Permeation data pervaporation

To produce permeation data for a simulation model, pervaporation experiments were carried out with a lab-scale permeation apparatus. Ethanol-water mixmres were used as a model mixture. The feed tank holds a volume of 4 L of feed mixmre and has a function of temperature control. A permeation cell made of stainless steel holds a fiat membrane of 154 cm. Feed composition ranged from 0.4 to 7 wt% water contents. Feed temperatures... 

Equation (2.79) expresses the driving force in pervaporation in terms of the vapor pressure. The driving force could equally well have been expressed in terms of concentration differences, as in Equation (2.83). However, in practice, the vapor pressure expression provides much more useful results and clearly shows the connection between pervaporation and gas separation, Equation (2.60). Also, the gas phase coefficient, is much less dependent on temperature than P L. The reliability of Equation (2.79) has been amply demonstrated experimentally [17,18], Figure 2.13, for example, shows data for the pervaporation of water as a function of permeate pressure. As the permeate pressure (p,e) increases, the water flux falls, reaching zero flux when the permeate pressure is equal to the feed-liquid vapor pressure (pIsal) at the temperature of the experiment. The straight lines in Figure 2.13 indicate that the permeability coefficient d f ) of water in silicone rubber is constant, as expected in this and similar systems in which the membrane material is a rubbery polymer and the permeant swells the polymer only moderately. 

Figure 2.17 Flux of n-hexane through a rubbery membrane as a function of the hexane concentration difference in the membrane. Data taken from both reverse osmosis ( ) and pervaporation (O) experiments. Feed-side and permeate-side membrane concentrations, Ci0 m) and Cie m), calculated from the operating conditions through Equations (2.26), (2.36) and (2.76). Maximum flux is obtained at the maximum concentration difference, when the permeate-side membrane concentration cit(m)), equals zero [19]. Reprinted from Driving Force for Hydraulic and Pervaporation Transport in Homogeneous Membranes, D.R. Paul and D.J. Paciotti, J. Polym. Sci., Polym. Phys. Ed. 13, 1201 Copyright 1975. This material is used by permission of John Wiley Sons, Inc.

A second method of determining the coefficient ( >,/5) and the intrinsic enrichment of the membrane Ea is to use Equation (4.11). The term ln(l — 1/E) is plotted against the permeate flux measured at constant feed solution flow rates but different permeate pressures or feed solution temperatures. This type of plot is shown in Figure 4.10 for data obtained with aqueous trichloroethane solutions in pervaporation experiments with silicone rubber membranes. 

Thus, the permeation of hydrocarbons in polymer membranes is governed by the basic regularities typical of permeation of low MW penetrants, modified however by certain peculiarities related to the stmcture and shape of hydrocarbon molecules. We will now discuss the physicochemical regularities of hydrocarbon separation and removal using polymer membranes, by trying to reveal the relationship between the chemical stmcture of polymers and their separation properties with respect to mixtures containing hydrocarbons. It follows from literary data that mbbery polymers are mainly used in gas/vapor separation processes for selective separation of hydrocarbon vapors from their mixtures with air as well as in pervaporation processes for the removal of hydrocarbons from their aqueous solutions. In practice, glassy polymers are used for separation of olefins and paraffins as well as for separation of aromatic, ahcyclic, and aliphatic hydrocarbons. 

Finally, it is important to notice the effect of the support in the pervaporation flux, analyzed by Bruijn et al. [117] who proposed a model and evaluated the contribution of the support layer to the overall resistance for mass transfer in the selected literature data. They found that in many cases, the support is limiting the flux the permeation mechanism through the support corresponds to a Knudsen diffusion mechanism, which makes improvements in the porosity, tortuosity, pore diameter, and thickness necessary for an increase in the pervaporation flux. 

These data are important in understanding vapor permeation through polymeric membranes, which occurs in the pervaporation process. 

Finally, it is important to notice the effect of the support in the pervaporation flux, analyzed by de Bruijn et al. [164] who proposed a model and evaluated the contribution of the support layer to the overall resistance for mass transfer in the selected literature data. They found that in many cases, the support is limiting the flux the permeation mechanism through the support corresponds to a Knudsen diffusion mechanism, which makes improvements in the porosity, tortuosity, pore diameter, and thickness necessary for an increase in the pervaporation flux. In fact, the researchers of Bussan Nanotech Research Institute Inc. (BNR), Sato et al. [165], designed and patented an appropriate asymmetric ceramic porous support to suppress pressure drop, and in this case, the water flux increased dramatically compared to previous reported results. Wang et al. [166] have clearly shown that the flux of the membranes increased with the porosity of the hollow fiber supports. In spite of the thin 1 pm zeolite layer, prepared by Zhou et al. [167], the flux enhancement compared to layers 10 times thicker [168] was not significant. 

The membranes were employed for the pervaporation separation of water-isopropanol mixtures. The experimental data demonstrated that membrane containing 40 mass% of ETMS showed the highest separation selectivity of 17,990 with a flux of 2.92 x 10 kg/m h at 30°C for 10 mass% of water. The activation energy values obtained for water permeation are two times lower than those of isopropanol... 

The design of any real pervaporation and vapor-permeation installation has thus to be based on experimental data measured in the laboratory under conditions as similar as possible to those of the subsequent full-size plant. These conditions include the flow regime of the feed mixture, the temperature and the geometry of the feed side, the composition and nature of the feed mixture, the permeate side geometry and partial vapor pressure. From the experimental data the partial transmembrane fluxes of all components of a mixture and thus the selectivity can be determined as a function of composition, temperature and permeate-side conditions for the respective mixture and geometry. In practice the permeate-side conditions (total pressure, condensation temperature) are kept as close as possible to those expected in the final plant, thus changes of these parameters do not need to be considered. Figure 3.3 depicts the partial fluxes of EtOH and water measured for a PVA-membrane. 

Hollow fibers or capillary modules have not yet found an industrial application in pervaporation or vapor-permeation processes. A few data have been reported where organic capillary structures with an outside diameter of 0.5 to 1 mm have been coated with silicon and used in organophilic separation. With the flow on the shell side permeate pressure losses inside the bore of the fiber control the process. For specific organophilic applications, these pressure losses may be tolerable. For hydrophilic processes, however, the useful length of a module would be of the order of 20 to 30 cm only, even at an inner diameter of the capillary of 1 mm. Such a module, including housing and connection in any industrial application, is more costly than a plate module. So far no potting material is available that combines the necessary chemical and mechanical stability at the operation temperature and pressure of a dehydration plant. 

An important dynamic parameter is the holdup in the pervaporation unit. Since the mass of the liquid retentate phase is much larger than that of the vapor permeate phase, it will dominate the dynamic response. The holdup depends on the volume. In this study, the data given in Geankoplis is used (328m /m ). Thus the 300 m module has a volume of 0.9 m that is split among the three cells in the module. 

As for the pervaporation separation of binary mixtures, the calculation was done with respect to the binary system of water(i) / ethanol(j) mixtures. The water mole fraction in the permeate vapor, K/j, and the relative weight flux. IVrei> were calculated from Equations 6.135 and 6.145, for different Bj/B, values. The saturation vapor pressure and the vapor composition of water/ethanol binary mixtures for 40 C were adopted from the literature [243], and therefore the calculated values should correspond to pervaporation data for 40 C. All the results are summarized in Figure 6.25. 

Figure 6.31 shows some experimental data for the pervaporation of water/ethanol mixtures by a silicone rubber membrane preferentially permeable to ethanol. The experiment was conducted at 23 C for downstream pressures of 667, 1200, and 2100 Pa (5,9, and 16 mmHg). As reported by Hoover and Hwang [250] and Tanigaki et al. [245], the silicone membrane showed preferential permeation to ethanol. Evidently, the downstream pressure has little effect on both permeate composition and permeation rate, supporting the calculated results shown in Figure 6.30. When the experimental data arc closely examined, however, the relative permeation rate decreases slightly with an increase in the downstream pressure. The calculated values in Figure 6.30 show exactly the same tendency, justifying the transport model on which the calculation is based. It has to be noted, however, that the saturation vapor pressure of water and ethanol at 60 C are 1.99 x l(f and 4.69 x lO Pa (149.4 and 351.9 mmHg), respectively. When the downstream pressure approaches the saturation vapor pressure, the assumption on which the theoretical calculation is based (i.e., (he vapor permeation prevails across the membrane cross-.section) becomes invalid, since liquid penetrates more deeply into the pore.

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