CATALYTIC OLIGOMERIZATION

This invention relates to an oligomerization process to produce oligomers that may be hydrogenated to provide thermally stable lubricants and lubricant additives, gasoline and diesel. Efforts to improve upon the performance of natural mineral oil based lubricants by the synthesis of oligomeric hydrocarbon fluids have been the subject of important research and development in the petroleum industry for a large number of years and have led to the introduction of a number of superior polyalpha-olefin (PAO) synthetic lubricants produced by the oligomerization of alpha-olefins or 1-alkenes. In terms of lubricant property involvement, the thrust of the industrial research effort on synthetic lubricants has been toward fluids exhibiting useful viscosities over a wider range of temperature, i.e., improved viscosity index (VI), while also showing lubricity, thermal and oxidative stability and pour point equal to or better than mineral oil. These new synthetic lubricants exhibit lower friction characteristics and are therefore capable of increasing mechanical efficiency of various types of equipment including engines, transmissions, worm gears and traction drives, doing so over a wider range of operating conditions than mineral oil lubricants. PAOs useful as synthetic base stocks or functional fluids may be synthesized by homogeneous catalysts, such as promoted BF ₃ or A1C1 ₃ catalysts. The synthesis of PAOs with a promoted BF ₃ catalyst is discussed in the Theriot et al. U.S. Patent No. 5,171,905. The PAO processes using homogeneous catalysts always include a complicated and tedious catalyst separation step. For example, the promoted BF ₃ or A1C1 ₃ catalyst is usually deactivated and

destroyed by washing with sodium hydroxide, dilute acid and water consecutively. This separation step generates waste and is tedious. Therefore, it would be advantageous to use a solid and regenerable catalyst which can be separated easily from the product and regenerated for reuse.

The present invention is particularly useful for upgrading C ₂ -C ₅ lower olefins to heavier hydrocarbons, such as C ₆ -C ₂ o+ gasoline and distillate fuel product. Numerous mono-olefins, including ethene, propene, n-butenes, isobutene, pentenes, and mixtures thereof, etc., can be reacted selectively in aliphatic hydrocarbon feedstocks. An advantage of the present process is reaction select¬ ivity, such that non-olefinic products can be avoided as reaction by-products, due to the substantial absence of dehydrogenation, cyclization and alkane formation.

However, the feedstocks may contain non-deleterious amounts of paraffins.

In a preferred embodiment, the catalyst of the present invention is employed in the conversion of propene by oligomerization to gasoline, diesel fuel and lube coproducts. Preferred propene feedstocks include propene and FCC propane/propene. N-butane and butenes can also be included in the feed. Optionally, hydrogen can be cofed into the reactor with the hydrocarbon feed. Oligomerization reaction temperature is in the range of from 25 to 400°C, e.g., in the range of from 120 to 350°C, e.g., in the range of 30 to 250°C. Pressure is in the range of from 0 to 2000 psig. The reaction may be conducted in the gas phase, liquid phase or dense phase with continuous or batch operation using, for example, a fixed bed or stirred-tank reactor. Generally, the liquid hourly space velocity, based on volume of liquid olefin per volume of catalyst per hour, is in the range of 0.1-10, preferably 0.5-3.

The product slates can be adjusted by varying the operation conditions. The coproduction of gasoline and diesel is favored by higher temperatures (120-350°C) and lower pressures (30-1000 psig) . The coproduction of gasoline, diesel and lube is favored by lower temperatures (30-250°C) and higher pressures (300-2000 psig) . In cases when an increased amount of lube fraction is desirable, part of the oligomer product (e.g., C ₁₉ - or C ₁₀ -) can be recycled to the reactor. The deactivated catalyst can be hydrogenatively regenerated in the presence of hydrogen or oxidatively regenerated in the presence of air.

The catalyst of the present invention comprises an oxide of a Group IVB metal, preferably zirconia or titania. This Group IVB metal oxide is modified with an oxyanion of a Group VIB metal, such as oxyanion of tungsten, such as tungstate. The modification of the Group IVB metal oxide with the oxyanion of the Group VIB metal imparts acid functionality to the material. Suitable sources of the Group IVB metal oxide include compounds capable of generating such oxides, such as oxychlorides, chlorides, nitrates, oxynitrates, etc., particularly of zirconium or titanium. Alkoxides of such metals may also be used as precursors or sources of the Group IVB metal oxide. Examples of such alkoxides include zirconium n-propoxide and titanium i-propoxide. These sources of a Group IVB metal oxide, particularly zirconia, may form zirconium hydroxide, i.e., Zr(OH) ₄ , or hydrated zirconia as intermediate species upon precipitation from an aqueous medium in the absence of a reactive source of tungstate. The expression, hydrated zirconia, is intended to connote materials comprising zirconium atoms covalently linked to other zirconium atoms via bridging oxygen atoms,

i.e., Zr-O-Zr, further comprising available surface hydroxy groups.

Suitable sources for the oxyanion of the Group VIB metal, preferably molybdenum or tungsten, include, but are not limited to, ammonium metatungstate or metamolybdate, tungsten or molybdenum chloride, tungsten or molybdenum carbonyl, tungstic or molybdic acid and sodium tungstate or molybdate.

The Group IVB metal (i.e., Ti, Zr or Hf) and the Group VIB metal (i.e., Cr, Mo or W) species of the present catalyst are not limited to any particular valence state for these species. These species may be present in this catalyst in any possible positive oxidation value for these species. In the present catalyst, of the Group IVB oxides, zirconium oxide is preferred and of the Group IVB anions, tungstate is preferred.

Qualitatively speaking, elemental analysis of the present acidic solid will reveal the presence of Group IVB metal, Group VIB metal and oxygen. The amount of oxygen measured in such an analysis will depend on a number of factors, such as the valence state of the Group IVB and Group VIB metals, the form of the optional hydrogenation/dehydrogenation component, moisture content, etc. Accordingly, in characterizing the composition of the present catalyst, it is best not to be restricted by any particular quantities of oxygen. In functional terms, the amount of Group VIB oxyanion in the present catalyst may be expressed as that amount which increases the acidity of the Group IVB oxide. This amount is referred to herein as an acidity increasing amount. Elemental analysis of the present catalyst may be used to determine the relative amounts of Group IVB metal and Group VIB metal in the catalyst. From these amounts, mole ratios in the form of

catalyst. From these amounts, mole ratios in the form of X0 ₂ /Y0 ₃ may be calculated, where X is said Group IVB metal, assumed to be in the form of X0 ₂ , and Y is said Group VIB metal, assumed to be in the form of Y0 ₃ . It will be appreciated, however, that these forms of oxides, i.e., X0 ₂ and Y0 ₃ , may not actually exist, and are referred to herein simply for the purposes of calculating relative quantities of X and Y in the present catalyst. The present catalysts may have calculated mole ratios, expressed in the form of X0 ₂ /Y0 ₃ , where X is at least one Group IVB metal (i.e., Ti, Zr, and Hf) and Y is at least one Group VIB metal (i.e., Cr, Mo, or W) , of up to 1000, e.g., up to 300, e.g., from 2 to 100, e.g., from 4 to 30.

The modified oxide material may be prepared by combining a first liquid solution comprising a source of a Group IVB metal oxide with a second liquid solution comprising a source of an oxyanion of a Group VIB metal. This combination of two solutions takes place under conditions sufficient to cause co-precipitation of the modified oxide material as a solid from the liquid medium. Alternatively, the source of the Group IVB metal oxide and the source of the oxyanion of the Group VIB metal may be combined in a single liquid solution. This solution may then be subjected to conditions sufficient to cause co¬ precipitation of the modified oxide material, such as by the addition of a precipitating reagent to the solution. Water is a preferred solvent for these solutions. The temperature at which the liquid medium is maintained during the co-precipitation may be less than 200°C, e.g., from 0°C to 200°C. This liquid medium may be maintained at an ambient temperature (i.e., room temperature) or the liquid may be cooled or heated. A particular range of such temperatures is from 10°C to 100°C.

The liquid medium from which the catalyst components are co-precipitated may optionally comprise a solid support material, in which case the catalyst may be co-precipitated directly onto the solid support material. Examples of such support materials include the material designated M41S, which is described in U.S. Patent No. 5,102,643. A particular example of such an M41S material is a material designated MCM-41, which is described in U.S. Patent No. 5,098,684. Support materials and/or co-catalyst materials may also, optionally, be co-precipitated from the liquid medium along with the Group IVB metal oxide and the oxyanion of the Group VIB metal. An example of a co-catalyst material is a hydrogenation/dehydrogenation component. The modified oxide material may be recovered by filtration from the liquid medium, followed by drying. Calcination of the resulting material may be carried out, preferably in an oxidizing atmosphere, at temperatures from 500°C to 900°C, preferably from 700°C to 850°C, and more preferably from 750°C to 825°C. The calcination time may be up to 48 hours, preferably for 0.1-24 hours, and more preferably for 1.0-10 hours. In a most preferred embodiment, calcination is carried out at 800°C for 1 to 3 hours. The modified oxide material may be contacted with hydrogen at elevated temperatures. These elevated temperatures may be 100°C or greater, e.g., 250°C or grater, e.g., 300°C. The duration of this contact may be as short as one hour or even 0.1 hour. However, extended contact may also be used. This extended contact may take place for a period of 6 hours or greater, e.g., 18 hours. The modified oxide material may be contacted with hydrogen in the presence or absence of a hydrocarbon cofeed.

The following examples illustrate the process of the present invention. Example 1

Five hundred grams of ZrOCl ₂ -8H ₂ 0 were dissolved with stirring in 7.0 liters of distilled water. A solution containing 263 ml of cone. NH ₄ 0H, 500 ml of distilled water, and 54 grams of (NH ₄ ) ₆ H ₂ W ₁₂ 0 ₀ "xH ₂ 0 was added dropwise over 30-45 minute period. The pH of the solution was adjusted to approximately 9 (if needed) by adding additional cone. NH ₄ 0H dropwise. This slurry was then placed in the steambox for 72 hours. The product formed was recovered by filtration, washed with excess H ₂ 0, and dried overnight at 85°C. The material was then calcined in dry air at 825°C for 3 hours. The resulting binary oxide catalyst contained 15.9 wt.% of W and 58.6 wt.% of Zr.

This co-precipitated Wo _x /Zr0 ₂ catalyst is called Catalyst A. Example 2

7.0 cc (9.28 g) of Catalyst A (30-60 mesh particles) was charged to a fixed-bed tubular reactor. The catalyst was calcined with flowing air at 500°C and 1 atm for an hour, and the purged with N ₂ for 30 minutes. The temperature was decreased to 350°C, and the catalyst was reduced with flowing H ₂ at 350°C and 1 atmosphere atm for one hour. The reactor was then purged with N ₂ for one hour and the temperature was reduced to 120°C. At this stage, the reactor was pressurized with N ₂ to 400 psig and propene was fed into the reactor at a rate of 7 ml/hour. Then, the reactor temperature was gradually increased to 160°C. After the system ws lined out overnight, a material balance was conducted over a 6 hour period. The product analysis by gas chromatography showed that the propene conversion was 94.4 wt.% with a C,+ selectivity of 94.9 wt.%. The C ₅ + products contained 29.5 wt.% of gasoline (C ₅ -330°F) , 61.6

wt.% of diesel fuel (330-650°F) , and 8.9 wt.% of lube (650 F+) . Example 3

One part by weight of zirconyl chloride, ZrOC ₁₂ -H ₂ 0, was added to 3 parts by weight of a 10 M NH ₄ 0H solution. The resulting slurry, Zr(0H) ₄ , was filtered and washed with 5 parts by weight of distilled deionized water, then air dried at 140°C for 8 hours. Approximately 5.5 parts by weight of the resulting Zr(OH) ₄ were impregnated via incipient wetness with 2.2 parts of an aqueous solution containing 1 part of ammonium metatungstate, (NH ₄ ) ₆ H ₂ W ₁₂ O ₄₀ -χH ₂ O. The resulting material was dried for 2 hours at 120°C and then calcined at 825°C in flowing air for 3 hours. This impregnated Wo _x /Zr0 ₂ catalyst, called Catalyst B, contained 15 wt.% W. Example 4

6.5 cc (9.30 g) of Catalyst B (30-60 mesh particles) were charged to a fixed-bed tubular reactor. The catalyst was calcined with flowing air at 500°C and 1 atm for an hour, and then purged with N ₂ for 30 minutes. The temperature was decreased to 350°C, and the catalyst was reduced with flowing H ₂ at 350°C and 1 atm for one hour. The reactor was then purged with N ₂ for one hour and the temperature was reduced to 120°C. At this stage, the reactor was pressurized with N ₂ to 400 psig and propene was fed into the reactor at a rate of 7 ml/hour. Then, the reactor temperature was gradually increased to 160°C. After the system was lined out overnight, a material balance was conducted over a 6 hour period. The product analysis by gas chromatography showed that the propene conversion ws 69.8 wt.% with a C ₅ + selectivity of 93.0

wt.%. The C _s + products contained 38.5 wt.% of gasoline (C ₅ - 330°F), 56.0 wt.% of diesel fuel (330-650°F) , and 5.5 wt.% of lube (650°F+) .

The comparison between the co-precipitated catalyst (Catalyst A) and the impregnated catalyst (Catalyst B) for propene oligomerization under the same operating conditions are given below:

Catalyst Catalyst A Catalyst B

Method of Preparation Co-Precipitation Impregnation

Propylene Conversion, % 94.4 69.8

C ₅ + Selectivity, wt.% 94.9 93.0

C ₅ + Distribution, wt.%

Gasoline ( C ₅ -330°F) 29.5 38.5

Diesel (330-650°F) 61.6 56.0

Lube (650°F+) 8.9 5.5

The above comparison clearly indicates that the co- precipitated catalyst is more active than the impregnated catalyst, as reflected by the higher propene conversion and higher yields of diesel and lube.