Solvent Extraction in the Petroleum Industry

Solvent Extraction in the Petroleum Industry G. C. GESTER, JR.

Downloaded by UNIV OF MICHIGAN ANN ARBOR on October 31, 2014 | http://pubs.acs.org Publication Date: January 1, 1951 | doi: 10.1021/ba-1951-0005.ch016

California Research Corp., Richmond, Calif.

Development of solvent extraction processes in the petroleum industry and theoretical aspects of solvent extraction are reviewed. Six extraction processes which have received industrial acceptance are described and performance characteristics of furfural, phenol, and Duosol processes are compared. Data are presented to demonstrate the applicability of adsorption analyses for stock evaluation and prediction of commercial extraction yields. Correlations for predicting solvent requirements and layer compositions and process design and engineering considerations are included. The desirability of further fundamental work to facilitate design calculations from physical data is suggested.

S o l v e n t extraction is used extensively i n the petroleum industry to refine lubricating oils, kerosene, and specialty oils for medicinal and agricultural purposes. I t is a process that separates hydrocarbons into two phases—a raffinate which contains substances of high hydrogen to carbon ratio and an extract which contains substances of low hydrogen to carbon ratio. This paper discusses developments i n solvent extraction with emphasis on applications to petroleum fractions i n the kerosene and heavier boiling range. M a n y of these techniques have been extended to the refining of gasoline, Diesel fuels, catalytic cracker feed and cycle stocks, and butadiene. E a r l y refiners utilized simple batch distillation to prepare kerosenes and lubricating oils. A s the demand for these materials expanded and new crude oils were found, certain desirable and undesirable characteristics became apparent. Crude oils were selected from which products possessing desirable characteristics could be distilled—for example, oxidation stability, low smoke tendency, low carbon-forming tendency, small viscosity change with change i n temperature (high viscosity index), light color, and attractive appearance were more likely to be found i n petroleum of the "paraffinic" or Pennsylvania type. A s Pennsylvania type crude oils became limited i n quantity and commanded premium prices, the desirability of upgrading stocks from other crudes became evident. One of the first chemicals used for this purpose was sulfuric acid, and even today sulfuric acid is used extensively to refine many petroleum products. A whole lubricating oil fraction consists of four major classes of hydrocarbons—namely (a) asphalts and resins, (6) aromatics, (c) naphthenes and branched paraffins, and (d) paraffin wax. Sulfuric acid is remarkably effective for removing undesirable constituents a and b b y a combination of reaction and extraction and has little or no effect on wax (which must be removed by other means) or the naphthenic-type materials which comprise a good lubricating oil. I n contrast to the action of sulfuric acid, the solvents utilized i n commercial solvent 177

In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.

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ADVANCES IN CHEMISTRY SERIES


extraction processes function b y physical means only. Undesirable constituents are removed by selective solvent action, and the solvent is recovered b y distillation for re-use. Asphaltic constituents behave like colloids instead of solutes; hence, as is described i n another paper in this symposium (Î3), precipitation is used or depeptization by lower molecular weight hydrocarbons such as propane or butane for deasphalting. Aromatics are extracted by the selective solvent. The action, then, of the selective solvent is to leave the desirable naphthenic-type compounds along with the wax i n the so-called raffinate layer. Dewaxing processes (13) are utilized to separate these two materials.

Figure 1.

First Edeleanu Continuous Extraction Towers

E a r l y i n the twentieth century, Edeleanu ascertained that the poor burning qualities and high smoke tendency of kerosene distillates from Rumanian crude oils resulted from the presence of large quantities of aromatic materials. L i q u i d sulfur dioxide was found to possess a high solvent power for aromatics but dissolved relatively little of the saturated, naphthenic-type oils. I n a paper (9) presented before the International Petroleum Congress i n 1907, Edeleanu described the basic principles which later resulted i n commercial development of the sulfur dioxide solvent extraction process which now bears his name. I n 1909 at Vega, Rumania, the process was operated successfully on a small scale and first commercial production took place at Rouen, France, i n 1911. T h e first commercial continuous solvent extraction was accomplished at the Standard O i l C o . (California) refinery at Richmond, Calif., i n 1924. This plant, prefabricated i n Germany, had been constructed the previous year and was designed to operate batch wise. T h e ingenuity of Crawford and his coworkers resulted i n a number of minor modifications which permitted operation to become continuous with an immediate threefold increase i n daily production (S). This plant, still i n operation at the Richmond Refinery, is illustrated i n Figures 1 and 2. ?



GESTER—SOLVENT EXTRACTION IN THE PETROLEUM INDUSTRY

Figure 2.

179

Sulfur Dioxide Compressors Imported from Germany in Early 1920's

Solvent extraction was limited to treatment of kerosenes with sulfur dioxide until early i n the 1930's. Prior to 1930 less than 2000 barrels of solvent-refined lubricating oil were produced daily. Sulfur dioxide below its atmospheric boiling point has only a slight solvent power for aromatic and unsaturated constituents of lubricating oil fractions. E a r l y attempts to operate at elevated temperatures and pressures gave rise to costly and annoying sulfur dioxide leakage and packing and sealing problems. I n 1933 Cottrell (4) described the use of benzene added to sulfur dioxide to increase its solvent power. The lack of satisfactory solvent recovery methods prior to 1930 prevented the use of selective solvents more suitable for lubricating oils. The major part of any solvent extraction plant is its complex solvent recovery system. Chemical engineering's contributions to distillation theory and process design resulted i n the development of efficient solvent recovery techniques. I n 1933, as illustrated b y Figure 3, large commercial plants were built employing as selective solvents furfural, nitrobenzene, Selecto (cresylic acid plus

Figure 3.

YEAR Growth of Lubricating Oil Solvent Extraction (World-Wide)


180


phenol), propane, Chlorex, and phenol (14). A n interesting combination of simultaneous extraction and reaction results from one refiner's use of nitrobenzene i n conjunction with sulfuric acid. Figure 3 also illustrates the rapid growth of lubricating oil solvent extraction proc esses. F r o m the point of view of popularity, furfural, phenol, and the Duosol processes are outstanding. M o s t of the recent developments have been improvements i n the sol vent recovery systems to minimize solvent consumption, to reduce costs, and i n the treater sections, to increase raffinate yield.


Theoretical Aspects Although theoretical considerations are helpful to an understanding of the principles involved and may be useful for studying and predicting simple extractions of pure sub stances, an empirical approach ultimately must be resorted to for cases involving such complex and undefinable mixtures as kerosenes and lubricating oils. The ideal distribu tion law which states that the ratio of concentrations of a component distributed between two mutually insoluble phases is a constant dependent only on the temperature (K = C i / C ) , is analogous to Henry's law for absorption and is rarely valid for commercial extraction problems. 2

SOLVENT

0.82 0.04 O.M O N

α90 0.92 0Λ4 0.9· 0.9·

VISCOSITY GRAVITY CONSTANT

Figure 4. Equilibrium Data for Light Lubri cating Oil Distillate from California WaxFree Crude in Aqueous Phenol Trilinear diagrams may be used for presentation of equilibrium data, for graphical calculation of solvent extraction problems, for evaluating efficiency of operations of existing equipment, and for comparison of various solvents. I f three pure components are involved, concentrations i n weight percentages (because of their additive nature) are used for coordinates to express behavior. I t is impractical to designate raffinate and extract concentrations for complex hydrocarbon mixtures by means of percentage values, but instead i t is desirable to utilize some nearly additive function such as specific gravity, refractive index, or viscosity-gravity-constant. Figure 4 presents such isothermal data for a system comprising low cold test California lubricating oil distillate and aqueous phenol. Three tie lines connecting compositions of phases i n equilibrium are shown. Figure 5, which represents commercial extraction of a California wax-free light lubricating oil distillate, is used to review graphical techniques for studying countercurrent multistage extractions. Point F represents the composition of the given feed stock. I f we assign as desired end conditions a raffinate and extract viscosity-gravity-constant ofD.870 and 0.955, re spectively (such values would represent a yield of 58.5% by weight), we can locate points R and Eu A n y addition of solvent to the feed stock must effect a composition which would fall on the line FA. Likewise, any mixture of final extract and raffinate layers must lie on the line r e i . A n over-all materials balance requires that the solvent-feed mixture n

n


181


be represented by the point of intersection, J. The solvent for this particular case would be 53.5% of the mixture (a dosage of 115 weight % or 100 volume % ) . A l l lines constructed to connect compositions of materials entering any stage with compositions leaving any stage intersect at a common point, 0 , called the operating point. The significance of this purely geometrical point has been discussed b y Skogen and Rogers {25). H a v i n g located 0 and assuming that tie line data are available, one may graphically determine the number of stages required for the desired separation. A tie line through 6i will intersect the binodal curve at n. Connecting η with 0 locates e , etc. T w o and a fraction theoretical stages would be required to effect the specified extraction.


2


Figure 5. Commercial Extraction of Light Lubricating Oil Distillate from California Wax-Free Crude in Aqueous Phenol D a t a could be obtained from such a graph to construct a McCabe-Thiele type dia gram. T i e line points on the binodal curve would provide data for an equilibrium curve, and data could be obtained to construct an operating line (from lines drawn from Ο through the binodal curve). The advantages that may justify the additional costs of reflux at the bottom of a solvent extraction tower are illustrated i n Figure 6. M i n i m u m reflux is represented b y the tie line from / to f . M a x i m u m reflux would be represented by the line / to / ' , the extract layer which would exist for infinite solvent to feed ratio. Practical operation of the equipment will fall between the limits of minimum and maximum reflux, as repre sented by fy. The operating point for the enriching section of the extraction column is located b y an intersection between the lines e'E and/2/, and the reflux ratio is the ratio of the distances k'e'/k'e (22). e

If, from a specified lube distillate we are required to make a given raffina te, R, and extract, Ε, and are given the reflux ratio for the enriching section of the column, operating points k' and 0 for each section may be constructed on the diagram. A line constructed In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.

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from / through k will intersect the Une from the raffinate composition through the pure solvent at 0 . T h e intersection of the extract portion of the curve with this line at point y represents the composition of the extract solution flowing below the feed point. The number of theoretical stages then required can be stepped off utilizing operating point k' and tie lines for the enriching section and operating point 0 and appropriate tie lines for the upper section of the tower. r


If, i n the foregoing example, there had been no reflux and the same solvent dosage had been applied, the final extract composition would have been limited to that shown at point /«. If, on the other hand, we had been required to make an extract of composition Εj i t would have been necessary to utilize a lower solvent dosage and a larger number of theoretical stages.


Figure 6. Construction on Trilinear Diagram to Illustrate Effect of Reflux It is also possible to impress reflux on the top section of the tower. This type of reflux would be represented by mixing a fraction of the overhead raffinate layer with i n coming solvent. A temperature gradient may be impressed across the length of a solvent extraction tower. The result of such temperature gradient is to have at each stage i n the tower a different temperature and therefore a different binodal curve. I t is more difficult to handle this situation i n construction on a trilinear diagram. If, for simplification, we assume that each theoretical stage has its own binodal curve, we may then proceed i n a manner similar to that described for the isothermal cases. When utilizing equilibrium line considerations, i t is necessary to move from one equilibrium condition to another and change equilibrium curves as the temperature changes. In the practice of solvent extraction certain general relationships exist which are independent of the solvent and treating temperature. F o r a particular base stock, rela tionships among the following physical properties of the raffinate will be about the same regardless of solvent or temperature : specific gravity, viscosity, viscosity-gravity constant, and viscosity index. Precipitative solvents, such as propane, provide an exception to this rule. Products obtained by the Duosol process, therefore, would not conform to the curves for single solvent processes. I n general, the best solvent will be that which gives the highest yield at lowest cost. Y i e l d relationships for any solvent i n a given process may be determined graphically on the basis of laboratory data as plotted on trilinear diagrams. Costs will be a function of solvent circulation and price, number of actual stages required, solvent recovery require ments, and special equipment to control water content and corrosion. In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


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The characteristics of an ideal solvent for extraction of lubricating oils may be summarized as follows: 1. H i g h selectivity 2. Good solvent power for the undesirable components 3. E a s y solvent recovery (wide difference i n volatility between solvent and oil) 4. L o w solubility for the desirable oil 5. Sufficient density difference to allow rapid separation of raffinate and extract by gravity 6. L o w interfacial tension i n order to permit rapid and complete separation into two liquid phases 7. L o w cost 8. Ready availability 9. Reasonable temperature of application (preferably treating temperatures such that the oils have low viscosities) 10. Adaptability to a wide range of oils 11. L o w toxicity 12. Chemical stability under plant conditions 13. L o w freezing point (to avoid complications i n cold weather) 14. Noncorrosivity to the usual metals of construction

Characteristics of Commercially Accepted Solvents Solvents used i n commercial operations at the present time are furfural, phenol, cresylic acid, Chlorex, nitrobenzene, and sulfur dioxide. The Duosol process utilizes a solvent called Selecto which is a mixture of phenol and cresylic acid. The propane used in the Duosol process precipitates the asphalt (13). The disadvantages of sulfur dioxide are its high vapor pressure and its low solvent power under normal operating conditions. Fortification of sulfur dioxide with benzene to increase solvent power has been successful i n some applications (4). O n the other hand, the high volatility of sulfur dioxide results in easy solvent stripping, and its low viscosity aids rapid settling. Nitrobenzene has the highest solvent power for extract oils and is best suited for Pennsylvania stocks. However, its high boiling point makes recovery difficult and its high solvency necessitates low temperature operation which requires costly refrigeration. Chlorex (/3,0'-dichlorodiethyl ether) also has a high solvent power which makes i t particularly suitable for Pennsylvania-type oils. The hydrogen chloride i n commercial Chlorex (5.5 mg. per 100 ml.) must be removed before use to prevent corrosion of common construction materials. Injection of ammonia into recovery system vapor lines is required to control corrosion. The selectivity and solvent power of phenol can be controlled by varying the water content. Phenol extraction temperatures are sufficiently high (100° to 200° F.) to effect a reduction i n viscosity, to permit imposition of a temperature gradient, and to allow treatment of waxy stocks. This results in savings i n dewaxing operations. Phenol, which has excellent chemical stability, can be recovered b y ordinary vacuum distillation, although steam stripping is often desirable for complete solvent removal. Drawbacks to phenol are its toxicity and relatively high freezing point. Recently, limited availability and increasing costs of phenol have caused concern. Furfural exhibits high selectivity at elevated temperatures (175° to 250° F . ) ; this characteristic results i n the benefits of reduced viscosity and high temperature gradient and permits operation on waxy stocks. A t ambient temperatures, kerosenes and gas oils may be extracted. Water i n furfural has a bad effect on its extraction efficiency; therefore, solvent recovery systems must include dehydration facilities. F r o m its chemical composition aldehyde furfural appears to be unstable. However, when handled i n accordance with standard operating procedures, decomposition and polymerization are negligible. The raw material for furfural is an agricultural waste. It is therefore readily available i n unlimited quantities (largest manufacturer at the present time is the Quaker Oats Co.). The price of furfural has remained surprisingly constant during the period of In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.

In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951. Information supplied by Ε. B. Badger & Sons Co.

S3

Figure 7. Flow Diagram of Edeleanu Treating Process

RAFFINATE

OUT


CHARGE FILTERS

WASTE

SOLVENT-WATER FRACTIONATOR

Ό

CO

η

Ζ

>


185

inflation when other chemical costs rose materially. During early plant operations i t was necessary to remove acidic materials from commercial furfural, but, as currently available, i t is completely satisfactory for immediate use.


Commercial Processes Edeleanu Process (Liquid Sulfur Dioxide). T h e forerunner of a l l commercial solvent extraction processes, the Edeleanu process, has been pre-eminently successful for treating low molecular weight stocks i n the kerosene-spray o i l boiling range. The earliest commercial plants were designed to operate batchwise and consisted of hydrocarbon and solvent coolers, a mixing vessel which could also be used for settling, and evaporators to remove sulfur dioxide from the raflinate and extract phases. Figure 7 illustrates a modern, continuous sulfur dioxide solvent extraction plant used for kerosenes and light lubricating oils. Dehydrated feed stock is pumped through a heat exchanger and distillate precooler to the bottom of the extraction tower, which is packed with about 30 feet of Raschig rings (or similar packing), and may range from 3 to 8 feet i n diameter. Liquid sulfur dioxide is precooled and pumped i n at the top of the tower. T y p i c a l extraction temperatures range from 20° for kerosenes to 50° to 75° F . for lubricating oils. I t often is desirable to hold the liquid interface at the top of the tower because the low viscosity of the extract layer permits higher feed rates than would be possible if the treater were filled predominantly with cold raflinate. Such operation is reversed for the phenol and furfural processes wherein bottom interface levels are maintained to take advantage of the extended surface area presented b y the packing which is preferentially wetted b y the selective solvent. The raflinate and extract layers pass through heat exchange equipment to their respective recovery systems. Modern improvements incorporating the use of multiple stage, high pressure evaporation for sulfur dioxide recovery, and the latest designs i n heat exchanger equipment have greatly improved the economics of the process. Vaporized sulfur dioxide from the recovery system is passed through the coolers and condensers (to separate any light hydrocarbons) and through drying towers (countercurrent to sulfuric acid) prior to compression. I t is essential that all water be removed from the system. Extreme corrosion difficulties result wherever water leaks into the system and contacts sulfur dioxide. A maximum water content of 0.05 weight % can be tolerated. A n interesting modification of the Edeleanu process was the sulfur dioxide-benzene process utilized by the Union Oil Co. at Oleum, Calif. (6). I n order to overcome the low solvent power of sulfur dioxide, benzene was added to the solvent. The immediate result of such addition is to reduce the selectivity of the process. However, selectivity can be improved b y reducing temperatures. Therefore, a combination of reduced temperature and increased quantity of solvent provides a reasonably wide range of operating conditions i n which satisfactory selectivity can be obtained. Experience at Standard of California's refineries has shown that selected lubricating oil stocks from California nonwaxy crude oils can be refined satisfactorily with sulfur dioxide without the addition of benzene. Lubricating oils above the range of S A E Grade 10 to 20 require increased temperatures and pressures. Nitrobenzene Process. C o m m e r c i a l solvent extraction utilizing nitrobenzene was developed b y the A t l a n t i c Refining C o . , and the following description pertains particularly to the plant operated b y that company (7). O i l feed is mixed with a quant i t y of nitrobenzene layer from N o . 2 extractor (to reduce viscosity), is brought to extraction temperature (30° to 100° F . ) , and is charged to the first of a series of five mixing and settling tanks. Nitrobenzene (100 to 300% of feed) is fed to the fifth tank. Heat exchangers on the feed stream are of the scraped type to permit operation at reduced temperatures on wax-bearing distillates and residual stocks. Raffinate and extract layers pass from N o . 5 and N o . 1 extractors to their respective recovery units. Recovery takes place under reduced pressure at temperatures of 215° to 350° F . Stripping steam is removed from the nitrobenzene b y distillation. The water content of nitrobenzene has little effect on its usefulness for solvent extraction; therefore, it is not necessary to resort to thorough dehydration methods for the solvent. A unique characteristic of nitrobenzene is that i t operates satisfactorily on residual In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


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oils. E v e n though the solvent extraction system operates at a temperature below the solidification of the wax, the presence of solid wax crystals seems to have no deleterious effect on its action. A modification of the nitrobenzene process, which utilizes nitrobenzene and sulfuric acid simultaneously, is i n operation at the Sinclair Refining C o . refinery at AVellsville. Ν. Y .

STOCK

O I L " "

Figure 8.

O

n

.

Flow Diagram of Anhydrous Phenol Extraction Plant

Information supplied by The M. W. Kellogg Co.

Chlorex Process (0,j3'-Dichloroethyl Ether). T h e Chlorex process developed b y Standard O i l (Indiana) (8) was patented i n 1934. T h efirstcommercial plant to be put into operation was at Casper, W y o . , i n June 1932. B y August 1936 there were seven commercial plants i n operation with a total throughput of 6150 barrels daily. Because of its high solvent power, Chlorex extraction takes place at 75° to 150° F . Such ambient temperature operation effects heating and cooling economies but necessi tates dewaxing prior to extraction or the use of scraped exchangers. The Chlorex process utilizes simple countercurrent mixing and settling tanks (four to seven stages) or modern vertical packed towers. Solvent recovery involves conven tional flash columns and strippers operated under vacuum (26 to 28 inches of mercury) at about 300° to 325° F . L o w temperatures are desirable to minimize decomposition and formation of hydrochloric acid. Phenol Processes—Anhydrous and Aqueous. T h e original patent describing utilization of phenol for solvent extraction is dated July 1908 (2). This was followed i n 1922 b y Polish patents (20) and i n 1926 by a British (11) and a German (24) patent. These processes describe the use of phenol, either aqueous or mixed, with a variety of In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.



Figure 9.

187

Flow Diagram of Aqueous Phenol Extraction Plant

materials. I n 1928 Stratford and coworkers with Imperial O i l , L t d . (a subsidiary of Standard of N e w Jersey), found that the phenol process, as disclosed i n earlier patents, gave poor results on the stocks under consideration and i n 1929 Stratford applied for a patent involving anhydrous phenol. This patent was granted December 10, 1931 {26). The first plant was put into operation b y Imperial i n Sarnia, Ont. A flow diagram of an anhydrous phenol solvent extraction plant is shown i n Figure 8. R a w distillate is passed through a tower i n which i t absorbs phenol from the recovery system vapor. The oil is then passed to the treating tower, generally a few sections above the bottom. Anhydrous phenol is introduced at the top of this tower. Phenolic water condensate from the solvent recovery system (about 9 . 5 % phenol) is introduced at the bottom of the tower to effect reflux. A temperature gradient of 10° to 75° F . may be

Figure 10. Aqueous Phenol Solvent Extraction Plant In PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


RAFFINATE

EXTRACT

Figure

WET FURFURAL VAPOR CONDENSER-

Flow Diagram of Furfural Extraction Plant

Information supplied by Texaco Development Co.

11.


REFINED OIL MIX HEATER

CO

η ζ m

I >

00

oo



189

impressed on the tower; this also effects internal reflux. The raffinate layer leaving the top of the solvent extraction tower contains from 10 to 2 0 % phenol. Raffinate and extract layers are freed of solvent i n their respective solvent recovery systems which involve exchangers, fired heaters, flash columns, and vacuum steam strippers as indicated. Extraction tower design involves a combination of vertical and cross flow which ensures intimate contact between the two phases without channeling. Whereas Stratford and his coworkers utilized anhydrous phenol to effect satisfactory solvent treatment of the oils under examination, aqueous phenol (up to 1 2 % water) is necessary for phase separation at 100° to 175° F . when treating California-type low cold test lubricating oil distillates. Such a plant designed and installed b y the Standard Oil Co. of California at its Richmond Refinery i n 1934 is represented by Figures 9 and 10. Distillate feed stock is pumped through a heat exchanger to the bottom of the packed solvent extraction treater; phenol containing a controlled amount of water is introduced into the top of the treater. The raffinate layer passes from the top of the tower through Appropriate heat exchange into a raffinate recovery tower where, under high vacuum and with the bottoms circulating through a fired heater, the raffinate layer is stripped of phenol. The extract layer solvent recovery system consists of a vacuum flash tower, atmospheric pressure flash tower, and d r y vacuum stripper. Use of these two high vacuum dry stripping columns provides constant water content i n the solvent during operations on one stock. Furfural Process. A patent i n v o l v i n g furfural for solvent extraction of l u b r i c a t i n g oils was issued to Eichwald of the R o y a l D u t c h Shell Co. i n 1925 (10). The first commercial application of the furfural process was at the Lawrenceville, 111., refinery of the Indian Refining C o . i n December 1933. Modern furfural solvent extraction plants utilize vertical counterflow towers packed with about 20 feet of Raschig rings with redistribution equipment at about 4-foot intervals. Charge rates average about 35 gallons of oil per hour per square foot of over-all tower cross-sectional area. However, rates as high as 70 gallons per square foot per hour have been observed. A flow diagram of a typical furfural solvent extraction plant for lubricati n g oil is illustrated i n Figure 11. R a w feed at a temperature of 110° to 220° F . , depending on the nature of the oil, is introduced at the center of the extraction tower. Furfural is fed into the top of the tower at about 200° to 290° F . Recycled extract is introduced into the lower section of the tower for reflux. Likewise, internal reflux is effected by the temperature gradient which is brought about b y introducing the solvent at an elevated temperature and b y intermediate cooling systems. Furfural is removed from raflinate a n d extract layers i n conventional distillation and stripping equipment. The water content of furfural is critical and must be kept to a minimum. I n order to dehydrate the solvent a pair of distillation columns is utilized. Condensed wet solvent separates into two layers i n a settling drum. The top layer, which is lean i n furfural, is separated into saturated solvent vapor and solvent-free water i n the furfural-from-water stripper. T h e solvent-rich layer is separated i n the water-from-furfural stripper into wet solvent vapor and dry furfural. Azeotropic distillation is used to recover furfural from light gas oils, kerosenes, and Diesel fuels. Duosol Process. T h e D u o s o l process developed b y the M a x B . M i l l e r C o . (28) is an outstanding example of commercial adoption of a double solvent extraction process. Patents (27) for this process date from M a y 1933 and cover numerous aspects of the problem including a variety of paraffinic solvents (ethane, propane, butane, petroleum ether) and "naphthenic" solvents (wood tar acids, cresols, creosote, and phenol). Present commercial application utilizes propane and Selecto (a mixture of phenol and •cresylic acid, normally ranging i n composition from 20 to 8 0 % phenol). Propane acts as a raffinate solvent and as a deasphalting agent. F o r normal operating temperatures of from 60° to 150° F . , wax is held i n solution i n the propane so that the advantage of solvent refining prior to dewaxing can be utilized. Distillation into cuts can be delayed until after refining and dewaxing. I n single solvent extraction processes, i t is generally preferable to solvent treat disIn PROGRESS IN PETROLEUM TECHNOLOGY; Advances in Chemistry; American Chemical Society: Washington, DC, 1951.


188 ADVANCES IN CHEMISTRY SERIES



GESTE*—SOLVENT EXTRACTION IN THE PETROLEUM INDUSTRY

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tillate cuts and a deasphalted residuum to recover all lubricating oil fractions from a given crude oil. Duosol operation permits treatment of the whole range of lubricating oil components i n a long residual stock of, say, 450° F . flash point. Distillate fractions may be treated separately i f desired. The practical range of operating temperatures is limited on the high side (about 150° F.) by the high vapor pressure of propane and on the low side b y the tendency for emulsion formation as wax precipitates or possibly as phenol freezes as i n high phenol content Selecto. Operation between these limits provides no difficulty because the solvent power of Selecto can be altered b y changing the ratio of phenol to cresylic acid. Solvent dosages i n the Duosol process are usually high, up to 400%. Experience has indicated that b y carefully coordinating the effects of treating temperature and Selecto composition, a point of optimum operation can be attained that will permit minimum solvent dosages and maximum throughput. The effects of operating variables on the Duosol extraction system are complex and are best determined empirically. A s would be expected, an increase i n Selecto results i n increased refinement and reduced yield. A n increase i n propane results i n improved raffinate color and decreased viscosity index, but its effect on carbon residue is unpredictable. Figure 12 presents the flow diagram of a modern Duosol solvent refining plant. Figure 13 shows the extraction system. Charge o i l enters the process through a mixer into the third of seven countercurrent extraction stages. Selecto is introduced into the N o . 7 extractor and pure propane into the N o . 1 extractor. The raffinate layer passes through the upper sections of the system, rising to the top i n each vessel and carrying with i t the refined oil. The extract layer which settles to the bottom is pumped countercurrently to the raffinate layer. Extractors 1 and 2 effect a form of reflux on the system. Additional reflux within the system is imposed b y a temperature gradient. Raffinate and extract layers pass, respectively, from separator N o . 7 and separator N o . 1 to their solvent recovery systems. The two recovery systems are similar: propane is recovered at elevated pressure (about 250 pounds per square inch absolute) i n the first column, and the bottoms pass to the second column (85 pounds per square inch absolute), from which Selecto is passed overhead to a drying tower. T w o stages of stripping are employed (85 pounds per square inch absolute and vacuum) to remove final traces of solvent from the raffinate and extract streams. Vapor from the top of the dryer (a constant boiling mixture) is separated into two layers, and the water layer is used to generate stripping steam in a closed system.

Process Selection Process selection for new solvent extraction facilities involves numerous considerations; some may be evaluated numerically and others can only be evaluated on the basis of judgment, experience, and personal taste. Approximate numerical evaluation may be computed for installation, operating and royalty costs, yields, and product quality. However, the important considerations of design and operating experience, availability of construction materials, solvent availability, toxicity, stability, and price trends, and process flexibility are not so tangible. E v e n more intangible, but likely to influence the refiner, are such considerations as the popular acceptance of various processes and the relation between the prospective builder and the licensor of a process. A s illustrated by Figure 3, the three processes which are used to refine the bulk of the nation's solventtreated lubricating oil are furfural, phenol, and Duosol. Consideration of installation costs cannot be restricted to the solvent refining step. A residuum of 460° F . flash point usually will contain a l l usable lubricating o i l . I n addition to the cost of solvent extraction facilities, the expense of equipment to fractionate the oil into usable cuts, to deasphalt, to dewax, and to clay filter, all must be •

Solvent Extraction in the Petroleum Industry

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