NM
NBI
XI, x2, x ,
flux of solute A2 through membrane for feed system [A2-Hz0], g. mole/ sq. cm. sec. flux of solvent water through membrane for feed system [AI-H20], g. moleisq. cm. sec. flux of solvent water through membrane for the feed system [A2-HzO], g. mole/ sq. cm. sec. product rate for feed system [Al-HzO], g. per hour per 7.6 sq. cm. of film area product rate for feed system [A2-H20], g. per hour per 7.6 sq. cm. of film area product rate for the mixed solute system, g. per hour per 7.6 sq. cm. of film area pure water permeability, g. per hour per 7.6 sq. cm. of film area fractional solute concentration ratios defined by Equations 1, 2, and 16, respectively
literature Cited
Agrawal, J. P., Sourirajan, S., IND.ENG.CHEM.PROCESS DESIGNDEVELOP. 8,439 (1969). Erickson, D. L., “Cellulose Acetate Membranes as a Means of Removing Scale Forming Ions of Natural Saline Waters,” Department of Engineering, University of California, Los Angeles, Rept. 66-7 (1966). Fischer, R. B., “Quantitative Chemical Analysis,” p. 239, W. B. Saunders Co., Philadelphia, 1956. Kimura, S.,Sourirajan, S., A.1.Ch.E. J . 13, 497 (1967).
Kimura, S., Sourirajan, S., IND. ENG. CHEM. PROCESS DESIGN DEVELOP.7,197 (1968a). Kimura, S., Sourirajan, S., IND. ENG. CHEM. PROCESS DESIGNDEVELOP.7, 548 (1968b). Ohya, H., Sourirajan, S., IND.ENG. CHEM. PROCESS DESIGNDEVELOP.8, 131 (1969). Slavin, W., “Atomic Absorption Spectroscopy,” Chap. IV, Interscience, New York, 1968. Sourirajan, S., Ind. Eng.Chem. Fundamentals 3,206 (1964). Sourirajan, S., “Separation of Some Inorganic Salts in Aqueous Solution by Flow under Pressure through Porous Cellulose Acetate Membranes,” National Research Council of Canada, N.R.C. 7498 (1963). Sourirajan, S., Agrawal, J. P., “Reverse Osmosis,” I& EC Symposium on Flow through Porous Media, Washington, D. C., June 1969. Sourirajan, S., Govindan, T. S., Proceedings of First International Symposium on Water Desalination, Washington, D. C., 1965, Office of Saline Water, U.S. Dept. Interior, Washington, D. C., Vol. 1, pp. 251-74. Sourirajan, S., Kimura, S., IND. ENG. CHEM. PROCESS DESIGNDEVELOP.6, 504 (1967). Van Nieuwenburg, C. J., Van Ligten, J. W. L., “Quantitative Chemical Micro-Analysis,” p. 121, Elsevier, Amsterdam, 1963. RECEIVED for review November 8, 1968 ACCEPTED August 11, 1969 Issued as N.R.C. No. 11110.
HYDROCARBON EXTRACTION OF SALINE WATERS PAUL
B A R T O N
A N D
M . R .
F E N S K E
Department of Chemical Engineering, The Pennsyluania State University, University Park, Pa. 16802 Potable water and saturated brines are produced from saline waters by extraction with liquid hydrocarbon at 650’ F. and 2600 p.s.i.0. High selectivity of hydrocarbon for water over salt enables use of a single extraction stage with sea water. Solvent and heat recovery are accomplished in fluidized particle heat exchangers to eliminate scaling problems. Vessels are constructed of high strength steel lined with titanium. Particles are alumina. Product water contains 120 p.p.m. of salt and 12 p.p.m. of CII-C~~paraffinic hydrocarbons. A proposed plan! producing 10,000,000 gallons per day of desalted water has heat exchange towers up to 18 feet in diameter and 24 feet in height and requires a capital investment of $18,000,000. Operating costs are 99 cents per thousand gallons of water.
THEexpanding need for water in certain population and industrial centers can be satisfied economically by desalting sea water or brackish inland waters. Evaporation is the conventional method for saline water conversion, wherein the large amount of energy required to vaporize water is sometimes recovered by multistage operation. Freezing processes for desalting water require less conversion energy, but brine occlusion has retarded their commer18
cial development. Membrane processes have low energy requirements, but low selectivity, plugging, and lack of large scale equipment currently limit their application. Liquid extraction processes are low energy consumers, whereas deficiencies in solvent selectivity and product potability exist with solvents reported previously. A new extraction process uses extremely selective hydrocarbon solvents and produces high quality potable water. Ind. Eng. Chem. Process Des. Develop., ‘401. 9,No. 1, January 1970
Pure hydrocarbon
~ - - [ - - r - i - - - r 1~- 3- Reference
doto are at
Solid lines
Fenske et ol., 1966 Guerront, 1964
A n-Decane o 1,2,4-TrimelhyIbenzene
conditions
I#Z,4-Trimethylbenzene: Brine(20.5 wl.%NoCI) Thisstudy ( 4 w t . x NaCI)
a 0.05 01
L
0) c
L
i
0.01
z .C_
C
0
c
Reterence Fenske et 01,1966
Solid l i n e s
0 0005
u
E
Guerront, I964 0
\2,4-Trimethylbenzene ,- 3 5 wt % NaCI) This study Benzene brine ( 4 i y f Yo N a C l ) ~
0
Id0
2d0
'
Temperature,
do
1 I
1
400 500600
F
Figure 1. Solubility of water in hydrocarbons
The hydrocarbon is used to extract desalted water from brines. Processes based on this phenomenom have been revealed recently in an independent investigation by Hess et al. (1967). Solubilities of Hydrocurbons and Water
Hydrocarbon extraction of desalted water from brines is based on the high solubility of water in liquid hydrocarbons at high temperatures and pressures and low mutual solubilities of hydrocarbons and water a t ambient temperature. Solubilities of water in selected hydrocarbons under vapor-liquid-liquid equilibrium conditions and in liquid kerosine are shown in Figure 1. Solubilities of some hydrocarbons in water under vapor-liquid-liquid equilibrium conditions are presented in Figure 2. Additional solubility data can be found in the work of Guerrant (1964). The effect of hydrocarbon type on water solubility is seen by comparing the curves of benzene, cyclohexane, and n-hexane. Water is equally soluble in the saturates, but more soluble in the aromatic. Comparison of the curves for benzene and xylenes indicates that a change in hydrocarbon molecular weight does not change the amount of water dissolved on a mole fraction basis. Mixtures of hydrocarbons are expected t o exhibit about the same affinity as their individual components for water. Kerosines with average molecular weights of 170 to 173 and characterization factors of 11.8 to 12.0 dissolve the same number of moles of water as the pure Ct, saturates, Ind. Eng. Chem. Process Des. Develop., Vol. 9, No. 1, January 1970
t
I"
'
Id0 ' Z d O ' 3 d O
'
4 A O ' 5 d O ' 6O :
T e m p e r a l u r e , OF
'
Figure 2. Solubility of hydrocarbons in water under vapor-liquid-liquid equilibrium conditions
except near the critical temperatures of the hydrocarhons. A three-phase critical point exists when mixtures of water and low molecular weight hydrocarbons are heated under vapor-liquid-liquid equilibrium conditions. The liquid hydrocarbon phase is vaporized as the critical temperature of the pure hydrocarbon is approached. With C6 hydrocarbons, these phase disappearances occur 20" t o 40" F. below their respective critical temperaturrs. Because critical temperatures increase in the order paraffin, naphthene, aromatic for hydrocarbon molecules with the same number of carbon atoms, liquid-liquid extraction conditions can be maintained up to maxiriium [emperatures increasing in that order. T h e solubility of water in hydrocarbons is more strongly dependent on temperature than on the nature of the hydrocarbon. With C6 hydrocarbons, liquid-liquid extraction conditions exist up t o 432" to 515°F. (n-hexane and benzene, respectively). The corresponding solubilities of water in the hydrocarbon phase are 7.4 to 26 weight ci; (28 t o 60 mole W water). Higher solubilities can be attained only by going to higher temperatures, and this usually requires that hydrocarbons with higher critical temperatures be employed. With n-decane and 1,2,4trimethylbenzene as liquid extraction solvents for water, solubility levels of 34 and 38 weight 0% (80 mole 7% water) are attained a t 624" and 606"F., respectively. These temperatures are 30" and 104" F. below the respective critical temperatures of the hydrocarbons. The effect of molecular type on the solubility of hydrocarbons in liquid water is seen in Figure 2 by comparing 19
the benzene, cyclohexane, and n-hexane curves. A t 120" F., the solubilities decrease from 470 to 25 to 15 p.p.m. (mole basis) hydrocarbon, respectively. The number of carbon atoms in hydrocarbon molecules affect the solubility significantly. At 120"F., water will dissolve only 15 p.p.m. of isopropylbenzene and 2.4 p.p.m. of 1,4-dimethylcyclohexane under vapor-liquid-liquid equilibrium conditions (mole basis). With 1,2,4-trimethylbenzene as solvent a t 600" F., the water phase at this temperature contains 1.4 mole % (8.5 weight %) hydrocarbon. At 624" F. with n-decane, the corresponding value is about 0.1 mole % (1 weight 70 hydrocarbon).
dard solution of silver nitrate to precipitate the chloride as silver chloride (white). The end point is determined by the precipitation of silver chromate (red) from sodium chromate indicator solution. The most sensitive method for low salt concentrations is atomic absorption spectrophotometry. The salt contents of the water layer and the hydrocarbon layer of each sample are determined separately by comparison with freshly prepared standard samples. I n this study, salt analyses of enriched brines are accurate to 0.1 weight 7%. The accuracies of the salt analyses of the extract and vapor samples range from 20 to 100% of the concentration level for the atomic absorption and gravimetric-titrimetric methods, respectively. Equilibrium data for several tests with benzene, n-decane, and 1,2,4-trimethylbenzene extracting desalted water from brine are given in Table I . I n a single extraction stage, extracts containing 9 to 300 p.p.m. of salt (by weight) are produced from sodium chloride feed solutions with the same salt content as sea water. Cooled decanted water from these extracts contains 40 to 2000 p.p.m. of salt. The brine by-products contain from 4 to 20.5 weight % sodium chloride. The effect of temperature and water solubility on the selectivity for water over salt is notable. The extract phase contains less salt as the temperature and consequently the solubility of water are increased. Above 600" F., desalted waters containing 120 p.p.m. (by weight) or less of salt are produced in a single extraction stage, well below drinking water standards of 250 mg. per liter of chloride and 500 mg. per liter of total dissolved solids. Below 550" F., it may be possible to produce potable water by using multistage extraction with desalted water reflux. Aromatic and paraffinic hydrocarbons are equally selective in dissolving water from brine. Increasing the salt content of the brine does not deleteriously affect this selectivity. In hydrocarbon extraction of sea water it should be possible to increase the yield of desalted water to the point that the residual brine is saturated with sodium chloride (ca. 27 weight %). With the autoclave used in these experiments, visual confirmation could not be made that two liquid phases and a vapor phase exist at high temperature-to exclude the possibility that a distillation process and not an extrac-
Hydrocarbon-Water-Salt Equilibria
Vapor-liquid-liquid equilibrium data were obtained on hydrocarbon-water-sodium chloride systems using a 1.3liter rocking autoclave as an equilibrium cell (described by Guerrant, 1964). Temperature was controlled with an insulated heating jacket. Temperatures were measured with an iron-constantan thermocouple which had been calibrated with a platinum resistance thermometer. Pressures were measured with a Bourdon tube gage calibrated with a dead weight tester. After charging 0.7 to 0.8 liter of fluid, the autoclave was brought to temperature equilibrium and rocked for 10 to 30 minutes. Then the phases were allowed to settle and equilibrate for 20 to 100 minutes before sampling. Two or three samples ranging from 25 to 55 ml. were withdrawn from each phase through ice-cooled sample taps with 3-ml. holdups. The first 5 to 12 ml. sampled from each phase were purged. The cooled samples were centrifuged and the hydrocarbon and water or brine contents were measured volumetrically. These values were adjusted for the solubility of water in the hydrocarbon layer and the solubility of hydrocarbon in the water layer at room temperature to calculate the sample composition. Accuracy of the analyses is estimated to range from 1 to 5% of the concentration level at high and low weight fractions, respectively. The salt content of each sample was determined by one of several analytical methods. One technique is to evaporate the fluids and measure the salt residue gravimetrically. Another is to titrate the sample with a stan-
~~~
~
~~
Table I. Vapor-liquid-liquid Equilibria for Hydrocarbon-Sodium Chloride-Water
Hydrocarbon Benzene Temperature, F. Pressure, p.s.i.a. Composition, wt. % Vapor phase Sodium chloride Water Hydrocarbon Liquid hydrocarbon phase Sodium chloride Water Hydrocarbon Liquid water phase Sodium chloride Water Hydrocarbon
20
505 1220
...
1,2,4- Trimethylbenzene
n-Decane 537 1010
602 1830
623 2100
...
625 2040 0.0001 26 74
0.002 36 64
0.001 29 71
0.03 14 86
0.02 8 92
0.002 26 74
0.005 41 59
0.0009 22.5 77.5
4 92 4
4 96
5 95
8 90 2
20.5 79.0 0.5
...
...
... ...
Ind. Eng. Chem. Process Des. Develop., Vol. 9,No. 1, January 1970
tion process is producing the separations. Evidence confirming the presence of three phases is noted by the differences in the analyses of phase samples removed from the autoclave a t the top (vapor phase) and below the middle (upper liquid or hydrocarbon layer) in the tests in Table I. The closest approach to the hydrocarbon critical temperatures in these experiments ranged from 46" to 83' F. Solubility data for hydrocarbon-brine solutions and hydrocarbon-water solutions are compared in Figures 1 and 2. Significant decreases in the solubility of water in the hydrocarbon layer and of hydrocarbon in the water layer are caused by the presence of 4 to 20.5 weight (C sodium chloride. The equilibrium pressures are also lowered. The presence of nearly saturated brine in an extraction process requires about a 30" F. increase in temperature to attain the same solubility level as in a saltfree system. Process for Extracting Fresh Water
Selection of the hydrocarbon solvent is the first step in devising a process t o extract fresh water from saline waters. The more volatile hydrocarbons such as benzene, hexane, cyclohexane, and light naphthas are not recommended because they have low critical temperatures and must be used well below 550°F. to ensure the existence of two liquid phases a t the system vapor pressure. The selectivity of hydrocarbon solvents for water over salt is not high enough to extract potable water from sea water in a single extraction stage a t such temperatures. I n addition, the solubility of Cs to Cg aromatics in the water and brine products is high enough to affect the cost of conversion, adding 50 to 6 cents per thousand gallons of feed water, respectively. With saturated petroleum fractions or less volatile aromatics the value of the hydrocarhon in the water products is reduced to 0.2 to 2 cents per thousand gallons of water. T o minimize the solvent to water ratio, it is desirable to operate at as high a water solubility level as possible. Inspection of Figure 1 shows that a solubility of 80 mole 5 water in the hydrocarbon phase is attainable. With brine raffinates containing in excess of 20 weight "C salt, extraction temperatures corresponding t o this solubility level are about 650°F. with paraffinic hydrocarbons and slightly lower with aromatic hydrocarbons. Hydrocarbons with sufficiently high critical temperatures for such operation include C9 aromatics, CIL1 naphthenes, CI1olefins, and C l l paraffins. It is preferable not to use hydrocarbons with higher molecular weights than necessary for extraction a t a specified solubility, because on a weight basis the solvent to water treat becomes greater and consequently the heat transfer duty increases. Preliminary taste tests indicated that product water is more palatable when saturated with a paraffinic hydrocarbon fraction than with kerosine containing hydrocarbons of other molecular types. Therefore, a paraffinic CII--C~.' hydrocarbon fraction is recommended as solvent for the extraction process. The concentration of these hydrocarbons in the product water is estimated as 12 p.p.m. by weight. If more stringent water purity is demanded, the desalted water can be treated with charcoal. The initial step in the conversion process is to heat the saline water to about 650°F. The brine is brought in contact with hydrocarbon solvent either while or after being heated. The vapor pressure of the system is about Ind. Eng. Chem. Process Des. Develop., Vol. 9,No. 1, January 1970
2500 p.s.i.a. The hydrocarbon extract and brine raffinate phases are then settled and separated. Two methods exist for precipitating the desalted water from the extract phase: by cooling the extract and by increasing the pressure on the extract. Both methods lower the solubility of water in the hydrocarbon. Inasmuch as water solubility is more sensitive to temperature changes than to changes in pressure (Hess et al., 1967) and the water must be cooled down anyway, the former method appears the simplest. Most of the water is released from the extract upon cooling from 650' to 500°F. The solubility of water in the hydrocarbon decreases from 31 to 4.3 weight 5 in this interval. The solvent is then recycled between 500" and 650°F. and only the water streams are heated and cooled between ambient temperature and 650°F. It is advantageous to heat the water and hydrocarbon streams as a mixture a t temperatures where appreciable water is dissolved by indirectly bringing it in contact with the cooling extract (and brine). I n this way, the heat of dissolution which is gradually released by the cooling extract can be recovered differentially to supply the heat of solution for dissolving the water. If the solvent and water were heated separately, the heat of solution would all have to be supplied a t the highest temperature level in the process from an external source. Hess et al. (1967) apparently have not taken into account heats of solution in their processes. Heat effects with hydrocarbon-water systems have been measured a t high temperatures and pressures in a specially designed continuous flow calorimeter. Hydrocarbon and water were heated separately to the same temperature and forced through two 0.0135-inch-diameter orifices spaced 0.015 inch apart, such that the jets impinged on each other. Temperature changes over the orifices were measured with both fluids flowing in proper proportioqs and with each fluid flowing alone. T h e differences b e t w e n the temperature changes represent the heat of solution effect. Gillen (1968) described the apparatus in detail and t,ested it with the acetone-water sj'stem. Results agreed with published data (Kister and Waldman, 1958) to within 25";. preliminary trials h a w shown that ternperat !ire decreases of 45' to 70°F. occur on niixing ri-decane and 1,2,4-trimethylbenzene with salt-free water at temperatures above 625" F. a t solvent to water weight ratios in the range of 2 to 4. T h e integral heats of solution are endothermic and are estimated to range from 200 to 300 B.t.u. per pound of water dissolved in the hydrocarbon, with the lower values corresponding to higher solubility levels. After mixing, the solubility of water in the hydrocarbon ranged from 15 to 35 weight % water. The effect of salt on heats of solution a t these high temperatures is unknown. Heats of solution were calculated for dilute solutions of hydrocarbons and water below 200°F. from solubility data. The transfer of 1 pound of water from the water layer to the hydrocarbon layer a t equilibrium is estimated to have an endothermic heat of solution of about 800 B.t.u. The transfer of 1 pound of hydrocarbon from the hydrocarbon layer to the water layer is also calculated to be endothermic, but exhibits only about one tenth of the above heat effect. Coupled with the low solubility of hydrocarbon in water a t the extraction temperature, the heat effect on dissolving hydrocarbon in the brine 21
(optlonoll
Figure 3. Raining solids-liquid heat exchange system with hydraulic particle recycle
layer is small compared to that of dissolving water in the hydrocarbon layer. Since most of the water is dissolved in the hydrocarbon layer above 500"F., the over-all energy related to heat of solution is estimated to be approximately 200 B.t.u. per pound of water dissolved in the total extraction process. The energy required to obtain water by hydrocarbon extraction is about one fifth of that required to obtain water by evaporation. Several methods can be used to transfer heat from the cooling extract and brine systems to saline water feed and solvent recycle. Indirect heat transfer in shelland-tube exchangers is one. The saline water and solvent would be passed through the tube side. With sea water, scaling would occur during the heating and extracting process. The tubes need t o be cleaned periodically to maintain the heat flux and avoid plugging. Chemical pretreatment and deaeration inhibit carbonate fouling. The calcium sulfate scale formed can be removed by chemical cleaning. The alternate set of heat exchangers required to keep the plant operational during cleaning imposes a cost disadvantage on this method of heat transfer. Another method of heat transfer is through the direct contact of immiscible fluids in spray towers. The metallic barrier through which heat is transferred and scale is formed is thus replaced by a mobile heat carrier such as vapor, solvent, mineral oil, or other liquid. The scaleforming components would be precipitated directly in the liquid and carried through the process as a dilute slurry. Among nontoxic materials available, significant solubility between the fluids is displayed a t the temperature levels required for the hydrocarbon extraction process. For this reason, the hydrocarbon solvent itself may be the most applicable heat carrier. This is the principle upon which Hess et al. (1967) have based their flow plans. 22
Several disadvantages become explicit. A hot hydrocarbon stream flowing countercurrently to a cold water stream for the purpose of heating the water will inadvertently dissolve some water and subsequently release it higher in the tower. Consequently considerable internal recycle of fluids within the towers is unavoidable and the towers must be made larger to account for this. Secondly, the low density difference between the fluids limits the capacity of the heat transfer towers. A plant processing more than a few million gallons per day of water would require parallel sets of heat exchanger towers to limit tower diameters to fabricative limits for high-pressure operation. The heat transfer efficiency of liquid-liquid spray towers is low and consequently the towers must also be tall. For example, Greskovich et al. (1967) found that the height per theoretical heat transfer stage ranged from 1 to 7 feet in towers 4 and 6 inches in diameter. The above disadvantages can be overcome by using immiscible, high density, solid particles in place of the fluid heat carriers. Fluidized Particle Heat Exchangers
Solid particles can serve as immiscible heat carriers in saline water conversion processes, using heat exchange towers such as those in Figure 3. The solid particles fall countercurrently to the fluid with which they are transferring heat. The high density difference available between fluids and solids produces high slip velocities, which allows for small tower diameters and high heat transfer coefficients, which in turn reduce tower height. The towers can be made compact because of the large heat transfer areas made avaliable per unit volume of tower by using small particles. Corrosion problems are eliminated by selecting solid pellets made of refractory Ind. Eng. Chem. Process Des. Develop., Vol. 9, No. 1, January 1970
650 "F Extractor Saline water and solvent heaters,
Settler
Y
Brine cooler
Hydrocarbon recycle W o r k recovery unit
solvent
water feed
Figure 4. Process for liquid hydrocarbon extraction of desalted water from brines using raining solids heat exchangers
solids such as alumina. Scaling is not a problem. The scale formed on the solid particles is continuously rubbed off and flows with the fluid as a dilute slurry. I n the heat exchange sections of the towers in Figure 3, the solid particles are closely packed as they fall through the liquid to keep the towers compact. A void fraction of about 0.55 should be maintainable. Vibrating screens can be used to distribute the solids evenly as they enter the heat exchange zone. Heat transfer coefficients of 100 to 3000 B.t.u./(hr.) (sq. ft.) (OF.) were reported by Holman et al (1965) in water-fluidized heat exchangers with %,to I (,-inchmetal spheres. The solid particles can be transferred mechanically, pnamatically, or hydraulically from the bottom of one heat exchange tower to the top of another. Hydraulic transfer as shown in Figure 3 appears most appropriate for the hydrocarbon extraction process on a large scale. Kopko (1969) has demonstrated the operability, lift fluid requirements, and lift line pressure drops for this type of heat exchanger. Dense phase lifting was accomplished with void fractions of as low as 0.74 in the lift line. Kopko's apparatus contained a 2.5-inch-i.d. lift line 13 feet high and a 10-inch-diameter solids return annulus. Water was used to recycle Yl ,,-inch and %$-inch-diameter iron shot and 4-inch-diameter alumina pellets. Each volume of water lifted from 0.03 to 0.17 volume of solids. Pressure drops ranged from 0.7 to 1.0 p.s.i. per foot of lift line. Most of the pressure drop was attributed to static head in the fluidized system. Washing of solids is necessary to prevent lift water depletion and contamination of fresh water with salt in fluidized heat exchangers. Preliminary measurements showed that 7 volume "c water is occluded by the iron shot after short drainage periods. Hydrocarbons can be used to displace this water and to wash the particles Ind. Eng. Chem. Process Des. Develop., Vol. 9,No. 1, January 1970
in transit between heat exchange towers, as shown in Figure 3. Flow Plan and Economics
One flow plan for hydrocarbon extraction of desalted water from saline waters using fluidized particle heat exchangers for energy and solvent recovery is presented in Figure 4. Saline water feed is pressurized to 1000 p.s.i.a. and heated to 456°F. in exchangers E-6 and E-7. It is then compressed to 2600 p.s.i.a. and heated with recycle CII-CI1 paraffinic hydrocarbon solvent to the extraction temperature of 650°F. in exchangers E-2 and E-3. The solvent to water weight ratio is 3.1 with sea water feed. The brine is extracted in a single stage to saturation, separated, and cooled to 122'F. indirectly by solvent and brine feed in exchangers E-4 and E-5. The extract is cooled to 500" F. indirectly by solvent and brine feed in exchanger E-1. The heats of solution of water and hydrocarbon are released in exchanger E-1 and regained in E-2. The solvent is decanted at 500°F. and recycled. The desalted water is cooled indirectly by brine to 122" F. in exchanger E-8. The product water streams are throttled through work recovery units to supply energy for compressing the feed water streams. With sea water feed, the yield of desalted water is 87"c and the yield of brine saturated with sodium chloride is 135. For a plant producing 10,000,000 gallons of desalted water per day from sea water, the heat exchange towers range from 3 to 18 feet in diameter and from 13 to 24 feet in height. Average heat transfer coefficients of 1000 to 2000 B.t.u./(hr.)(sq. ft.)('F.) were used to size the exchangers, employing ?/,j 2-inch and '/-inch alumina pellets. Lift lines 1.5 to 9 feet in diameter recycle the pellets a t rates of ?ito 9 tons per second, or a total
23
160
I
I
8
1
,
I
I
I
1
I
I
I
i
I
1
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1
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-
700
----_
Figure 5 . Enthalpy diagram for sea water, desalted water, and solvent heat exchangers
700 140
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li
600
;120 0
m
100
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-
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s
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- Bottom of seo water 80
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'
3d0 ' 3;O
' 4d0 ' 450 ' 5 d 0 ' 5O ;
Enthalpy of Fluid, B t u / Ib Solvent Plus Woter or Brine
Figure 6. Enthalpy diagram for process with minimum number of heat exchangers
of 12 tons per second. The amount of lift water cycled daily is 614,000,000 gallons. The heat exchanger vessels are constructed of high strength maraging steel fitted with titanium liners. Steel wall thicknesses range from 0.5 to 6 inches. The energy required t o provide the temperature differentials to operate the heat exchangers is supplied by furnaces through which the lift fluids flow. These fluids then heat the solids in the lift lines of exchangers E-1 and E-4. Average temperature differentials between the solids and the fluids in the countercurrent heat exchange zones range from 3" to 26"F., most of the heat being transferred across differentials of 8" to 12" F. An enthalpy diagram in Figure 5 depicts the operation of the heat exchangers in the flow plan. The departure 24
between the equilibrium and operating lines indicates the thermal driving forces a t various locations in the exchangers. The slope of the operating lines equals the ratio of fluid to solids flow in each exchanger. The spread between the operating lines a t the hot end of the process shows how much the solids are heated by the furnace. The closer these operating lines approach each other (at constant slope), the lower the energy requirements are for the process. Likewise the products will emerge cooler. A simpler hydrocarbon extraction process with lower capital investment uses only two sets of heat exchange towers (instead of four) and recycles the solvent over the complete temperature range of ambient to 650" F. The enthalpy diagram for this process in Figure 6 shows why this scheme is not recommended. The operating lines Ind. Eng. Chem. Process Des. Develop., Vol. 9,No. 1, January 1970
Table II. Cost of Hydrocarbon Extraction of Sea Water
Capacity. 10,000,000 gallons desalted water per day 1,200,000gallons saturated brine by-product per day Capital costs, millions of dollars Raining solids heat exchangers 2.4 Furnace 1.0 Extractor and decanters 3.1 Pumps and drives 3.2 Total equipment 9.7 Total plant investment (TPI)“ 18 Operating costs, cents/lOOO gal. desalted water Amortization, 5% TPI/annum 27 Energy, 25t/million B.t.u., 7 mills/kw.-hr. 30 Materials 1 Operating labor 2 Maintenance, 2% TPI/ annum 11 Overhead 14 14 Finance charge (4%) Product water cost 99 “Based on calculation procedure of Oftice of Saline Water, 1956.
by-product. Of this, 41% is for depreciation and financing. The total energy cost is 30 cents per thousand gallons, 12 cents of which is for operating the pumps. Solvent cost is negligible. I n conclusion, hydrocarbon extraction with fluidized particle heat exchange provides a simple, compact plant for desalting corrosive and scaling saline waters and for producing saturated brines useful for the preparation of chemicals. Future studies should include experimental demonstrations in a continuous pilot plant and detailed engineering design of commercial scale plants to predict more accurately the capital investment and operating costs. Acknowledgment
The authors are thankful for the assistance of C. C. Peiffer, R . P. Guerrant, K. K. Dye, and the Esso Research and Engineering Co. Literature Cited
are forced far apart throughout the process by the curvatures of the equilibrium lines. Temperature differentials between the fluids and the solids average 20°F. and the products emerge at 170”F. Consequently the furnace energy costs for the simplified flow plan are relatively high a t 33 cents per thousand gallons of fresh water produced (at 25 cents per million B.t.u.). The curvature of the equilibrium lines is caused by heats of solution as the water is differentially dissolved and by varying heat capacity ratios. I n the flow plan in Figure 4, the operating lines can follow the equilibrium lines more closely because the solids recycle rates (and fluid to solids flow ratios) are made different in the ambient to 500OF. and the 500° to 650°F. heat exchangers. As a result, the furnace energy costs are reduced to 18 cents per thousand gallons of desalted water. Capital and operating cost estimates for a hydrocarbon extraction plant producing 10,000,000 gallons per day of desalted water and 1,200,000 gallons per day of saturated brine are given in Table 11. The total equipment costs are about $10,000,000 divided about equally between heat exchangers, pumps and drives, and extractor plus decanters. The total plant investment is $18,000,000. Operating costs are about one dollar per thousand gallons of desalted water, taking no credit for the brine
Ind. Eng. Chem. Process Des. Develop., Vol. 9,No. 1, January 1970
Fenske, M. R., Braun, W. G., Thompson, W. H., Eds., “Technical Data Book. Petroleum Refining,” American Petroleum Institute, New York, 1966. Gillen, J. A., Jr., B.S. thesis, Pennsylvania State University, University Park, Pa., 1968. Greskovich, E. J., Barton, Paul, Hersh, R. E., A.1.Ch.E. J . 13, 1160 (1967). Guerrant, R. P., M.S. thesis, Pennsylvania State University, University Park, Pa., 1964. Hess, H. V., Guptill, F. E., Jr., U. S. Patents 3,308,063; 3,318,805: 3,350,299; 3,350.300 (1967). Hess, H. V., Guptill, F. E., Jr., Carter, N. D., U.S. Patent 3,325,400 (1967). Holman, J. P., Moore, T. W., Wong, V. M., Ind. Eng. Chem. Fundamentals 4, 2 1 (1965). Kister, A. T., Waldman, D . C., J . Phys. Chem. 62, 245 (1958). Kopko, R . J., Ph.D thesis, Pennsylvania State University, University Park, Pa., 1969. Office of Saline Water, U.S. Dept. of Interior, “Standardized Procedure for Estimating Costs of Saline Water Conversion,” PB161375 (1956). RECEIVED for review December 16, 1969 ACCEPTED August 4, 1969 Division of Water, Air and Waste Chemistry, 158th Meeting, ACS, New York, N. Y., September 1969.
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