CA2319035C - Hydroconversion process for making lubricating oil basestocks - Google Patents

Abstract

A process for producing a lubricating oil basestock having at least 90 wt % saturates and a VI of at least 105 by solvent extracting a feedstock in unit (20) to produce a raffinate, solvent dewaxing the raffinat e, selectively hydroconverting the solvent dewaxed raffinate in a two step hydroconversion zone which comprises a first hydroconversion unit (42) and to a second hydroconversion unit (52) followed by a hydrofinishing zone in hydrofinishing unit (60) and a dewaxing zone in unit (74).

Description

HYDROCONVERSION PROCESS FOR
MAKING LUBRICATING OIL BASESTOCKS
FIELD OF THE INVENTION

This invention relates to a process for preparing lubricating oil basestocks having a high satiuates content, high viscosity indices and low volatilities.
BACKGROUND OF THE INVENTION

It is well known to produce lubricating oil basestocks by solvent refining. In the conventional process, cnuie oils are ffiactionated under atmospheric pressure to produce atmospheric resids which are fiirther fiactionated under vacuum.
Select distillate fractions are then optionally deasphalted and solvent extracted to produce a paraffin rich raffinate and an aromatics rich extiact. The raffnate 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 ciude 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 chaiacteristics. Solvent refining can produce lubricating oil basestocks have 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 isoparaffnic character, i.e., those with VI's of 105 or greater.

Solvent refining alone cannot economically produce basestocks having a VI of with typical ciudes. Nor does solvent refining alone typically produce basestocks with high saturates contents. Two alternative approaches have been developed to produce high quality habricating oil basestocks; (1) wax isomerization and (2) hydrocracking. Both of the methods involve high capital investments. In some locations wax isomerization economics can be adversely impacted 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 Hydrotreatnnenf', 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 hydrotreeaating step involves a single hydrotreating zone. U.S. Patent 3,732,154 discloses hydrofinishing the extract or rafftnate from a solvent extraction process. The feed to the hydro-finishing 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.

U.S. Patent 4,636,299 discloses a process for reducing the pour point of a feedstock containing nitrogen and sulfur-containing compounds wherein the feedstock is solvent extrwted with N-me8ryl-2-pyrrolidone to produce a raffinate, the raffinate is hydrotreated to convert the nitrogen and sulfur containing compounds to ammonia and hydrogen sulfide, stripped of ammonia and hydrogen sulfide and stripped effluent cat dewaxed 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 thermat stability, fuel economy and cold start properties without incurring any significant yield debit which process requires much lower investment costs than competing technoiogies such as hydrocracking.

SU1ViMARY OF THE INVENTIQN

Thi.s invention relates to a process for producing a lubricating oil basestock meeting at least 90% saturates by selectively hydroconverting a raffinate produced from solvent refining a lubricating oil feedstock which comprises:

(a) conducting the hibricating oil feedstock to a solvent extmction zone and separating therefrom an aromatics rich extract and a paraffns rich raffinate;
(b) solvent dewaxing the raffmate under solvent dewaxing conditions to obtain a dewaxed oil feed;

(c) passing the dewaxed oil feed to a first hydroconversion zone and processing the dewaxed oil feed in the presence of a non-acidic hydroconversion catalyst at a temperahue of from 340 to 420 C, a hydrogen partial pressure of from 1000 to 2500 psig, space velocity of 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a first hydroconverted dewaxed oil;

(d) passing the hydroconverted dewaxed oil from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted dewaxed oil in the presence of a non-acidic hydroconversion catalyst at a temperature of from 340 to 400 C provided that the teanperatm in second hydroconversion is not greater than the temperature in the first hydroconvelsion zone, a hydrogen partial pressure of from 1000 to 2500 psig, a space velocity of from 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a second hydroconverted dewaxed oil;

(e) passing the second hydroconverted dewaxed oil to a hydrofinishing zone and conducting cold hydrofinishing of the second hydroconverted dewaxed oil 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, a space velocity of from 0.2 to 5 LHSV and a hydrogen to feed ratio of from 500 to 5000 ScfB to produce a hydrofinished dewaxed oil;

(f) passing the hydrofinished dewaxed oil to a separation zone to remove products having a boiling less than about 250 C; and (g) passing the hydrofinished dewaxed oil from step (f) to a dewaxing zone and catalytically dewaxing the hydrofinished dewaxed oil under catalytic dewaxing conditions in the presence of hydrogen and a catalytic dewaxing catalyst comprising a metal hydrogenation component and a crystalline 10 or 12 ring molecular sieve.

In another embodiment this invention relates to a process for producing a lubricating oil basestock meeting at least 90% saturates by selectively hydroconverting a laffinate 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 extiact and a pamffins rich raffinate;
(b) solvent dewaxing the raffinate under solvent dewaxing conditions to obtain a dewaxed oil feed;

(c) passing the dewaxed oil feed to a first hydroconversion zone and processing the dewaxed oil feed in the presence of a non-acidic hydroconversion catalyst at a temperature of from 340 to 420 C, a hydrogen partial pressure of from 1000 to 2500 psig, space velocity of 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a first hydroconverted dewaxed oil;

(d) passing the hydroconverted dewaxed oiI from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted dewaxed oil in the presence of a non-acidic hydroconversion 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, a space velocity of from 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a second hydroconverted dewaxed oil;

(e) passing the second hydroconverted dewaxed oil to a separation zone to remove products having a boiling less than about 250 C;

(f) passing the stripped second hydroconverted dewaxed oil from step (e) to a dewaxing zone and catalytically dewaxmg the stripped second hydroconverted dewaxed oil under catalytic dewaxing conditions in the presence of hydrogen and a catalytic dewa)dng catalyst comprising a metal hydrogenation component and a crystalline 10 or 12 ring molecular sieve to produce a catalytically dewaxed oil; and (g) passing the catalytically dewaxed oil to a hydrofinisbing zone and conducting cold hydrofinishing of the catalytically dewaxed oil in the presence of a hydrofinishing catalyst at a temperatim of from 260 to 360 C, a hydrogen partial pressure of from 1000 to 2500 psig, a space velocity of from 0.2 to 5 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B.

The process according to the invention produces in good yields a basestock which has VI and volatility properties meeting futare industry engine oil standards while achieving good oxidation stability, cold stark 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 DF,SCRIPTION OF THE DRAWINGS

Figure 1 is a plot of NOACK volatility vs. viscosity for a 100N
basestock.

Figure 2 is a schematic flow diagram of the hydroconversion process.
Figure 3 is a graph showing VI HOP vs. conversion at different pressures.

Figure 4 is a graph showing temperature in the first hydroconversion zone as a fiinction of days on oil at a fixed pressure.

Figure 5 is a graph showing sahuates concentration as a function of reactor temperature for a fixed VI product.

Figure 6 is a graph showing toxicity as a function of temperature and pressure in the cold hydrofinishing step.

Figure 7 is a graph showing control of saunabes concentration by vaiying conditions in the cold hydrofinishing step.

Figure 8 is a graph showing the cornelation between the DMSO
screener test and the FDA (c) test Figure 9 is a graph showing the catalytic dewaxing of dewaxed oil and total liquid products.

Figure 10 is a graph showing the comparison catalytic dewaxing a total liquid product vs. solvent dewaxing to the same pour point.

DETAII,ED DESCRIPT'ION OF THE INVENTION

The solvent refining of select crude oils to produce lubricating oil basestocks typically involves atmospheric distillation, vacuwn distillation, extraction, dewaxing and hydrofinishing. Because basestocks having a high isoparaffin content are characterized by having good viscosity index (VI) properties and suitable low tE~nperahnc 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 wl% 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 extcacted The residue from vacuam distillation which may be deasphalted is sent to other processing. Other feeds to solvent extraction include waxy streams such as dewaxed oils and foots oils.

The solvent extraction process selectively dissolves the aromatic components in an extiract 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, fiufural and N-methyl pyrrolidone. By controlling the solvent to oil ratio, extraction temperature and method of contacting distillate to be e:cftcted 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 fibm 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 I'3i, 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, deallcylation of aromatics rings, ring opening, straight chain and side-chain cracking, and wax isomerization depending on operating conditions. In view of these reac-tions, separation of the aromatics rich phase that occurs in solvent extraction is an unnecessary step since hydrocracldng reduces aromatics content to veiy low levels.

By way of contrast, the process of the present invention utilizes a three step hydroconversion of the solvent dewaxed oil produced from the raffinate from the solvent extraction unit under conditions which minimizes hydrocracking and passing waxy components remaining in the dewaxed oil tllrough the process without wax isomerization. 'Thus, dewaxed oil (DWO) and low value foots oil streams can be added to the raffinate feed to the solvent dewaxer whereby hard waxes are removed fiom the solvent dewaxer and the residual wax molecules in the solvent dewaxed oil pass unconverted through the hydroconversion process.
Removing hard wax from the raffinate feed to the hydroconversion units lessens the load on the hydroconversion units and preserves the wax as a valuable by-product.
Moreover, unlike hydrocracking, the present hydroconversion 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'/o, 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-extra.cted, i.e., the extraction is cairied out under conditions such that the raffinate yield is maximized while still removing most of the lowest quality molecules fiom the feed.
Raffinate yield may be maximized by controlling extraction conditions, for exainple, by lowering the solvent to oil treat ratio and/or decreasing the extaction temperature.
The raffinate from the solvent extraction unit is solvent dewaxed under solvent dewaxing conditions to remove hard waxes from the raffinate from the solvent extmction unit.

Solvent dewaxing typically involves mixing the raffinate feed from the solvent extraction unit with chilled dewaxing solvent to form an oil-solvent solution and precipitated wax is thereafter separated by, for example filtratioa The tempera-ture and solvent are selected so that the oil is dissolved by the chilled solvent while the wax is precipitated.

A particularly suitable solvent dewaxing process involves the use of a cooling tower where solvent is prechilled and added incrementally at several points along the height of the cooling tower. The oil-solvent mixhire is agitatied during the chilling step to permit substantially instantaneous mixing of the prechilled solvent with the oil. The prechilled solvent is added incrementally along the length of the cooling tower so as to maintain an average chilling rate at or below 10OF/minute, usually between about I to about 5 F/minute. The final temperatiue of the oil-solvendprecipitated wax mixtnre in the cooling tower will usually be between 0 and 50OF (-17.8 to 1M. The mixtiae may then be sent to a scraped smface chiller to sepazate precipitated wax from the mixhme.

In general, the amount of solvent added will be sufficient to provide a liquid/solid weight ratio between the range of 511 and 20/1 at the dewaxing tempera-ture and a solvent/oil volume ratio between 1.511 to 511. The solvent dewaxed oil is typically dewaxed tio an mteamediabe pour point, preferably less than about +10 C.

Representative dewaxing solvents are aliphatic ketones having 3-6 carbon atoms such as methyl ethyl ketone and methyl isobutyl ketoney low molecular weight hydrocarbons such as propane and butane, and mixhures thereof. The solvents may be mixed with other solvents such as benzene, toluene or xylene.
Further descriptions of solvent dewaxing process useful herein are disclosed in U.S.
Patents 3,773,650 and 3,775,288.

The dewaxed oil feed is then sent to a first hydroconversion unit containing a hydroconversion catalysL This dewaxed oil feed has a viscosity index of from about 85 to about 105 and a boiling range not to exceed about 650 C, prefer-ablY less thm 60(rC, as deteimined by ASTM 2887 and a viscosity of from 3 to 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 mixhues thereof. These metals or mixtures of metals are typically present as oxides or sulfides on reffiactory 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 methyl2-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 temperahue, typically 200 C. In the presence of catalyst sites, 2 me8ry12 pe,ntene fonns 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 sirongly 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 4Ø
Very wealdy acidic sites will have values near 0.0 whereas very strongly acid.ic sites will have values approaching 4Ø The catalysts usefal 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 dopant.s include halogen, especially fluorine, phosphorus, boron, ydria, rare-earth oxides and magnesia. Promoters such as halogens generally increase the acidity of metal oxide supports while mildly basic dopants such as ydria or magnesia tend to decrease the acidity of such supports.

Suitable metal oxide supports include low acidic oxides such as silica, alumina or titmiia, preferably alumina. Preferred alimninas are porous aluminas such as gamma or eta having average pore sizes from 50 to 200J~ preferably 75 to 150.A, a surface area from 100 to 300 m~/g, preferably 150 to 250 m2/g and a pore volume of from 0.25 to 1.0 cm3/g, preferably 0.35 to 0.8 cm3/g. The supports are preferably not promoted with a halogen such as fluorine as this generally increases 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% Co as oxide) or nickel/hmgsten (1-5% Ni as oxide, 10-3(rW as oxide) on alumina.
Especially preferred are nickel/molybdennm catalysts such as KF-840.

Hydroconversion conditions in the first hydroconversion unit include a temperature of from 340 to 420 C, preferably 350 to 4009C, a hydrogen partial pressure of finm 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 1.0 LHSV, and a hydrogen to feed ratio of from 500 to 5000 Scf/B (89 to 890 m3/m3), preferably 2000 to 4000 Scf/B (356 to 712 m3/m3).

The hydroconverted dewaxed oil from the first hydroconversion unit is conducted to a second hydroconversion unit. The hydroconverted dewaxed oil 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 temperateues, if desired. Tempeivtues in the second hydroconversion unit should not exceed the temperature used in the first hydroconversion unit Conditions in the second hydroconversion unit include a temperatare 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, and a hydrogen to feed ratio of from to 5000 Scf/B (89 to 890 m3/m), preferably 2000 to 4000 Scf/B (356 to 712 m3/m).
'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 hydroconverbed dewaxed oil from the second hydroconversion unit may then conducted to a cold hydrofinishing unit Alternatively, cold hydro-finishing may be deferred until afler the catalytic dewaxing step. A heat exchanger is preferably located between these units. Reaction conditions in the hydrofinishing unit are mild and include a tempeiature of from 260 to 360 C, preferably 290 to 350 C, more preferably 290 to 330 C, a hydrogen partial pressare 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, preferably 0.7 to 3.0 LHSV, and a hydrogen to feed ratio of from 500 to 5000 SCFB (89 to 890 m3/m), preferably 2000 to 4000 Scf/B (356 to 712 m3/m). 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 ahimina, zimonia and the like may be used in the cold hydrofinishing unit In order to prepare a finished basestock, the hydrofinished oil from the hydrofinishing unit is conducted to a separator, e.g., a vacuum stripper (or fractiona-tion) to separate out low boiling products if the separator is followed by a catalytic dewaxing step. Such products may include hydrogen sulfide and ammonia fonned in the first two reactors. If desired, a stripper may be situated between the second hydroconversion unit and the hydrof nnshing unit, but this is not essential to produce basestocks according to the invention.

The hydrofinished dewaxed oil seperated from the separator is then conducted to a dewaxing unit. Catalytic dewaxing, solvent dewaxing or a combina-tion may accomplish dewaxing thereof.

The catalysts useful in the catalytic dewaxing step include crystalline and 12 ring moleciilar sieves and a metal hydrogenation component. Crystalline molecular sieves include alumino silicates and aluminum phosphates. Examples of crystalline alumino silicates include zeolites such as ZSM-5, ZSM-11, ZSM-12, theta 1(ZSM-22), ZSM-23, ZSM-35, felrierite, ZSM-38, ZSM-48, ZSM-57, beta, mordenite and offretite. Examples of crystaliine aluminum phosphates include SAPO-11, SAPO-41, SAPO-3 1, MAPO-11 and MAPO-3 1. Preferred molecular sieves include ZSM-5, theta 1, ZSM-23, ferri.erite and SAPO-1 1.

The dewaxing catalyst may also contain an amorphous component.
The acidity of the amorphous component is preferably from 0.3 to 2.5, preferably 0.5 to 2.0 on the Kramer/McVicker acidity scale described above. Examples of amorphous materials include silica-alumina, halogenated alumina, acidic clays, silica-magnesia, yqria silica-alumina and the like. Especially preferred is silica-alumina.

If the dewaxing catalyst contains an amorphous component, the crystalline molecular sieve,lmetal hydrogenation component(amorphous component may be composited together. The hydrogenation metal can be deposited on each component separately or can be deposited on the composited catalyst. In the altema-tive, the crystalline molecular sieve and ffinorphous component can be in a layered configuration. Preferably, the top layer in the reaction vessel is the amorphous component and the lower layer is the crystalline molecular sieve, although the configuration can be rewersed with the top layer as the molecular sieve and the bottom layer as the amorphous component. In the layered configuration, the hydrogenation metal should be deposited on both the molecular sieve and the amorphous component The metal hydrogenation component of the dewaxing catalyst may be at least one metal from the Group VIB and Group VIII of the Periodic Table (published by Sargent-Welch Scientific Company). Prefeired metals are Group VIII
noble metals, especially palladium and platinum.

The decwaxing catalyst may contain, based on the weight of total catalyst, fmm 5 to 95 wt% of crystalline molecular sieve, from 0 to 90 wt% of amorphous component and from 0.1 to 30 wt'/ of metal hydrogenation component with the balance being matrix material.

The dewaxing catalyst may also include a matrix or binder which is a.
material resistant to process conditions and which is substantially non-catalytic under reaction conditions. Matrix mateTials may be synthetic or naturally occurring materials such as clays, silica and metal oxides. Matrix matetials which are metal oxides include single oxides such as ahmnina, binary compositions such as silica-magnesia and ternary compositions such as silica-alumina-zirconia.

Process conditions in. the catalytic dewaxing zone include a tempera-ture of from 240 to 420 C, preferably 270 to 400 C, a hydrogen partial pressure of from 3.45 to 34.5 mPa (500 to 5000 psi), preferably 5.52 to 20.7 mPa, a liquid hourly space velocity of from 0.1 to 10 v/v/br, preferably 0.5 to 3.0, and a hydrogen circulation rate of from 89 to 1780 m3/m3 (500 to 10000 scf/B), preferably 178 to 890 m3Im3.

The final catalytic dewaxing step may be followed by a second cold hydrofinishing step under the cold hydrofinishing conditions described above.
This second cold hydrofinishing step would be used in those instances where needed to meet product quality requirements such as color or light stability.

In an alteinative embodiment, hydroconverted dewaxed oil from the second hydroconversion unit is conducted to a separator to separate low boiling components such as ammonia and hydrogen sulfide. The stripped hydroconverted dewaxed oil is then sent to a catalytic dewaxing unit and catalytically dewaxed under the conditions set fordl above. The catalytically dewaxed oil from catalytic dewax-ing can then be cold hydrofinished as described above.

The lubricating oil basestock produced by the process according to the invention is chara.cterized by the following properues: viscosity index of at least about 100, preferably at least 105 and sabnates of at least 90 /g preferably at least 95 wt%, NOACK volatility improvement (as measured by DIN 51581) over solvent dewaxed oil feedstock of at least about 3 wt%, preferably at least about 5 wt%y 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. 7 332(1984) for further details]. The present process reduces these polycyclic aromatic compounds to such levels as to pass carcino-genicity tiests. 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 Figure 1, NOACK volatility is related to VI for any given basestock. The relationship shown in Figure 1 is for a light basestock (about 100N). If the goal is to meet a 22 wt'/o 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., to 5(roff 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 desh oy 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 fiuther illustrated by Figure 2. The feed 8 to vacuam pipestill 10 is typically an atmospheric reduced c:ude 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 6509C. The bottoms from vacunm 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 gneater 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 exhu-tion 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 exhact. The present process permits the extrac-tion unit to operate to an "under extraction" mode, i.e., a greater amount of aromatics in the paiaffins rich raffnate 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 39 to solvent dewaxing unit 40.

Solvent dewauing unit 40 is a oooling tower wherein chilled solvent is added at several points along the height of the unit 40 through line 41.
Precipitated wax is removed through line 45 while dewaxed oil is sent to first hydroconveision unit 42 through line 43.

TM
The first hydroconversion unit 42 contains KF-840 catalyst which is nickel/molybd.enum 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 dewaxed oil indic,ate 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 occius in this stage is not lcnown with certainty, it is known that heteroatoms are being removed, aromatic rings are being satwated and naphthene rings, particularly multi-ring naphthenes, are selectively eliminated.

Hydroconverted dewaxed oil from hydroconversion unit 42 is conducted through line 46 to heat exchanger 48 where the hydroconverted dewaxed oil stream may be cooled if desired. The cooled hydroconverted dewaxed oil stream is conducted ttirough line 50 to a second hydroconversion unit 52. Additional hydrogen, if needed, is added through line 53. This second hydroconversion unit is operated at a lower temperature (when required to adjust product quality) than the first hydroconversion 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 hydro-carbon species back towards increased saturates concentration. In this way, the concentration of sat<ntes can be maintained at greater than 90 wC% by appropriately controlling the combination of temperature and space velocity in second hydro-conversion unit 52.

Hydroconverted dewaxed oil from unit 52 is conducted through line 54 to a second heater exchanger 56. Alteinatively, hydroconverted dewaxed oil from unit 52 can be sent directly through line 55 to separator 68. After additional heat is removed through heat exchanger 56, cooled hydroconverted dewaxed oil 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 hydro~nishing 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 poly-nuclear aromatics to very low levels.

Hydrofinished dewaxed oil is then conducted through line 64 to separatior 68. Light liquid products and gases are separated and removed through line 72. The reroaining hydrofini' shed dewaxed oil is conducted through line 70 to catalytic dewaxing unit 74. Catalytic dewaxing involves selective hydrocracking with or without hydroisomerization as a means to create low pour point lubricant basestocks. Finished lubricant basestock is removed through line 76. If hydro-converted raffinate from unit 52 is sent directly to separator 68 through line 55, then basestock removed through line 76 can be sent to cold hydrofinishing (not shown).
While not wishing to be bound by any theory, the factors affecting saturates, VI and toxicity are discussed as follows. The tenm "saturates"
refers to the sum of all saturated rings, paraffins and isoparaffins. In the present raffinate hydro-conversion process, under-extiracted (e.g., 92 Vl) light and medium raffinates including isoparaffns, n paraffms, naphthenes and aromatics having from 1 to about 6 rings are processed over a non-acidic catalyst which prinlarily 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 satiu ates content of a dewaxed oil product is a fimction 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 37(rC-) which is nearly linearly related to the VI increase, irrespective of pressure. This is shown in Figure 3 relating the VI increase (VI HOP) to conversion.
For a fixed pressure, the satnrates content of the product depends on the conversion, i.e., the VI achieved, and the temperattnre 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 comesponding saturates of the dewaxed oil will normally be in excess of 90 wV/o, for processes operating at or above 1000 psig (7.0 mPa) H2.
However, the catalyst deactivates with time such that the tempeiature required to achieve the sanae conversion (and the same VI) must be increased. Over a 2 year penod, 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 Figure 4 which shows temperatare as a fimction 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,1he sahaates associated with the ring species left in the product are detezmined only by the reactor temperature, i.e., the naphthene population reaches the equilibrium value for that temperattnm.

Thus as the reactor temperature increases from about 350 C, saturates will decline along a smooth curve defining a product of fixed VI. Figure 5 shows three typical curves for a fixed product of 112 VI derived from a 92 VI feed by operating at a fixed conversion. Satiu-ates are higher for a higher pressure process in accord with simple eqnilibrium considerations. Each curve shows saturates falling steadily with temperatures increasing above 350 C. At 600 psig (4.24 mPa) H2, the process is incapable of sinmultaneously meeting the VI target and the required saturates (90+ wt%). The projected temperatm needed to achieve 90+ wt'/o 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%atalyst charge. However, at 1000 psig H2 and above, the catalyst can simultaneously achieve 90 wt% satinaties and the target VI.

An impoitant aspect of the invention is ttiat a temperatnre staging strategy can be applied to maintain saturates at 90+ wt% for process pressures of 1000 psig (7.0 mPa) H2 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 tempeiatiu+es and pressures of the conventional hydrocracldng process.
This is accomplished by separating the fimctions to achieve VI, saturates and toxicity using a cascading temperature profile over 3 reactors without the expensive inserdon of stripping, recompression and hydrogenation steps. API Group II and III base-stocks (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 illushated in Figure 6 which shows that higher pressures allows a greater teanperature range to correct toxicity.

The invention is further illusuated by the following non-limiting examples.

This example summarizes fimctions of each reactor A, B and C.
Reactors A and B affect VI though A is controlling. Each reactor can contribute to sabuutes, but Reactors B and C ma.y be used to control satiuates. Toxicity is controlled primarily by reactor C.

Product Paratneter Reactor A Reactor B Reactor C
VI X X
Saturates X X
Toxicity x 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 nm (EOR) are swnmarized 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 adjmtnent of the temperatare of reactor C, it is possible to keep saturates above 90 wl'/o for the entire run length without compromising toxicity (as indicated by DMSO screener result; see Example 6). A combination of higher ~ and lower sps.ce velocity in reactor C (even at end of run conditions in reactors A and B) produced even higher satiuates, 96.2%. For a 100N feed stock, end-of-nn product with greater tU.an 90% satUrabes may be obtained with reactor C
operating as low as 290C at 2.5 v/v/h (Table 3).

SOR ---- EOR ------ EOR EOR ----Temp. LHSV Temp. LHSF Temp. LHSV Temp. LHSV
Rea~or C m m ~ v co A 352 0.7 400 0.7 400 0.7 400 0.7 B 352 1.2 400 1.2 400 1.2 400 1.2 C 290 2.5 290 2.5 350 2.5 350 1.0 * Other Conditions: 1800 psig (12.5 mPa) H2 inlet pressure, 2400 scf/b (427 m3/m3) Dewaxed Oil Properties 250N(1) Feed Q F.R EOR EOR
N p 100 C Viscosity, cSt 7.34 5.81 5.53 5.47 5.62 40 C V'iscosity, cSt 54.41 34.28 31.26 30.63 32.08 Viscosity Index 93 111 115 115 114 PourPoint, C -18 -18 -16 -18 -19 Saturates, wt% 58.3 100 85.2 91 96.2 DMSO Screener for 0.30 0.02 0.06 0.10 0.04 toxicity(2) 370 C Yield, wt. on 100 87 81 81 82 raffinate feed (1) 93 VI under extracted feed.
(2) Maximum ultra-violet absorbance at 340 to 350 nm.
Ir SOR ------- -EOR --------Temp. LHSV Temp. LHSV
Reactor C v/v/hr C v/v/hr A 355 0.7 394 0.7 B 355 1.2 394 1.2 C 290 2.5 290 2.5 * Other Conditions: 1800 psig (12.5 mPa) H2 inlet pressure, 2400 scf/B (427 m3/m3).
Dewaxed Oil Pro,perties 100N (1) Feed SOR EOR
100 C Viscosity, cSt 4.35 3.91 3.83 40 C Viscosity, cSt 22.86 18.23 17.36 Viscosity Index 95 108 112 Pour Point, C -18 -18 -18 Satiuites, wC'/o 64.6 99 93.3 DMSO Screener for 0.25 0.01 0.03 toxicity (2) 370 C+ Yield, wt% on 93 80 75 raffinabe feed (1) 95 VI under extracted feed.
(2) Iviaximum ultra-violet absorbance at 340 to 350 nm.

The effect of tetnperatare 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) pressure. Process conditions were 0.7 LHSV (reactor A + B) and 1200 to 2400 SCFB (214 to 427 m3/m). Both reactors A and B were operating at the same temperatm~e (in the range 350 to 415 C).

As can be seen from the figure it is not possible to achieve 90 wt%
sahuates at 600 psig (4.14 mPa) hydrogen partial pressure. While in ffieory, one could reduce the temperature to reach the 90 wt'/o 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 metallwgy typically used for hydrocrackers, which the process of the invention can avoid.

The catalyst deactivaxion 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 versvs time on stream. KF840 catatyst was used for reactors A
and C. Over a two year period, reactor A tempeiazares could increase by about 50 C. This will affect the product satarates content. Strategies to offset a decline in product sahumtes as reactor A temperature is increased are shown below.

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) HZ process with a 93 VI
raffinate feed.

Base Temperature Reactor Sequence: Case Staged Case Rea~or T LHSV T LHSV
cci v/v/h C
A 390 0.7 390 0.7 B 390 1.2 350 0.5 C 290 2.5 290 2.5 Dewaxed Oil Viscosity 114 115 Index Dewaxed Oil Satumates, wt% 80 96 A comparison of the base case versus the temperature staged case demonstrates the merit of operating reactor B at lower tempecature and space velocities. The bulk sahuates content of the product was restored to the thermodynamic equilibrium at the temperature of reactor B.

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 between 300-359 nm in the FDA (c) test as shown in Figure S. The upper limit of acceptable toxicity using the scre.ener 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 parkial pressure allows the use of a much broader temperature range (e.g., 290 to -360 C
versus a maximum of only about 315 C when operating at 1200 psig H2 (8.35 Mpa)) in the cold hydrofinisher to achieve a non-toxic product. The next example demonstrates that higher satmmtes, non-toxic products can be made when reactor C
is operabed at higher temperature.

This example is directed to the use of the cold hydrofinishing (reactor C) unit to optimize satutates content of the oil product. Reactois A and B
were operated at 1800 psig (12.7 mPa) hydrogen partial pressure, 2400 Scf/B (427 m3/m) treat gas rate, 0.7 and 1.2 L.HSV respectively and at a near end-of -nm (EOR) temperature of 400 C on a 92 VI 250N raffinate feed. The effluent from reactors A
and B contains just 85% sahnates. 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, sattuabes in excess of 95% are achieved.

RUNS
Run Number 1 2 3 4 Tempe>nhue, C 290 330 350 350 LHSV, v/v/br 2.5 2.5 2.5 1.0 H2 Press, pslg 1800 1800 1800 1800 Treat Gas Rate, SCF/B 2400 2400 2400 2400 DWO Saturates, wt% 85 88 91 96 DMSO Screener for Toxicity(1) 0.06 0.05 0.10 0.04 (1) Maximum ultra-violet absorbance at 340-350 nm Figure 7 further illustrates the flexible use of reactor C. As shown in Figure 7, opti.mization of reactor C by controlling temperathm and space velocity gives Group II basestocks This example demonstiates tllat 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 exftwted raffulates or dewaxed oils. In the example below (Table 6), foots oil feeds were upgraded at psig (4.58 mPa) H2 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.

Feed Product Feed Product Temperahue, C (Reactor A/B) - 354 - 354 Treat Gas rate (TGR), Saf/B, (m3/m3) - 500(89) - 500 (89) Hydrogen PartiW 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 Insoections 400C 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(2) 111/-8 129/-9P) Sat rates, w't% 73.4 82.8() 79.03 88.57(1) GCD NOACK, wt% 4.2 8.0 19.8 23.3 Dr'y Wax, ~/o 66.7 67.9 83.6 83.3 DWO Yield, wt'/o of Foots Oil Feed 33.2 31.1 16.2 15.9 Saturates improvement will be higher at higher hydrogen pressures (2) Excellent blend stock Table 6 shows that both a desirable basestock with significanfly higher VI and sabuates content and a valuable wax product can be recovered from foots oil.

In general, since wax molecules are neither consumed or fonmed in this process, inclusion of foots oil streams as feed blends provides a means to recover the vaiuable wax while improving the quality of the resultant base oil product.

This example illustrates the advantage of catalytic trim dewaxing a solvent dewaxed hydrotreated raffnate. The trim catalytic dewaxed products, even though they have lower VI, have much better low temperature properties ( products as defined by lower Brookfield Viscosity ) than the corresponding solvent dewaxed feed. Trim dewaxing refers to the process of solvent dewaxing followed by catalytic dewaxing.

A raffinate product made under the conditions in Table 7 was topped at 370 C to give a 370 C+ product which was solvent dewaxed using 1VIIBK in a 3:1 solvent to raffmate product ratio and a filter temperature of -21 C to make a dewaxed oil having the properties shown in Table 8.

Process Conditions Rl Conditi ons Pressure, psig 1800 (12.4 mPa) TGR, scf/B 2500 (445 m3/m) Space Velocity, v/v/h 0.7 Temperature, C 375 R2 Conditions Pressute, psig 1800 TGR, scf/B 2400 (427 m3/m) Space Velocity, v/v/h 2.5 Temperahue, C 290 Product Pmperties Viscosity, cSt at 100 C, 4.182 Viscosity, cSt at 40 C, 20.495 SUS, cP at 100 F 107.7 vi 106 Pour Point, C -19 Brookfield Viscosity, at -40 C 39900 This dewaxed oil was then catalytically dewaxed over a 0.5 wt% Pt TON (zeolite) / Pt Silica-alumina (25:75 wdwk zeolite: silica-alumina ) mixed powder c,omposite catalyst under the conditions shown in Table 9 and to produce the products, after fisctionation at 370 C, shown in Table 9.

Process Conditions Pressure, psig 1000 1000 (7.0 mPa) TGR, scf/B 2500 2500 (445 m3/m3) Space Velocity, v/v/h 1.0 1.0 Temperatcue, C 295 303 Yield, wt% 67 60 Product Pro~
Viscosity, cSt at 100 C, 4.150 4.122 Viscosity, cSt at 40 C, 20.634 20.441 SUS, cP at 100 F 108.4 107.5 vi 101.7 101.3 Pour Point, C -33 -40 Brookfield Viscosity, cP at -40 C 32100 22900 The dewaxed oils, both feed and products from the catalytic dewaxer were formulated as Automatic Transmission Fluids using a Ford type ATF ad pack (22 wt% treat rate of ATF ad pack, 78 wt% dewaxed oil) and Brookfield Viscosities at -40 C measured. The Brookfield Viscosities for both feed and products are shown in Tables 8 and 9 respectively.

This example illustrates the advantage of catalytic dewaxing a total liquid product produced from hydrotreating a raffinate over the process described in Example 9. Catalytic dewaxing is shown to give a product with improved VI over that obtained by solvent dewaxing at the same pour points. In addition, the catalytic dewaxed products have much better low temperature properties (as defined by lower Brookfield Viscosity ) than the corresponding solvent dewaxed product.

A hydrotreated raffinate product was made under the conditions listed in Table 10.

Process Conditions Rl Conditions Presmre, psig 1800 (12.4 mPa) TGR, scf/B 2400 (427 m3/m) Space Velocity, v/v/h 0.7 Temperature, C 382 R2 Conditions Pressare, psig 1800 TGR, scf/B 2400 Space Velocity, v/v/h 2.5 Temperature, C 290 The hydrotreated raffinate total liquid product made under the condi-tions in Table 10 was topped at 370 C to give a 370 C+ product which was solvent dewaxed using MIBK in a 3:1 solvent to raffinate product ratio and a filter tempera ture of -21 C. to make a dewaxed oil having the properties shown in Table 11.

Product Properties Viscosity, cSt at 100 C, 3.824 Viscosity, cSt at 40 C, 17.5 SUS, cP at 100 F 93.5 VI 109.3 Pour Point, C -19 Yield on TLP, wN/o 65.5 Brookfield Viscosity, cP at -40 C 26800 The total liquid product fiam this step was then catalytically dewaxed over a 0.5 wt'/o Pt TON (zeolite) / Pt Silica-alumina (25:75 wr/wt zeolite:silica-alumina) mixed powder composite catalyst under the conditions shown in Table and to produce the products, after topping at 370 C, shown in Table 11.

Process Conditions Pressm, psig 1000 1000 1000 (7.0 mPa) TGR, scf/B 2500 2500 2500 (445 m3/m3) Space Velocity, v/v/h 1.0 1.0 1.00 Temperature, C 304 306 314 Yield, wt% 48.2 46.3 33.5 Product Prooerties Viscosity, cSt at 100 C 3.721 3.672 3.593 Viscosity, cSt at 40 C, 16.511 16.256 15.925 SUS, cP at 100 F 89.0 87.8 86.4 vi 112.6 111 107.0 Pour Point, C -20 -23 -39 Brookfield Viscosity, at -40 C 13640 12740 10600 The dewaxed oils, both solvent dewaxed and the products from the catalytic dewaxer, were fonnulated as Automatic Transmission Fluids using a Ford type ATF ad pack (22 wt% treat rate of ATF ad pack, 78 wt% dewaxed oil) and Brookfield Viscosities at -40 C measured. The Brookfield Viscosities for both feed and products are shown in Tables 5 and 6 respectively.

Figure 9 shows the benefit of catalytic dewaxing both the DWO and total liquid products. Comparing the data in Examples 9 and 11(Tables 9 and 12) shows a fiirther benefit for dewaxing a TLP vs. a DWO in that the fonner results in products having a higher VI at the same pour point Catalydc dewaxmg also improves the VI of the products from dewaxing a TLP over that obtained by solvent dewaxing.

This example further illustrates the advantage of catalytic dewaxing a total liquid product veisos solvent dewaxing to the same pour point Catalytic dewaxing is shown to give a product with improved VI over that obtained by solvent dewaxing at the same pour points. In addition, the catalytic dewaxed products have much better low temperature properties (as defined by lower Brookfield Viscosity ) than the conesponding solvent dewaxed product A hydrotreated raffinate product was made under the conditions listed in Table 10.

Process Conditions Ri Conditions Pressure, psig 1800 (12.5 mPa) TGR, scf/B 2400 (427 m3/m) Space Velocity, v/v/h 0.7 Temperature, C 382 R2 Conditions Pressare, psig 1800 TGR, scf/bbl 2400 Space Velocity, v/v/h 2.5 Temperature, C 290 The hydrotreated raffinate totai liquid product made under the condi-tions in Table 4 was topped at 370 C to give a 370 C+ product which was solvent dewaxed using MIBK in a 3:1 solvent to raffinate product ratio and a filter tempera-ture of -21 C to make a dewaxed oil having the properties shown in Table 14.

Product Properties Viscosity, cSt at 100 C, 5.811 Viscosity, cSt at 40 C, 34.383 SUS, cP at 100 F 177 vi 110.6 Pour Point, C -21 Yield on TLP, wt% 64.6 Brookfield Viscosity, cP at -40 C 148200 The total liquid product from this step was then catalytically dewaxed over a 0.5 wt% Pt TON ( zeolite )/ Pt Silica-alumina ( 25:75 wdwt~ zeolite:
silica-alumina ) mixed powder composite catatyst under the conditions shown in Table and to produce tbe products, affter topping at 370 C, shown in Table 11.

Process Conditions Pressure, psig 1000 1000 1000 (7.0 mPa) TGR, scf/B 2500 2500 2500 (445 m3/m) Space Velocity, v/v/h 1.0 1.0 1.00 Temperature, C 304 306 314 Yield, wt% 48.2 46.3 33.5 Product Pr4Per~'es Viscosity, cSt at 100 C, 5.309 5.261 5.115 Viscosity, cSt at 40 C, 28.899 28.552 27.364 SUS, cP at 100 F 148.9 147.2 141.2 VI 117.6 117.0 116.4 Pour Point, C -13 -20 -18 Brookfield Viscosity, at -40 C 47150 35650 38150 The dewaxed oils, both solvent dewaxed and the products from the catalytic dewaxer, were formulated as Automatic Transmission Fluids using a Ford type ATF ad pack (22 wt% treat rate of ATF ad pack, 78 wt% dewaxed oil) and Brookfield Viscosities at -40 C measured. The Brookfield Viscosities for both feed and products are shown in Tables 14 and 15 respectively.

Figure 10 is a graphical illustration of the results from Example 11.
This example also illustzates the benefit of catalytic dewaxing versus solvent dewaxing in that the VI of the products from catalytic dewaxing are higher than that obtained by solvent dewaxing.

Claims (18)

CLAIMS:

1. A process for producing a lubricating oil basestock meeting at least 90% saturates 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) solvent dewaxing the raffinate under solvent dewaxing conditions to obtain a dewaxed oil feed;

(c) passing the dewaxed oil feed to a first hydroconversion zone and processing the dewaxed oil feed in the presence of a non-acidic hydroconversion catalyst at a temperature of from 340 to 420°C, a hydrogen partial pressure of from 1000 to 2500 psig, space velocity of 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a first hydroconverted dewaxed oil;

(d) passing the hydroconverted dewaxed oil from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted dewaxed oil in the presence of a non-acidic hydroconversion 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, a space velocity of from 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a second hydroconverted dewaxed oil;

(e) passing the second hydroconverted dewaxed oil to a hydrofinishing zone and conducting cold hydrofinishing of the second hydroconverted dewaxed oil 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, a space velocity of from 0.2 to 5 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a hydrofinished dewaxed oil;

(f) passing the hydrofinished dewaxed oil to a separation zone to remove products having a boiling less than about 250°C; and (g) passing the hydrofinished dewaxed oil from step (f) to a dewaxing zone and catalytically dewaxing the hydrofinished dewaxed oil under catalytic dewaxing conditions in the presence of hydrogen and a catalytic dewaxing catalyst comprising a metal hydrogenation component and a crystalline 10 or 12 ring molecular sieve.

2. A process for producing a lubricating oil basestock meeting at least 90% saturates 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) solvent dewaxing the raffinate under solvent dewaxing conditions to obtain a dewaxed oil feed;

(c) passing the dewaxed oil feed to a first hydroconversion zone and processing the dewaxed oil feed in the presence of a non-acidic hydroconversion catalyst at a temperature of from 340 to 420°C, a hydrogen partial pressure of from 1000 to 2500 psig, space velocity of 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a first hydroconverted dewaxed oil;

(d) passing the hydroconverted dewaxed oil from the first hydroconversion zone to a second hydroconversion zone and processing the hydroconverted dewaxed oil in the presence of a non-acidic hydroconversion 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, a space velocity of from 0.2 to 3.0 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B to produce a second hydroconverted dewaxed oil;

(e) passing the second hydroconverted dewaxed oil to a separation zone to remove products having a boiling less than about 250°C;

(f) passing the stripped second hydroconverted dewaxed oil from step (e) to a dewaxing zone and catalytically dewaxing the stripped second hydroconverted dewaxed oil under catalytic dewaxing conditions in the presence of hydrogen and a catalytic dewaxing catalyst comprising a metal hydrogenation component and a crystalline 10 or 12 ring molecular sieve to produce a catalytically dewaxed oil; and (g) passing the catalytically dewaxed oil to a hydrofinishing zone and conducting cold hydrofinishing of the catalytically dewaxed oil 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, a space velocity of from 0.2 to 5 LHSV and a hydrogen to feed ratio of from 500 to 5000 Scf/B.

3. The process of claim 1 wherein there is no disengagement between the first hydroconversion zone, the second hydroconversion zone and the hydrofinishing reaction zone.

4. The process of claim 1 or 2 wherein the basestock contains at least 95 wt% saturates.

5. The process of claim 1 or 2 wherein the raffinate is under-extracted.

6. The process of claim 1 or 2 wherein the non-acidic hydro-conversion catalyst is cobalt/molybdenum, nickel/molybdenum or nickel/tungsten on alumina.

7. The process of claim 1 or 2 wherein the hydrogen partial pressure in the first hydroconversion zone, the second conversion zone or the hydrofinishing zone is from 1000 to 2000 psig.

8. The process of claim 1 or 2 wherein the non-acidic hydro-conversion catalyst has an acidity less than about 0.5, said acidity being determined by the ability of the catalyst to convert 2-methyl-2-pentene to 3-methyl-2-pentene and 4-methyl-2-pentene and is expressed as the mole ratio of 3 methyl-2 pentene to 4-methyl-2 pentene.

9. The process of claim 1 or 2 wherein the dewaxing catalyst is a ZSM-5, ZSM-11, ZSM-12, Theta-1, ZSM-23, ZSM-35, ferrierite, ZSM-48, ZSM-57, beta, mordenite or offretite zeolite.

10. The process of claim 1 or 2 wherein the dewaxing catalyst is a SAPO-11, SAPO-31 or SAPO-41 aluminum phosphate.

11. The process of claim 1 or 2 wherein the dewaxing catalyst is a composite of a crystalline molecular sieve and an amorphous component.

12. The process of claim 1 or 2 wherein the dewaxing catalyst is layered catalyst containing a first layer of amorphous component and a second layer of crystalline molecular sieve.

13. The process of claim 1 or 2 wherein the metal hydrogenation component of the dewaxing catalyst is at least one of a Group VIB and Group VIII
metal.

14. The process of claim 1 wherein the catalytic dewaxing step is followed by a cold hydrofinishing step.

15. The process of claim 11 wherein the amorphous component of the dewaxing catalyst is silica-alumina, silica magnesia, halogenated alumina, yttria silica-alumina or a mixture thereof.

16. The process of claim 13 wherein the metal hydrogenation component is at least one of Pt or Pd.

17. The process of claim 1 or 2 wherein solvent dewaxing comprises mixing the raffinate with a chilled solvent to form an oil-solvent solution mixed with precipitated wax, separating precipitated wax from the oil-solvent solution, and separating the solvent from the solvent-oil solution thereby forming a solvent dewaxed oil.

18. The process of claim 17 wherein the solvent is at least one of propane, butane, methyl ethyl ketone, methyl isobutyl ketone, benzene, toluene and xylene.