JP4406541B2 - Dehydrogenation process for olefins - Google Patents

Description

【００４３】
表２：イソブタン脱水素結果

【００４４】
実施例１、３、４および６についての結果を選択し、図１から図５に図示した。図１、２および３から、供給原料中の酸素のイソブタンに対する比が最大Ｏ₂／イソブタンモル比０．０２７５まで増加するに従って、イソブタン転化が低下する一方、イソブチレン選択率および収率は増加することが分かる。更に酸素濃度が増加すると、転化および選択率共に減少する。前述したＯ₂濃度での操業では炭化水素供給原料であるイソブタンを約５％節約するので、この結果は非常に重要である。加えて、当該発明のように酸素の添加は継続的なイソブタン節約およびより長い反応時間におけるイソブチレン生産性向上ができ（実施例２および５）、コークス生成減少により反応時間を長くできる。
【００４５】
図４は、Ｏ₂／イソブタンのモル比が約０．０２７５において最も低いコークス選択率が生じることを示した。より高いまたはより低いＯ₂／イソブタンのモル比は、より高いコークス選択率を生じるようである。図５は、実施例１，３および４で（または実施例２および５における）与えた水素ピーク面積は、酸素がオレフィン生成を促進する範囲においては酸素量に関わらず殆ど同じであることを示す。当該水素ピーク面積は追加の酸素を添加すると減少し、イソブタン転化およびイソブチレン選択率は降下する（実施例６参照）。更なる酸素の添加は反応平衡により予想されるように、当該反応を生成物側に移行する。
【００４６】
当該実施例は明らかにクロムレドックス特性の選択的制御剤としての酸素の機能を示す。クロムの酸素分子による酸化は、水素または触媒表面におけるコークス活性種の酸化より相当早く進行することを示す。
【００４７】
より低いＷ／Ｆ比におけるイソブタン脱水素の運転は、実施例７，８および９における生産性並びに供給原料消費速度の数字から分かるように当該発明のプロセスを用いる際でも魅力的ではない。
【００４８】
本発明の好ましい実施態様をかくして詳細に説明したが、添付した特許請求の範囲で定義した発明は、本発明の本質または範囲から外れることなくそれを多く変形することができるので、本明細書において述べた特定の詳細に限定されることはないことは理解すべきである。
【図面の簡単な説明】
【図１】図１は、当該発明のプロセスにおいて、酸素濃度の関数としてイソブタン転化を表した図である。
【図２】図２は、当該発明のプロセスにおいて、酸素濃度の関数としてイソブチレン選択率を表した図である。
【図３】図３は、当該発明のプロセスにおいて、酸素濃度の関数としてイソブチレン収率を表した図である。
【図４】図４は、当該発明のプロセスにおいて、酸素濃度の関数としてコークス選択率を表した図である。
【図５】図５は、当該発明のプロセスにおいて、酸素濃度の関数として水素ピーク面積カウントを表した図である。[0001]
BACKGROUND OF THE INVENTION
The present invention relates to a process in the production of olefins by the normal dehydrogenation of the corresponding paraffin. More particularly, the present invention relates to an improved dehydrogenation process for the production of olefins using a chromium-based dehydrogenation catalyst in the presence of oxygen.
[0002]
[Prior art]
Hydrocarbon dehydrogenation is an important industrial process. This is because there is a great demand for dehydrogenated hydrocarbons as feedstock for industrial processes. For example, dehydrogenated hydrocarbons are used inter alia for the production of surfactants, high octane gasoline, pharmaceutical products, plastics and synthetic rubbers. An example of one specific dehydrogenation process is the dehydrogenation of isobutane to produce isobutylene, which is then polymerized to provide tackifiers for adhesives, viscosity index additives for motor oils, and anti-impact additives for plastics Provided as an agent and antioxidant additive. Isobutylene is likewise used in the production of methyl tertiary butyl ether (MTBE) or can be dimerized to isooctane. MTBE and isooctane are used as anti-knock agents in gasoline fuel.
[0003]
The main traditional sources of alkenes have been steam cracking and fluid catalytic cracking. Both processes simultaneously provide a limited range of other broad products. However, there is an increasing demand for certain alkenes such as isobutylene or propylene, and it is not cost effective to use expensive crackers to satisfy the demand. The best technical choice for supplying alkenes was the normal dehydrogenation of alkanes.
[0004]
The main factors affecting the technical implementation of alkane dehydrogenation are the thermodynamic equilibrium that limits the single pass conversion and the endothermic properties of the dehydrogenation reaction. In addition, the temperature required for conventional dehydrogenation causes thermal decomposition, which reduces alkene selectivity, particularly in propane dehydrogenation.
[0005]
Two types of catalysts based on chromium oxide for the dehydrogenation of lower alkanes have been described in the scientific and patent literature: oxidation supported on γ, δ /, θ-alumina doped with alkali metals Chromium oxide supported on chromium and ZrO ₂ . Chromium oxide supported on γ, δ /, θ-alumina with added alkali metal is used in the fluidized bed dehydrogenation process of Lumnus Catofin and Snamprogeti-Yarsintez. On the other hand, chromium oxide supported on ZrO ₂ was studied for higher thermal stability than alumina.
[0006]
The nature of the active site of chromium oxide supported catalysts has been discussed for many years. Cr ₂ O ₃ is the most stable form of all chromium oxides.
[0007]
Two types of Cr ⁶⁺ species have been found in chromium oxide supported on alumina, after calcination and before reaction. After reduction with hydrogen, it is generally considered that all Cr ⁶⁺ is reduced to Cr ³⁺ . However, according to Grunert et al. (J. Catal. 110 (1986), 138), the reduction occurs in two steps: Cr ⁶⁺ to Cr ³⁺ is a very fast step, followed by lower oxidation states from Cr ^3+. A slower process follows (Cr ²⁺ ). The active site of the dehydrogenation reaction is described by Grunert, W. et al. (J. Catal. 99 (1986), 149; Delmon, B. et al. (J. Catal. 24 (1972), 336)) and Konig, P. et al (Acta Chim. Acad. Sci. Hung. 76 (1976). , 123) according to Cr ⁺³ , Ashmawy, FM (J. Chem. Soc., Faraday Trans. 76 (1980), 2096), both Cr ²⁺ and Cr ³⁺ , Lunsford, H (J. Catal. 91 (1985), 155), it was presumed to be coordinatively unsaturated Cr ²⁺ .
[0008]
It has been found that the activity in the dehydrogenation of isobutane and also ethane is proportional to the chromium content whatever the species is present. Similarly, in the case of propane dehydrogenation with ZrO ₂ supporting chromium oxide (Indovina et al., Appl. Catal. 81 (1992), 113), the activity per atom of chromium is the same as the total amount of chromium. It was found. The activity was thought to be due to the presence of mononuclear Cr ³⁺ species.
[0009]
However, since hydrogen is present as part of the dehydrogenation reaction product, further reduction of Cr ³⁺ will continue. In the present invention, it is shown that the lower oxidation chromium species causes undesired decomposition reactions during dehydrogenation and causes coke formation. In order to minimize the formation of cracked products and coke, the oxidation state of the chromium catalyst sites must be controlled during the reaction cycle by a functional redox system. The redox system according to the present invention is believed to be Cr ²⁺ / Cr ³⁺ / O _{2 as} shown by oxygenation of hydrocarbons in the feedstock.
[0010]
According to the literature, oxygen can change the thermodynamic equilibrium of the dehydrogenation reaction in the direction of increasing olefin production by reaction with the hydrogen product. It is believed that oxygen likewise burns coke and eliminates coke deposition from the catalyst surface. Since coke does not accumulate and as a result the catalyst is not deactivated, regeneration of the catalyst can be avoided. Such a reaction is called an oxidative dehydrogenation reaction. Until recently, there was no commercial process available for producing light alkanes, especially C ₃ and C ₄ -olefins, by oxidative dehydrogenation as there was no suitable catalyst. Known oxidative dehydrogenation processes are described in several patent publications.
[0011]
For example, U.S. Pat. No. 4,996,387 contacts a portion of the catalyst with an oxygen mixture containing regeneration gas and diluent, and contacts the remainder of the catalyst with a hydrocarbon feedstock and diluent. A dehydrogenation process with continuous regeneration of the dehydrogenation catalyst is disclosed which is carried out in a circulating manner. In this disclosure, oxygen is added for the purpose of catalyst regeneration in a manner separate from the dehydrogenation medium.
[0012]
In Canadian Patent 912,051, paraffins and olefins are dehydrogenated with oxygen and halogen in the presence of a solid catalyst containing an alkali metal or alkaline earth metal compound and a promoter. Describes the vapor phase process. U.S. Pat. No. 3,697,614 uses molten alkali metal hydroxide-containing alumina, preferably dichromate, molybdate, tungstate, manganate, permanganate, ferric acid. It relates to the production of olefins by oxidative dehydrogenation by solution transition metal oxygenation consisting of salts and metavanadate.
[0013]
U.S. Pat. No. 4,046,833 discloses a vapor phase process for dehydrogenating paraffinic hydrocarbons containing from 3 to 6 carbon atoms to the corresponding monoolefin, the process comprising oxygen And in the presence of an oxidative dehydrogenation catalyst containing vanadium and aluminum. Effective paraffin to oxygen ratios are claimed in the range of 1: 0.04 to 1:10. The examples in the mentioned invention only show the effect of several catalyst compositions tested under different conditions with different oxygen concentrations.
[0014]
U.S. Pat. No. 4,788,371 discloses another oxidative dehydrogenation process using a dehydrogenation catalyst comprising at least one noble metal component. According to the disclosure, an alkane feed containing oxygen-containing gas is introduced into a dehydrogenation reactor containing the noble metal catalyst. The mole ratio of oxygen added to the reactor feed to alkane is very similar to that of U.S. Pat. No. 4,046,833 (and the ratio used in the present invention intended to be described below). Is different). The effect of oxygen on the description system is to burn hydrogen while minimizing the burning of valuable hydrocarbons.
[0015]
Numerous processes involving removal of hydrogen from a mixture of hydrogen and one or more organic compounds have been disclosed as well: for example, US Pat. No. 4,788,371 is incorporated herein by reference. Discloses a process for hydrocarbon dehydrogenation, in which hydrogen obtained by hydrocarbon dehydrogenation is catalytically reacted with oxygen. The disadvantage of all of these processes is that some of the oxygen gas reacts chemically with organic compounds, not just the hydrogen, and is converted to undesirable products.
[0016]
European patents A1-0219271 and A1-0219272 likewise disclose a process for hydrocarbon dehydrogenation, in which the hydrogen obtained by hydrocarbon dehydrogenation is removed. In these processes, the dehydrogenation occurs in the presence of a zeolite catalyst, and the hydrogen is removed by chemical reaction with an acidic oxide such as sulfur dioxide or nitrous oxide. These processes do not share the disadvantages of processes using oxygen gas, since sulfur dioxide and nitrous oxide are not as reactive as oxygen is to organic compounds. These processes also appear to be less effective at removing hydrogen.
[Problems to be solved by the invention]
[0017]
The oxidative dehydrogenation process described above is totally different from the process of the present invention described below. They generally have the disadvantage of low olefin selectivity and have not been used on a commercial scale to date due to the lack of suitable activity and selective catalysts, and the oxygen concentrations utilized so far in these processes are It is not consistent with the invention.
[0018]
In the production of olefins by catalytic dehydrogenation of paraffin, it is of course desirable to obtain as high an olefin yield as possible in a single conversion pass. Also, in order to minimize the hot spot effect and consequently increase the life of the catalyst, it is desirable to perform the reaction under conditions that produce a minimum amount of coke on the catalyst. It is also desirable to increase the duration of the dehydrogenation reaction by reducing coke formation.
[0019]
The advantages of the present invention are to improve olefin yield and selectivity while reducing thermal cracking, reducing feed consumption, and maintaining high paraffin conversion while extending catalyst life.
[0020]
The object of the present invention is to provide a dehydrogenation process for the production of olefins with high alkane conversion and improved olefin selectivity.
[0021]
Another object of the present invention is to provide a process for hydrocarbon dehydrogenation for the production of olefins with high yield, less feedstock consumption and less coke production.
[0022]
Another object of the present invention is to increase the cycle time of the dehydrogenation reaction.
[0023]
Other objects and advantages of the present invention will be apparent from the following description.
[0024]
[Means for Solving the Problems]
The present invention is an improved process for the production of olefins from the corresponding alkanes by conventional dehydrogenation and in particular for improved hydrocarbon dehydrogenation with oxygen and chromium-based dehydrogenation catalysts. Provide a processed process. The chromium-based dehydrogenation catalyst useful in this process has been found to be surprisingly selective for hydrogen oxidation in the presence of small amounts of oxygen, increasing the yield of the desired dehydrogenated product. . Accordingly, the process of the present invention broadly aims at a hydrocarbon conversion process, particularly a hydrocarbon dehydrogenation process with a chromium-based dehydrogenation catalyst using a small amount of oxygen.
[0025]
In the hydrocarbon dehydrogenation process of the invention, a feedstock comprising a dehydrogenable hydrocarbon and a small amount of pure oxygen in a molar ratio of the hydrocarbon to oxygen between about 1: 0.0001 and 1: 0.04. In order to promote dehydrogenation, a dehydrogenation catalyst containing 10 to 20% by weight of chromium is used. According to an example of a preferred embodiment of the invention, the dehydrogenable hydrocarbon is isobutane and provides the amount of oxygen between 0.0001 and 0.001 mole per gram of chromium in the catalyst; The reaction temperature is between 560 ° C and 600 ° C.
[0026]
In further examples of embodiments of the present invention, air or other oxygen-containing gas is used instead of pure oxygen.
[0027]
DETAILED DESCRIPTION OF THE INVENTION
The present invention relates to an improved process in the production of olefins and in particular isobutylene by the usual dehydrogenation of the corresponding alkanes. The process aims to minimize side reactions occurring during the dehydrogenation reaction by controlling the oxidation state of the chromium sites in the dehydrogenation catalyst. This was achieved by adding a small amount of oxygen. Oxygenation such as the present invention provides an in situ heat source for the reaction, reduces coke formation, increases olefin selectivity, and lengthens the catalytic cycle of dehydrogenation. In other embodiments of the invention, air or oxygen-containing gas can be used instead of oxygen.
[0028]
The dehydrogenation catalyst useful in the present invention comprises 10-20% by weight of chromium, and preferably 12-18% by weight of chromium. The chromium-based dehydrogenation catalyst may be spray-dried, pelletized and molded, and is preferably spray-dried. The catalyst used in this process can be used alone or supported or immersed in a support material. Suitable carrier materials include alumina, silica, tria, zirconia, titania, boron phosphate, silicon carbide, pumice, diatomaceous earth, clay and the like. The catalyst of the invention is preferably calcined to obtain desirable physical properties such as frictional resistance, optimal surface area and particle size.
[0029]
Dehydrogenation of a dehydrogenable paraffinic hydrocarbon such as isobutane is achieved in the present process by contacting the dehydrogenable paraffinic hydrocarbon with the above-described catalyst under conditions suitable for normal dehydrogenation. The conditions include a temperature in the range of about 400 ° C. to about 700 ° C., preferably about 500 ° C. to about 650 ° C. and more preferably about 560 ° C. to about 600 ° C. and a reaction pressure in the range of about 0.1 to about 10 atmospheres. Including.
[0030]
The oxygen concentration in the process of the invention is very important to obtain the explained benefits of the invention. In one embodiment, the amount of oxygen required for the dehydrogenation of isobutane is between 0.0001 and 0.001 moles of oxygen per gram of chromium in the catalyst, preferably 0.0002 to 0.0008. Between moles. In addition, the molar ratio of dehydrogenable hydrocarbon to oxygen is desirably between 1: 0.0001 and 1: 0.04, preferably between 1: 0.001 and 1: 0.04, More preferably, it is between 1: 0.01 and 1: 0.035.
[0031]
Oxygen can be added to the process in various ways, such as mixed into a relatively cool alkane hydrocarbon feedstream or steam diluent or added directly to the reactor separately from the feed alkane hydrocarbon. Can do.
[0032]
An inert diluent can be used in the dehydrogenation process of the present invention. Any inert material that does not adversely affect the dehydrogenation process may be used as a diluent. Examples of suitable diluents are methane, ethane, propane, nitrogen or steam.
[0033]
The process of the invention can be applied to any dehydrogenable paraffinic hydrocarbon and can be applied to any type of dehydrogenation reactor including fixed bed, moving bed and fluidized bed. The preferred embodiment of the invention utilizes a fixed bed catalyst system or a dense phase moving bed system as shown in US Pat. No. 3,725,249.
[0034]
One fixed bed catalytic reaction system for the present invention consists of a single reaction zone in a single reactor with one inlet to the reactor and a reactor outlet that discharges the products and by-products of the process. Composed.
[0035]
In another embodiment of the invention, the fixed bed catalyst system has a plurality of separate oxygen inlets along the reaction zone, and oxygen can be injected at different points into the catalyst bed of the reaction zone. This type of arrangement mimics a reaction system having multiple catalyst beds.
[0036]
In an alternative embodiment, the process of the present invention is performed in a moving bed catalyst system such as the system described in US Pat. No. 3,725,249. This embodiment of the invention is quite effective when used in reactions known to deactivate the catalyst by coke deposition on the catalyst. In this embodiment, the catalyst of this process moves continuously through the reaction zones of the process, and once deactivated, it is transferred to a continuous catalyst regeneration system. Once regenerated, the catalyst is returned to the reaction system.
[0037]
【Example】
The invention will be described and understood in more detail with reference to Examples 1-9. Although these examples are shown by way of illustration, the claimed invention is not limited to these examples.
[0038]
The catalyst used in the following examples is a commercially available 14% chromium-containing chromium-alumina based dehydrogenation catalyst produced by spray drying (United Catalyst, Inc.). Catalyst T-2715 obtained from
[0039]
Charge 5 or 1 gram of the catalyst into a standard SS fixed bed reactor consisting of 1 inch tubes. A 1/8 inch thermowell was installed in the axial direction of the reaction tube to enable temperature measurement. The reactor is placed in an electrically heated oven to maintain the desired temperature. The effluent from the reactor is directed to a GC HP 6890 equipped with an on-line connected FID and TCD detector. Nitrogen is introduced into the effluent from the reactor at a constant rate to provide an internal standard for accurately calculating the mass balance. The catalyst was pretreated with 33.3 cc / min helium for 5 minutes at the reaction temperature before introducing a feed of a given composition into the reactor. Analysis of the product in the reactor was performed and reported in Tables 1 and 2. After analysis of the reaction product, the reaction was terminated, the catalyst was exhausted at once with nitrogen and regenerated using an oxygen flow of 33.3 cc / min at the same reaction temperature. The catalyst pretreatment can also be hydrogen, nitrogen or methane. The purpose of the catalyst pretreatment is to reduce the catalyst and prepare it for the reaction. Regeneration with oxygen is not limited and can be performed using air.
[0040]
The percent conversion obtained, shown in the table below, represents total isobutane converted to isobutylene and other products. The reported selectivity to isobutylene was the percentage of isobutylene obtained based on the converted butane, all percentages calculated on a molar basis. Productivity (Prod.g _HC ⁼ / g _cat · h) is defined as grams of isobutylene ( _HC ⁼ ) produced per hour per gram of catalyst. Feedstock consumption (Cons.g _HC / g _cat · h) is defined as grams of isobutane consumed per hour per catalyst gram. W / F is S.W. T. T. It is defined as the catalyst gram weight divided by the flow rate of the reactant stream measured in P / ml.
[0041]
Table 1 shows the experimental conditions performed in isobutane dehydrogenation. All examples in Table 1 were carried out with 5 g of catalyst. Examples 1 to 6 were carried out at 580 ° C., and Examples 7 to 9 were carried out at 575 ° C. using 25 ml / min of isobutane without oxygenation. The reaction temperature was 580 ° C. The results are shown in Table 2. The reported results represent the catalytic activity measured after 5 minutes of reaction, while Examples 2 and 5 were measured after 10 minutes of reaction. In Table 2, X represents conversion; S represents selectivity; Y represents yield in mole percent; iC ₄ represents isobutane; iC ₄ ⁼ represents isobutylene, and H ₂ area represents the hydrogen peak area. Represents a count (hydrogen peak area count).
[0042]

[0043]
Table 2: Results of isobutane dehydrogenation

[0044]
The results for Examples 1, 3, 4 and 6 were selected and illustrated in FIGS. From FIGS. 1, 2 and 3, as the ratio of oxygen to isobutane in the feed increases to a maximum O ₂ / isobutane molar ratio of 0.0275, isobutane conversion decreases while isobutylene selectivity and yield increase. I understand. As the oxygen concentration further increases, both conversion and selectivity decrease. This result is very important because the O ₂ concentration operation described above saves about 5% of the hydrocarbon feedstock isobutane. In addition, as in the present invention, the addition of oxygen can continuously save isobutane and improve isobutylene productivity over longer reaction times (Examples 2 and 5), and can increase reaction time by reducing coke formation.
[0045]
FIG. 4 shows that the lowest coke selectivity occurs at an O ₂ / isobutane molar ratio of about 0.0275. Higher or lower O ₂ / isobutane molar ratios appear to result in higher coke selectivity. FIG. 5 shows that the hydrogen peak areas given in Examples 1, 3 and 4 (or in Examples 2 and 5) are almost the same regardless of the amount of oxygen in the range where oxygen promotes olefin production. . The hydrogen peak area decreases with the addition of additional oxygen and the isobutane conversion and isobutylene selectivity decreases (see Example 6). Further oxygen addition shifts the reaction to the product side as expected by reaction equilibrium.
[0046]
This example clearly demonstrates the function of oxygen as a selective control agent of chromium redox properties. It is shown that the oxidation of chromium by oxygen molecules proceeds much faster than the oxidation of hydrogen or coke active species on the catalyst surface.
[0047]
The operation of isobutane dehydrogenation at lower W / F ratios is not attractive when using the process of the invention as can be seen from the productivity and feed consumption rate figures in Examples 7, 8 and 9.
[0048]
Having thus described the preferred embodiment of the invention in detail, the invention as defined in the appended claims can be modified in many ways herein without departing from the spirit or scope of the invention. It should be understood that the invention is not limited to the specific details set forth.
[Brief description of the drawings]
FIG. 1 is a diagram representing isobutane conversion as a function of oxygen concentration in the process of the invention.
FIG. 2 is a graph showing isobutylene selectivity as a function of oxygen concentration in the process of the invention.
FIG. 3 shows the isobutylene yield as a function of oxygen concentration in the process of the invention.
FIG. 4 is a diagram showing coke selectivity as a function of oxygen concentration in the process of the present invention.
FIG. 5 shows a hydrogen peak area count as a function of oxygen concentration in the process of the invention.

Claims (17)

In the presence of molecular oxygen at a temperature of 400 ° C. to 700 ° C. and a pressure of 0.1 to 10 atm, the molar ratio of alkane to oxygen is 1 : 0.0001 to 1: 0.04, and the alkane is 10 An alkane dehydrogenation process comprising contacting with a chromium-based dehydrogenation catalyst comprising -20 wt% chromium.   The method of claim 1, wherein the alkane is isobutane.   The process of claim 1, wherein the catalyst comprises 12 to 18 wt% chromium. The alkane is isobutane, and the temperature of 5 00 ° C. or al 6 50 ° C., The method of claim 3. The alkane is isobutane, and the temperature of 5 60 ° C. or al 6 00 ° C., The method of claim 3.   The process according to claim 1, wherein the alkane dehydrogenation is carried out in a fixed bed catalyst, a moving bed catalyst or a fluidized bed catalyst.   The method of claim 1, wherein the contacting is performed in the presence of an inert diluent.   8. The method of claim 7, wherein the alkane is isobutane and the inert diluent comprises methane, ethane, propane, nitrogen or steam.   The method according to claim 4, wherein the contacting is performed in the presence of methane, ethane, propane, nitrogen or steam.   The method of claim 5, wherein the contacting is performed in the presence of methane, ethane, propane, nitrogen or steam. The process according to claim 1, wherein the molar ratio of alkane to oxygen is from 1 : 0.001 to 1: 0.04. The process according to claim 1, wherein the molar ratio of alkane to oxygen is from 1 : 0.01 to 1: 0.035. The method of claim 4, wherein the molar ratio of alkane to oxygen is from 1 : 0.01 to 1: 0.035. Oxygen abundance relative to the amount of chromium in said catalyst per gram oxygen 0 chromium of the catalyst. 4. The method of claim 3, wherein the amount is from 0001 to 0.001 mole. Oxygen abundance relative to the amount of chromium in said catalyst per gram oxygen 0 chromium of the catalyst. The process of claim 4 wherein the process is from 0002 to 0.0008 moles. Oxygen abundance relative to the amount of chromium in said catalyst per gram oxygen 0 chromium of the catalyst. 6. A process according to claim 5, which is between 0001 and 0.001 mol. Oxygen abundance relative to the amount of chromium in said catalyst per gram oxygen 0 chromium of the catalyst. 13. A method according to claim 12, wherein the amount is from 0001 to 0.001 mol.

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