Hydroconversion process and production of lubricating oil basestocks

Abstract: A paraffins-rich raffinate from the solvent extraction of a lubricating oil feedstock is treated by two-stage hydroconversion followed by hydrofinishing. A preferred subsequent step comprises dewaxing. A lubricating oil basestock having at least 90 wt.% saturates and a VI of at least 105 may be obtained.

Description: This invention relates to a hydroconversion process and production of lubricating oil basestocks having a high saturates content, high viscosity indices and low volatilities. It is well known to produce lubricating oil basestocks by solvent refining. In the conventional process, crude oils are fractionated under atmospheric pressure to produce atmospheric resids which are further fractionated under vacuum. Select distillate fractions are then optionally deasphalted and solvent extracted to produce a paraffin rich raffinate and an aromatics rich extract. The raffinate is then dewaxed to produce a dewaxed oil which is usually hydrofinished to improve stability and remove color bodies. Solvent refining is a process which selectively isolates components of crude oils having desirable properties for lubricant basestocks. Thus the crude oils used for solvent refining are restricted to those which are highly paraffinic in nature as aromatics tend to have lower viscosity indices (VI), and are therefore less desirable in lubricating oil basestocks. Also, certain types of aromatic compounds can result in unfavorable toxicity characteristics. Solvent refining can produce lubricating oil basestocks having a VI of about 95 in good yields. Today more severe operating conditions for automobile engines have resulted in demands for basestocks with lower volatilities (while retaining low viscosities) and lower pour points. These improvements can only be achieved with basestocks of more isoparaffinic character, i.e., those with VIs of 105 or greater. Solvent refining alone cannot economically produce basestocks having a VI of 105 with typical crudes. Nor does solvent refining alone typically produce basestocks with high saturates contents. Two alternative approaches have been developed to produce high quality lubricating oil basestocks; (1) wax isomerization and (2) hydrocracking. Both of the methods involve high capital investments. In some locations wax isomerization economics can be adverselyimpacted when the raw stock, slack wax, is highly valued. Also, the typically low quality feedstocks used in hydrocracking, and the consequent severe conditions required to achieve the desired viscometric and volatility properties can result in the formation of undesirable (toxic) species. These species are formed in sufficient concentration that a further processing step such as extraction is needed to achieve a non-toxic base stock. An article by S. Bull and A. Marmin entitled "Lube Oil Manufacture by Severe Hydrotreatment", Proceedings of the Tenth World Petroleum Congress, Volume 4, Developments in Lubrication, PD 19(2), pages 221-228, describes a process wherein the extraction unit in solvent refining is replaced by a hydrotreater. U.S. Patent 3,691,067 describes a process for producing a medium and high VI oil by hydrotreating a narrow cut lube feedstock. The hydrotreating step involves a single hydrotreating zone. U.S. Patent 3,732,154 discloses hydrofinishing the extract or raffinate from a solvent extraction process. The feed to the hydrofinishing step is derived from a highly aromatic source such as a naphthenic distillate. U.S. patent 4,627,908 relates to a process for improving the bulk oxidation stability and storage stability of lube oil basestocks derived from hydrocracked bright stock. The process involves hydrodenitrification of a hydrocracked bright stock followed by hydrofinishing. EP-A-O-323-092 describes the production of lubrication oil by severe hydrotreating of Fischer-Tropsch wax followed by hydroisomerization and dewaxing. US. Patent 5,110,445 describes the production of lubricating oil by hydroisomerization followed by aromatics extraction of a slack wax. It would be desirable to supplement the conventional solvent refining process so as to produce high VI, low volatility oils which have excellent toxicity, oxidative and thermal stability, fuel economy and cold start properties without incurring any significant yield debit which process requires much lower investment costs than competing technologies such as hydrocracking. Summary of the Invention This invention relates to a process for producing a lubricating oil basestock meeting at least 90% saturates and VI of at least 105 by selectively hydroconverting a raffinate produced from solvent refining a lubricating oil feedstock which comprises: (a) conducting the lubricating oil feedstock to a solvent extraction zone and separating therefrom an aromatics rich extract and a paraffins rich raffinate; (b) stripping the raffinate of solvent to produce a raffinate feed having a dewaxed oil viscosity index from 85 to 105 and a final boiling point of no greater than about 650° C; (c) passing the raffinate feed to a first hydroconversion zone and processing the raffinate feed in the presence of a non acidic catalyst at a temperature of from 340 to 420° C, a hydrogen partial pressure of from 7,0 to 17,3 MPa (1000 to 2500 psig), space velocity of 0.2 to 3.0 LHSV h -1 and a hydrogen to feed ratio of from 89 to 890 m 3 /m 3 (500 to 5000 Scf/B) to produce a first hydroconverted raffinate; (d) passing the hydroconverted raffinate from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted raffinate in the presence of a non acidic catalyst at a temperature of from 340 to 400° C provided that the temperature in second hydroconversion is not greater than the temperature in the first hydroconversion zone, a hydrogen partial pressure of from 7.0 to 17.3 MPa (1000 to 2500 psig), a space velocity of from 0.2 to 3.0 LHSV h -1 and a hydrogen to feed ratio of from 89 to 890 m 3 /m 3 (500 to 5000 Scf/B) to produce a second hydroconverted raffinate; (e) passing the second hydroconverted raffinate to a hydrofinishing zone and conducting cold hydrofinishing of the second hydroconverted raffinate in the presence of a hydrofinishing catalyst at a temperature of from 260 to 360° C, a hydrogen partial pressure of from 7,0 to 17.3 MPa (1000 to 2500 psig), a space velocity of from 0.2 to 5 LHSV h -1 and a hydrogen to feed ratio of from 89 to 890 m 3 /m 3 (500 to 5000 Scf/B) to produce a hydrofinished raffinate; (f) passing the hydrofinished raffinate to a separation zone to remove products having a boiling less than about 250° C; and (g) passing the hydrofinished raffinate from the separation zone to a dewaxing zone to produce a dewaxed basestock having a viscosity index of at least 105 provided that the basestock has a dewaxed oil viscosity index increase of at least 10 greater than the raffinate feed, a NOACK volatility improvement over raffinate feedstock of at least about 3 wt.% at the same viscosity in the range of viscosity from 3.5 to 6.5 cSt viscosity at 100° C, and a saturates content of at least 90 wt.%; wherein said non acidic catalysts have an acidity value less than 0.5 on the scale for acidity based on the isomerization of 2-methyl-2-pentene with said catalysts at a fixed temperature. The basestock also has a low toxicity (passing the IP346 or FDA(c) tests). In another embodiment, this invention relates to a process for selectively hydroconverting a raffinate produced from solvent refining a lubricating oil feedstock which comprises: (a) conducting the lubricating oil feedstock to a solvent extraction zone and separating therefrom an aromatics rich extract and a paraffins rich raffinate; (b) stripping the raffinate of solvent to produce a raffinate feed having a dewaxed oil viscosity index from 85 to 105 and a final boiling point of no greater than about 650° C; (c) passing the raffinate feed to a first hydroconversion zone and processing the raffinate feed in the presence of a non acidic catalyst at a temperature of from 340 to 420° C, a hydrogen partial pressure of from 7.0 to 17.3 MPa (1000 to 2500 psig), space velocity of 0.2 to 3.0 LHSV h -1 and a hydrogen to feed ratio of from 89 to 890 m 3 /m 3 (500 to 5000 Scf/B) to produce a first hydroconverted raffinate; (d) passing the hydroconverted raffinate from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted raffinate in the presence of a non acidic catalyst at a temperature of from 340 to 400°C provided that the temperature in the second hydroconversion is not greater than the temperature in the first hydroconversion zone, a hydrogen partial pressure of from 7.0 to 17.3 MPa (1000 to 2500 psig), a space velocity of from 0.2 to 3.0 LHSV h -1 and a hydrogen to feed ratio of from 89 to 890 m 3 /m 3 (500 to 5000 Scf/B) to produce a second hydroconverted raffinate; (e) passing the second hydroconverted raffinate to a hydrofinishing reaction zone and conducting cold hydrofinishing of the second hydroconverted raffinate in the presence of a hydrofinishing catalyst at a temperature of from 260 to 360°C, a hydrogen partial pressure of from 7.0 to 17.3 MPa (1000 to 2500 psig), a space velocity of from 0.2 to 5 LHSV h -1 and hydrogen to feed ratio of from 89 to 890 m 3 /m 3 (500 to 5000 Scf/B) to produce a hydrofinished raffinate. The process according to the invention produces in good yields a basestock which has VI and volatility properties meeting future industry engine oil standards while achieving good oxidation stability, cold start, fuel economy, and thermal stability properties. In addition, toxicity tests show that the basestock has excellent toxicological properties as measured by tests such as the FDA(c) test. Brief Description of the Drawings Fig. 1 is a plot of NOACK volatility vs. viscosity for a 100N basestock. Fig. 2 is a schematic flow diagram of the hydroconversion process. Fig. 3 is a graph showing VI HOP vs. conversion at different pressures. Fig. 4 is a graph showing temperature in the first hydroconversin zone as a function of days on oil at a fixed pressure. Fig. 5 is a graph showing saturates concentration as a function of reactor temperature for a fixed VI product. Fig. 6 is a graph showing toxicity as a function of temperature and pressure in the cold hydrofinishing step. Fig. 7 is a graph showing control of saturates concentration by varying conditions in the cold hydrofinishing step. Fig. 8 is a graph showing the correlation between the DMSO screener test and the FDA (c) test. Detailed Description of the Invention The solvent refining of select crude oils to produce lubricating oil basestocks typically involves atmospheric distillation, vacuum distillation, extraction, dewaxing and hydrofinishing. Because basestocks having a high isoparaffin content are characterized by having good viscosity index (VI) properties and suitable low temperature properties, the crude oils used in the solvent refining process are typically paraffinic crudes. One method of classifying lubricating oil basestocks is that used by the American Petroleum Institute (API). API Group II basestocks have a saturates content of 90 wt.% or greater, a sulfur content of not more than 0.03 wt.% and a viscosity index (VI) greater than 80 but less than 120. API Group III basestocks are the same as Group II basestocks except that the VI is greater than or equal to 120. Generally, the high boiling petroleum fractions from atmospheric distillation are sent to a vacuum distillation unit, and the distillation fractions from this unit are solvent extracted. The residue from vacuum distillation which may be deasphalted is sent to other processing. The solvent extraction process selectively dissolves the aromatic components in an extract phase while leaving the more paraffinic components in a raffinate phase. Naphthenes are distributed between the extract and raffinate phases. Typical solvents for solvent extraction include phenol, furfural and N-methyl pyrrolidone. By controlling the solvent to oil ratio, extraction temperature and method of contacting distillate to be extracted with solvent, one can control the degree of separation between the extract and raffinate phases. In recent years, solvent extraction has been replaced by hydrocracking as a means for producing high VI basestocks in some refineries. The hydrocracking process utilizes low quality feeds such as feed distillate from the vacuum distillation unit or other refinery streams such as vacuum gas oils and coker gas oils. The catalysts used in hydrocracking are typically sulfides of Ni, Mo, Co and W on an acidic support such as silica/alumina or alumina containing an acidic promoter such as fluorine. Some hydrocracking catalysts also contain highly acidic zeolites. The hydrocracking process may involve hetero-atom removal, aromatic ring saturation, dealkylation of aromatics rings, ring opening, straight chain and side-chain cracking, and wax isomerization depending on operating conditions. In view of these reactions, separation of the aromatics rich phase that occurs in solvent extraction is an unnecessary step since hydrocracking reduces aromatics content to very low levels. By way of contrast, the process of the present invention utilizes a three step hydroconversion of the raffinate from the solvent extraction unit under conditions which minimizes hydrocracking and passing waxy components through the process without wax isomerization. Thus, dewaxed oil (DWO) and low value foots oil streams can be added to the raffinate feed whereby the wax molecules pass unconverted through the process and may be recovered as a valuable by-product. Moreover, unlike hydrocracking, the present process takes place without disengagement, i.e., without any intervening steps involving gas/liquid products separations. The product of the subject three step process has a saturates content greater than 90 wt.%, preferably greater than 95 wt.%. Thus product quality is similar to that obtained from hydrocracking without the high temperatures and pressures required by hydrocracking which results in a much greater investment expense. The raffinate from the solvent extraction is preferably under-extracted, i.e., the extraction is carried out under conditions such that the raffinate yield is maximized while still removing most of the lowest quality molecules from the feed. Raffinate yield may be maximized by controlling extraction conditions, for example, by lowering the solvent to oil treat ratio and/or decreasing the extraction temperature. The raffinate from the solvent extraction unit is stripped of solvent and then sent to a first hydroconversion unit containing a hydroconversion catalyst. This raffinate feed has a viscosity index of from about 85 to about 105 and a boiling range not to exceed about 650° C, preferably less than 600° C, as determined by ASTM 2887 and a viscosity of from 3 to 15 cSt at 100°C. Hydroconversion catalysts are those containing Group VIB metals (based on the Periodic Table published by Fisher Scientific), and non-noble Group VIII metals, i.e., iron, cobalt and nickel and mixtures thereof. These metals or mixtures of metals are typically present as oxides or sulfides on refractory metal oxide supports. It is important that the metal oxide support be non acidic so as to control cracking. A useful scale of acidity for catalysts is based on the isomerization of 2-methyl-2-pentene as described by Kramer and McVicker, J. Catalysis, 92 , 355(1985). In this scale of acidity, 2-methyl-2-pentene is subjected to the catalyst to be evaluated at a fixed temperature, typically 200° C. In the presence of catalyst sites, 2-methyl-2-pentene forms a carbenium ion. The isomerization pathway of the carbenium ion is indicative of the acidity of active sites in the catalyst. Thus weakly acidic sites form 4-methyl-2-pentene whereas strongly acidic sites result in a skeletal rearrangement to 3-methyl-2-pentene with very strongly acid sites forming 2,3-dimethyl-2-butene. The mole ratio of 3-methyl-2-pentene to 4-methyl-2-pentene can be correlated to a scale of acidity. This acidity scale ranges from 0.0 to 4.0. Very weakly acidic sites will have values near 0.0 whereas very strongly acidic sites will have values approaching 4.0. The catalysts useful in the present process have acidity values of less than about 0.5, preferably less than about 0.3. The acidity of metal oxide supports can be controlled by adding promoters and/or dopants, or by controlling the nature of the metal oxide support, e.g., by controlling the amount of silica incorporated into a silica-alumina support. Examples of promoters and/or dopants include halogen, especially fluorine, phosphorus, boron, yttria, rare-earth oxides and magnesia. Promoters such as halogens generally increase the acidity of metal oxide supports while mildly basic dopants such as yttria or magnesia tend to decrease the acidity of such supports. Suitable metal oxide supports include low acidic oxides such as silica, alumina or titania, preferably alumina. Preferred aluminas are porous aluminas such as gamma or eta having average pore sizes from 50 to 200Å, preferably 75 to 150Å, a surface area from 100 to 300 m 2 /g, preferably 150 to 250 m 2 /g and a pore volume of from 0.25 to 1.0 cm 3 /g, preferably 0.35 to 0.8 cm 3 /g. The supports are preferably not promoted with a halogen such as fluorine as this generallyincreases the acidity of the support above 0.5. Preferred metal catalysts include cobalt/molybdenum (1-5% Co as oxide, 10-25% Mo as oxide) nickel/molybdenum (1-5% Ni as oxide, 10-25% Mo as oxide) or nickel/tungsten (1-5% Ni as oxide, 10-30% W as oxide) on alumina. Especially preferred are nickel/molybdenum catalysts such as KF-840. Hydroconversion conditions in the first hydroconversion unit include a temperature of from 340 to 420° C, preferably 350 to 400° C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 Pa), preferably 1000 to 2000 psig (7.0 to 13.9 MPa), a space velocity of from 0.2 to 3.0 LHSV h -1 , preferably 0.3 to 1.0 LHSV h -1 , and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ), preferably 2000 to 4000 Scf/B (356 to 712 m 3 /m 3 ). The hydroconverted raffinate from the first hydroconversion unit is conducted to a second hydroconversion unit. The hydroconverted raffinate is preferably passed through a heat exchanger located between the first and second hydroconversion units so that the second hydroconversion unit can be run at cooler temperatures, if desired. Temperatures in the second hydroconversion unit should not exceed the temperature used in the first hydroconversion unit. Conditions in the second hydroconversion unit include a temperature of from 340 to 400° C, preferably 350 to 385° C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 Mpa), preferably 1000 to 2000 psig (7.0 to 13.9 Mpa), a space velocity of from 0.2 to 3.0 LHSV, preferably 0.3 to 1.5 LHSV h -1 , and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ), preferably 2000 to 4000 Scf/B (356 to 712 m 3 /m 3 ). The catalyst in the second hydroconversion unit can be the same as in the first hydroconversion unit, although a different hydroconversion catalyst may be used. The hydroconverted raffinate from the second hydroconversion unit is then conducted to cold hydrofinishing unit. A heat exchanger is preferably located between these units. Reaction conditions in the hydrofinishing unit are mild and include a temperature of from 260 to 360° C, preferably 290 to 350° C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), preferably 1000 to 2000 psig (7.0 to 13.9 MPa), a space velocity of from 0.2 to 5.0 LHSV h -1 , preferably 0.7 to 3.0 LHSV h -1 , and a hydrogen to feed ratio of from 500 to 5000 SCF/B (89 to 890 m 3 /m 3 ), preferably 2000 to 4000 Scf/B (356 to 712 m 3 /m 3 ). The catalyst in the cold hydrofinishing unit may be the same as in the first hydroconversion unit. However, more acidic catalyst supports such as silica-alumina, zirconia and the like may be used in the cold hydrofinishing unit. In order to prepare a finished basestock, the hydroconverted raffinate from the hydrofinishing unit is conducted to a separator e.g., a vacuum stripper (or fractionation) to separate out low boiling products. Such products may include hydrogen sulfide and ammonia formed in the first two reactors. If desired, a stripper may be situated between the second hydroconversion unit and the hydrofinishing unit, but this is not essential to produce basestocks according to the invention. The hydroconverted raffinate separated from the separator is then conducted to a dewaxing unit. Dewaxing may be accomplished by catalytic processes or by using a solvent to dilute the hydrofinished raffinate and chilling to crystallize and separate wax molecules. Typical solvents include propane and ketones. Preferred ketones include methyl ethyl ketone, methyl isobutyl ketone and mixtures thereof. The solvent/hydroconverted raffinate mixture may be cooled in a refrigeration system containing a scraped-surface chiller. Wax separated in the chiller is sent to a separating unit such as a rotary filter to separate wax from oil. The dewaxed oil is suitable as a lubricating oil basestock. If desired, the dewaxed oil may be subjected to catalytic isomerization/dewaxing to further lower the pour point. Separated wax may be used as such for wax coatings, candles and the like or may be sent to an isomerization unit. The lubricating oil basestock produced by the process according to the invention is characterized by the following properties: viscosity index of at least about 105, preferably at least 107 and saturates of at least 90%, preferably at least 95 wt%, NOACK volatility improvement (as measured by DIN 51581) over raffinate feedstock of at least about 3 wt.%, preferably at least about 5 wt.%, at the same viscosity within the range 3.5 to 6.5 cSt viscosity at 100° C, pour point of-15° C or lower, and a low toxicity as determined by IP346 or phase 1 of FDA (c). IP346 is a measure of polycyclic aromatic compounds. Many of these compounds are carcinogens or suspected carcinogens, especially those with so-called bay regions [see Accounts Chem. Res. 17 , 332(1984) for further details]. The present process reduces these polycyclic aromatic compounds to such levels as to pass carcinogenicity tests. The FDA (c) test is set forth in 21 CFR 178.3620 and is based on ultraviolet absorbances in the 300 to 359 nm range. As can be seen from Fig. 1, NOACK volatility is related to VI for any given basestock. The relationship shown in Fig. 1 is for a light basestock (about 100N). If the goal is to meet a 22 wt. % NOACK volatility for a 100N oil, then the oil should have a VI of about 110 for a product with typical-cut width, e.g., 5 to 50% off by GCD at 60° C. Volatility improvements can be achieved with lower VI product by decreasing the cut width. In the limit set by zero cut width, one can meet 22% NOACK volatility at a VI of about 100. However, this approach, using distillation alone, incurs significant yield debits. Hydrocracking is also capable of producing high VI, and consequently low NOACK volatility basestocks, but is less selective (lower yields) than the process of the invention. Furthermore both hydrocracking and processes such as wax isomerization destroy most of the molecular species responsible for the solvency properties of solvent refined oils. The latter also uses wax as a feedstock whereas the present process is designed to preserve wax as a product and does little, if any, wax conversion. The process of the invention is further illustrated by Fig. 2. The feed 8 to vacuum pipestill 10 is typically an atmospheric reduced crude from an atmospheric pipestill (not shown). Various distillate cuts shown as 12 (light), 14 (medium) and 16 (heavy) may be sent to solvent extraction unit 30 via line 18. These distillate cuts may range from about 200° C to about 650° C. The bottoms from vacuum pipestill 10 may be sent through line 22 to a coker, a visbreaker or a deasphalting extraction unit 20 where the bottoms are contacted with a deasphalting solvent such as propane, butane or pentane. The deasphalted oil may be combined with distillate from the vacuum pipestill 10 through line 26 provided that the deasphalted oil has a boiling point no greater than about 650° C or is preferably sent on for further processing through line 24. The bottoms from deasphalter 20 can be sent to a visbreaker or used for asphalt production. Other refinery streams may also be added to the feed to the extraction unit through line 28 provided they meet the feedstock criteria described previously for raffinate feedstock. In extraction unit 30, the distillate cuts are solvent extracted with N-methyl pyrrolidone and the extraction unit is preferably operated in countercurrent mode. The solvent-to-oil ratio, extraction temperature and percent water in the solvent are used to control the degree of extraction, i.e., separation into a paraffins rich raffinate and an aromatics rich extract. The present process permits the extraction unit to operate to an "under extraction" mode, i.e., a greater amount of aromatics in the paraffins rich raffinate phase. The aromatics rich extract phase is sent for further processing through line 32. The raffinate phase is conducted through line 34 to solvent stripping unit 36. Stripped solvent is sent through line 38 for recycling and stripped raffinate is conducted through line 40 to first hydroconversion unit 42. The first hydroconversion unit 42 contains KF-840 catalyst which is nickel/molybdenum on an alumina support and available from Akzo Nobel. Hydrogen is admitted to unit or reactor 42 through line 44. Gas chromatographic comparisons of the hydroconverted raffinate indicate that almost no wax isomerization is taking place. While not wishing to be bound to any particular theory since the precise mechanism for the VI increase which occurs in this stage is not known with certainty, it is known that heteroatoms are being removed, aromatic rings are being saturated and naphthene rings, particularly multi-ring naphthenes, are selectively eliminated. Hydroconverted raffinate from hydroconversion unit 42 is conducted through line 46 to heat exchanger 48 where the hydroconverted raffinate stream may be cooled if desired. The cooled raffinate stream is conducted through line 50 to a second hydroconversion unit 52. Additional hydrogen, if needed, is added through line 54. This second hydroconversion unit is operated at a lower temperature (when required to adjust product quality) than the first hydroconverion unit 42. While not wishing to bound to any theory, it is believed that the capability to operate the second unit 52 at lower temperature shifts the equilibrium conversion between saturated species and other unsaturated hydrocarbon species back towards increased saturates concentration. In this way, the concentration of saturates can be maintained at greater than 90 wt.% by appropriately controlling the combination of temperature and space velocity in second hydroconversion unit 52. Hydroconverted raffinate from unit 52 is conducted through line 54 to a second heater exchanger 56. After additional heat is removed through heat exchanger 56, cooled hydroconverted raffinate is conducted through line 58 to cold hydrofinishing unit 60. Temperatures in the hydrofinishing unit 60 are more mild than those of hydroconversion units 42 and 52. Temperature and space velocity in cold hydrofinishing unit 60 are controlled to reduce the toxicity to low levels, i.e., to a level sufficiently low to pass standard toxicity tests. This may be accomplished by reducing the concentration of polynuclear aromatics to very low levels. Hydrofinished raffinate is then conducted through line 64 to separator 68. Light liquid products and gases are separated and removed through line 72. The remaining hydrofinished raffinate is conducted through line 70 to dewaxing unit 74. Dewaxing may occur by the use of solvents introduced through line 78 which may be followed by cooling, by catalytic dewaxing or by a combination thereof. Catalytic dewaxing involves hydrocracking or hydroisomerization as a means to create low pour point lubricant basestocks. Solvent dewaxing with optional cooling separates waxy molecules from the hydroconverted lubricant basestock thereby lowering the pour point. In markets where waxes are valued, hydrofinished raffinate is preferably contacted with methyl isobutyl ketone followed by the DILCHILL® Dewaxing Process developed by Exxon. This method is well known in the art. Finished lubricant basestock is removed through line 76 and waxy product through line 80. While not wishing to be bound by any theory, the factors affecting saturates, VI and toxicity are discussed as follows. The term "saturates" refers to the sum of all saturated rings, paraffins and isoparaffins. In the present raffinate hydroconversion process, under-extracted (e.g. 92 VI) light and medium raffinates including isoparaffins, n-paraffins, naphthenes and aromatics having from 1 to about 6 rings are processed over a non-acidic catalyst which primarily operates to (a) hydrogenate aromatic rings to naphthenes and (b) convert ring compounds to leave isoparaffins in the lubes boiling range by either dealkylation or by ring opening of naphthenes. The catalyst is not an isomerization catalyst and therefore leaves paraffinic species in the feed largely unaffected. High melting paraffins and isoparaffins are removed by a subsequent dewaxing step. Thus other than residual wax the saturates content of a dewaxed oil product is a function of the irreversible conversion of rings to isoparaffins and the reversible formation of naphthenes from aromatic species. To achieve a basestock viscosity index target, e.g. 110 VI, for a fixed catalyst charge and feed rates, hydroconversion reactor temperature is the primary driver. Temperature sets the conversion (arbitrarily measured here as the conversion to 370° C-) which is nearly linearly related to the VI increase, irrespective of pressure. This is shown in Fig. 3 relating the VI increase (VI HOP) to conversion. For a fixed pressure, the saturates content of the product depends on the conversion, i.e., the VI achieved, and the temperature required to achieve conversion. At start of run on a typical feed, the temperature required to achieve the target VI may be only 350° C and the corresponding saturates of the dewaxed oil will normally be in excess of 90 wt.%, for processes operating at or above 1000 psig (7.0 MPa) H 2 . However, the catalyst deactivates with time such that the temperature required to achieve the same conversion (and the same VI) must be increased. Over a 2 year period, the temperature may increase by 25 to 50° C depending on the catalyst, feed and the operating pressure. A typical deactivation profile is illustrated in Fig. 4 which shows temperature as a function of days on oil at a fixed pressure. In most circumstances, with process rates of about 1.0 v/v/hr or less and temperatures in excess of 350° C, the saturates associated with the ring species left in the product are determined only by the reactor temperature, i.e., the naphthene population reaches the equilibrium value for that temperature. Thus as the reactor temperature increases from about 350° C, saturates will decline along a smooth curve defining a product of fixed VI. Fig. 5 shows three typical curves for a fixed product of 112 VI derived from a 92 VI feed by operating at a fixed conversion. Saturates are higher for a higher pressure process in accord with simple equilibrium considerations. Each curve shows saturates falling steadily with temperatures increasing above 350° C. At 600 psig (4.24 MPa) H 2 , the process is incapable of simultaneously meeting the VI target and the required saturates (90+ wt.%). The projected temperature needed to achieve 90+ wt.% saturates at 600 psig (4.24 MPa) is well below that which can be reasonably achieved with the preferred catalyst for this process at any reasonable feed rate/catalyst charge. However, at 1000 psig (7.0 MPa) H 2 and above, the catalyst can simultaneously achieve 90 wt.% saturates and the target VI. An important aspect of the invention is that a temperature staging strategy can be applied to maintain saturates at 90+ wt.% for process pressures of 1000 psig (7.0 MPa) H 2 or above without disengagement of sour gas and without the use of a polar sensitive hydrogenation catalyst such as massive nickel that is employed in typical hydrocracking schemes. The present process also avoids the higher temperatures and pressures of the conventional hydrocracking process. This is accomplished by separating the functions to achieve VI, saturates and toxicity using a cascading temperature profile over 3 reactors without the expensive insertion of stripping, recompression and hydrogenation steps. API Group II and III basestocks (API Publication 1509) can be produced in a single stage, temperature controlled process. Toxicity of the basestock is adjusted in the cold hydrofinishing step. For a given target VI, the toxicity may be adjusted by controlling the temperature and pressure. This is illustrated in Fig. 6 which shows that higher pressures allows a greater temperature range to correct toxicity. The invention is further illustrated by the following non-limiting examples. EXAMPLE 1 This example summarizes functions of each reactor A, B and C. Reactors A and B affect VI though A is controlling. Each reactor can contribute to saturates, but Reactors B and C may be used to control saturates. Toxicity is controlled primarily by reactor C. PRODUCT PARAMETER Reactor A Reactor B Reactor C VI x x Saturates x x Toxicity x EXAMPLE 2 This example illustrates the product quality of oils obtained from the process according to the invention. Reaction conditions and product quality data for start of run (SOR) and end of run (EOR) are summarized in Tables 2 and 3. As can be seen from the data in Table 2 for the 250N feed stock, reactors A and B operate at conditions sufficient to achieve the desired viscosity index, then, with adjustment of the temperature of reactor C, it is possible to keep saturates above 90 wt.% for the entire run length without compromising toxicity (as indicated by DMSO screener result; see Example 6). A combination of higher temperature and lower space velocity in reactor C (even at end of run conditions in reactors A and B) produced even higher saturates, 96.2%. For a 100N feed stock, end-of-run product with greater than 90% saturates may be obtained with reactor C operating as low as 290°C at 2.5 v/v/h (Table 3). EXAMPLE 3 The effect of temperature and pressure on the concentration of saturates (dewaxed oil) at constant VI is shown in this example for processing the under extracted 250N raffinate feed. Dewaxed product saturates equilibrium plots (Figure 5) were obtained at 600, 1200 and 1800 psig (4.24, 8.38 and 12.5 MPa) H2 pressure. Process conditions were 0.7 LHSV (reactor A + B) and 1200 to 2400 SCF/B (214 to 427 m 3 /m 3 ). Both reactors A and B were operating at the same temperature (in the range 350 to 415°C). As can be seen from the figure it is not possible to achieve 90 wt.% saturates at 600 psig (4.14 MPa) hydrogen partial pressure. While in theory, one could reduce the temperature to reach the 90 wt.% target, the space velocity would be impractically low. The minimum pressure to achieve the 90 wt.% at reasonable space velocities is about 1000 psig (7.0 MPa). Increasing the pressure increases the temperature range which may be used in the first two reactors (reactor A and B). A practical upper limit to pressure is set by higher cost metallurgy typically used for hydrocrackers, which the process of the invention can avoid. EXAMPLE 4 The catalyst deactivation profile as reflected by temperature required to maintain product quality is shown in this example. Figure 4 is a typical plot of isothermal temperature (for reactor A, no reactor B) required to maintain a VI increase of 18 points versus time on stream. KF840 catalyst was used for reactors A and C. Over a two year period, reactor A temperatures could increase by about 50°C. This will affect the product saturates content. Strategies to offset a decline in product saturates as reactor A temperature is increased are shown below. EXAMPLE 5 This example demonstrates the effect of temperature staging between the first (reactor A) and second (reactor B) hydroconversion units to achieve the desired saturates content for a 1400 psig (9.75 MPa) H 2 process with a 93 VI raffinate feed. Reactor Sequence: Base Case Temperature Staged Case Reactor T LHSV T LHSV (C°) (v/v/h) (°C) (v/v/h) A 390 0.7 390 0.7 B 390 1.2 350 0.5 C 290 2.5 290 2.5 Dewaxed Oil Viscosity Index 114 115 Dewaxed Oil Saturates, wt% 80 96 A comparison of the base case versus the temperature staged case demonstrates the merit of operating reactor B at lower temperature and space velocities. The bulk saturates content of the product was restored to the thermodynamic equilibrium at the temperature of reactor B. EXAMPLE 6 The effects of temperature and pressure in the cold hydrofinishing unit (reactor C) on toxicity are shown in this example. The toxicity is estimated using a dimethyl sulphoxide (DMSO) based screener test developed as a surrogate for the FDA (c) test. The screener and the FDA (c) test are both based on the ultra-violet spectrum of a DMSO extract. The maximum absorbance at 345 +/- 5 nm in the screener test was shown to correlate well with the maximum absorbance bewteen 300-359 nm in the FDA (c) test as shown in Figure 8. The upper limit of acceptable toxicity using the screener test is 0.16 absorbance units. As shown in Figure 6, operating at 1800 psig (12.7 MPa) versus 1200 psig (8.38 MPa) hydrogen partial pressure allows the use of a much broader temperature range (eg. 290 to ∼360°C versus a maximum of only about 315°C when operating at 1200 psig H 2 (8.35 MPa)) in the cold hydrofmisher to achieve a non-toxic product. The next example demonstrates that higher saturates, non-toxic products can be made when reactor C is operated at higher temperature. EXAMPLE 7 This example is directed to the use of the cold hydrofinishing (reactor C) unit to optimize saturates content of the oil product. Reactors A and B were operated at 1800 psig (12.7 MPa) hydrogen partial pressure, 2400 Scf/B (427 m 3 /m 3 ) treat gas rate, 0.7 and 1.2 LHSV h -1 respectively and at a near end-of -run (EOR) temperature of 400° C on a 92 VI 250N raffinate feed. The effluent from reactors A and B contains just 85% saturates. Table 5 shows the conditions used in reactor C needed to render a product that is both higher saturates content and is non-toxic. At 350°C, reactor C can achieve 90+% saturates even at space velocities of 2.5 v/v/hr. At lower LHSV, saturates in excess of 95% are achieved. Figure 7 further illustrates the flexibile use of reactor C. As shown in Fig. 7, optimization of reactor C by controlling temperature and space velocity gives Group II basestocks EXAMPLE 8 This example demonstrates that feeds in addition to raffinates and dewaxed oils can be upgraded to higher quality basestocks. The upgrading of low value foots oil streams is shown in this example. Foots oil is a waxy by-product stream from the production of low oil content finished wax. This material can be used either directly or as a feed blendstock with under extracted raffinates or dewaxed oils. In the example below (Table 6), foots oil feeds were upgraded at 650 psig (4.58 MPa) H 2 to demonstrate their value in the context of this invention. Reactor C was not included in the processing. Two grades of foots oil, a 500N and 150N, were used as feeds. 500 N 150 N Feed Product Feed Product Temperature, °C (Reactor A/B) - 354 - 354 Treat Gas rate, Scf/B, (m 3 /m 3 ) - 500 (89) - 500 (89) Hydrogen partial pressure, psig (MPa) - 650 (4.58) - 650 (4.58) LHSV, v/v/hr (Reactor A+B) - 1.0 - 1.0 wt.% 370° C - on feed 0.22 3.12 1.10 2.00 370° C+ DWO Inspections 40° C viscosity, cSt 71.01 48.80 25.01 17.57 100° C viscosity, cSt 8.85 7.27 4.77 4.01 VI / Pour Point, ° C 97 / -15 109 / -17 111 / -8 129 / -9 Saturates, wt. % 73.4 82.8 79.03 88.57 GCD NOACK, wt. % 4.2 8.0 19.8 23.3 Dry Wax, wt. % 66.7 67.9 83.6 83.3 DWO Yield, wt. % of Foots Oil Feed 33.2 31.1 16.2 15.9 Table 6 shows that both a desirable basestock with significantly higher VI and saturates content and a valuable wax product can be recovered from foots oil. In general, since wax molecules are neither consumed or formed in this process, inclusion of foots oil streams as feed blends provides a means to recover the valuable wax while improving the quality of the resultant base oil product.

Claims: A process for selectively hydroconverting a raffinate produced from solvent refining a lubricating oil feedstock, which process comprises: (a) employing a paraffins-rich raffinate product from the solvent extraction of a lubricating oil feedstock and separation of an aromatics rich extract; (b) stripping the raffinate of solvent to produce a raffinate feed having a dewaxed oil viscosity index from 85 to 105 and a final boiling point of no greater than about 650°C; (c) passing the raffinate feed to a first hydroconversion zone and processing the raffinate feed in the presence of a non acidic catalyst at a temperature of from 340 to 420°C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), space velocity of 0.2 to 3.0 h -1 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ) to produce a first hydroconverted raffinate; (d) passing the hydroconverted raffinate from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted raffinate in the presence of a non acidic catalyst at a temperature of from 340 to 400°C provided that the temperature in the second hydroconversion is not greater than the temperature in the first hydroconversion zone, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), a space velocity of from 0.2 to 3.0 h -1 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ) to produce a second hydroconverted raffinate; (e) passing the second hydroconverted raffinate to a hydrofinishing reaction zone and conducting cold hydrofinishing of the second hydro-converted raffinate in the presence of a hydrofinishing catalyst at a temperature of from 260 to 360°C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), a space velocity of from 0.2 to 5 h -1 LHSV and hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ) to produce a hydrofinished raffinate; and, wherein said non acidic catalysts have an acidity value less than 0.5 on the scale for acidity based on the isomerization of 2-methyl-2-pentene with said catalysts at a fixed temperature. A process for producing a lubricating oil basestock having at least 90 wt.% saturates and a VI of at least 105 by selectivity hydroconverting a raffinate produced from solvent refining a lubricating oil feedstock, which process comprises: (a) employing a paraffins-rich raffinate product from the solvent extraction of a lubricating oil feedstock and separation of an aromatics rich extract; (b) stripping the raffinate of solvent to produce a raffinate feed having a dewaxed oil viscosity index from 85 to 105 and a final boiling point of no greater than about 650°C; (c) passing the raffinate feed to a first hydroconversion zone and processing the raffinate feed in the presence of a non acidic catalyst at a temperature of from 340 to 420°C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), space velocity of 0.2 to 3.0 h -1 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ) to produce a first hydroconverted raffinate; (d) passing the hydroconverted raffinate from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted raffinate in the presence of a non acidic catalyst at a temperature of from 340 to 400°C provided that the temperature in second hydroconversion is not greater than the temperature in the first hydroconversion zone, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), a space velocity of from 0.2 to 3.0 h -1 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ) to produce a second hydroconverted raffinate; (e) passing the second hydroconverted raffinate to a hydrofinishing reaction zone and conducting cold hydrofinishing of the second hydroconverted raffinate in the presence of a hydrofinishing catalyst at a temperature of from 260 to 360°C, a hydrogen partial pressure of from 1000 to 2500 psig (7.0 to 17.3 MPa), a space velocity of from 0.2 to 5 h -1 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m 3 /m 3 ) to produce a hydrofinished raffinate; (f) passing the hydrofinished raffinate to a separation zone to remove products having a boiling point less than about 250°C; and (g) passing the hydrofinished raffinate from the separation zone to a dewaxing zone to produce a dewaxed basestock having a viscosity index of at least 105 provided that the basestock has a dewaxed oil viscosity index increase of at least 10 greater than the raffinate feed, a NOACK volatility improvement over raffinate feedstock of at least about 3 wt.% at the same viscosity in the range of viscosity from 3.5 to 6.5 cSt viscosity at 100°C, and a saturates content of at least 90 wt.% and a basestock with low toxicity by passing the IP346 or FDA(c) tests; and, wherein said non acidic catalysts have an acidity value less than 0.5 on the scale for acidity based on the isomerization of 2-methyl-2-pentene with said catalysts at a fixed temperature. The process of claim 1 or claim 2, wherein there is no disengagement between the first hydroconversion zone, the second hydroconversion zone and the hydrofinishing reaction zone. The process of any preceding claim, wherein the raffinate is under-extracted. The process of any preceding claim, wherein the non acidic catalyst is cobalt/molybdenum, nickel/molybdenum or nickel/tungsten, on alumina. The process of any preceding claim, wherein the hydrogen partial pressure in the first hydroconversion zone, the second hydroconversion zone or the hydrofinishing zone is from 1000 to 2000 psig (7.0 to 12.5 MPa). The process of any preceding claim, wherein the temperature in the hydrofinishing zone is from 290 to 350°C. The process of any preceding claim, wherein the non acidic catalyst also includes at least one silica, alumina or titania.

Inventors: ALWARD SANDRA J, BOATE DOUGLAS R, CODY IAN A, DEMMIN RICHARD A, GALLAGHER JOHN E, MURPHY WILLIAM J, RAVELLA ALBERTO, ALWARD, SANDRA J., BOATE, DOUGLAS R., CODY, IAN A., DEMMIN, RICHARD A., GALLAGHER, JOHN E., MURPHY, WILLIAM J., RAVELLA, ALBERTO
Applicants: EXXONMOBIL RES & ENG CO, EXXONMOBIL RESEARCH AND ENGINEERING COMPANY
Application year: 1997
Main classification: C10G67
IPC codes: C10G65, C10M177, C10G45, C10G67, C10G73