface area, the metal loading, and the oxidation state of vanadium.201 Far below the theoretical monolayer coverage isolated vanadium species are present, at intermediate surface coverage polymeric species or sheets of vanadium oxide are favored, and above monolayer coverage V2O5 crystallites become dominant. Therefore, a combination of V−O-support, V−O−V, or V O moieties may be obtained, each having different catalytic activity. In addition, depending on the support, mixed oxides with Al, Zr, Ti, or Mg may be formed. The surface area-to-metal loading ratio determines which species will form. Indeed, Wu et al. reported that monomeric species are predominantly present below 1.2 V/nm2, polymeric vanadium species prevail between 1.2 and 4.4 V/nm2, and crystalline V2O5 species preponderate at even higher loadings. On the one hand, the presence of hydroxyl groups on the surface of the support is important to obtain a well-dispersed 10631
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Table 4. Summary of the Catalytic Data of the Cr-Based Dehydrogenation Catalysts Discussed catalyst (1) mesoporous 9 wt % CrOx/ Al2O3 (2) mixed oxide 40% Cr2O3/60% Al2O3 (3) 73 wt % Cr2O3-pillared α-ZrP
reaction temp (°C)
space velocity (h−1)
580 588 ± 2 550
(4) Ga−Cr mixed oxide on α-ZrP (20.1 wt % Cr, 9.3 wt % Ga) (5) 14−16 wt % CrOx/Al2O3
540
(6) 5 wt % CrOx/SBA-1
550
(7) 4 wt % CrOx/ZrO2
550
(8) (1 wt % Cr) Cr−Si−Zr Xerogel (9) Cr(2)Mg(12)Al(4)Ox (mixed metal oxide) (10) 5 wt % Cr, 10 wt % Ce/ SBA-15 (11) 20 wt % Cr, 1 wt % Na/ Al2O3
450
550
700 700 550
feed composition
conversion (%)a
selectivity (%)a
C3H8 = 5, He = 95
15.7−10.0
98
=
iC4H10 = 100
∼23−14
99.4
=
C3H8 = 7, He = 93
18.1−6.3
=
C3H8 = 7, He = 93
=
iC4H10 = 10, N2 = 90
WHSV = 1.2 WHSV = 0.3 WHSV = 0.86 WHSV = 3.2 GHSV = 3600 WHSV = 0.12
C3H8 = 6.67, CO2 = 33.3, He = 60 C3H8 = 2.5, N2 = 97.5 C3H8 = 1.9, N2 = 0.2, He = 97.1 C2H6 = 28.6, N2 = 71.4 C2H6 = 25, CO2 = 75 C3H8 = 10, N2 = 90
WHSV 3.3 WHSV 1.7 WHSV 2 WHSV 46.6
specific activity (s−1)b
kd (h−1)
catalyst lifec (h)
ref
0.52
1
197
6.87 × 10−4
0.48
1.25
154
87.1
1.23 × 10−5
0.24
5
162
27.6−4.4
80
5.05 × 10−5
0.42
5
196
16 (start first cycle), 8 (30th cycle) 37−26
98 ± 1
1.21 × 10−3
0.026
30
163
85−92
1.66 × 10−4
0.14
3.75
165
60.9−24.3
76.2−87.0
2.86 × 10−5
0.26
6
170
37.8−22.7 (35 and 400 m) 29.6
91.1−94.7
1.85 × 10−4
0.12
6.08
191
71
6.54 × 10−4
1
195
55−46
96−97
0.059
5
175
47−37
80−89
0.069
6
198
7.41 × 10−6
a First value is obtained at the start of the cycle, second at the end. bSpecific activity is defined as (mol olefin formed)/(mol Cr*t(s)). cCatalyst life = total time single cycle/experiment.
Figure 21. Various vanadium oxide species present on a catalyst support (S): isolated vanadium oxide species (a), dimers (b), two-dimensional chains of vanadium oxide species (c), and V2O5 crystallites (d). Reprinted from ref 201, copyright 2003, with permission from Elsevier.
Ogonowski et al. reported that the effect of CO2 depends strongly on the catalyst support. Relatively weak basic sites on carriers, such as those present on an activated carbon support, facilitate the reverse water gas shift reaction, which lifts the thermodynamic limitation of the reaction and increases alkane conversion. When these basic sites are too strong, as is the case on Al2O3 and ZnO2, CO2 irreversibly adsorbs and poisons these sites. This increases the acidity of the support, which in turn increases coke formation and lowers selectivity. If these basic sites are absent, as in the case of SiO2, no beneficial effects of CO2 are observed. Alumina is commonly used as a support for vanadium oxide catalysts, although Sokolov et al. concluded through testing a wide range of Si- and Al-based supports that using silica to dope the alumina support provides beneficial effects on catalyst stability and activity.214 Furthermore, as was shown by Harlin et al., the performance of supported vanadium catalysts can be improved through the addition of MgO, either as a promoter or as a component of a mixed oxide. According to these workers, the increase in dehydrogenation selectivity stems from the basic nature of MgO, resulting in reduced coke formation.208 An overview of the different VOx catalysts used for the dehydrogenation of light alkanes is shown in Table 5.
more easily reduced than V−O-support bonds. Harlin et al. studied the effects of vanadium oxide reduction with H2, CO, and CH4 prior to the dehydrogenation reaction. In all cases, the reduction was incomplete, and yielded a mix of vanadium oxide species, including V3+, V4+, and V5+ (see Figure 22). Higher initial catalytic activities were obtained in all cases, although the effect was strongest after reduction with CO. The authors suggested that when the catalyst was reduced with H2 and CH4, detrimental −OH groups were formed. V4+ was determined to be the most active species in butane dehydrogenation. Initially, deactivation is caused by the strong adsorption of reactants on the active sites and to a lesser extent to coke deposition. When the coke is removed through the calcination of the catalyst, the latter may deactivate further as vanadium oxide species sinter and form larger V2O5 crystallites or mixed oxides such as AlVO4, both of which are less active.206 V2O5 crystallites can revert to active polymeric vanadium species by treating the catalyst under an oxygen atmosphere at 600 °C.210 In several studies, CO2 is added to the feed as a mild oxidant to increase alkene yields.211−213 Different reasons for this beneficial effect of CO2 have been suggested, such as the removal of acid sites that catalyze unwanted side reactions, coke gasification, oxidation of overly reduced vanadium centers, and hydrogen removal by the reverse water−gas shift reaction. 10632
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10633
600
As some of these catalysts require an activation period, the maximum (as opposed by the initial) conversion is considered to determine the deactivation rate. bFirst value is obtained at the start of the cycle, second at the end. cSpecific activity is defined as (mol olefin formed)/(mol Cr*t(s)). dCatalyst life = total time single cycle/experiment. eCatalyst is prereduced by CO prior to the dehydrogenation reaction. a
215 3 0.86 12−30 to 1- and 1−3 butene 70−15
212 0.72
i-C4H10 = 14.3, CO2 = 85.7 GHSV = 72 (at a pressure of C4H10 = 3, N2 = 97 0.5 bar) WHSV = 1.9 600
45.5−26.5
76.4−85.0
5.77 × 10−4
1.17
208 1.34; 3.12 9.32 × 10−4; 9.40 × 10−4 85−90 77−88 (to ibutene) 44−36; 54−35 WHSV = 5 580
(4) 4.9 wt % VOx−4.7 wt % Mg/Al2O3e (5) 5.4 wt % VOx−4.5 wt % Zr/Al2O3e (6) 4 wt % V−36 wt % MgOx/ Cact (7) 3.5 wt % VOx/Al2O3
62−38 WHSV = 5
i-C4H10 = 10, N2 = 90 i-C4H10 = 10, N2 = 90 580 (3) 5.2 wt % VOx/Al2O3c
32.9−8.5 (15−60 m) GHSV = 14 400 600
WHSV = 2.6, N2 = 90
(2) 3.5 wt % VOx/θ-Al2O3
nC4H10 = 100
0.25
207 0.085 11.5 9.46 × 10−4
205 1 2.22
202 1.33 1.04 11−3 nC4H10 = 10
55−76 (to all butene isomers) 56.1 (all butene isomers) 65−85 (to i-butene)
1.75 × 10
−4
520
conversion (%)b feed composition
(1) 2.2 wt % VOx/γ-Al2O3
Molybdenum oxides are also frequently used as catalysts in hydrocarbon conversion reactions, including dehydrogenation. To our knowledge, the first report of the dehydrogenation activity of a molybdenum oxide-based catalyst was of MoO3/ Al2O3 in the dehydrocyclization of n-heptane, which was published in 1946.216 Since that first report, the nonoxidative dehydrogenation of alkanes on MoOx has been studied scarcely, as emphasis has been given to the study of oxidative dehydrogenation on these materials. The chemistry of molybdenum oxide is comparable to that of vanadium oxide, as Mo can be present as MoOx monomers, polymers, or MoO3 crystallites, depending on the molybdenum loading, the type of support used, and catalyst preparation conditions.217 Harlin et al. noted that optimal activity was achieved with monolayer coverage of molybdenum oxide on an alumina support. In the fresh catalyst, molybdenum is present as Mo6+, but analogously to catalysts comprising vanadium and chromium oxide, MoO3 is reduced by the hydrocarbon feed during reaction to create the active species.218 These reduced sites are believed to be Mo4+ and Mo5+. Further reduction results in sites that are more active, but the latter catalyze cracking and deactivate fast due to coke deposition. In fact, at
space velocity (h−1)
4.4. Molybdenum Oxide-Based Catalysts
reaction temp (°C)
Figure 22. Vanadium 2p3/2 XPS spectra of a VOx/Al2O3 catalyst calcined (top) and reduced with CO (bottom). Reduction yields a mixture of V5+, V4+, and V3+. Reprinted from ref 207, copyright 2000, with permission from Elsevier.
catalyst
Table 5. Summary of the Catalytic Data of the V-Based Dehydrogenation Catalysts Discusseda
selectivity (%)b
specific activity (s−1)c
kd (h−1)
catalyst lifed (h)
ref
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loadings above monolayer coverage, the catalyst is more easily reduced, which results in a drop in catalyst activity.218,219 The same authors also point out that such a catalyst deactivates due to the formation of catalytically inactive Al 2 (MoO 4 ) 3 crystallites. To prevent the formation of this mixed oxide, they added magnesium oxide to obtain a molybdenum− magnesium mixed oxide that also shows dehydrogenation activity. It was found that an additional beneficial effect of magnesia addition is that the resulting mixed oxide has lower acidity, which increases alkene selectivity and reduces coke deposition. The reduction of the catalyst with hydrogen prior to reaction increases conversion, although selectivity decreases due to the formation of acid sites.220 Overall, MoO3 catalysts deactivate rapidly due to coking and generally afford a relatively low conversion and selectivity. The formation of molybdenum oxycarbides (MoOxCy) has been reported to occur during butane dehydrogenation on unsupported molybdenum oxide. These materials are also active in the dehydrogenation of paraffins, and a high activity close to the thermodynamic equilibrium is initially obtained, although over a period of a few hours the selectivity drops to zero and the reaction yields only methane, as shown in Figure 23. During this period, the oxycarbide becomes depleted from oxygen and a wide range of carbides active in hydrogenolysis are formed.222 To improve the stability of the oxide, steam can be added to the feed (which hampers the growth of these carbides) or the catalyst can be doped with vanadium (which increases the oxophilicity of the catalyst). X-ray diffractograms showing this transition are shown in Figure 24. However, the introduction of steam is not free of drawbacks, as it induces sintering. Hydrogen addition prevents coke formation and increases conversion, albeit it is believed that H2 competitively adsorbs on the active sites of the catalyst.221−223 Consistent with the fact that MoOxCy is a catalyst active in butane dehydrogenation, Harlin et al. showed that MoO3 deposited on SiC displays significant catalytic properties when hydrogen and steam are present in the feed.224 As compared to bulk MoO3 and MoO3 supported on SiO2 and Al2O3, higher selectivities and stability are obtained over MoO3/SiC, albeit at the cost of a lower conversion. This was explained by the increased reducibility of MoO3/SiC as compared to the MoO3/Al2O3 catalyst. A summary of the Mo-based catalysts used for the dehydrogenation of light paraffins is shown in Table 6. 4.5. Gallium Oxide-Based Catalysts
Figure 23. Yield of different carbon molecules during butane dehydrogenation on a MoOxCy catalyst at different steam partial pressures. When no steam is added, the selectivity to C4 alkenes drops rapidly and the methane yield reaches 100% after 5 h. The hydrogenolysis reaction is suppressed when steam is added to the feed. Reprinted from ref 221, copyright 1999, with permission from Elsevier.
The dehydrogenation activity of gallium oxide supported on ZSM-5, which was first reported in the late 1980s in studies dealing with conversion of propane to aromatics, has been the target of renewed attention in recent years.225−228 In the interim, both bulk and supported Ga2O3 have been used as dehydrogenation catalysts. The exact nature of the active sites in Ga-based dehydrogenation catalysts is still under debate. Akin to Al2O3, Ga2O3 has different polymorphs of which β-Ga2O3 is not only the most stable, but also the most active Ga species in PDH according to some authors.229 β-Ga2O3 has a monoclinic structure, in which gallium atoms are equally distributed between tetrahedral and octahedral configurations. This crystal structure has a high concentration of weak Lewis acid sites (coordinatively unsaturated tetrahedral ones on the surface of the material), which are deemed to be centers of dehydrogenation activity.230 Indeed, a heterolytic dissociation reaction mechanism has been
proposed where H− adsorbs on a Ga+ Lewis acid site and C3H7+ on a neighboring oxygen, although these species can also be adsorbed reversed.33,231 The assignment of Lewis acid sites as the active species was substantiated by Chen et al.,232 who, by relating the results of NH3-TPD measurements with activity data for different GaxAl10−xO15 mixed oxides, showed that a high concentration of Lewis acid groups in the form of coordinatively unsaturated tetrahedral Ga3+ cations is a prerequisite for dehydrogenation activity (see Figure 25). Xu et al. reported dehydrogenation activity on both Lewis and Brønsted acid sites on GaO3/H-ZSM-5. A zeolitic support displaying a low concentration of medium and strong acid sites and a relatively high concentration of weak acid sites was found to lead to the most active and stable catalyst.233,234 In addition, Michorczyk et al. concluded that acidic groups catalyze the 10634
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Figure 24. XRD diffractograms of the fresh Mo oxycarbide catalyst before dehydrogenation, but after pretreatment (top left), as well as after dehydrogenation reaction with no steam (top right), with 5 Torr of steam (bottom left), and with 18 Torr of steam (bottom right). In the absence of steam, the well-defined MoO2 and MoOxCy species disappear during reaction and are replaced by more amorphous MoOxCy species, which are enriched with carbon and presumably responsible for hydrogenolysis. The addition of steam prevents this transformation and with 18 Torr of steam, MoO2 and ordered MoOxCy species are still observed on the spent catalysts. Reprinted from ref 221, copyright 1999, with permission from Elsevier.
Table 6. Summary of the Catalytic Data of the Molybdenum Oxide-Based Dehydrogenation Catalysts Discusseda catalyst
reaction temp (°C)
space velocity (h−1)
feed composition
(1) 13.4 wt % MoO3/Al2O3
560
WHSV = 2
n-C4H10 = 10, N2 = 90
(2) 12.8 wt % MoO3−8.4 wt % Mg/Al2O3 (3) MoO3
550
WHSV = 2
(4) 14.2 wt % MoOx/SiC
560
WHSV = 2
n-C4H10 = 9.3, N2 = 83.9, H2O = 6.8 n-C4H10 = 10, N2 = 90, H2O = 3.1
conversion (%)b 21−18 (second min 25) 13−11
specific activity (s−1)c
kd (h−1)
catalyst lifed (h)
ref
50
7.21 × 10−5
2.78
0.17
218
82
7.66 × 10−5
1.12
selectivity (%)b
−5
42−8
55−84
2.96 × 10
0.28
7.5
221
22−17
38−71
5.41 × 10−5
0.32
1
224
a
As some of these catalysts require an activation period, the maximum (as opposed to the initial) conversion is considered to determine the deactivation rate. bFirst value is obtained at the start of the cycle, second at the end. cSpecific activity is defined as (mol olefin formed)/(mol Mo*t(s)). dCatalyst life = total time single cycle/experiment.
high energy barrier. Nevertheless, theoretical calculations by Pidko et al. indicate that these relatively stable [H−Ga−OH]+ sites are preferentially regenerated by water desorption resulting in a reduced active site.236 Murata and co-workers have observed an increase in activity after a prereduction step with hydrogen at 550 °C, while a simultaneous increase of Gaδ+H and Ga−OH sites was observed via NH3 TPD. These authors proposed that these sites also form during reaction and may be responsible for the dehydrogenation activity of the silica supported Ga2O3 catalyst.237,238 Contradictory information has been reported regarding the reducibility of Ga2O3. Rodriguez et al. reported that Ga2O3/ Al2O3 is not reduced by hydrogen at 700 °C, but the presence of Pd on the catalyst can assist the reduction of Ga3+ to Ga+ and
reaction after observing that promotion with a base, such as potassium, resulted in a decrease in activity.235 Several authors have performed DFT calculations in an effort to understand the dehydrogenation reaction mechanism on Ga2O3. Studies by Liu et al. regarding propane dehydrogenation on a perfect Ga2O3 (100) surface revealed that the C−H bond is activated by a surface oxygen that abstracts an hydrogen from the propane molecule, thus forming a propyl species, which coordinates with a Ga surface site and a hydroxyl group. A second hydroxyl group is formed via β-hydrogen elimination, but for these hydroxyl groups to recombine, either as H2O or as H2, a high energy barrier needs to be overcome. This barrier is much lower over a gallium hydride−hydroxide species, although the formation of the gallium hydride also involves a 10635
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with CrOx (vide supra), where a weak metal−support interaction results in an increased reducibility. Indeed, Meitzner et al. showed that Ga2O3/H-ZSM-5 could be reduced by hydrogen or hydrocarbons at dehydrogenation reaction conditions. Finally, regarding the reduction of Ga2O3, it is important to note that Ga2O is gaseous and metallic gallium is liquid at dehydrogenation reaction conditions.240 A number of studies have shown that the addition of CO2 to the reactant feed improves the olefin yield.241−245 Nakagawa et al. found that the addition of steam and CO2 decreased coke deposition on a Ga2O3/TiO2 catalyst by the Boudouart reaction, while also assisting ethylene desorption.241,242 However, other authors claimed that the increase in activity was caused by the reverse water−gas shift reaction (RWGS).243,244 Additionally, Michorczyk et al. noted that this effect could only be seen at relative low levels of coke deposition.235 It was suggested that CO2 could also poison the basic sites required for the dissociative adsorption of the alkane. Additionally, Hensen et al. studied the promoting effect of water with experiments and theory and observed an increase in the conversion of propane by cofeeding water.246 The authors attributed this promoting effect to the formation of partially hydrolyzed gallium species, more specifically, to the formation of binuclear hydroxyl-bridged Ga3+ cations. Different carriers have been used in gallium oxide-based dehydrogenation catalysts, the inclusion of acid and basic sites on the support being important to achieve a high conversion. Xu et al. reported that in the absence of an oxidizing agent, the conversion obtained using Ga2O3 supported on ZrO2, Al2O3, and TiO2 was high, medium, and low, respectively, while SiO2 and MgO-supported Ga2O3 were inactive in alkane dehydrogenation.233 The authors proposed the activity to be dependent on the presence of acid sites of medium strength on the catalyst surface. However, a different trend was observed when CO2 was present, as conversion was observed to be high on Ga2O3/ TiO2 and low on Ga2O3/ZrO2 and Ga2O3/Al2O3. This was attributed to the two contrary roles of CO2, which on the one
Figure 25. Concentration of Lewis acid sites versus propane conversion measured 15 min after the start of the PDH reaction on GaxAl10−xO15. A clear correlation between the conversion and the surface density of weak Lewis acid sites is observed. For these experiments, CO2 is used to improve conversion. Reprinted from ref 232, copyright 2008, with permission from Elsevier.
Ga0 at temperatures between 300 and 550 °C and form an alloy (in which Ga0 is immobilized). Indeed, it has been reported that the reduction of Ga2O3 can be promoted by the proximity of other, more easily reduced metals, such as Pt and Pd.131−133,238 For instance, on a Pt−Ga/H-beta catalyst, reduced gallium hydrides were observed after hydrogen reduction at 550 °C.239 It has also been proposed that the reducibility of these species depends on the interaction of gallium oxide with the catalyst support, which is also the case
Table 7. Summary of the Catalytic Data of Gallium Oxide-Based Dehydrogenation Catalysts Discusseda reaction temp (°C)
space velocity (h−1)
(1) 1.7 wt % Ga2O3/SiO2
550
(2) Ga2O3/MTS (mesoporous silica) ratio Ga/Si = 0.05 (3) β-Ga2O3
550
(4) Ga2O3
600
(5) 5 wt % Ga2O3/ZrO2
600
WHSV = 0.97 h−1 WHSV = 0.33 h−1 WHSV = 0.15 h−1 WHSV = 1.2 h−1 WHSV = 0.3 h−1
catalyst
500
(6) 5 wt % Ga2O3/TiO2 (7) Ga8Al2O15
(8) 5 wt % Ga2O3/H-ZSM-5 (w Si/Al = 97) (10) 0.001 wt % Pt, 1.2 wt % Ga/Al2O3
500
650 590
WHSV = 0.3 h−1
WHSV = 0.36 h−1 GHSV = 400 h−1
feed composition
selectivity (%)b
conversion (%)b
specific activity (s−1)c
kd (h−1)
catalyst lifed (h)
ref
C3H8 = 10, Ar = 90
23−20
86−80
4.97 × 10−4
0.036
5
249
i-C4H10 = 10, He = 90 C3H8 = 2.5, N2 = 97.5 C3H8 = 17, CO2 = 83 C3H8 = 2.5, N2 = 97.5 C3H8 = 2.5, CO2 = 5, N2 = 92.5 (a) C3H8 = 2.5, N2 = 97.5 (b) C3H8 = 2.5, CO2 = 5, N2 = 92.5 C2H6 = 3, CO2 = 15, Ar = 82 C3H8 = 100
46−40
58.3 to i-C4H8
5.52 × 10−5
0.092
2.67
230
33−12
95−95
6.95 × 10−7
0.21
6
229
31−3
95−82
3.55 × 10−5
0.67
4
235
39−6; 32−8
74
1.90 × 10−5
0.77
3
233
73
1.54 × 10−5
0.56
51.7−22.5
91.6−98.2
2.39 × 10−6
0.168
8
232
49.7−33.1e
91.7−98.0e
2.30 × 10−6
0.089
23.7−14.5
85.9−93.7
2.84 × 10−5
0.0086
70
244
39
86
0.25
247
a
As some of these catalysts require an activation period, the maximum (as opposed to the initial) conversion is considered to determine the deactivation rate. bFirst value is obtained at the start of the cycle, second at the end. cSpecific activity is defined as (mol olefin formed)/(mol Ga*t(s)). dCatalyst life = total time single cycle/experiment. eValues obtained at 15 min and 8 h. 10636
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Chemical Reviews
Review
activity. The authors hypothesized that the carbon species formed on the catalyst surface were in fact catalyzing the reaction. The highest activity was observed at 700 °C, a temperature at which the carbon deposits display a very ordered structure. In view of this, the authors decided to test their hypothesis by using carbon nanofibers (CNFs) as a catalyst under similar conditions. Tellingly, CNFs proved to be catalytically active. It was concluded that radicals formed on the surface of coke facilitate the dehydrogenation reaction, albeit at elevated temperatures the noncatalytic conversion of propane is also relatively high. However, elevated temperatures also lead to a drop in selectivity, as cracking becomes the dominant reaction. Interestingly, the latter was not the case in the work of McGregor et al., who reported a selectivity toward butenes of ∼80%. Furthermore, hybrid nanocarbon, comprising a diamond core and a graphitic shell, displayed activity for the dehydrogenation of propane, which was attributed to the presence of either ketone groups or structural defects in the graphitic material.251 In previous work, CNFs had been shown to be active in the oxidative dehydrogenation of n-butane, although in this case surface oxygen groups were believed to be the active sites.252 Furthermore, nitrogen doped carbon nanotubes (CNTs) showed some activity in the oxidative dehydrogenation of propane, although conversion and selectivity were relatively low.253 Liu et al. prepared carbon-based ordered mesoporous systems that showed activity in the dehydrogenation of propane in the absence of steam or oxidants.254 The active sites were proposed to be oxygen-containing groups (CO), which are formed during a pretreatment step involving nitric acid. The catalyst shows remarkable long-term stability, displaying conversion and selectivity values of 45% and 85%, respectively, after 100 h of time on stream at 600 °C. Experiments in which carbon nanotubes were tested under similar conditions resulted in low activity. Neylon et al. studied the activity of supported early transition metal (groups 5 and 6) nitrides and carbides in butane dehydrogenation,255,256 as the catalytic properties of these materials are believed to be similar to those of platinum. Interestingly, Mo, V, and W carbides and nitrides showed activity in both dehydrogenation and hydrogenolysis, displaying specific activities similar to those of a commercial Pt−Sn/Al2O3 catalyst. Niobium carbide and nitride did not show any activity. The group 6 compounds tested were found to be more active, albeit they primarily catalyzed hydrogenolysis. In Table 8 are shown the catalytic data of the carbon-based dehydrogenation catalysts as reported by McGregor et al. and Liu et al. In addition, silica supported metal sulfides, such as ZnS, FeS, CuS, CoS, MnS, NiS, and MoS, were tested for the dehydrogenation of isobutane, and were found to perform significantly better than their respective oxides. The materials show higher initial yields than commercial CrOx and Pt−Sn catalyst, but they deactivate over a few hours due to sulfur leaching.257
hand removes H2 from the surface of the catalyst by the reverse water−gas shift reaction, while on the other hand absorbs strongly on the basic sites of the support. Thus, the low conversion observed when zirconia or alumina were employed as the support can be ascribed to the relatively high amount of basic sites on these oxides. Notably, the structure of the support is also important, as a mesoporous Ga−Al mixed oxide exhibited an increased resistance to catalyst coking.245 Work on the promotion of GaOx-based dehydrogenation catalysts appears to be limited, although a patent by Iezzi et al. claims that a catalyst comprising 1.2 wt % Ga2O3,
