1220
Ind. Eng. Chem. Process Des. Dev. 1905, 2 4 , 1220-1225
Mixer-Settter-Extraction Column: Mass-Transfer Efficiency and Entrainment Mllos Horvath and Stanley Hartland" Swiss Federal Institute of Technology, Department of Industrial and Engineering Chemistty, Universitatsstrasse 6. CH-8092Zurich
-
A new design of a mixer-settler column with mean stage efficiencies up to 100% for solute (s) transfer from the d) and up to 170% for the opposite direction is presented. The performance continuous to the disperse phase (c of a 152-mm inside diameter column with five stages was investigated by using the test system toluene(c)acetone(s)-water(d) at 20 ' C . In the rotor speed range of 200-400 rpm, the entrainment of the dispersed phase was extremely low. No interstage mixing of the dispersed phase was detectable. The column is capable of handling most liquid-liquid systems, being especially suitable for extractions where many theoretical stages are needed or long residence times per stage are required, e.g., when slow chemical reaction is involved, as in the extraction of metals.
Power input generally improves the efficiency of extractors since it increases the mass-transfer rate by providing (a) a large interfacial area between the two immissible liquids and (b) new surfaces and high turbulence which increases the mass-transfer coefficients. The power is usually supplied by pulsation or stirring. A variety of stirred extractors, as summarized by Hanson (1968) and Lo et al. (1983), is known. These may be mixer-settlers or stirred columns, e.g., Scheibel, MIXCO, Kuhni, RDC, ARD, EC, and QVF which utilize gravity for transport of the phases. When the ground area is expensive, columns are preferable to mixer-settlers. Stirred columns are divided into stages by perforated plates or baffles in order to reduce the axial mixing caused by agitation. Through the free area, there is inevitably some undesirable axial mixing which reduces the efficiency of the column. Reduction or even elimination of the axial mixing is the key to higher column efficiency and is therefore of interest in column design. In addition to the physical mass transfer, chemical reaction may be involved in some systems. If the physical transfer is fast enough, the chemical reaction becomes the rate-controlling step and the residence times of the phases per stage become important. Mixer-settlers are usually applied if high residence times and intensive mixing are required. They do however occupy a disproportionally large ground area and require a large solvent inventory. In addition, some mixer-settlers must be equipped with a complicated pipe system with pumps to transport the liquids from stage to stage. On the other hand, in a compact type of mixer-settler, the interface-level control becomes complicated. A column extractor having the advantages of the mixersettler with high stage efficiency and low interstage mixing, yet requiring a small ground area and no additional pipes and pumps, has obvious advantages. The Lurgi-Tower, Wirz, Treybal, and Kuhni MS columns approximate to these requirements, but their design is rather complicated. Mixer-Settler Column The proposed mixer-settler-extraction column (MIXET) combines in a relatively simple way the advantages of columns (Table I) and mixer-settlers. The MIXET column is built vertically like other columns, but its high stage efficiency is comparable with that obtained in mixer-settlers. This is due to repeated complete coalescence 0196-430518511124-1220$01.50/0
and redispersion in each stage and to the large interfacial area produced by intensive mixing combined with low back-mixing. No additional pipes and pumps are needed since both phases flow countercurrently by the action of gravity alone. The MIXET column can contain any number of stages placed vertically one above the other. A typical stage, as shown in Figure 1,contains one mixer and a settler. The dispersed phase, after coalescing in the settler, flows into the mixer of the next stage over a weir (1)in the center of the column into the eye of a stirrer(2) where it is dispersed. All stirrers are attached to one common shaft (3) along the axis of the column. The light, continuous phase enters the mixer from a settler below, through a perforated plate (4),separating the mixer and settler of the same stage. Both phases are intensively mixed inside the mixer. The drops of the dispersed, heavy phase fall through the perforated plate from the mixer into the settler and coalesce completely before entering the mixer of the next stage via the weir. The light, continuous phase rises from the mixer along the shaft through a cylinder (5) in the center of the column into the settler of the next stage. This cylinder forms a hydraulic seal, preventing backflow of both the dispersed and the continuous phases, forcing the dispersed phase to remain in the settler until complete coalescence is achieved. After coalescing, the heavy phase flows via the weir into the next stage mixer to be redispersed. If the light phase, rather than the heavy phase, is dispersed, the mixer-settler shown in Figure 1 should be inverted. Coalescence aids could be used in the settlers in order to enhance coalescence of the dispersed phase. Alternatively, perforated plates, separating the mixers and settlers, could be replaced by EC plates similar to those introduced by Steiner and Hartland (1980). Experimental Section (a) Mass-Transfer Experiments. A mixer-settler column of 0.152-m inside diameter and 0.150-m stage height with five stages was tested at 20 "C for separation performance using the system toluene(c)-acetone(s)water(d). Each stage contained 70-mm high-mixer and 75-mm high-settler compartments. The plate thickness was 5 mm, giving a total stage height of 150 mm. The stirrers used were standard 6-flat-blade turbines, 53 mm in diameter, placed in the middle of the mixer which was equipped in addition with four 15-mm-widebaffles. The 0 1985 American Chemical Society
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985
1221
Table I. Comparison of Performance of Stirred Columns‘“ ~~~
~
column tvDe
~
diam. m
stage ht. m
stage effic. %
ARD
0.150
0.037
6.3
SCH 1
0.289
0.305
SCH 2
0.289
MIXCO
NTS/m, throughput l/m m3/(m2/hl 1.7
8-22
45/93
1.5-3.1
2-22
0.102
40/100
3.9-9.8
7-19
0.152
0.076
55/75
7.2-9.8
12-22
Kuhni
0.150
0.070
56
5-8
4-24
QVF
0.150
0.075
60
5-8
9-15
EC
0.072
0.060
45
3-7.5
Wirz
0.500
0.100
86
8.6
KMS
0.060
system used and mass-transfer direction
data source Marek, 1970
o-xylene-acetone-water (c) (dl MIBK-acetic acid-water (c)(d) (d)(c) MIBK -ace tic acid-w ater (d) + (c) o-xylene-acetone-water (d) (c) toluene-acetone-water (d) (c) toluene-acetone-water (d) (“1. toluene-propionic acid-water +
+ -
+
10-60
+
8.8
--f
Scheibel and Karr, 1 9 5 0 Scheibel, 1956
Oldshue and Rushton 1 9 5 2 Kumar, 1985
Zuhlke, 1978
v. Fischer, 1983
Wirz, 1967
toluene-ace tone-water Buhlmann, 1981 (d) (c) PC 0.07 toluene-ace tone-water 1.8-2.5 15-30 Stichlmair , 1980 (d) (c) PPC 0.07 3.8-5.8 18-20 toluene-ace tone-water Stichlmair, 1980 (d) (c) PSE 0.05 3.5-6.0 20-30 0.10 60 toluene-acetone-water Stichlmair, 1980 (d) (c) KARR 0.05 3.5-6.0 30-40 0.025 15 toluene-acetone-water S tichlmair, 1980 (d) (c) 2.8-3.5 15-35 RDC 0.07 toluene-acetone-water Stichlmair , 1980 (d) (c) MIXET 0.152 0.150 97 6.5 2-6 toluene-acetone-water present work (c) (d) MIXET 11.3 2 -4 0.152 0.150 170 toluene-ace tone-water present work (c) (d) a Abbreviations of column types: ARD, axisymmetric rotating disk; SCH, Scheibel; MIXCO, mixed column; EC, enhanced coalescence; KMS, Kuhni mixer-settler; PC, packed column; PPC, pulsed packed column; PSE, pulsed sieve plate extractor; RDC, rotating disk column; MIXET, mixer-settler-extraction column. (c) = continuous, (d) = disperse. 0.350
100/130
2.9-3.7
4-10
+
-
+ -
+
+
+
+
+
Figure 1. Design of the MIXET column.
column and pipework were built of industrial glass, the stirrers of stainless steel, and the plates of Teflon. Deionized water was the dispersed phase, toluene of technical grade the continuous phase, and acetone of technical grade purity the solute. Both mass-transfer directions were investigated (from the continuous to the dispersed phase and in the reverse direction). The volumetric flow rate ratio remained the same in most runs: ( V J V , = 1.5), to keep the operating line approximately
parallel to the slightly curved equilibrium line. Total throughputs of 2, 4,and 6 m3/(m2h) were used. Typical Run. The feed solutions of either the water or the toluene phase were made up beforehand. The well-mixed feed streams, containing less than 6% by weight acetone, were fed into the column at given flow rates and stirrer speed. All adjustments to the operating conditions such as flow rates, interface level, and stirrer speed were made during the early stages of the run. The acetone concentrations in both streams leaving the column were monitored continuously during the run by means of a density meter in the water phase and of a refractometer in the toluene phase. When no changes of density and refractive index were noticeable, steady state was assumed and samples of inlet and outlet streams of both phases were taken in all runs. In some runs, additional samples of both phases were withdrawn from some or all stages along the column. Clear samples were sucked from the mixer and settler through porous caps into glass syringes which were plugged immediately after having been filled. The samples were immediately analyzed on a gas chromatograph. From the known inlet and outlet concentrations, the number of theoretical stages (NTS) was calculated for the separation obtained; dividing by the actual number of stages (in this case five) gave the mean stage efficiency E. Its variation with stirrer speed for different throughputs is shown in Figure 2 and the variation in NTS per unit length of active column with total throughput in Figure 3, other types of stirred columns also being shown for comparison. In some earlier runs, acetone was injected into the toluene feed stream which enabled experiments to be per-
1222
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985
.
r o t o r speed
(rpm)
Figure 2. Variation of stage efficiency with rotor speed in MIXET column: (0) 2, (*) 4, (#) 6 m3/(m2h).
t
MlXt I ‘\
5t i
t
ARD
01 1
i1
\
I
2
4
Total throughput
6
per unit
IO
20
40
60
(00
column-cross-section (m3/m2h)
Figure 3. Comparison of performance of stirred columns.
formed with a much smaller toluene inventory. Pure solvents were then circulated through the column at preset flow rates and stirrer speed until the temperature of both phases stabilized at 20 f 0.5 “C. The flow rates were easy to control manually and the chosen feed temperatures were attained by cooling the inlet streams. For the masstransfer direction c d, the acetone pump was then started and solute was injected into the toluene feed from a separate tank via rotameter. Immediately after solute injection was started, the water outlet was switched to flow to a separate tank. The toluene was recirculated during the run until the same amount of acetone was extracted in the column by the water, as injected by the pump. The total amount of toluene in the circuit was kept as low as possible to ensure that the steady state was reached as soon as possible. (b) Back-mixing Experiments. Stator plates are usually inserted in stirred columns to reduce axial mixing caused by the agitation. However, some axial mixing still occurs, the actual back-flow often being bigger than the net forward flow. In the present mixer-settler column, the back-flows of the phases are virtually eliminated since the settlers contain hydraulic seals which prevent the contin-
-
uous and dispersed phase from flowing backwards. The dispersed phase can reach an upstream stage by entrainment, which was measured by the steady-state method. The tracer used was aqueous KC1 continuously injected into the mixer of stage 3, where it was immediately dispersed. The tracer concentrations were monitored in the settlers above and below the injection point. Typical Run. Solvents without solute were recirculated at the preset flow rates and a given stirrer speed until a constant temperature of 20 f 0.5 OC inside the column was reached. The temperatures of the feed streams were manually adjusted for this purpose. Then, by the use of a peristaltic pump, a tracer of 1kmol/m3 KC1 in water was steadily injected into the mixer of stage 3. A sample of the dispersed aqueous phase after coalescence was pumped from the settler of stage 2 through a conductivity cell at the same volumetric rate as the tracer solution was pumped in. Due to this arrangement, no change in the volumetric flow of the dispersed phase along the column was created. During the first three runs, the tracer concentrations inside the mixer 3 were compared with those in settler 3 and were found to be the same. In all following runs, the concentration in settler 3 was measured and used for calculation, sampling of the clear dispersed phase from settlers being more convenient than from mixers. After a certain time, the conductivities did not change. The steady state was then assumed to have been reached and the conductivities were noted. For the next run, with a new rotor speed, the column was given time to reach the steady state. The rotor speed was measured by means of a digital tachometer. The organic continuous phase was recirculated, and the coalesced aqueous (dispersed) phase, containing KC1, was discarded.
Results (a) Mass Transfer. The performance of a separation device with well-defined stages can be expressed in terms of stage efficiency:
E = -NTS x 100 N The number of theoretical stages (NTS) needed for a given separation was calculated numerically from the inlet and outlet solute concentrations on a desk computer. The slightly curved equilibrium line was approximated by a third-order polynomial based on experimental data of Hackl et al. (1978). The stage efficiency was then obtained from the NTS needed for the separation attained in a given run, divided by the actual number of stages (N), in this case 5. For the investigated mixer-settler column, the stage efficiency was strongly dependent on the stirrer speed (see Figure 2). A t low speeds, the efficiency was as low as 50-60% due to inadequate mixing, resulting in large drops and a small interfacial area. The stage efficiency increased with increasing rotor speed and reached a maximum of close to 100% for the mass-transfer direction c d at speeds between 325 and 375 rpm (see lower part of Figure 2). After having reached its maximum, the stage efficiency fell with further increase in stirrer speed. The fall in stage efficiency was probably a result of a substantial decrease of the mass-transfer coefficients for small drops, behaving as rigid spheres, present in the mixer. There was found to be a remarkable difference in stage efficiency with mass-transfer direction. Whilst for the solute transferred from the continuous (toluene) into the dispersed (water) phase (c d), the stage efficiencies reached maximum values of about 100% at approximately 350 rpm (see Figure 2); the mean stage efficiencies for the opposite mass-transfer direction (d c) reached values -+
-
-
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985 1223
well above 100% (in fact up to 170%) at the same stirrer speed. Stage efficiencies of over 100% are possible as reported earlier (Scheibel). In the MIXET column, equilibrium (or close to equilibrium) concentrations were attained in the mixer of each stage and additional mass transfer took place in the settler. This fact was proved by analyzing samples of both phases taken from the mixer and the settler. The solute concentrations inside a mixer were very close to equilibrium, and there was still a substantial difference in solute concentration of the dispersed phase between that measured inside a mixer and that in a settler. The solute concentration in the heavy phase leaving the settler of a given stage was greater than the equilibrium solute concentration in the light phase leaving the corresponding mixer. This was due to additional mass transfer taking place during the fall and coalescence of the drops inside the settler, which resulted in stage efficiencies of over 100%. The high mass-transfer rates which allow the phases to reach equilibrium inside the mixer within relatively short times result from the new surface created by enhanced coalescence and redispersion of the dispersed phase. Coalescence of the drops falling in the settler is also enhanced, creating large drops. The mass transfer is correspondingly increased, being greater than in the case of the opposite mass-transfer direction (c d). The enhanced coalescence due to mass transfer was reported by Groothius and Zuiderweg (1960) and has been observed in other types of mass-transfer equipment (Steiner et al., 1978). The throughput capacity of the column was however reduced for the mass-transfer direction d c. Maximum throughputs of 4 m3/(m2h) were achieved compared with 6 m3/(m2h) for c d for the present column design. The mean stage efficiency observed in runs with higher solute concentrations was in general slightly higher, other parameters (like rpm, throughput, and mass-transfer direction) being the same, than in runs with lower solute concentrations. Higher solute concentration infers a lower interfacial tension, so for the same power input, a larger interfacial area per unit volume is created, resulting in higher mass-transfer rates. For this reason, it is essential for scale-up purposes that the range of solute concentration be kept similar. There was no dependence of stage efficiency on throughput found for the mass-transfer direction c d. Within the range of throughputs investigated (2-6 m3/(m2 h) total flow of both phases, corresponding to mean residence times of 2.14.7 min per mixer), the residence times were sufficiently large not to affect the mass transfer achieved. For all the investigated total throughputs ( V , + VJ, the stage efficiency was the same for the same rotor speed within the experimental error. For the opposite c, there is a noticeable demass-transfer direction d pendence of stage efficiency on throughput. For lower throughputs (higher residence times), higher efficiencies were observed. A maximum stage efficiency of 170% was measured for a mean residence time of 2.1 min, falling to 150% at a mean residence time of 1.05 min per mixer. A comparison of the separation performance of the present mixer-settler column with mechanically agitated columns of existing design was made. The criterion was, as suggested by Stichlmair (1980), the number of theoretical stages per unit length of the active part of the column (NTS/m) vs. to+ throughput per unit area of the column cross section ( ( V ,+ V,)/Q). The present design of the mixersettler column reached values of between 6.5 and 11NTS/m at low to medium loads (2 to 6 m3/(m2h)). (See Figure 3.) At higher throughputs, the column flooded.
-
-
-
-
-
I
II Injection
MIXER
1.
cM(n)
Figure 4. Tracer balance for steady injection in MIXET column.
The smaller values (-6.5) were obtained by using the optimal stage efficiency for the mass-transfer direction c d (approximately 97%); the higher values (- 11)were calculated by using the stage efficiency of 150% measured for mass-transfer direction d c. After a certain time of column operation, the originally clear interface of each settler became pearly, indicating that impurities had gathered at the interfaces. The interface together with the adjacent liquids was then sucked out. After separation in a funnel, the toluene phase was filtered and returned to the tank. The water was drained. After porous caps were installed into the air-vent system, the interfaces stayed clean for a longer period of time, indicating a likely source of impurities to be dust in the air. The other possible source of impurities was the deionized water which was renewed after each run. Toluene was recirculated in a closed system. Surface active impurities were most likely to be responsible for the relatively large scatter of results, since they affect the mass-transfer rates in an uncontrollable way. The mixer-settler column used was designed for large residence times (1-2 min per stage), with corresponding low throughputs. Once an appropriate redesign of the plates and hydraulic seals is made, the column should be able to handle loads of over 10 m3/(m2h) at the same efficiency. There is also no reason to believe that the efficiency would be reduced for even smaller throughputs than those used here (2 m3/(m2h)). Since the residence times of the phases for the given system toluene-acetone-water, within the range of variables investigated, do not seem to be important, the overall efficiency of the column can be improved by reducing the height of each stage. More stages of the same efficiency could then be accommodated within the same length of the column, improving the overall separation performance. (b) Back-mixing. When tracer solution was injected into a mixer of a given stage, no tracer could be detected upstream other than in the adjacent settler. The backmixing coefficient f was therefore derived from a balance on the settler (n - l ) , shown in Figure 4,as
-
-
1224
Ind. Eng. Chem. Process Des. Dev., Vol. 24, No. 4, 1985
higher continuous-phase velocities in the cylinder, bigger drops are carried upstream, the total amount of tracer carried backwards being larger. The sharp increase of 008 L entrainment at 6 m3/(m2h) over 400 rpm indicates that flooding is approached. At this throughput, the mean i linear velocity of the continuous phase in the hydraulic-seal cylinder is about 1 cm/s which should not be exceeded if stable operation and low entrainment are to be achieved. This must be considered when designing the plates and hydraulic-seal geometries at higher throughputs. Conclusion The present mixer-settler column design proved suc002 4 4 cessful. When using the toluene (c)-acetone (s)-water (d) system at 20 "C,mean stage efficiencies close to 100% for the mass-transfer direction c d and up to 170% for the opposite direction were obtained in a 152 mm inside diameter MIXET column. These stage efficiencies correrotor speed (rpm) spond to 6.5 and 11.3 theoretical stages per meter at throughputs between 2 and 6 m3/(m2h). Figure 5. Variation of dispersed-phase entrainment with rotor speed in MIXET column: (0) mean value, ( X ) maximum value at The stage efficiency depends mainly on stirrer speed, 4 m3/(m2h), (*) mean value at 6 m3/(m2h) total throughput. reaching a maximum value at about 350 rpm for both mass-transfer directions. where n is counted from the dispersed-phase inlet. The amount of entrainment of the dispersed phase, The conductivity of the inlet water stream was finite, represented by the back-mixing coefficientf , proved to be so an equivalent tracer concentration cinwas included. The extremely low. Less than 1% of the dispersed-phase tracer concentration below the injection point is constant, volumetric flow rate flows backwards from a mixer into since there are no other sources or sinks. The concenthe adjacent settler. tration in the mixer (n)could therefore be replaced by that No tracer could be detected in the adjacent upstream in the settler ( n ) ,so c M ( ~ )= c S ( ~ ) . As mentioned in the mixer, indicating that the interstage back-mixing of the Experimental Section, sampling from settlers is easier than dispersed phase was zero. from mixers. Because of the absence of back-mixing in the present The variation of the back-mixing coefficientf with rotor mixer-settler column, the stage efficiency should be inspeed for different total throughputs is shown in Figure dependent of column diameter, providing that the power 5. It must be stressed that these coefficiencts represent input per unit volume is kept constant, so no problems the amount of entrainment of the dispersed phase between with scale-up are anticipated. The solute concentration a mixer and the adjacent upstream settler. Circles (0) range used in both sizes must, however, be the same. represent the mean values of the coefficient f and crosses The compact construction makes the mixer-settler (X) the maximum values calculated from the least favorcolumn especially suitable for purposes where a large able conditions measured during the run at steady state, corresponding to the maximum recorded value of c ~ ( ~ - ~ ) number of theoretical stages is required for separation and also for systems in which chemical reactions are involved and the minimum recorded values of c ~ ( and ~ ) ch = 0. No and large residence times per stage are necessary. tracer could be detected in the upstream mixer, indicating Acknowledgment that drops entrained by the continuous phase leaving the mixer where the tracer was injected coalesced in the upWe thank Beili Zhu, Peoples Republic of China, for her stream settler and returned without reaching the upstream assistance in the experimental work and preparation of the mixer. Although some back-mixing then occurred between graphs. a mixer and the upstream settler, no measurable backNomenclature mixing occurred between adjacent mixers. The values of f increase with rotor speed, particularly close to the c = concentration, kmol/m3 flooding point, indicating an increase in the amount of f = back-mixing coefficient E = stage efficiency, % entrainment due to high turbulence inside the mixing N = actual number of stages compartment. NTS = number of theoretical stages The hypothesis that entrainment increases with the V = volumetric flow rate, m3 s fraction of small drops is supported by von Fischer (1983) Q = column cross section, m who measured the drop size distribution for stirrers of various sizes in the kerosene-water system. He found that Indexes at tip velocities between 0.6 and 1.2 m/s, the mean drop c = continuous phase size inside a mixer compartment of a stirred 150 mm inside d = dispersed phase M = mixer diameter column decreased noticeably and that at tip ( n )= stage number velocities around 0.8 m/s, the proportion of the smallest S = settler drops, of 0-0.2 mm in diameter, suddenly increased. AlT = tracer together, the amount of small drops and higher turbulence inside the mixer leads to higher entrainment. Literature Cited The influence of flow rates on the amount of backBuhlmann, U. Chem. Tech. 1961, IO, 769-722. von Fischer, E. Ph.D. Dissertation No. 7220, ETH Zurich, Switzerland, 1983. mixing is small at low rotor speeds. As expected, at higher Groothuis, H.; Zuiderweg, F. J. Chem. Eng. Sci. 1960, 72,288-289. flow rates of the continuous phase, the entrainment is Hackl, A.; Solar, W.; Ziebland, G. In "Recommended Systems for Liquid Extraction Studies"; Misek, T. Ed.; European Federation of Chemical Engilarger due to higher linear velocities in the hydraulic-seal neering, The Institution of Chemical Engineers: London, 1978. cylinder. The increase of entrainment for 6 m3/(m2h) of Hanson, C. Chem. Eng. 1968, 75, 76-98. over 400 rpm is sharper than for 4 m3/(m2h). Also, at Kumar, A. Ph.D. Dissertation No. 7806, ETH Zurich, Switzerland, 1985.
-
I
Ind. Eng. Chem. Process Des. Dev. 1985, 2 4 , 1225-1230 Lo, T. C.; bird, M. H. I.; Hanson, C. "Handbook of Solvent Extraction"; Wiley: New York, 1983. Marek. J. Report Luwa AG, Zurich, 1970. Oldshue, J. Y.; Rushton, J. H. Chem. Eng. Prog. 1952, 4 8 , 297-306. Scheibel, E. G.; Karr, A. E. Ind. Eng. Chem. 1950, 4 2 , 1048-1057. Scheibei, E. G. AIChEJ. 1956, 2 , 74-78. Steiner, L.; Horvath, M.; Hartland, S. Ind. Eng. Chem. Process Des. Dev. 1978, 17, 175-182.
1225
Steiner, L.;Hartland, S. Chem. Eng. Prog. lg80, 7 6 , 60-82. Stichlmair, J. Chem. h g . Tech. 1980, 52, 253-255. Wlrz, W. Swiss Patent 424 723, 1967. Zuhike, G. Graduierungsarbeit, Fachhochschule Koln, 1978.
Received for review September 11, 1984 Accepted January 22, 1985
Measurement and Prediction of Vapor-Liquid Equilibrium for an H-Coal and an SRC Coal Liquid with and without Hydrogen Ho-Mu Lin, William A. Leet, Hwayong Kim, and Kwang-Chu Chao" School of Chemical Engineering, Purdue Universi~,West Lafayette, Indiana 47907
Vapor-liquid equilibrium (VLE) has been experimentally determined for an Woal coal liquid produced from a Wyodak, WY, coal and for an SRC coal liquid from an Illinois No. 6 coal. The VLE measurements of the coal liquids with and without hydrogen were performed at temperatures up to 710 K and pressures to 25 MPa. Inspections are reported for the normal boiling point, molecular weight, and specific gravity of the VLE overhead and bottoms fractions as well as for the true boiling point (TBP) fractions of the total coal liquids. The VLE data are correlated with the Cubic Chain-of-Rotators equation of state.
Knowledge of vapor-liquid equilibrium (VLE) of coal liquid by itself and in a mixture with hydrogen is basic to the technology of coal conversion. Experimental VLE data of highly asymmetric mixtures of light gases and model compounds have been summed up by Sebastian et al. (1981 a-c) and Radosz et al. (1982). Correlation of the data has been reported by the same authors and by El Twaty and Prausnitz (1980), Wilson et al. (1981), Watanasiri et al. (1982), Gray et al. (1983), and Kim et al. (1985). Experimental data on VLE of coal liquids are scarce and limited to reports by Henry (1980),Lin et al. (1981),Sung (1981), Wilson et al. (1981), and Lin et al. (1985). We report in Chart I experimental data and correlation of the VLE of the two coal liquids. The first coal liquid was produced from the H-Coal PDU (Process Development Unit) in the syncrude mode by using Wyoming subbituminous coal from the Wyodak mine. The PDU was operated by Hydrocarbon Research, Inc. A detailed description of the process and operation conditions was given by Merdenger (1982). The second coal liquid was obtained from the operation of the Wilsonville Advanced Coal Liquefaction R & D Facility, Wilsonville, AL, on Illinois No. 6 coal. The process is a combination of a Solvent Refined Coal (SRC) unit, a Critical Solvent Deashing (CSD) unit, and a Hydrotreater (HTR) unit. Lewis (1982) provided information in detail on the operation conditions and product analyses. The VLE measurements were made in a flow apparatus at temperatures up to 710 K and pressures to 25 MPa. All condensate samples from cell effluents were collected and inspected for the boiling point, molecular weight, and density at 298.2 K. The feed coal liquids were fractionated by distillation under vacuum, and the resulting fractions were also inspected. The experimental VLE data are correlated with the Cubic Chain-of-Rotators (CCOR) equation of state (Kim et al., 1983). Experimental Apparatus a n d Results A flow apparatus was used in this work for the VLE measurements to minimize thermal decomposition of the coal liquid at the high temperatures of interest. A detailed description of the apparatus and experimental procedure 0196-4305/85/1124-1225$01.50/0
T a b l e I.
TBP D i s t i l l a t i o n and I n s p e c t i o n of Wyoming Coal
Liquid
cut feed 1 2
3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33
34
residue
wt % distilled
Tb,K
~(298.2K), g/cm3
MW
1.08 1.23 2.01 1.76 1.42 2.01 2.05 2.58 2.05 2.23 2.43 2.51 2.04 2.75 2.13 1.57 2.86 2.02 2.21 2.13 1.94 2.33 2.04 2.06 2.45 2.11 2.40 2.36 2.35 2.11 2.11 1.79 1.88 1.32 29.68
562 467 475 483 492 494 498 501 508 512 515 517 521 525 529 533 535 538 542 547 550 553 556 558 561 563 568 571 575 581 585 588 592 597 601 652
0.9477 0.8964 0.9095 0.9068 0.9052 0.9053 0.9089 0.9145 0.9162 0.9112 0.9076 0.9074 0.9134 0.9214 0.9225 0.9190 0.9170 0.9180 0.9248 0.9310 0.9300 0.9295 0.9335 0.9438 0.9540 0.9544 0.9426 0.9298 0.9364 0.9514 0.9518 0.9514 0.9556 0.9645 0.9654 1.0098
206 142" 151 152 165 165 170 171 176 173 177 180 181 180 183 187 191 193 194 194 196 199 197 198 198 199 199 203 203 204 206 211 213 212
212 268
"Estimated from boiling point; MW too low for osmometer measurement.
was presented by Lin et al. (1985). The hydrogen gas used was purchased from Airco with a minimum purity of 99.95%. Neither coal liquid nor hydrogen gas was recycled back to the system. 0 1985 American Chemical Society
