"Present Addre.. : Creole Petroleum Corp., Amuay, Venezuela

R. N. Watkins, Sugar Land, Texas*

SPECIAL REPORT

December 1969

How toDesignCrude

Whether you are making preliminary

or detailed designs for atmospheric

crude distillation columns, this

process design guide should aid you

Distillation

THE DESIGN and operation of crude oil distillation unitsare still done almost exclusively on an empirical basis.This is because crude petroleum and its products, asidefrom light ends products, are made up of a large numberof discrete hydrocarbons-each compound being presentin relatively small amount. It is simpler then to talk ofboiling ranges when describing the gross properties of acrude and its fractions.

BASIC PRINCIPLESFirst let us consider some of the basic steps for separat­

ing crude oil into fractions without regard for the effectof side strippers or stripping steam. The separation dis­cussed here will involve a vapor overhead product, fivesidestream products and a residual liquid bottom product.A small amount of extra vaporization called overflash willbe used to insure the desired vaporization occurs and toprovide some reflux for the bottom section of the tower.

gYDROCARBON PROCESSING

HOW TO DESIGN CRUDE DISTILLATION •.•

• • • Investment costs must be weighed against operating

The Art of Crude DistillationThe art of process design for crude distil­

lation relies heavily upon' empirical correla­tions. This article uses the methods proposedby Packie1 and Edmister5 as a backgroundfor making preliminary calculations. Latercorrelations can be substituted ,into this de­sign procedure.

Many designers standardize their pro­cedure so that it can be. programed for acomputer. Then they can give more atten­tion to the establishment of an optimum de­sign. In any event the method usually is simi­lar to the one given here.

Pa[Je• Basic Principles . . . . . . . . . . . . . . . . . .. 93• Design Procedure ..." . . . . . . . . . . . .. 95• Crude Oil Evaluation 101• Charge Conditions 102• Column Conditions 104• Summary 106

Usually it is not feasible to use a reboiler on a crudetower. Thus the charge usually is heated to its maximumtemperature before being fed into the column. All distil­late products are vaporized as they enter the tower. Thusthe total heat required by the process must be containedin the feed as it leaves the charge furnace.

Various ways of producing distillate products from thefeed vapor are shown in Fig. 1. The symbols used todcsignate some of these methods were chosen for easyassociation with the type of separation: Type U desig­nates that the sidedraw points are unrefluxed even thoughthere is reflux at the top of the last column. Type A usesan externally-circulated fluid which behaves as an absorp­tion oil for heat removal purposes. Type R designates theuse of an external reflux which is an equilibrium liquidwith respect to the tray on which it re-enters the tower.

Series of Flash Drums. Before discussing the foregoingthree methods of separation, consider how fractions hav­ing the desired boiling ranges can be made using a series offlash drums. Consider the top train of Fig. 1 in which thefurnace effluent enters the first drum where the residualliquid W is separated from the vapor. The first drum ef­fluent vapor is cooled just enough to condense the over­flash L o, so that it will be separated as a liquid in thesecond drum. The overflash leaves the system as part ofthe residual liquid W.

The vapor from the second drum is cooled again tosome predetermined temperature and enters a· third flash.There the heaviest distillate product Di is removed as a·liquid. The process is repeated to pl'od~ce successively

lighter fractions D2, Ds, D4 and DB, The overhead fromthe seventh drum is a vapor product VB.

Each of these separations occurs from cooling andequilibrium condensation so that the equilibrium liquidcontains small concentrations of components lighter thanthe desired distillate. Also the equilibrium vapor containssome of the component~ which belong in the liquid so thatthe next fraction to be condensed will have some undesir­able heavy ends. However it would be possible to use liquidreflux to wash back the heavy components from the vaporand to use stripping vapors to strip light componentsfrom the liquid.

Type U. The first approach to genuine fractionation isshown in the second train of Fig. 1. Each of the vessels inthis train represents one separation section. This type ofunit is uneconomical since heat is removed only in thefinal condenser and is available at a temperature levelwhich is too low for recovery and use elsewhere. Further­more the vapor and liquid traffic throughout the columnincreases markedly from the bottom to the top stages.

A plot of the vapor and liquid traffic in each of thethree designated types of fractionation is shown in Fig. 2.Sizing a tower based on conditions at the top tray resultsin a much larger tower diameter for Type U than re­quired for either a Type A or a Type R unit. These lattertwo types of units accomplish heat recovclY by withdraw­ing one of the internal streams of the tower, cooling it andreturning it to the column. Thus the vapor and liquidtraffic is diminished above these withdrawal sections.

Fig. I-Ways of separating a crude into fractions are com·pared with a series of flash drums. ' .

Bec b\ll':.J:969 HYDROGARBON PROCESSING

~ cost's to select the best type of crude distillation . ..

Type A. The more common way of attaining intermedi­a te heat recovery is shown by Type A. Heavy liquid iswithdrawn from the tray above the draw tray, cooledexternally, and returned to the column three trays abovethe draw tray. This type is called a pumparound refluxsystem. The removal of heat in the side cooler results ina smaller requirement for overhead reflux. This techniqueis generally employed at a low and at an intermediatepoint in the column which makes heat available at hvodifferent temperature levels for exchange against crudeoil charge.

This method has the advantage of stabilizing vapor andreflux liquid flow throughout the section where it is used.I t materially reduces vapor and liquid traffic throughoutthe entire column. These two points can be seen in Fig. 2.

The disadvantage of this method is that the three traysused for heat removal must be considered as only one trayfor fractionation purposes. This is because the pump­around liquid is foreign to the zone in which it is intro­duced. This liquid is good for heat removal purposes butoffers no mass transfer advantages.

Type R. A rectifying section employing subcooled refluxis designated Type R. This is a more efficient method forrecovering heat from the system because it can be usedat every section of the column. It has the additional ad­van.tage of being true reflux which contributes to frac­tionation on all trays. As shown in Fig. 2, a fairly uniformvapOI' and liquid traffic exists in all sections of the tower,thus making it possible to design the trays to be more

Fig. 2-Vapor and liquid loadings at the top of a column canbe reduced by using interstage coolers, ,

efficient. This is the type of separation which will be usedas the basis for the design method to be discussed here.

This method has the severe disadvantage, however, ofrequiring higher capital investment. Note that the designexample given later is set up to withdraw the total down­flow liquid and route it to the side stripper, followed bycrude-ta-product heat exchange. This method incurs alarge stripping cost. The alternate method of strippingonly the product while exchanging against both productand reflux requires two sets of pumps and exchangers ateach draw point. Thus, an evaluation of a Type R unitversus a Type A unit can become very complex and mustconsider investment and operating costs for the plant siteunder study.

DESIGN PROCEDUREIn the sense that engineers understand distillation of

discrete compounds, crude oil fractionation is still a blackart. Packie's work1 at least provides an elementary basisfor analyzing distillation problems. The reader also isreferred to NeJson2 and HengstebeckS for background in­f0n11ation. Other referencesG-

7 are used to derive the prop­erty data used in this discussion.

In the separation of chemically similar materials, it isgenerally true that the greater the difference in boilingpoints of individual fractions, the easier it is to separatethe fractions. The sharpness of separation usually is de­fined as the gap or overlap of the boiling ranges ofadjacent fractions. For a given separation, a large numberof trays will result in a low reflux requirement. Likewise,

Fig. 3-These pressures a!ld stages are recommended formaking preliminary calculatIons.

HYDROCARBON PROCESSING' December 1969 95

HOW TO DESIGN CRUDE DISTIlLATION •••

a higher reflux rate will require a lower number of trays.

Product Specifications. The properties of each fractioncan be varied as required by sales demands, but only atthe expense of the adjacent fractions. Generally the pro­cess designer must estimate the material balance on thebasis of the specifications given for the desired products.The basis for most product specifications for a crudecolumn are derived from the method proposed by Ameri­can Society for Testing Materials (ASTM). This methodreports the temperatures at which certain portions of thematerial are vaporized.

ASTM End Points. One method for setting productspecifications ~s to state the maximum allowable endpoints for the fractions. The heaviest distillate product,called atmospheric gas oil, is excluded from this specifi­cation since it is withdrawn from the column to provideadequate fractionation between the other liquid distillatesand the reduced crude.

If the crude processing facilities contain a vacuum col­umn, economics favor the production of the maximum

feasible amount of all distillates in the atmospheric col­umn. Then the size of the vacuum column can be mini­mized. A side benefit is the maximization of the crudepreheat prior to entering the furnace.

For the example discussed here, the identity of thecrude oil fractions starting with the lightest are: overheadvapor, light naphtha, heavy naphtha, light distillate, heavydistillate, gas oil and reduced crude. Isolation of thesefractions should be computed on two representativecrudes, one light and one heavy.

The material balance will be based on alternatelymaximizing the production of gasoline (naphthas), lightdistillate and heavy distillate. Some typical ASTM boilingranges for these alternatives are given in Table 1.

Key Temperatures. Intermediate temperatures usuallyare specified relative to the product streams. These set thegap between the temperature at 5% ASTM for theheavier fraction and the temperature at 95% ASTM forthe lighter fraction. For preliminary designs, the gapsshown in Table 2 can be used when specificationsare not given.

Fig. 4--ASTM VS. TBP (0% to 0%. 100% to 100%).

.. .Jl

'..;'" :96'~... ~.;.:.:~~,:' "'1 December 1969 HYDROCARBON PROCES~lNO

--rII

TABLE 2-Typlcol Gas Specifications

TABLE l-Typlcal Product Spa<lficatlons

liminary calculations. Other recommendations are givenby Nelson.2 Packie's method1 gives an estimate of therelation between number of trays and reflux ratio.

Operating pressures are estimated by assuming a pres­sure of 0.5 to 1.0 psig. in the final accumulator, i.e., the

I HeayYDistillate

20 toU25 to 50

105 to 10

Temp. for Mal<. Product Given, OF

i~ ;f~ 1

250l~~t~S EP

1i~~m

---Set by allowable furnace temp.---

Product

Light Naphtha .Heavy Naphtha .Light Distillate .Heavy DisUllate .Atm. Gas Oll ,

Hvy. Naph.-Lt. Naph .Lt. DlsUI.-Hvy. Naph .Hvy. DlsUI.-Lt. Distil .Gal Oll-Hvy. Distil .

Separation

These specifications must be transformed into ASTMend point specification before continuing with a design.The method for making this extrapolation will dependon the way in which the specifications are written. Onemethod uses probability graph paper. On this type ofgraph, the plot of ASTM temperature versus volumedistilled for a well fractionated product will appear asa straight line, or at least nearly straight. Thus it ispossible to develop fraction properties from a minimumof information.Material Balance. The material balance is estimatedas a function of the desired characteristics of the frac­tions and the amount of crude that can be vaporized atthe outlet conditions of the furnace. It is necessary toestimate the number of actual trays in the column andthe operating pressures of the column in order to set theconditions at the furnace outlet.

A simplified version of the crude column is shown inFig. 3. The number of trays that have been observed inexisting installations are shown along with some recom­mendations for tray numbers to use when making pre-

FIe. 5-ASTM VS. TBP (5% to 0%. 95% to 100%.)'

December 1969 97

HOW TO DESIGN CRUDE DISTILLATION ...

. . . The desired product specifications determine thevessel which serves as a suction drum for the compre~sor

which feeds the gas recovery plant. Ordinarily, this \I,illbe the only accumulator in the system although unitshave been designed using a primary condenser to generatereflux followed by a secondary exchanger for condensingoverhead product. In this latter case, this final drum alsoserves as a suction drum for the recovery plant feed com­pressor: this drum will operate at a pressure of 0.5 to 1.0psig. For pressure drop through the condensers, use 2.5psig. per shell which will also take care of piping. Use 5psig. across the tower itself. Use 5 psig. drop between thefurnace outlet and the entrace to the tower flash zone.

Product specifications determine the cut points betweenfractions. These temperatures, in turn, determine thequantity of each fraction which will be produced. Themethod to calculate the material balance is as follows:

1. Convert the TBP curve of the whole crude into anequilibrium flash vaporization (EFV) curve for the pres­sure at the charge furnace outlet.

2. Calculate the heat content for both vapor and

liquid phases of the charge at the furnace outlet tempera­ture and pressure.

3. Repeat the preparation of the EFV curve for thetotal pressure at the flash zone inlet.

4. By trial-and-error, find the crude vaporization atthe flash zone total pressure which results in the sameheat content as at the furnace outlet.

5. Plot a line of constant heat content as determinedby the two conditions of Steps 2 and 4. The volume offlashed crude leaving the flash zone will lie on this line ata flash zone temperature of 1_20 F lower than the flashzone inlet temperature of Step 4.

6. The net distillate is equal to the total crude flashminus the overflash. Overflash is taken to insure adequatevaporization and to make sure there will be sufficientreflux in the gas oil and heavy distillate fractionatingzone.

7. Set the stripping steam rate to the bottoms strippingzone at 10 pounds per barrel of net bottoms. Net bottomsis the sum of flash zone liquid and condensed overflashfrom Step 6.

Fig. ~ASTM VS. TBP (10% to 0%.90% toi .'0Q.9Io).

e cut points to use for separating crude fractions . ..8. Calculate the hydrocarbon partial pressure at flash

zone conditions.9. The volumetric yields of the distillate fractions are

determined by converting their ASTM specifications intoTBP initial and end points using Figs. 4-6. The TBP cutpoints on either side of a fraction establish the volu­metric yield ·of that fraction.

10. The ASTM, TBP and EFV data for aU fractionsare determined with the aid of Figs. 4-6 and the correla­tions of Edmister.5 The ASTM and EFV curves areplotted and checked for consistency with specified keytemperature data.

11. Extrapolate the EFV curves from their initialpoints to the points representing minus 11.1 volume per­cent to account for the removal of 10 percent front endsfrom the product streams by the product strippers.

12. The separation of the light naphtha in the con­denser-accumulator is established by vapor-liquid equilib­rium calculations. The front end analysis of the crudeand the TBP estimates for the light naphtha are usedfor this step.

Heat Balance. Aside from heat accompanying the strip­ping steam, the only heat to the column is added by thecharge furnace. For most cases, the furnace outlet tem­perature is assumed to be 7000 F. The furnace outlet tem­perature should not exceed this temperature without spe­cific reasons to the contrary. Some refiners limit furnaceoutlet temperatures to 650° F for kerosine and jet fuelproductions.

The influence of charge heat balance on the flash zonealready has been discussed in the section on materialbalance. The bottoms product is assumed to leave thecolumn at a temperature 30° F lower than the tempera­ture of the flash zone. The other fractions are withdrawnfrom the column at dra\-I' tray temperatures determinedby Fig. 7.

Note the conditions for using Fig. 7 since it correlatesdraw tray temperature against the EFV temperature ofthe product on the tray. Gas oil and bottoms are assumedto be steam stripped while all other products are reboilstripped. For side draw products, bubble point tempera­tures for the unstripped liquid are used. For the overheadproduct, the dew point temperature is used.

For Fig. 7, the flash zone pressure is assumed to be 24.7psia. If a higher flash zone pressure is used, higher tem­peratures will result. If steam stripping is used instead ofreboiling, lower temperatures will result. If only the prod­uct portions are stripped, lower temperatures result.

The number of trays and the position of the draw traysis estimated on the basis of the data given earlier in Fig.3. These estimates will be confirmed by calculations. Esti­mated temperatures and pressures are plotted versus traynumber to approximate a column profile.

Some investigators have stated that 40° F temperaturerises are experienced between -the draw tray and the firstand second trays below the draw and a 15° F drop occursbetween the draw tray and the one above. This does notseem logical in many cases, particularly in the lower

sections of the tower. It seems particularly suspicious inlight of stating that heat removal is more effective on oneof two adjacent trays. Until more operating data is avail­able, the assumption of linear temperature drops betweendraw trays is recommended since it results in a highervalue of circulating reflux, cooler size, and tower size.Thus, any error inherent in the assumption is on the safeside.

Draw tray temperature assumptions are checked by cal­culating the bubble-point temperature of the unstrippedproduct at the partial pressure of the product vapor inthe total vapor leaving the draw tray. This is done bymaking a heat balance around the draw tray and calcu­lating the amount of internal reflux required to absorbthe excess heat. This first requires making a heat balanceat the first tray below the draw tray in order to establishcirculating reflux requirements.

Side strippers are calculated in different ways for thesteam stripper than for the reboiled stripper. The steamstripper is tl:e simpler of the two cases. The steam rateis arbitrarily set at 10 pounds per barrel of stripper net

Fig. 7-The EFV of a product sets draw tray temperature.

HYDROCARBON PROCESSING December 1969 99

HOW TO DESIGN CRUDE DISTILLATION ...

bottoms. For distillate fractions, this will result in a strip­ping rate of between 6 and 10 volume percent of the grossstripper charge. This is shown by Envelope III of Fig. 8.

The trend in recent years has been to specify reboiledstrippers wherever thermally possible. This is due to twofactors: Primarily, stripping steam increases the size ofthe crude tower, the size of the condenser, and increasesthe load on the effluent water treating facilities. Secon­darily, in jet fuel production, it is mandatory that thisfinished product be completely water-free.

Care should be exercised in specifying reboiled strip­pers to insure that an economical heat source is available.For the reason that a high temperature oil source is sel­dom plentiful, the light gas oil is usually steam stripped.Heavy distillate and lighter streams can be reboiled inmost refineries.

For steam strippers, the temperature of the strippedmaterial leaving the bottom of the stripper is assumedto be 30° F lower in temperature than the correspondingdraw tray temperature. For reboiled strippers, this tem­perature is assumed to be 30° F higher than the corre­sponding draw tray temperature.

TABLE 3-Narrow Fractions of 36.3* API Gravity Crude 011

TBP Cum. % on Crude Cut Mol. Cum.Cut Cut Pt., GraTlty, Wt. Mola perNo. OF Wt. Vol. °API Celc. 100 Bbl.

1-'67

0.59 1.00 42.1\ .412-- 8.12 4.70 114:i 68.1 13.253 108 5.62 8.04 93.0 72 23.494 145 6.93 9.70 81.1 85 28.035 172 8.5.~ 11.63 71.0 87 33.45

6 200 10.56 13.98 62.9 94 39.827 212 12.31 15.96 5$.2 96 45.208 231 14.18 18.03 64.6 99 00.779 2Ii7 15.96 20.01 55.1 104 55.82

10 269 17.89 22.16 5U 113 60.86

11 288 19.72 24.13 50.0 116 65.5112 aoo 21.16 25.67 48.7 119 69.0813 317 23.04 27.69 48.8 125 73.5214 335 24.95 29.73 47.8 132 77.7915 350 26.29 31.15 46.0 136 SO.70

16 373 28.26 33.23 45.7 142 84.7917 382 30.13 35.18 43.1 148 88.5218 400 31.35 36.45 42.3 154 90.8519 425 33.28 38.44. 41.1 162 94.3620 437 35.24 40.44 39.9 167 97.82

21 460 87.24 42.45 87.2 176 101.1722 480 39.28 44.48 36.1 186 104.4023 496 41.50 46.69 36.2 193 107.7924 500 42.24 47.42 34.7 198 108.8925 620 44.27 49.42 84.1 202 111.85

26 543 46.38 61.49 38.4 210 114.8127 664 48.48 63.55 33.3 221 117.6128 681 50.61 65.65 33.6 236 120.2729 699 52.63 57.64 83.2 244 122.7130 616 54.84 59.79 81.7 250 125.32

31 625 65.SO 60.72 SO.8 258 126.4232 650 67.94 62.76 29.1 254 126.8183 670 60.08 64.SO 28.8 276 131.1034 695 62.33 66.94 27.9 287 133.4186 700 63.36 67.92 27.7 298 134.~

36 718 66.42 69.87 27.5 ... ......37 737 67.48 71.82 27.4 ... ......88 762 69.66 78.78 26.9 ... ......39 777 71.66 75.76 26.3 ... ......40 804 78.76 77.72 25.8 ... ......41 821 76.87 79.69 25.4 ... ......42 837 77.98 81.66 25.1 ... ......43 870 80.27 83.78 24.2 ... ......44 918 82.66 85.97 23.1 ...Residuum 99.64 100.44 1U

......... ......

• Cut No. lis C. throulb Ca',.... Cut No.2 II Co·...

The temperature of the liquid on the top tray is cal­culated as discussed previously by making a heat balance.On this tray however, instead of using the bubble pointtemperature, the dew point temperature is adjusted tothe partial presure of the condensible hydrocarbons inthe total vapor leaving the top tray. The applicable en­velopes for top tray calculations are Envelopes V and VIon Fig. 8.

FIg. 8--Envelopes used for making heat balances.

100 December 1969 HYDROCARBON PROCESSING

Example Calculations

Basis: Packie's methodDRL alope= (TBP70% -TBP 10%)/60 =9.55 0 F/%DRL 50% =]44 + 40(9.55) = 526 0 FFRL50% = DRL50% - 35 = 491 0 FFRL slope =6.4 0 F/% (from Packic curve1)

= (MeABP, °R)1./3/Sp. Gr.= (920)1./3/0.843 = 11.5

Fig. 9--The TBP of the crude is converted to EFV data.

TBP Slope = (TrO% - T,o'if,) /60= (717 -144)/60 = 9.55°F/%

For volumetric average boiling point,

VAPB = (T2 0% + T 6o% + TsO$) /3= (255 + 531 + 819) /3 = 535° F

MeABP = VABP - 75= 460° F or 920 0 R

From a generalized correlation~:

The characterization factor is defined as

Equilibrium Flash Vaporl%ationo When a crude isheated and allowed to vaporize so that all the vapor isin equilibrium with the remaining liquid for anyone tem­perature, a plot of percent vaporized versus temperaturegives an equilibrium flash vaporization (EFV) curve.These data are laborious to obtain experimentally. There­fore, they are usually derived from TBP data.

The method of Packie' assumes a straight distillation

0.010.Q70.921.442.261.901.800.151.150.751.100.660.670.343.770.951.131.881.17

Vo]••%

-269-128-44

11318297

121136141146166161177176194209211214231

Characterization Factor. Enthalpy data needed for heatbalance calculations usually are correlated with the char­acterization factor for gross hydrocarbon fractions. Thecharacterization factor for each fraction is assumed to bethe same as for the whole crude. To estimate the char­acterization factor, it is first necessary to convert TBPdata into a mean average boiling point (MeABP) asfollows:

TABU 4-Llght Ends Analysis of Crucle

Compo

Cl. .•.........................C2 •..........................Ca ..IC .nC .lC .nC .CP ••........................2. SDMB ..2MP .SMP .n.Co .MCP .CH .Benz ..IC ..nC .DMP .MCH .Tot .

CRUDE OIL EVALUATIONThe design of a crude oil distillation unit requires a

knowledge of the type of crude to be charged. The crudeoil analysis usually is given in the form of a true boilingpoint (TBP) curve and a light ends analysis. For theexample to be discussed here, Table 3 gives the boilingpoint data from which the TBP curve of Fig. 9 is plotted.The light ends analysis is given in Table 4. The otherproperties of the crude needed for subsequent calcula­tions are derived from these tabular data.

To elaborate on the generalizations made in the earliersection, an example problem will be calculated for a col­umn using subcooled reflux (Type R). The charge ca­pacity will be 55,000 barrels per day of a 36.3° API crudeoil having an analysis as given .in the following section.

TABLE S--Crude EfV Calculations

EFV-FRLVol.

TBP-DRL EFV-FRL EFV% TBP DRL FRL TBP-DRL1- ----

-13 tQ,J6 62 96 203 -34 0.3910 t« 144 235 0

0:86 '5 23520 256 240 299 ]5 30430 344 335 363 9 0.34. 3 S6640 437 431 427 6 0.34 2 42950 631 626 491 1\ 1I.34 2 49360 623 622 US 1 0.34 0 US70 717 717 619 0 '2

61980 819 8111 683 6 0:84 68685 897 861 715 36 0.34 12 727

HYDROCARBON PROCESSING December 1969 101

HOW TO DESIGN CRUDE DISTILLAnON ...

. . . The conditions in the flash zone are determined byreference line (DRL l through the 10 percent and 70 per­cent TBP points. Then the cun'es reported by MaxwellG

are used to perIOlm the calculations depicted in Table5. The results are plotted in Fig. 9 as EFV at 14.7 psia.

CHARGE CONDITION

Charge Furnace. Fig. 3 gives a furnace pressure rangeof 14.9-15.4 psig. For this example, a furnace outlet pres­sure of 15 psig (29.7 psia) will be used. Vaporization ofthe charge is dete1111ined by converting the atmosphericEFV curve to an EFV curve at the pressure of the furnaceoutlet. This can be done by extending the atmosphericEFV points on a generalized plot of vapor pressure versustemperature (Cox chart). The results are shown in Fig.10 which is an enlarged view of the upper portion ofFig. 9.

The furnace outlet temperature is assumed to be 700° F.The EFV curve fOl' 29.7 psia and 700° F in Fig. 10 shows72.5 volume percent crude will be vaporized. Referringto Table 3, this volume percent is equivalent to 68.2weight percent vaporized.

Since the total crude has a specific gravity of 0.843, theequivalent of 36.3° API, the density of the vapor andliquid portions is determined by

(sp. gr.)v = (sp. gr.) crude (wt frac) v/ (vol frac) V

(sp. gr.h = (sp. gr.) crude(wt frach/ (vol frach

For 72.5 volume percent vaporized

(sp. gr.) v = 0.843 (0.682/0.725)= 0.792 or 46.8° API

(sp.gr.)r. = 0.843 (1- 0.682)/(1 - 0.725)= 0.976 or 13.5° API

Calculate the total heat content of the column charcicat the conditions of the furnace outlet.

F = 675,800 lb/hrVpo = 0.682(675,800)

= 460,800 lb/hr @ 46.8° APIL po = 675,800 - 460,800

= 215,000 lb/hr @ 13.5° API

H po @ 700° F & 46.8° API = 591 Btu/lbhpo @ 700° F & 13.5° API = 457 Btu/lb

Qpo = VFoHpo + L1I'ohPO

= 460,800(591) + 215,000(457)= 370,550,000 Btu/hr

Flash Zone. Assume the pressure in the flash zone will be5 psi less than that at the furnace outlet; i.e., 29.7 - 5.0= 24.7 psia. Construct an EFV curve for 24.7 psia asshown in Fig. 10.

By trial-and-error, find the amount of crude vapori­zation at the flash zone pressure which results in the sametotal heat content as at the charge furnace outlet. ThiswiII occur normally at a temperature of 4-60 F less than

the furnace outlet temperature. For this example, makethe first estimate 40 F less.

QFZ = Q.ro = 370,550,000 Btu/hr

T(FZ)l = 700 - 4 = 696° FV(J'Z)l = 75.0 vol %or 70.9 wt %

= 0.709(675,800)= 479,100 lb/hr @ 46.1 0 API

L C1I'Z)1 = 675,800 - 479,100= 196,700 lb/hr @ 12.40 API

H(FZ)l @ 696° F & 46.1 0 API = 589 Btu/lbh CFZ)l @ 696° F & 12.40 API ~ 453 Btu/lb

Q(J'Z)l = V(FZ)lH UI'Z)l + L(J'z)lhcJ.z)l= 479,100(589) + 196,700(453)

= 371,300,000 Btu/hr. Too high.

Make the second estimate at 5° Flower.

T(Fz)2 = 700 - 5 = 695° FV(FZ)~ = 74.8 vol %or 70.6 wt %

= 477,100 lb/hr @ 46.2° APIL(n)! = 198,700 lb/hr @ 12.5° APIQ(J'Z)'2 = 370,500,000 Btu/hr. OK

Therefore, 695 0 F is OK.

Fig. IG--The EFV i,s corrected for column pressure.

102 December 1969 HYDROCARBON PROCESSING

the conditions at the outlet of the charge furnace . ..Sf'ripplng Steam Effect. Construct a line of constantheat content, Q,o. The volume of flashed crude leavingthe flash zone will lie on this line at a temperature ofapproximately 2° F lower than the temperature computedfor the non-steam condition. Then

T(J'Z)ll = 695 - 2 = 693° F.V(J'Z)S = 75.2 \'01 % or 71.0 wt %

= 480,500 lb/hr

Next set overflash at 2 volume percent of the total crudecharge. Then the net product distillates including over­head vapor are

~D = 73.2 vol % or 69.0 wt %= 466,300 lb/hr @ 46.5° API= 3259 mol/hr

LOF = 480,500 - 466,300= 14,200 Ib/hr @ 28.6° API= 42 mol/hr

W = 675,800 - 466,300= 209,500 lb/hr @ 13.6° API= 614.2 Bbl/hr

Fig. ll-Material and heat' balance In the flash zone.

Set steam rate at 10 pounds per barrel of bottoms rate.

S = 10(614.2) = 6,142 lb/hr= 341 mol/hr

The partial pressure of the hydrocarbon portion is esti­mated from a ratio of the mole rates.

PHO = P (vapor) / (vapor + steam)= 24.7 (3301/3642) = 22.4 psia

Using the EFV curve for 14.7 psia, the point for 75.2volume percent vaporized occurs at 653° F. Using a gen­eralized chart of vapor pressure versus temperature (Coxchart), the atmospheric conditions are equivalent to693° F at 22.4 psia-eonfirming the estimated tempem­ture and partial pressure for the hydrocarbons. Thus, theflash zone conditions are

PFZ = 24.7 psiaTFz = 693° FQFZ = 370,550,000 Btu/hr

Bottoms Stripping. Assume bottoms temperature is30° F lower than that of the flash zone.

T w = 693 - 30 = 663° Fhw @ 653° F & 13.6 ° API = 433 BtuJlbQw = 209,500(433) = 90,710,000 Btu/hr

Fig. 12-The ASTM and EFV data are computed from TBP.

HYI:?ROCARBON- PRoOES'SING December 1969 103

HOW TO DESIGN CRUDE DISTILLATION •..

With stripping steam available at 50 psig saturated(65 psia & 298° F), the heat added is

H sw = 1179 Btu/lbQsw =6,142( 1179) = 7,240,000 Btu/hr

Summary. The results of the foregoing calculations areshown in Fig. 11.

COLUMN CONDITIONSThe usual specifications for crude oil products are

ASTM end points. and the gaps between the 5% ASTMand 95% ASTM. These specifications are listed in Table6 for this example. From these data, TBP cut points areestimated by using Figs 4-6. The estimates are shown inTable 6.

The TBP curve for each product is reproportioned toextend from °to 100 volume percent as shown in Table7. Intermediate TBP data are read from the curve forthe whole crude. Then Figs. 4-6 and Edmister's methodll

are used to estimate the final ASTM data for each prod­uct. These results are added to Table 7 and plotted inFig. 12. These calculated data are consistent with thespecified key temperature data.

Light Naphtha and Gas. Combine the light ends analy­sis of Table 4- with TBP. data for the lighter boiling por­tions of Table 3 to determine the composition distributionshown in Table 8. Assume the overhead accumulator pres­sure is in the middle of the range suggested by Fig. 3, or15.2 psia.

Use the calculation method shown in Table 8 to de­termine the distribution between vapor and liquid. Thecalculation is made by trial-and-error, but only one trialis shown as an example. These data will permit the com­pletion of the TBP curve for the liquid portion, or lightnaphtha.

Remember that the vapor portion is water-saturated.For the conditions of 100° F and 15.2 psia

p....aler = 0.95 psiaPHO = 14.25 psia

The calculations of Table 8 showVII = 361 mol/hr

Then total vapor including water is(361) 15.2/14.52) = 385 mol/hr

Water in the overhead vapor is385 - 361 = 24 mol/hr or 4341b/hr

Column Profile. Note Fig. 12 shows an extension of theEVF curves to minus 11.1 volume percent. Since theproducts will be stripped by removing 10 percent of thefront ends, this extension is made to estimate the lighterproperties of the products existing on the trays.

Draw tray temperatures are estimated from Fig. 7.These temperatures along with estimates of hydrocarbonvapor pressure are shown in Table 9. Draw tray locationsare estimated by Fig. 3. Now it is possible to constructthe temperature and pressure profile of the column asshown in Fig. 13.

FIRST DRAW TRAYThe draw tray which is used to withdraw gas oil D 1 is

calculated first. Make a heat balance about the bottomof the column which includes the tray below the drawtray. This is Envelope I as shown in Fig. 8 and repeatedin Fig. 14. Tray temperature is read from Fig. 13 andcooled reflux is at 400° F.

hRo1 @ 4000 F & 26.6° API = 287 Btu/lbHR01 @ 623 0 F & 26.6° API = 514 Btu/lbHxn @ 623 0 F & 46.50 API = 541 Btu/lbHS1 @ 623 0 F & 24.3 psia = 1,347 Btu/lb

Heat in = Heat outQ..z + Qsw + Qr..no1 = Qxn + QV.R01 + QS1 + Qw370,550,000 + 7,240,000 + R o1 (287) = 466,300(541)

+ Ro1 (514) + 6,142(1,347) + 90,710,000R01 = 116,900 lb/hr

TAILE 6--Ptoduct Specification.

Gben Eatimated---,- ------T100% TS-9S% T100%

TO%TBPTBP $:;1~ ~i.'

Stream ASTM "8TM TBP oal % %----LI.Naph.... ~5 298 ... 22.lJ 22.0

20 273Hvy.Naph.. 380 399 248 11.8 33.8

35 379Lt. Dlatll.... 560 587 359 19.2 63.0

10 558Hvy.D1stll.. 650 584 529 10.4 63.4

5 654GuOll ..... 735 780 624 9.8 73.2

5 750Bolloms.... ... ... 720 .. . 100.0

TABle 7--Prodvct Et'V Calculations

CIUIl.Vol. Wt.

Stteam % % TBP toTBP toASTM toEFV ASTM EFV1----------

Lt. Naph. ... 0 0 -65 - 3 6610 2.2 -3 62 35 14 32 7030 6.6 41 44 24 13 56 8350 11.0 156 115 100 67 156 15070 15.4 206 50 36 18 192 16890 19.8 255 49 39 18 231 186

100 22.0 298 43 40 18 271 204- --Hvy.Naph.. 0 22.0 248

29 ia '(, 283 31310 23.18 277 296 31830 25.54 304 27 13 7 309 32550 27.9 326 22 12 6 321 33170 80.26 345 19 12 6 333 . 33790 32.62 370 25 18 7 351 344

100 33.8 399 29 26 7 377 351

Lt.Dietil.... 0 33.8 359a6 ii '6 398 447

10 35.72 395 415 45330 S9.56 432 87 19 10 434 46350 43.40 467 35 21 10 455 47370 47.24 504 37 25 12 480 48990 61.08 542 58 28 12 508 501

100 53.0 587 45 42 14 550 515r---------

Hvy. Distil.. 0 53.0 52928

542 59610 64.04 657 13 5 555 60180 56.12 587 80 15 7 670 60850 58.20 606 19 10 5 580 61370 60.28 628 22 14 6 594 61990 62.36 655 27 19 8 613 627

100 63.4 684 29 26 7 639 634------Guou..... 0 63.4 6U

26 :i.2 '4 624 69310 64.58 650 636 69780 66.34 680 30 15 7 651 70450 68.30 700 20 11 5 662 70970 70.26 718 18 11 6 673 71490 72.22 745 27 19 8 692 722

100 78.2 780 85 32 9 724 731

1.04 December .1969 HYDROCARBON PROOESSING

-'--:-r, ~,

!I I

~'.

Another Heat Balance. Make a second heat balancearound the bottom of the column which includes the drawtray (Envelope II).

hL8 @ 5900 F & 26.6° API = 401 Btu/lbhLs @ 5790 F & 26.60 API = 394 Btu/lb

HLs @ 5900 F & 26.60 API = 495 Btu/lbH(~D.Dl) @ 5900 F & 50.00 API = 529 Btu/lb

H S8 @ 5900 F & 24.15 psia = 1,331 Btu/lb

To establish flow rates, make a hydrocarbon materialbalance around the stripper assuming VLa is ten percentof L, (Envelope III.)

La = ROT + VL8 + D1

= 116,900 + 0.1 La + 70,300= 187,200/0.9 = 208,000Ib/hr

Then the heat balance becomes

Heat in = Heat outQn + Qsw + QaOT + Qr.s + QV>L8 = Q(~Dl)

Qr.s + QV,LS + Qsa + Qw .370,550,000 + 7,240,000 + 116,900(287) + L 9 (394)

+ 20,800(394) = (466,300 - 70,300) (529)+ 208,000(401) + L s (495) + 6,142(1,321)+ 90,710,000

La =: 274,800 lbjhr or 903 mol/hr

Check original estimate of column temperature by com­puting the partial pressure of La using a mole ratio

PXO = 24.1(902/4271)= 5.09 psia

T o%TBP @ atmospheric is the same as595 0 F @ 5.09 psia. Thus, OK.

Stripping Steam. Steam is used at a rate of 10 poundsper barrel of total stripper bottoms which is

1:11- 13--Column temperature and pressure profile.

H'£DROCWUmN ,PROCESSING December 1969

TABLE a-Separation of Overhead Into light Naphtha and Gas

TABP, Vol. Mol.CODIP. 'p 9'0 Gal/Dt Lb/Gal Lb/1Ir MW M.I/I1r %---I-----C, ........... -259 om 10 2.6 so 16 1.8 0.13C............ -128 0.07 87 3.144 210 80 7.0 0.40Ca ........... - 44 0.92 888 4,220 3,740 H 85.0 5.98ic. ........... 11 1.44 1.386 4.886 6,400 68 111.9 7.84.c. .......... 81 2.28 2.178 4,865 10,680 68 182.4 12.78iC............ 82 1.00 1,829 6.199 9.600 72 13U 9.24DC........... 97 1.80 1,733 6,253 0,100 72 120.4 8.86116-1U°1'...•. 130 1.6 1,640 4.388 6,760 86 78.6 5.60144-176'1'.... , 160 2.5 2,406 6.806 14,180 80 160.3 11.16176-203'1'.... , 189 2.5 2,~06 6.336 12.840 05 136.l 9.47203-229"1'..... 216 2.5 2,~06 6.470 1MOO 00 157.1 11.0122ll-256"F.... , 243 2.5 2,406 5.616 13,510 103 181.1 0.10266-298"J..... 277 2.0 1,926 7.020 13.610 113 119.6 8.37------I-----

Tola! OR... ... 22.00 21,175 .... 116,000 .. 1,427.0 100.00

TABLE a cantlnuecl--Vapor·L1quld Equlllbrium' Calculations

(OB) (K)(Lm +](---

X.b"id. DsK@ FUit Loat14.25 Trial, Trial, Vapor, Va (]Iy illereD'.)p.ia La- Ln- Mol/ll1 LbjH.r

----,---Compo 10000F 5 .0 2 .3 Mol/Ilr 1./111 Gal/Hr Vol. '1l>1-----0, 197.0 0.18 0.13 1.8 80 o 0 0 ....c. ".0 0.49 0.46 6.6 100 .4 20 0 Nil.0. 11.& 6.40 4.71 67.2 2,~ 17.8 790 187 1.1~ U 6.42 4.76 67.9 3,930 44.0 2,MO 548 8.3DC. 3.8 9.76 6.71 96.7 6,840 86.7 6,040 1,030 8.2lC. U14 U2 UO 41.4 2,080 90.5 6,520 1,254 7.6DC. 1.10 4.67 2.37 33.8 2,430 92.6 6,670 !,i70 7.7130"11 6.603 2.06 0.93 18.3 1.140 65.2 6,610 1,278 7.7180'11 0.887 2.80 1.14 16.3 1,480 143.0 12,730 2.159 13.0180'11 0.182 1.44 0.56 7.8 740 127.3 12.100 2.208 13.7216"F 0.10ll 1.00 0.39 6.6 650 151.6 Is.o10 2,320 14.0243"11 0.068 Q.l5O 0.18 2.6 280 128.5 18,250 2,369 lU'D'M! 0.026 0.21 0.07 1.0 110 118.5 18.400 1,909 11.6

ToW 40.24 26.30 881.0 22,800 1,066.0 93.700 16.502 100.0Lo... OX:

TAIIlE 9--P...Umlnary EstImates of Column Condition.

Prusure, Eat.Temp.• OF Millimeters ab•. D...."

Stream Te~••UllItdppecl O%EFV 100% EF'V p PRO

Lt. Nallh. ......... 56 204 1.044 916 216

i:t'?'Dl:tie~:::::::: 308 361 1.096 860 316436 816 1.168 670 400

Hvy. Dl.tU, ....... 685 634 1.205 840 525Atm. Gas Oil...... 686 731 1,251 2W 590

105

,.~

TABU: 1~roduct Stream Summary

lndex!.o.J Terms: Boiling Poinu-6,7. Columns/Process/·910, Crudes· I , De.sjgn-4,8, Distillation.4.8,9, Equations·IO. Heat·6. Oils-I,2. Pbysical Prop_erties.6,7. Properties/Charactemtic:s/.6,7, Proportioning.7, Sepa,·ation-4,8,9, Vapori.zabon-7,

NOMENCLATURE

distillate ratedistillation reference lineequilibrium flash evaporizationfeed ratevapor enthalpyliquid enthalpyvapor·liquid equilibrium constantliquid ratemolecular weightmean avera~e boiling pointnormal boiling pointoverhead ratetotal pressurepartial pressureheat contentreflux ratesteam ratespecific gravitytemperaturetrue boiling pointvapor ratevolumetric average boiling pointresidual rateSubscriptscooledfurnace outletflash zonehydrocarbon portioninternaloverflash

DDRLEFV

FHh

KL

MWMeABP

NBPOH

P

&RS

SpGrT

TBPV

VABPW

cFOFZ

HCi

OF

ASTM TBP IStream EP,oF ~VoJ. % BbJ/D "API Lb/llr

Gas.V•..........Zii 273 22.0 2.U20

7i22,300

Lt. N~b., D•..... 9,~SO 93.700

r~))lstu.'~?~: : :377 379 11.8 6,490 46 75,400650 558 19.2 10,5flO 34 131,600

Hvy. Distil. D•... 6a9 654 10.4 5.720 30 73,000Atm. Ga. Oil, D•.. 724 750 9.8 5.390 26.6 70,aoo--

I I 1'466,300Total Stream., 1:D ... ... I 73.2 40,260 46.5Reduced Crude. , .. ." ... 26,8 14,740 la.6 200.500

Crude Cbatle ..... .. , - ...-!loo.O I 55,000 I 36.3 I 675,800

Liquid loading is

~. = Rc• + R I •

= 116,900 + 165,600= 282,500 lb/hr@ 26.60 .-\PI

Vapor loading is

4V = V~D + R!, + Sw= 466,300 + 282,500 + 6,14-2= 754,942 lb/br or 4,529 mol/hr

SUMMARY OF CALCULATIONSEach of the draw trays are computed in a manner

similar to the first draw tray. Note ho\\"e\'er that the re­boiled strippers are assumed to have temperatures 300 Fhigher than their associated draw trays, instead of 300 Flower as in the steam stripper. A summary of the finalproduct is given in Table 10.

LITERATURE CITED1 PaclI;i';t J. w.~ "Diltillation Equipmcnt in the Oil Relining Industry," 1'ra.,.AlCM. 'Vol. ,7. l!*l'. p. 5lo.

• Ne1Ion. W. L,. "Petroleum Refinery Engineering," <lth Ed., McGraw·HiII,1958.

• Hengstebeck R. J., "Petroleum Procesung," McGraw·Hill, 1959.• Prater, N, H. and BoYd. C. W,. "How to Calculate Multidraw Towers,"

Oil anil Gas Jo~",al May 2, 1955. p. 72.• Edmister. w: C.• "Applied Hydrocarbon Tbermodynantk.," Gul£ Publish.

jnJf Co:. 19H.• Maxwell. J. ll" "Data llook on Hydrocarbons," D. van Nuotrand Co.• 1965,, American "'petrOleum Institute, '\Technical Data Book-Petrolcum Re6mng,"

1966.

Qnl = 70,300(275)= 19,330,000 Btu/hr

Re• + D1 = 116,900 + 70,300= 137,200 lb/hr @ 26.6 0 API

or 598 Bbl/hrS1 = 10(598) = 59801b/hr

or 332 mol/hrQSl = 5980( 1179) = 7,050,000 Btu/hr

HOW TO DESIGN CRUDE DISTILLATION ...

Then the stripper bottoms cooler duty is

QOl = 187,200(381-275)= 19,840,000 Btu/hr

Product heat content is

Tray 7 Conditions. The induced reflux is

Rtf '= Ref [ (h8'23°1i'- h,oo°li') /

(H838o F - hG230p) h6.00 API

= 116,900(428 - 275) /(536 - 428)= 165,600 lb/hr

Operating conditions are

T 7 = 6230 F and P7 = 24.15 psia

About the authorR. N. WATKINS joined Creole Peflro­leum Corp., Amuay, Venezuela, afterp"eparing this article. He has wewkedas a process engineer few Monsanto Co.,Celanese Chemical Co., Fluo'1' CO'fp.Ltd. and Bechtel Corp. He holds B.S.degrees in chemistry and chemical en­gineering f'1'om University of Kentuckyand is a '1'egistered p'1'ofessional engi.neer. Mr. Watkins specializes in theprocess design of petroleum and petro­

chemical plants with pa'1'ticula'1' emphasis on distillation.

December 1969 HYDROCARBON PROCESSING