c
FIGURE 1.
0
PREDISTILLAPION AUD
RECOVERY UUITS
GASOLINE
Applications of Distillation in Modern Petroleum Refining
ISTILLATION is one of our major unit operations and is so ranked because of its wide application in the entire field of chemical engineering. For sheer magnitude in the total quantity of raw material processed and in the size and capacity of individual process units, the largest scale applications of distillation are found in the refining of petroleum. The distillation operations employed in the general chemical industries usually have as their aim the isolation of fairly pure individual compounds from mixtures consisting of several well-defined components. I n petroleum processing, all degrees of fractionation, as dictated by current economic considerations, are to be found, varying from the relatively rough segregation of raw distillates to the isolation of pure compounds from mixtures of the lower boiling hydrocarbons.
A general outline is given covering types of units utilized in the major bulk-distillation operations. In each of the process units described, distillation with fractionation is the primary objective. While distillation and rectification principles have long been established and applied, their recognition in petroleum refining has been relatively recent. All of the operations described have been widely adopted and developed into their present form within the past ten years.
JAMES S. CAREY Alco Products Incorporated, New York, N. Y.'
The purpose of this paper is to outline briefly the general features of process units illustrative of present practice in the major refinery bulk-distillation operations. Reference will be made either to actual operating units or to units typical of present design and practice. Many variations in detail are possible in the various processing schemes, but an attempt has been made to use illustrations on the whole representative of current procedure. For convenience in presentation, the various major distillation operations will be outlined as nearly as possible in the order of the flow of crude oil from the field and through the refinery.
Stabilization of Crude Oil Starting with crude, either a t its source or in the refinery, the first distillation operation which is being performed in some instances is stabilization. Stabilization of crude usually consists in the substantially complete removal of all light hydrocarbons through butanes, together with portions of the pentanes and hexanes. The major results achieved by stabilization of crude are as follows: 1. There is a decreased loss from crude in storage, and a minimum load to the refinery vapor recovery system from crude storage. 2. Hydrogen sulfide is eliminated with consequent diminished corrosion to tankage and primary distillation equipment.
-1
Present address, The M. W. Kellogg Company, New York, N. Y.
1XDUSTRI-41. .4XD EN GlrUEERIiXG CHEAIISTRL'
796
3. A stable straight-run gasoline is recovered from the primary distillation units reducing treating losses for straight-run gasoline. 4. Vapor losses from the straight-run gasoline condensers and storage tanhs are greatly reduced. 5 . A settling operation under temperature and pressure can be conveniently combined with the crude stabilization, thus permitting "bottom sediment" and water to beremoved at this time.
Stabilization of crude can be carried out in the field, the refinery, or a t some intermediate bulk storage point. Crude stabilization in the field would appear to have certain economic advantages, especially from the point of view of conservation. Field stabilization of crude can be combined with the field gas recovery system, resulting in three major rav materials: 1. Stabilized crude. 2. Stabilized natural gasoline, as recovered from the natural gas and the crude. 3. Dry gas. The chief obstacle to field stabilization is that the separated natural gasoline, as well as the stabilized crude, must bear higher transportation rates to the refinery than the original crude. In some instances, however, import and transportation regulations are such that natural gasoline as recovered in the field is blended with the crude before transportation to the refinery. A crude stabilization operation a t the refinery then becomes both a technical and economic necessity. As an illustration of crude stabilization, we will consider a plant designed to stabilize a crude blend containing approximately 5 per cent by volume of unstabilized natural gasoline The analyses of the blended crude, the original raw crude, and the natural gasoline in mole per cent are as f o l l o ~ ~ s : Constituent Methane Ethane Propane Butanes Pentanes Hexanes Heptanes Octanes plus
Blended Crude
Raw Crude
0.05
. . I n
0.43 1.42
0.3 0.6 1.8
6.48 5.21 4.44 4.61
2,5
3,0
3.8 88.0
78.36
Unstable Gasoline 0,s
1.5 8.0 38.0
27.0 16.0
11.0 1.0
From the crude stabilization proper, stabilized crude and unstable light gasoline of the approximate compositions shown by the following table Till be recovered (in mole per cent) : Constituent Methane Ethane Propane Butanes Pentanes Hexanes ulus
Stabilized Crude
Unstable Light Gasoline
...I
0.49 4.03 13,30 51.34 19. s5 10.99
.... ~.._ ~...
3.47 96.53
__
100 00
100 00
In addition, the complete plant ia designed for the further separation of the lighter fractions into : ethane and lighter, commercial propane, commercial butane, and stabilized light gasoline. This design of plant is illustrated by Figure 1, The blended crude is taken from storage and preheated by exchange against the outgoing stable crude. The preheated crude is further heated to a temperature of approximately 300' F., a t which temperature it will be introduced into the pressure settler. From the pressure settler the charge will be introduced into the fractionating tower operating a t a top pressure of 75 pounds per square inch gage. Stripping of the crude bottoms will be aided by open &earn injected a t the base of the tower. The overhead vapors from the tower will be condensed and cooled, and will flow to a distillate receiver and water separator. Reflux ill be pumped to the torvcr top while the net yield of condensate will pass through a caustic soda washing bystem for the removal of hydrogen sulfide and thence
VOL. 2 7 , NO. 7
t o the additional towers for recovery of the sereral fractioirated products. General process conditions and size. of equipment are gken in Table I.
Charge bb1./24 hr. Gravit; of charge '4.P. I. Stable crude bottArns, bh1./34 hr, Gravity of bottoms 'A. E'. I. Proceas steam, lh./ir, Tube still duty, B. t. u,/hr. Tube-still heating surface, sq. f t , Tower diam. (crude stabilizer!, ft. No. plates In tower (18-in. spacing) (6 plater above inlet, 10 plates below inlet) moles reflux ratio ( x o x d ) Temp., F.: Tower top Tower bottom Tower inlet Pressure, lb./sq. in. gage: Tower top Settler
11,100 38 10,610
35 4
2,000
5,~00.000 1,050 5
I6 1.5/1 240
289
300
75
200
Primary Distillation of Crudes The next distillation operation in the complete refinery proc-, essing scheme is the initial separation of the crude into the major raw fractions, From the distillation. standpoint, crudes can be divided into two general groups: 1. Those containing appreciable quantities of asphalt: A , Mixed-base crudes (Midcontinent).
B. Asphalt-base crudes (Gulf Coast and California).
2 . Crudes containing relatively small quantities of asphaltic
material as represent'ed by Pennsylvania-grade crudes.
If the complete range of primary raw products is required, crudes of group 1 are processed by recovering all products down to asphalt as distillates, This requires the use of vacuum for the vaporization of the lubricating stocks a t teniperatures low enough to avoid appreciable thermal decornposition. The usual procedure is to recover the lighter distillates through kerosene, together with a portion of the gas oil, by distillation a t at'mospheric pressure. The residue from the atmospheric operation is then transferred to a vacuum stage in which the balance of the gas oil and several lubricating-stock fractions are recovered as distillates, and asphalt, as a final residuum. Figure 2 shows the flow diagram of an integrated two-stage atmospheric and vacuum-crude distillation unit which processes 10,000 barrels of crude per 24 hours. Figure 3 is a general photographic view of this installation. The general operation of this unit is as follows: Crude is taken from storage and pumped through the atmospheric vapor heat exchanger and thence to a settling tank for remwal of solids and water. The crude then flows through the vacuum-stage exchangers, acquiring a final preheat t'emperature of approxiniately 350" P. The preheated crude is Oransferred to the atmospheric stage tube still and heat,ed to a temperature sufficient to effect the desired vaporixation in the atmospheric fractionating tover. I n the atmospheric tower four distillate streams are recovered: gasoline as the top vapor st'ream, and naphtha, kerosene, and gas oil. as liquid side streams. The side streams are removed from respective plates in the main tower to external stripping sections for steam stripping to remove the lighter ends, The stripping steam and vapors are returned to the main tower a t points above the respective stream take-offs, The atmospheric stage bott'oma is subjected t o steam stripping on plates below the t,ower inlet. The reduced crude is pumped hot from the base of the atmospheric tmver directly to the vacuum-stage tube still for further heating.
JULY, 1935
INDUSTRIAL AND ENGINEERING CHEMISTRY
197
FIGURE 2. TWO-STAGE ATMOSPHERICAND VACUUX DISTILLATION UNIT
The vacuum tube still discharges into the vacuum fractionating tower, which is maintained at a low absolute pressure. Three distillate streams are recovered from the vacuum stage; heavy gas oil as the top vapor stream and para& distillate and cylinder stock as liquid side streams. Asphalt bottoms are recovered as a residue from the base of the tower. Steam is introduced into the base of the tower for stripping the asphalt and for partial-pressure effect. Vacuum is maintained on the system by a barometric condenser backed by steam jet ejectors. The overhead vapor stream, consisting of gas oil and process steam, passes through vapor heat exchangers and water-cooled surface condensers. Because of the temperature and pressure conditions maintained a t the condenser outlet, the process steam remains in the vapor state and flows to the barometric condenser, while practically complete condensation of the gas oil is obtained. The desirability of obtaining sharp phase separation a t the vacuum condenser outlet with a low partial pressure of gas ail vapor in the uncondensed steam, dictates the quantity of gas oil to be removed in the atmospheric stage; i. e., theinitial boiling point of the vacuum-stage gas oil must be sufficiently high t o insure substantially complete condensation. Table I1 shows representative A. S. T. M. distillations of the lighter products recovered in this operation.
cate the enormous capacity which can be built into single process units of this type. The two-stage units which have been described effect a separation between asphalt and lube stocks by distillation. Because of overlapping in boiling point range between the heaviest lube fraction and the lower boiling asphaltic material, TABLE111. TWO-STAQE DISTILLATION UNIT Crude charge, bb1./24 hr. Vacuum-stage charge bb1./24 hr Final asphalt residuuh, bb1./24 hr Temp., F.: Atm. tower inlet Atm. tower top Atm. naphtha. draw-ofi Atm. kerosene draw-off Atm. gas oil draw-off Atm. bottoms draw-off Vacuum tower inlet Vacuum tower top Vacuum tower light lube draw-off (wax dist.) Vacuum tower heavy lube draw-off (cylinder stock) Vacuum tower bottoms draw-off Abs. pressure at top of vacuum tower, mm. Atm. tower: Diam. ft. No. p k e s (30 above inlet, 4 below) Vacuum tower Diam it. No prates 16 above inlet, 4 below) Process 'steam IL./hr.: Atm. towLr and strippers Vacuum tower Vapor velocity: Atm. tower: Ft./sec. Lb./ss. ft./hr. Vacuum tower: Ft./sec. Lb./sq. ft./hr.
TABLE 11. A. s. T. M. DISTILLATIO~ O F LIGHTER PRODUCTS RECOVERED (IN F.) Date Nov. 9 Nov. 10 Nov. 11 Nov. 12 Nov. 13 Nov. 14 Nov. 15 Nov. 16 Nov. 17 Nov. 18
Gasoline --Naphtha-End Initial E n d Point b. p. point 320 327 387 327 333 405 309 325 396 315 327 392 307 324 390 316 333 399 316 322 387 316 331 390 315 327 394 309 327 396
Gas Oil Btm. and -KeroseneVacuum Composite Initial E?d Initial b. p. point b. p. 34% 414 581 536 572 417 597 . ... 406 588 .. ... 421 588 , .. .., 408 576 .. .,, 419 592 ... .. , 406 579 .., .,, 403 583 . .. .,, 403 574 ,. . ,,, 405 576 .,
... . . ..
.
Process conditions and sizes of the major items of equipment for this installation are shown in Table 111. Two-stage crude distillation units with individual unit capacities up to 50,000 barrels per day are now under construction or are being actively considered. Based on 330 stream days per year, 50,000 barrels per day correspond t o 16,500,000 barrels of crude oil per year. These figures indi-
I
10,000 5,000 837
599 224 328 420 537 530 772 364 504 690 709 60 11 34
17 20 3600 4500 2.21 1360 7.05 600 l
a sharp separation is difficult. Hence it has been found advisable in some instances t o remove a clean, heavy, lubricating stock several trays above the flash zone and a small quantity of material containing the lighter asphaltic materials as a total draw-off from the tray immediately above the tower inlet. This stream can be treated separately from the clean bulk heavy lube stock, or if the heavy lube available in the crude exceeds requirements, it can be routed t o cracking. Considerable progress has been made in the separation of asphalt from lube stocks by solvent refining rather than by distillation. Developments in this field may modify future primary distillation procedure. The primary distillation of crudes which are relatively free from asphalt differs from the two-stage process just described, in that the heaviest lube fraction need not be re-
198
INDUSTRIAL AND ENGINEERING CHEMISTRY
VOL. 27, NO.
a
Distillation in Conjunction w i t h Cracking Cracking is l o g i c a l l y c o n s i d e r e d together wit'h primary distillation in import'ance and position in the general refinery scheme. Economic considerations have d i c t a t e d that' increasing quantities of c r u d e be complet'ely cracked to yield three p r o d u c k gasoline, fuel oil (or coke), and gas. The introduction of fractionation into cracking has played an important role in the development of modern cracking units. Cracking unit's are now universally equipped for the recovery of a raw motor fuel dist'illate controllable as to end point. I n addition, the fractionating equipment must s e g r e g a t e fresh feed and cycle stocks with respect to boiling point ranges to furnish the charging stocks for two or more separately controlled cracking coils. The design of fractionating equipment for cracking units is largely dicATMOSPHERIC AND VACUUM FIGURE 3. PHOTOGRAPH O F TWO-STAGE tated by considerations bearing on the DISTILL-ATION UNIT i n t e g r a t i o n of such equipment into specific o v e r - a l l processing schemes. covered as a distillate product. The small asphalt content Generalizations as to fractionation in cracking ope&tions are makes it possible to refine a heavy lube residuum. I n terms difficult without detailed analysis of the respective cracking of the distillation methods employed, this means thats, in processes, which is beyond the scope of this paper, general, the primary distillation of paraffin-base crudes can be accomplished at a considerably lower maximum temperature Rerunning of Treated Motor Fuel Distillates than for crudes containing asphalt. Acid-treated distillates are customarily redistilled after Because of the lower boiling point range of the heaviest distreating for the purpose of producing a finished gasoline with tillate product, the majority of Pennsylvania refiners conduct a minimum gum content and the maximum stability with rethe primary crude distillation in a single-stage atmospheric spect to the development of both gum and color in storage. unit of the type shown in Figure 4. Representative product Sulfuric acid treatment of motor fuel distillate effects dedata for this operation are as follows: sulfurization and the elimination of the most undesirable comInitial End Boiling 90% pounds with respect to gum formation and instability. The Point, Over, Point, major portion of the materials eliminated is removed in the Crude Unit Gravity Color F. F. F. sludge which is separated after contacting the distillate with Overhead 66.5 26 94 292 320 .. 315 358 390 Solvent side stream 52.9 acid. However, appreciable quantities of sulfur compounds . .. 395 ~ j . 539 Kerosene 46.5 from the treating reactions are dissolved in the distillate. 300 dist." 40.7 503 639 .. 600 712 Gas oil 36.4 Fortunately these compounds have boiling points generally Long residuum bottoms, 82 Saybolt Seoonds Universal a t 210' F. above the motor fuel range and can be eliminated from the General process conditions and equipment sizes for the finished gasoline by distillation, yrovided sufficient residual Pennsylvania crude unit in Figure 4 are given in Table IV. liquid remains to hold these materials in solution. For these reasons, raw didillates which are to be acid-treated are "overcut" in end point; for example, if the finished gaaoline is to TABLEIV. PENKSYLYANIA CEUDE DIsTILL.4TION UNIT hare an end point of 400" F., the raw distillate may be cut to Crude charge, bb1./24 hr. 1,500 an end point of 450" F. or higher. The material boiling beLong reeiduum bottoms, bb1./24 hr. 380 Temp., F.: tween 400" and 450" l?. is separated as a bottoms in the rerun Tou-er inlet 685 216 Inlet top operation and serves as a medium in which dissoll-ed sulfur Solvent draw-off 309 compounds are segregated. Kerosene drau--off 399 300 dist." draw-off 494 To accomplish the desired results in rerunning, it is necesGas oil draw-off 575 558 Bottoms draw-off sary to minimize polymerization of unsaturated hydrocarbons 1,200 Proceas steam, Ib./hr. and thermal decomposition of the sulfur compounds. This 5 Tower diam. Et. 48 No. of trays '(40 above inlet, 8 belou inlet) requires low process temperatures and short time under temTube still duty, R . t. u./hr. 5,000,000 perature. N o general criteria have been established ac: to the limiting temperatures to which treated distillates may be heated without detrimental effects. The limiting temperaComparing the maximum process temperature of 686 " F. tures are dependent upon the source of the original crude. in this operation with the highest temperature of the twotype of cracking operation, method of acid treating, etc. stage distillation of i i 2 " F. as shown in Table 111, it will be Temperature limitations imposed by various major refiners noted that the separation of overhead lube from asphalt, vary from 250' to 400" F.or higher. The niaximuni peleven with the use of vacuum, requires appreciably higher procmissible process temperature has a marked bearing upon the ess temperatures than the Pennsylvania operation a t atmospheric pressure with the recovery of the lube stock as a reeconomics and design of the rerun equipment. The rerun operation is thus intimately associated with treating, and siduum. ~
.
~
~
.
I
JULY, 1935
INDUSTRIAL AND ENGINEERING CHEMISTRY
the two are preferably considered together in projecting new designs. Studies of this nature have shown in some instances that a complete revision of the acid-treating procedure, combined with a rerun operation under proper conditions, will pay out more rapidly than a new rerun unit alone. Based on the recovery of 400 " F. end point finished gasoline overhead distillate, atmospheric distillation may be used for limiting process temperatures of 325" F. and above by utilizing process steam for partial-pressure effect. If the limiting process temperature is considered to be between 300" and 325" F., the choice between an atmospheric unit employing large quantities of open steam and a two-stage unit is dictated by specific conditions in the particular refinery-for example, availability of low-pressure steam for process and its valuation, valuation of high-pressure steam, and cooling water valuation. I n general, for limiting process temperatures below 300" F., two-stage rerun units are dictated by both economic and technical considerations. For a limiting temperature as low as 250" F., three-stage units have been proposed and one or more installations have been made. The foregoing approximate temperature ranges dictating choice of design are modified, of course, if the finished distillate is to have an end point lower than 400" F. The flow diagram of a typical two-stage rerun unit is shown in Figure 5. I n operation the charge is pumped through the atmospheric-stage vapor heat exchanger and thence t o a steam-heated reheater for partial vaporization. The entire charge is then transferred to the atmospheric tower. A steam-heated reboiler in the base of the tower supplies sufficient vaporization to strip the bottoms to an initial boiling point which will give satisfactory condensation in the vacuum stage. Process steam is admitted a t the base of the tower as required to maintain the process temperature within the imposed limitations. Reflux is supplied to the tower top to control the overhead end point. The atmospheric tower bottoms is transferred directly to the vacuum tower. The vacuum tower is designed for dry distillation without the use of process steam, and reboil effect is supplied by a steam-heated reboiler. The overhead is condensed in water-cooled condensers and the condensate collected in a vacuum run-down tank from which a portion of the total condensate is used as tower reflux and the net yield is pumped to storage for blending with the atmospheric overhead.
799
FIGURE4. PENNSYLVANIA CRUDEDISTILLATIOX UNIT
Vacuum is maintained on the system by a set of two-stage steam jet ejectors. With reference to general design considerations for a unit of this type several points are of interest. The function of the atmospheric stage is to produce a bottoms product having a n initial boiling point sufficiently high so that appreciable loss to the vacuum jets of uncondensed light ends will not occur when this material is charged to the vacuum stage. At the same time, materials having boiling Doints above the requyreb. end point of the finished gasoline must be kept out of the atmospheric o v e r head. A high degree of separation between atmospheric overhead and bottoms is not required, h o w e v e r . These considerations d i e t a t e a minimum reflux to the atmospheric stage with o n l y sufficient r e b o i l effect to secure the desired stripping. Reflux ratios of 0.2 to 1 or lower are employed on atmospheric stages, and typical fractionation between atmospheric overhead and b o t t o m s is represented by an overhead end point of 375" F. and a bottoms initial boiling point of FIGURE 5. TWOSTAGE PRESSURE DISTILLATE RERW UNIT 300" F.
INDUSTRIAL AND EKGINEERING CHEMISTRY
800
TABLEv.
TRESTED
DISTILLATE RERUNU K I T
Charge. bb1./24 hr. Total overhead diet., bb1./24 hr. (95%) Characteristics of charge: Gravity O A. P. I. A. S. T: '>f. distn., O E'.: Initial b. p. 10%
32
40%
50% 60% 70% 80%
E2
point Residue, Loss, % Limiting process temp., F. Process steam, atm. stage, Ib./hr. Abs. pressure at top vacuum tower, mm Overhead dist., atm. stage, yo of charge Atm. tower: Diam., f t . h-o. plates Vacuum tower: Diam., Et. No. plates
TABLE
VI.
DEBUTdNIZlNG,
54
360
415
2 0
1.5 260 3,000 25 60
Quantity (normal) per 24 hr. Pressure, Ib./sq. in. Temp. O F. Sp. gr. Compn., mole % : Methane Ethane Propane Isobutane n-Butane Pentane Hexanes plus
10 12
&'ABIEIZIKG
.4ND
Raw 13ubbs Pressure Refinery Gas Dist. Gae 1,600,000 SCF'I 2,000 bbl. 200,000 8CF 40~65 40-65 Atm. 80 60-80 80 1.12 (air = I) 0.755 fend 1.08 (air 1) point 400500' F.) (62 Ib., 80' F.) (521b.,80' F.) (.Atm., 80" F.) 38.6 1 0 52 25.9 3 0 16 23.0 10 8 14 3.0 3.7 1 6.0 10 9 0 2.1 12 1 5 i.4 58 5 2 .
____
-
io0.o
100 0
100
Final Product Streamsb
13 12 Quantity per 24 hi,. SP. El".
The vacuum tower must perform the separation between the final residue and the heary naphtha overhead with good fractionation in order to obtain a higher yield of finished motor fuel distillate. Fractionation such that no A. S.T. AI. overlap is obtained between vacuum overhead and bottoms is fairly common practice. The unit shown in Figure 5 is designed for the process conditions given in Table V. The installation described i s designed to utilize steam as the heating medium, which gives definite assurance against high film temperatures. Successful installations have been heated by gas oil circulated through a tube still and thence to the shell and tube reboilers and preheaters. With the present knowledge of heat transfer, combined with intelligent operation, this heating means avoids excessive film temperature.. .
dBSoRBING,
OPEEATIONPEEDSTOCKS
14,000 13,300
115 160 195 220 245 260 280 300 335
S'OL. 27, NO '2
Stable Pressure Diet. 1400 bbi. 0.810
Stable Light Gasoline 555 bbl. 0.625
Residue Gaa 2,165.000 ijCF
..;
Sorinal Operating Coiiditioiis a i d Equipment Sizes Debutanizer Operatiiig pressure, lb./sq. in. Temw.. F.: To'w'er ton
Reheat &let Tower bottom Reflux quantities, gal /min Vapor velocity, ft./sec Reflux ratios Stripping steam, ib./hr. Tower diam., in. No. plates 5
b
SO
Stabilizer
4bsorber
125-175
50
160
130
350 18.1
220 20.5 0.42
...
0.53
111
....
36 26
Still 65
..
....
1800 3500 ibb 5 5 c 3600 19.8
3.2/1
. .. . .. ...
.,..
30 30
,
36 26
0.57 2/1 850 36 16
Standard cubic feet, measured at 80' F.and 760 mm. pressure. Butane recovery is 80% minimum, Butane recovery varied t o obtain
7 t o 13 pounds Reid vapor preseure of blended stable pressure distillate and
Lubricating- Stock Rerun Operations Several lubricating-stock rerun operations are commonly practiced. The modern equipment in general use consists of atmospheric or vacuum pipe stills similar in design to those used for crude units. Major operations are as follows. RERCXKI~VG OF PRESSED AX DISTILLATE. This operation consists in the redistillation of stock vhich has been dewaxed by filter pressing. The rerun operation eliminates nonviscous gas oil, and the balance of the dewaxed oil is fractionated into two or more neutral oils for use in blending for auto-
light gasoline. C Circulating laan oil rate.
motive lubricant or in the nianufacture of light machine oils Atmospheric distillation is custoiiiarily employpd with tuhe still outlet temperatures of 700" to 730" F. CYLINDERSTOCKSOLUTIOK RERUSSING.The cylinder stock-naphtha mixture froin centrifuge deTvaxing is rerun iii atmospheric pipe stills for removal of the naphtha dilueiit. Tube-still outlet temperatures of 375" t o 550" F.are requirrti RESIDLE GAS
ABSCRBER
FIGL HE 6 .
h 3 S O R P T I O N D F B U T 4 h I Z I V G 4UD STABILIZING UUITS
JULY, 1935
INDUSTRIAL AND ENGINEERING CHEMISTRY
801
RERUN OF TREATED ASD DEWAXEDLOKGRESIDUUM ucts consist of stable pressure distillate, stable light gasoline, and residue gas, S. A. E. MOTOROIL GRADES. A vacuum operation is The complete installation consists of: a debutanizer in dictated by the fact that lube fractions of high boiling point which the raw pressure distillate is stabilized; a stabilizer are taken overhead. for fractionating the overhead condensate from the debutaStabilization and Gas Recover) nizer and the condensate recovered in the absorption and stripping towers; an absorber in which the cracking still gas, the The gas recovery system is a necessary part of the modern refinery gas, and the uncondensed vapors from the stripping refinery. Wet gas as collected from the various operations is column are treated; and a steam stripping and rectifying processed for the recovery of stable light gasoline for blendtower for stripping the circulating absorption oil. ing. The gas recovery operation is conveniently combined Details of the debutanizing, absorbing, and stabilizing with the stabilization of cracked distillate. A combination operation are shown in Table VI. unit of this type which has been recently installed is shown in Figure 6. RECEIVED March 29, 1935. Presented as part of the Symposium on DistilThis plant receives cracking still gas and raw pressure dislation held under the auspices of the Division of Industrial and Engineering tillate directly from the cracking unit receiver. I n addition, Chemistry of the Smerican Chemical Society at the Massachusetts Institute refinery gas from the collecting system is processed. Prodof Technology, Cambridge, Mass., December 28 and 29, 1934.
FOR
Production of Potassium Sulfate from Poly halite and Svlvinite J
XPLORATIOSS by government and private agencies have disclosed the presence of large quantities of sylvinite, a mixture of potassium chloride and sodium chloride, and of polyhalite, a complex sulfate of potassium, magnesium, and calcium, occurring in bedded saline deposits in the Permian basin of Kern; Mexico and Texas (12, 16). Exploitation of sylvinite near Carlsbad, K.Mex., has already been undertaken actively by the U. S. Potash Company and the Potash Company of America, the former having produced refined potassium chloride since 1932 (16,dO). To date no attempt has been made either in this country or elsewhere to recover potassium salts from polyhalite on a commercial scale, but as a result of the prevalence of this mineral in the Texas-New Mexico deposits, this station of the U. S. Bureau of Mines has been investigating methods for its treatment. One of the possibilities considered has been the use of polyhalite and sylvinite together to produce potassium sulfate, which has always commanded an appreciably higher price than potassium chloride. As indicated by its formula, K2S04.MgSOc2CaS04.2H20,pure polyhalite might be made to supply three sulfate radicals in excess of the one equivalent to its potassium content. This paper presents a method of utilizing the sulfate content of polyhalite to convert refined potassium chloride into potassium sulfate by employing ammonia and carbon dioxide in a cyclic process.
Process Outline The complete process, which will be designated as Bureau of Mines process 8, may be divided arbitrarily into four parts:
I. Raw polyhalite is washed to remove sodium chloride and
is calcined. The material is then decomposed by water at 25' C.
to roduce a solid consisting of a mixture of syngenite (KzS04Ca804.HZO) and gypsum (CaS04.2HzO)and a liquor containing essentially all of the magnesium sulfate and approximately 8 per cent of the potassium sulfate from the solid. 11. The syngenite-gypsum mixture is decomposed at 25' C. by ammonia and carbon dioxide in aqueous solution. This treatment precipitates calcium carbonate and yields a liquor containing potassium and ammonium sulfates.
J
ALTON GABRIEL
AND
EVERETT P. PARTRIDGE
Nonmetallic Minerals Experiment Station, U. S. Bureau of Mines, Rutgers University,
New Brunswick, N. J.
111. Solid refined potassium chloride is added to the otassium causing sulfate-ammonium sulfate liquor and agitated at 25" deposition of part of the potassium sulfate. The potassium sulfate remaining in solution is preci itated with ammonia. IV. The mother liquor, whick contains essentially ammonia and ammonium chloride with small amounts of potassium chloride and ammonium carbonate, is digested with lime and distilled t o recover ammonia, which is recirculated.
&,
The magnesium sulfate liquor derived from I, the calcium carbonate residue from 11, and the calcium chloride liquor from IV may be processed further as discussed subsequently. TABLEI. PERCENTAGE ANALYSISOF AVERAGESAMPLEFROM CARLOAD OF NEWMEXICO POLYHALITE" Polyhalite: 22.15 %SO4 15.25 MgSOa 34.64 Cas04 4.6 H20 Anhydrite Halite Magnesite RzOa Si02 Analysis by F. Fraas of this station.
+
5
76.64
8.17 12.81 0.73 2.29
Figure 1, which presents the process outline in greater detail, is based on the following considerations: (1) Crude polyhalite of the composition indicated in Table I, based on the analysis of an average sample from a carload shipment. (2) A loss of 10 pounds of potassium sulfate and 7 pounds of magnesium sulfate per 100 pounds of sodium chloride removed during washing of crude polyhalite to a sodium chloride content of 1.5 per cent (4). (3) A concentration of 32 pounds of magnesium sulfate and not more than 4 pounds of potassium sulfate per 100 pounds of
