Membrane Area Requirements

An important characteristic of pervaporation that distinguishes it from distillation is that it is a rate process, not an equilibrium process. The more permeable component may be the less-volatile component. Perv oration has its greatest iitihty in the resolution of azeotropes, as an acqiinct to distillation. Selecting a membrane permeable to the minor corTiponent is important, since the membrane area required is roughly proportional to the mass of permeate. Thus pervaporation devices for the purification of the ethanol-water azeotrope (95 percent ethanol) are always based on a hydrophihc membrane. 

Because membrane equipment, capital costs, and operating costs increase with the membrane area required, it is highly desirable to maximize membrane flux. 

Applying this information to a typical diaphragm-cell tail gas, Fig. 7.4 shows the logarithm of the amount of unrecovered chlorine versus the relative membrane area required. Recovery of chlorine is not far from a first-order process. As chlorine selectively passes through the membrane, the partial pressures of the impurities increase in the remaining gas. This causes their rates of permeation to increase. The membrane area required for permeation of, say, 30% of the nitrogen is less than twice that required for 15%. 

The total membrane area required is 4.00 m. Thus, this oxygenator with Im membrane area can oxygenate 5lmin" of blood, as assumed. 

For electrical demineralization, the amount of electric current, the membrane area required, and the costs of the process depend on the amount of salt removed. The electric membrane process is currently most attractive for treatment of so-called brackish waters containing from 1000 to 10,000 p.p.m. of total dissolved solids. 

Another important factor in membrane gas separation is the pressure differential, P1 - P11 the greater this difference the less membrane area required. The membrane area is exactly proportional to the inverse of the pressure differential, i.e.,... 

Several field test studies have been undertaken utilizing the SEPAREX process in a 2-in. diameter element size Due to the modular configuration of membrane systems, a full size system can be directly designed from the test results with a small pilot plant. Although the flow rates for a pilot unit are considerably lower than might be encountered in a full-size system, all process parameters such as product purities, pressure drop, product recoveries, optimum pressure and temperature, membrane area required and series/parallel arrangement of the elements can be directly determined. 

In the case of the counter-flow/sweep membrane module illustrated in Figure 4.18(c) a portion of the dried residue gas stream is expanded across a valve and used as the permeate-side sweep gas. The separation obtained depends on how much gas is used as a sweep. In the calculation illustrated, 5 % of the residue gas is used as a sweep even so the result is dramatic. The concentration of water vapor in the permeate gas is 13 000 ppm, almost the same as the perfect counter-flow module shown in Figure 4.18(b), but the membrane area required to perform the separation is one-third of the counter-flow case. Mixing separated residue gas with the permeate gas improves the separation The cause of this paradoxical result is illustrated in Figure 4.19 and discussed in a number of papers by Cussler et al. [16]. 

A multistep design of this type can achieve almost complete removal of the permeable component from the feed stream to the membrane unit. However, greater removal of the permeable component is achieved at the expense of increases in membrane area and power consumption by the compressor. As a rule of thumb, the membrane area required to remove the last 9 % of a component from the feed equals the membrane area required to remove the first 90%. 

Here, R is the average resistance and A the area of a cell pair, Cfd and C<1 are the salt concentrations of the diluate at the inlet and outlet of the cell, Cfc and C are the salt concentrations of the concentrate cell at the inlet and outlet, F1" and rcm are the area resistances of the anion- and cation-exchange membranes. The membrane area required for a certain plant capacity as a function of the feed and product concentration of a single mono-valent salt is obtained by combination of Equations 5.22-5.26 and rearranging ... 

Figure 8.4 The compression power used and membrane area required for nitrogen membrane production as a function of membrane selectivity. The membrane permeability used for each selectivity is taken from the Robeson upper-bound trade-offline shown in Figure 8.3. All numbers are shown relative to a membrane with a selectivity of 6 and an oxygen permeability of 0.8 Barren...

Flux determines the overall size of the RO system in terms of membrane area required to achieve the desired separation. As discussed in Chapter 9.1.1, the water flux for a given application should be based on the feed water source. "Cleaner" source water allows for higher flux, which, in turn, means less membrane area is required to achieve the desired separation. 

Set out the basis for solving the problem. The problem involves determination of the membrane area required to produce 0.010 m3/s of water and the TDS concentration of the permeate. If the permeate TDS concentration is well below 200 g/m3, blending of feed and permeate will reduce the required membrane area. 

The percent additional membrane area required because of concentration polarization is... 

How well economically a gas separation membrane system performs is largely determined by three parameters. The first parameter is its permselectivity or selectivity toward the gases to be separated. Permselectivity affects the percentage recovery of the valuable gas in the feed. For the most part, it is a process economics issue. The second is the permeate flux or permeability which is related to productivity and determines the membrane area required. The third parameter is related to the membrane stability or service life which has a strong impact on the replacement and maintenance costs of the system. 

If a transmembrane pressure difference is imposed at a given constant temperature, the reaction zone will be shifted toward the lower pressure side. The mole fraction of the reactant entering the lower pressure side of the membrane surface drops to a level lower than that in the absence of a pressure difference. It has been shown [Sloot et al., 1990] that the molar fluxes of, say, hydrogen sulfide increases as the pressure on its side increases, thus potentially reducing the membrane area required. A serious drawback with this mode of operation, however, is the amount of inert gas introduced. 

Like all other chemical processes, the separation processes by inorganic membranes have two major cost issues capital investments and operating costs. Capital costs are affected by the membrane area required (which in turn arc determined by the permeability and permselectivity), compression or recompression energy requirements as dictated by the operating pressure, piping and vessels, instrumentation and control, and any pretreaunent requirements (depending on the nature of the feed material and the membrane). The operating cosLs arc determined by the required membrane replacement... 

Polymer-grade ethylene is produced in the second membrane stage at a permeate pressure of 103 kPa. Permeate pressure can be increased to lower the ethylene product compression cost, but this can only be achieved when the membrane feed composition is above its minimum value. When the permeate pressure is increased, the required feed composition is increased. Also, the membrane area required increases as the permeate pressure is increased. Finally, an increase in permeate pressure requires an increase in the recycling rate with an increase in compression costs. The second-stage performance is more efficient than the first stage with respect to both the selectivity and flux rate. This performance is affected by the feed and p meate pressures as well... 

Feed concentration in organic acid salt greater than 1 equiv/L this specification allows a reduction of the membrane area required for the conversion because the current density can remain high enough during most of the conversion. This leads to a decrease of both the investment and operating costs. 

Pilot plant smdied have also been performed by Larsen et al. [37], who obtained stable operation and more than 95% SO2 removal from flue gas streams with a gas-side pressure drop of less than 1000 Pa. The importance of the membrane structure on the SO2 removal has been studied by Iversen et al. [6], who calculated the influence of the membrane resistance on the estimated membrane area required for 95% SO2 removal from a coal-fired power plant. Authors performed experiments on different hydrophobic membranes with sodium sulfite as absorbent to measure the SO2 flux and the overall mass-transfer coefficient. The gas mixture contained 1000 ppm of SO2 in N2. For the same thickness, porosity, and pore size, membranes with a structure similar to random spheres (typical of stretched membranes) showed a better performance than those with a closely packed spheres stmcture. 

FIGURE 42.11 Variation of membrane area requirement per stage with average flux. 

The membrane areas required for the exit feed CO concentration of <10ppm in the H2 product were calculated with five different C02 permeabilities ranging from 1000 to 8000 Barrer, while the other parameters for the reference case were kept constant. As demonstrated in Figure 9.15, the required membrane area or hollow-fiber number dropped rapidly as permeability increased from 1000 to 4000 Barrer. Beyond that,... 

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