128 Ind. Eng. Chem. Process Des. Dev. 1985, 24, 128-132

Table X. Results for Raw Coal and for Pure Toluene and Methanol

extr pyridine solub supercrit solvent conditions wt 5% extr of extr coal, 5%

raw coal 11.2

toluene 400 “C 36.0 0.8 6.75 mol/L

6.75 mol/L methanol 400 “C 20.6 2.4

dissolves in the supercritical solvent during the high-tem- perature extraction.

It is also important to report here that part of the su- percritical extract recovered in the toluene phase after cooling the reactor’s contents to room temperature is in the form of a precipitate, fully soluble in pyridine, and amounting to approximately 8 wt 5% on the basis of raw dry coal. Therefore, in contrast to the results of Blessing and Ross, a significant part of the coal material which is soluble in the solvent under supercritical conditions be- comes insoluble when the solvent is brought back to am- bient conditions. Summary

The specific physical and chemical characteristics of supercritical solvents and solvent mixtures that can affect the yield and the properties of coal extracts were experi- mentally investigated. Strong nonideal interactions, such as polar bonding and hydrogen transfer, as well as syn- ergistic interactions in multicomponent solvent mixtures, were shown to produce large deviations from the simple, density-driven supercritical solubility. These interactions can be manipulated to optimize supercritical coal extrac-

tion by reducing the severity of the extraction conditions (pressure in particular). Physical and chemical changes occurring in the coal structure during supercritical solvent extraction were also examined. Acknowledgment

This work was supported by the United States De- partment of Energy under Grant No. DEFG22-81PC40801 and by the Center for Energy Studies, Austin, TX. The support of the two agencies is gratefully acknowledged.

Registry No. Pentane, 109-66-0; hexane, 110-54-3; heptane, 142-82-5; octane, 111-65-9; nonane, 111-84-2; decane, 124-18-5; undecane, 1120-21-4; dodecane, 112-40-3; toluene, 108-88-3; acetone, 67-64-1; methanol, 67-56-1; ethylene glycol, 107-21-1; methane, 74-82-8; ethane, 74-84-0; ethanol, 64-17-5; propane, 74-98-6; 1-propanol, 71-23-8; 1-hexanol, 111-27-3; 1-octanol, 111-87-5; pyridine, 110-86-1.

Literature Cited Adams, R. M.; Knebel, A. H.; Rhodes, D. E. Chem. Eng. Rog. 1979, 75, 44. Blessing, J. E.; Ross, D. S. ACS Symp. Ser. 1978, No. 71, 171-185. Fmg, W. S., et al. ”Experimental Observatlons on a Systematic Approach to

Supercritical Extractbn of Coal”, paper presented at the 89th AIChE Na- tlonal Meeting, New Orleans, LA, Nov 1981.

Gangdl. N.; Thodos, G. Ind. Eng. Chem. Prcd. Res. D e v . 1977, 16, 208. Giiddings, J. C., et el. Sc/enca 1968, 762, 87. Jezko, J.; Gray, D.; Kershaw, J. R. Fuel Process. Techno/. 1982, 5 , 229. Modell, M.; ReM, R. C.; Amin, S. I . U.S. Patent 4113446, 1978. Paul, P. F. M.; Wise, W. S. “The Principles of Gas Extraction”: Mllis and Boon

Ross, D. S.; Blessing, J. E. Fuel 1979, 58, 433. Rowllnson, J. S.; Richardson, M. J. A&. Chem. Fhys. 1959, 2 , 85. Whitehead, J. C. “Development of a Process for the Supercritical Gas Ex-

traction of Coel”, paper presented at the 88th AIChE National Meeting, Philadelphia, June 1980.

Received for review June 27, 1983 Revised manuscript received October 17, 1983

Accepted February 23, 1984

Ltd.: London, 1971.

Optimization of a DPstiltation Column with a Direct Vapor Recampremion Heat Pump

Josep A. Ferr6, France= Castells,’ and Joaquh Flores

Departament de O&nica T6cnIca. facultat de Ouimica, Unlversitat de Barcebna, Tatragona, Catalunya. Spain

Heat pumps can afford important savings of energy In distillation processes. However, the lack of knowledge of the profttabllity of their applicatlon in every case restricts thek use as an alternative for saving energy in distillation cdumns. In this paper we present the results of a simulation and optknizatlon computer program we developed to analyze the economical profitability of substhit@ the conventknal reboiler and condenser of an existing distillation column by a vapor recompression heat pump. Two examples of calculatlons are presented which correspond to the separation of organic compounds of mediwn molecular welght of close boiling points. Results obtained show that the POT for the substltution of the conventional distillation scheme by an overhead vapor recompression heat pump is about two years for these kinds of separations.

Introduction Distillation is one of the most popular methods of sep-

aration in the chemical and petrochemical industries. However, this process consumes a great deal of energy

*Department de Qdmica TBcnica, Facultat de Qui’mica, P h p

which is later degradated to the surroundings as con- densing heat. One of the most efficient ways of saving energy is through the we of heat P ~ P S in the distillation process. This technique is already used in Some of the more difficult separations: isobutaneln-butane (Barnwell and Morris, 1982), propane/propylene (Quadri, 1981a,b), ethane/ethylene (Menzies and Johnson, 1971), owing to the specific advantages involved (Finelt, 1979). Apart from Imperial Tarraco s/n, 43005 Tarragona, Catalunya, Spain.

0196-4305/85/ 1124-0128$01.50/0 0 1984 American Chemical Society

Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 1, 1985 129

C O N'v ENTIONAL HEAT PUMP APPLlCATiOE DlSTILLAiION external heat pump [ bottoms flash

I I

Figure 1. Heat pump applications.

tooting water E

apor recompression

~

feed J" 1 steam

E-2 TL

lL$ w- bottom *

Figure 2. Flow sheet of the vapor recompression heat pump.

these typical cases, its application is still very limited. At present, we can observe that heat pumps can afford

positive savings of energy, but there is a general inertia of the engineering and production departments to use heat pumps as an alternative production method, mainly due to the lack of knowledge of their profitability in every case.

Figure 1 shows three different ways of applying heat pumps in distillation processes. The most advantageous is the direct overhead vapor recompression (Danziger, 1979) (Null, 1976). In this work a simulation and opti- mization computer program for the heat-pump system is developed and applied to an existing distillation column whose energy and economic balance is to be improved. In brief, the following questions should be answered: What energy savings can be obtained in an existing distillation column by substituting the conventional reboiler and condenser by an overhead vapor recompression heat pump? What is the maximum profitability of this sub- stitution? Flow Sheet of the Vapor Recompression System

The two examples of the vapor recompression we are presenting correspond to the separation of binary and pseudobinary close-boiling organic mixtures of medium molecular weights. This type of separation requires a large number of equilibrium stages as well as high reflux ratios and, consequently, high energy consumption.

Theee components have characteristics which differ from other light components to which heat pumps can also be applied: the adiabatic compression of a saturated vapor does not produce a superheated vapor as happens, for instance, with propylene, but it ahows ita condensation due to the fact that adiabatic lines penetrate into the vapor- liquid equilibrium zone in a temperatureentropy diagram.

Table I. Overall Heat Transfer Coefficients for Heat Exchange Equipment

overall H.T. location H.E. equipment coeff, W/m2 K

E-1 preheater 150 E-2 reboiler-condenser

desuperheating zone 200 condensing zone 600 subcooling zone 430

E-5 cooler 350 E-6 secondary reboiler 1100

Table 11. Cost of Utilities steam 0.0095 $/kg electrical energy 0.06 $/kW h cooling water 0.05 $/m3

This effect is caused by the gradual deformation of the vapor-liquid equilibrium curve when the molecular weight is increased in the same homologous series. It is typical for these compounds to superheat their saturated vapors above their dew point before being compressed in order to avoid condensation.

Figure 2 shows the flow sheet of the above-mentioned vapor recompression system. Overhead vapors of the column are superheated above their dew point in the preheater E-1, countercurrent with the condensate a t the outlet of E-2. Later on, the vapors are compressed in K-1 up to the necessary pressure for their condensation into the reboiler-condenser E-2 giving the latent heat. After passing through E-1, the condensate is cooled in E-5 with cooling water up to its bubble point a t the column oper- ating pressure. We might need an additional reboiler to start up the distillation, owing the thermal state of the feed (saturated liquid), and to complete the energy balance and control the internal vapor rate at the bottom of the column. Once the overhead vapors condensing in E-2 are fixed, it might be necessary to make this control by an additional heat input through reboiler E-6. Simulation and Optimization

The simulation computer program of the vapor recom- pression system is carried out keeping the operating con- ditions of the column fixed. After estimating the pressure drop in E-1, and neglecting other pressure losses in pipes and equipment, the material and energy balances can be completed taking into account the three independent variables: T2 = superheated vapor temperature, P3 = compressor outlet pressure, and MTD, = mean temper- ature difference, calculated as a weighted temperature for the desuperheating, condensing, and subcooling zones in the reboiler-condenser E-2. These variables correspond to three degrees of freedom of the system.

The next step is to calculate the exchanger areas, using the overall heat transfer coefficients given in Table I. The installed cost of the main equipment has been calculated according to the data of the I.F.P. (Institut Franqais du Petrole) (Chauvel et al., 1981) and other values published for exchangers (Null, 1976). The cost of the compressor has been calculated considering the possibility of using special turboblowers, which operate up to pressure ratios of 3 and are cheaper than conventional compressors (Danziger 1979). The cost of equipment and installation, as well as the energy prices (Table II), has been updated for Western Europe in 1982.

For every set of values of the independent variables, different operating conditions for the heat pump will be obtained. The Complex algorithm (Kuester and Mize, 1973) has been used for the optimization of the operating conditions that give maximum energy savings and mini-

130 Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 1, 1985

b -

5 -

#

2 = 3 -

2 -

1 -

Table 111. Two Exampleso of Operating Conditions for the Conventional Design without Heat Pump

example example code name units 1 2

feed, saturated liquid kg/h 15000 25000 overhead vapors kg/h 190284 110000 distillate kg/h 1884 11747 bottom kg/h 13116 13253 reflux ratio 100 8 distillate compn % EB 99.5 98.0

T I top temperature "C 136 98 P, top pressure atm 1.0 0.34 r b bottom temperature "C 157 115

temp difference, top - "C 21 17

reboiler duty kW 17734 11029 condenser duty kW 17734 10975 normal bp, light key compn "C 136.2 136.2 normal bp, heavy key "C 140.7 145.2

bottom

compn

Example 1: separation of ethylbenzene (EB)/xylenes blend (XYL); example 2: separation of ethylbenzene (EB)/styrene (ST- Y).

Table IV. Operating Conditions at the Optimum with the Vapor Recompression Heat Pump

code or example example location name units la 2*

T , top temperaturec "C 136.0 98.0 P, top pressurec atm 1.0 0.34 E-1 preheater duty kW 2106 303 T, inlet compressor temp "C 159.1 104.3 P2 inlet compressor press. atm 0.9 0.33 T3 outlet compressor temp "C 165.0 124.4 P3 outlet compressor press. atm 2.04 0.72 K-1 pressure ratio 2.27 2.18 K-1 polytropic efficiency % 75.0 75.0 K-1 compression power kW 1971 964 E-2 dew point at P3 "C 165.0 124.4 E-2 bubble point at P3 "C 165.0 124.4 T4 E-2 outlet temperature "C 165.0 117.2 T b bottom temperature' "C 157.0 115.0 E-2 weighted temp "C 8.0 9.1

difference E-2 reboiler/condenser duty kW 16968 11029 T, inlet cooler temperature "C 148.8 112.7 T6 outlet cooler "C 136.0 98.0

E-5 cooler duty kW 2737 910

E-6 reboiler duty kW 766 0.0 E-6 steam consumption kg/h 1390 0.0

temperature

E-5 cooling water m3/h 196.0 65.0 consumption

Example 1: separation of ethylbenzene (EB)/xylenes blend (XYL). Example 2: separation of ethylbenzene (EB)/styrene (STY). Fixed operating conditions.

mum additional investment required for the substitution. The objective function is defined as

saving (M$/year) investment (M$) simple rate of return (SRR) = x 100

This corresponds to the inverse of the pay out time (POT) by the equation

100 POT = - SRR Presentation of the Results

The operating conditions for the two studied examples are listed in Tables I11 and IV for the conventional design and that with vapor recompression, respectively. The

- 10 LT L2I m - Z LT 3 35 + W [1I

$ 30

PRESSURE RATIO 2 3 2 2 2 1 2 6 2 8 3 0 3 2 3 L

1 \ I I 6

5

L # 5

- 3

TEMPERATURE DIFFERENCE - 'C Figure 3. Behavior of the objective function at the optimum for the (EB/XYL) example.

PRESSURE RATIO 2 0 2 2 2 L 2 6 2.8 3 0 3.2 3 1

\ ,m \ total investment

\ \

TEMPERATURE DIFFERENCE - 'C Figure 4. Equipment cost and total investment for the (EB/XYL) example.

operating conditions for the first example (ethylbenzene- /xylenes blend) a t the optimum is found to have neither desuperheating nor subcooling in E-2. At the same time, the preheating in E-1 is just enough to prevent conden- sation during the compression process. Thus we can re- duce the three independent variables to one: the pressure ratio, or its equivalent, the mean temperature difference in the reboiler-condenser E-2. In this case the behavior of the objective function around the optimum, shown in Figure 3, indicates that the total investment represents a relative stable value when the temperature difference is in the range of 7-30 "C and that the annual saving di- minishes when the pressure ratio or the temperature dif- ference is increased. In t h i s case, ,the maximum SRR (i.e., the optimum) appears a t relatively low temperature dif- ferences where the total investment is almost stable.

Figure 4 shows the equipment cost and total investment for the (EB/XYL) example. Figure 5 shows the annual saving, the electrical consumption of the compressor, and the savings in steam and cooling water. The savings in steam and cooling water are maintained above 90% and 75 % , respectively, but when the pressure ratio is increased,

Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 1, 1985 131

w a t e i saving - -. - - - - - - - - - _ _ -

Table V. Energy and Economic Balance for tne Two Studied Examples Example 1 : Separation of Ethylbenzene/Xylenes Blend

100 75 50

energy consumption ~~

without heat pump with heat pump flow rate M$/war % flow rate M $ /Y ear %

steam 32130 kg/h 2442 100 1390 kg/h 106 4 cooling water 1272 m3/h 509 100 196 m3/h 78 15 electricity 2318 kW 1113 100 total 2951 100 1297 44

net saving total investment simple rate of return (SRR) pay out time (POT)

1654 M$/year 3865 M$ 43% 2.3 years

Example 2 : Separation of Ethylbemenelstyrene energy consumption

without heat pump with heat pump

flow rate M$/vear % flow rate MBlsear %

steam 19980 kg/h 1518 cooling water 787 m3/h 315 electricity total 1833 net saving total investment simple rate of return (SRR) pay out time (POT)

PRESSURE RATIO , 2; 2.; 2.; 2.; 2 8 3.; 3; 3: , 100 100 65 m3/h 26 8

1134 kW 544 100 100 570 31

1263 M$/year 2194 M$ 58% 1.7 years

sized more compression power is needed as a result of the unnecessary increase of the inlet temperature of the com- pressor.

The preheater duty per unit flow rate for the first ex- ample is found to be much bigger than that for the second example. This difference is caused by the higher operating pressure in the first case, which produces a greater need for preheating in order to avoid condensation. Thus, the operating conditions at the optimum in example 1 indicate that the condensate does not subcool in the reboiler-con- denser E-2, and the heat balance at the bottom of the column must be closed by an additional steam consump- tion in E-6, which is not the case for example 2. If in the first case the condensate is subcooled in E-2, the inlet temperature to the preheater will be lowered and the area of E-1 will also be very big, due to the greater need of preheating.

Therefore, in example 2, where the need of preheating is lower than that in example 1, the operating conditions at the optimum indicate that it is more favorable, as well as possible, to eliminate steam consumption in the addi- tional reboiler E-6 by subcooling the condensate and even increasing the preheating area in E-1, because energy saving is greater than the increase in the preheater cost, which is the opposite of what happens in the first example.

For the first example a modified set of operating con- ditions can be checked. Upon withdrawing the distillate just before the throttle valve and allowing a flash in the reflux to increase the condensing flow rate in E-2 and to eliminate the vapor consumption in the additional reboiler E-6, the SRR increases only from 43% to 45%.

In order to observe the variation of the optimum and of the objective function we have reduced the column capacity in the first example to 60% of the original value. The results obtained indicate that the maximum SRR is reached for a mean temperature difference in the reboil- er-condenser of 7.2 OC, in comparison with the original calculated value of 8.0 "C. This shows the stability of the parameters which determine the optimum position vs. the

132 Ind. Eng. Chem. Process Des. Dev. 1985, 24, 132-140

variation of the column capacity. Nevertheless the SRR diminishes from 43% to 37%.

If these results are compared with those obtained from the ethylbenzene/styrene separation (example 2) in which the column rates are 60% of example 1, we come to the following conclusions.

1. The parameters having the most influence on the determination of the profitability of the vapor recom- pression heat pump applied to an existing distillation column are: (i) the temperature difference between top and bottom, (ii) the column capacity, and (iii) the need of preheating.

2. The smaller the temperature difference between top and bottom is, the more profitable the heat pump appli- cation will be, for the following reasons: (i) small tem- perature differences indicate difficult separations with high reflux ratios, and consequently, high steam and cooling water consumptions, and (ii) the pressure ratio and the compression power which are required increase with the increase of temperature difference.

3. The larger the column is, the more favorable will be the cost reduction of the main equipment referred to the proportional increase in energy saving.

4. The smaller the preheating is, the more interesting the heat pump application will be, because: (i) there is a reduction in the size and cost of E-1, (ii) the compression power decreases as the inlet temperature of the compressor decreases, and (iii) the steam consumption in E 6 decreases if the condensate is allowed to give more heat in the re- boiler-condenser E-2 by subcooling. Concluding Remarks

Calculation programs of simulation and optimization have been developed which allow us to analyze the eco-

nomical viability of substituting the conventional reboiler and condenser of a distillation column by a vapor re- compresion heat pump. For the two cases presented the POT is lower than 2.5 years.

The comparison among the above presented cases and other similar ones which have been studied by us (Flores, 1983), indicates that the parameters showing the position of the optimum for the same system are relatively stable in front of the variation of the column capacity, the ratio of energy prices/equipment costs, and the relation between cooling water and electricity cost. However, the value of the POT in the optimum is very sensitive to the variation of these factors, and therefore from this point of view it is necessary to take into the account the possible re-using of the conventional equipment to be substituted, with which the required additional investment would be re- duced. Literature Cited Barnwell. J.; Morris, C. P. Hy&ocerbon Process. 1982, 67(7), 117-199. Chauvel, A.; Leprince, P.; Barthel, Y.; Raknbault, C.; Arb, J. P. “Manual of

New York,

Danziger, R . Chem. Eng. Prog. 1979, 75(9), 58-64. Flnelt, S. Hydrocarbon Process. 1979, 58(2), 95-98. Flores, J. Hydrocarbon Process. 1984, 63(7), 59-62. Kuester, J. L.; Mlre, J. H. “Optimlzatlon Techniques wlth Fortran”; McGraw-

Menzies, M. A.; Johnson, A. I . Can. J. Chem. Eng. 1971, 49, 407-411. Null, H. R . Chem. fng. frog. 1978, 72(7), 56-64. Quedrl, G. P. Hydrocarbon Process. 198la, 60(2), 119-126. Quadri, G. P. Hydrocarbon Process. 1981b, 60(3). 147-151.

Economic Analysis of Chemical Processes”; McGraw-Hill: 1961.

Hill: New York, 1973.

Received for review July 6, 1983 Accepted March 26, 1984

This work was financially supported by Dow Chemical Iberica, S.A.

Design and Control of a Two-Column Azeotropic Distillation System

Samlr I . Abu-Elshah and Wllllam L. Luyben’

Department of Chemical Engineering, Lehlgh University, Bethlehem, Pennsylvania 180 15

The steady-state design of a two-column azeotropic distillation system operating at two different pressures was studied with the objective of reduclng energy consumption. The minimum-boiling, homogeneous binary azeotropic system tetrahydrofuran-water was used as a specific example. Energy consumption was reduced by a factor of 2 from conventional designs by using heat integration and feed preheat and by optimizing column pressures and overhead purities. The dynamics and control of the two-column, heat-integrated system was also explored. A control system was developed that effectively handled a variety of disturbances without any severe interactions between columns.

Introduction The design of azeotropic distillation systems has been

the subject of many papers and books (Hoffman, 1964). The use of two columns, operating a t two different pres- sures, is one of the simplest and most economical tech- niques for separating binary azeotropes (Van Winkle, 1967), provided that a substantial shift in the composition of the azeotrope occurs when pressure is changed.

For a minimum-boiling homogeneous binary azeotrope the distillate from the low-pressure column is fed to the high-pressure column (see Figure 1). The distillate from the high-pressure column is recycled back to the low-

pressure column. These distillates have compositions that are close to the azeotropic compositions at their respective pressures.

More or less pure products are removed from the bottom of each column. In the tetrahydrofuran (THF)-water system, water is produced from the base of the low-pres- sure column. THF, the low-boiling component, is pro- duced from the bottom of the high-pressure column.

While the steady-state design of these systems has been discussed, at least qualitatively, in the literature, there has been very little reported concerning the dynamics and control of the two-column system. Shinskey (1977) gave

0196-4305/85/1124-0132$01.50/0 0 1984 Amerlcan Chemical Society