Permeation membrane module

Permeate Membrane module permeate stream referred to... 

Fig. 22 shows a reaction scheme enhanced by continuously removing water directly from the reactor. In this case, the water is removed from the vapor phase. A vapor stream is sparged from the reactor and circulated through a vapor permeation membrane module, where water is selectively permeated through the membranes. The membrane unit is sized, such that all the reaction water can be removed with the water/ alcohol ratio just below the azeotropic composition. 

The mass balances for a gas permeation membrane module that is conpletely mixed on both sides of the membrane will have the sinplest form since the mole fracs and pressures on each side of the membrane are constant. Equations (17-2) and tl7-3b can then be used in algebraic mass balances with yp = y and integration is not required to find the average retentate and permeate mole fracs. 

If the pressure drop across a tubular membrane is 2.8 bars, determine the permeat velocity across the membrane module. The thickness and the porosity of the deposit are 2 mm and 40 %, respectively. The average diameter of the partices is 5 microns. The initial membrane resistance is estimated to be 1.7 X 10 1/m. 

Fluid Stream Designations For the generalized membrane module shown in Fig. 20-45, a feed stream enters a membrane module while both a permeate and a retentate stream exit the module. The permeate (or filtrate) stream flows through the membrane and has been depleted of retained components. The term filtrate is commonly used for NFF operation while permeate is used for TFF operation. The retentate (or concentrate) stream flows through the module, not the membrane, and has been enriched in retained components. 

Flow Flux, Permeability, Conversion The productivity of a membrane module is described by its flux J = volumetric permeate flow rate/membrane area with units of volume per area per time. Relatively high flux rates imply that relatively small membrane areas are required. The permeate volume is usually greater than the feed volume for a given process. Flux is also the magnitude of the normal flow velocity with units of distance per time. 

The efficiency of a membrane module is characterized by the recovery or conversion ratio CR = permeate flow rate/feed flow rate. Low conversion means that fluid has to be repeatedly cycled past the TFF module to generate permeate. High-efficienCT NFF has CR = 1. 

Pervaporation. Pervaporation differs from the other membrane processes described so far in that the phase-state on one side of the membrane is different from that on the other side. The term pervaporation is a combination of the words permselective and evaporation. The feed to the membrane module is a mixture (e.g. ethanol-water mixture) at a pressure high enough to maintain it in the liquid phase. The liquid mixture is contacted with a dense membrane. The other side of the membrane is maintained at a pressure at or below the dew point of the permeate, thus maintaining it in the vapor phase. The permeate side is often held under vacuum conditions. Pervaporation is potentially useful when separating mixtures that form azeotropes (e.g. ethanol-water mixture). One of the ways to change the vapor-liquid equilibrium to overcome azeotropic behavior is to place a membrane between the vapor and liquid phases. Temperatures are restricted to below 100°C, and as with other liquid membrane processes, feed pretreatment and membrane cleaning are necessary. 

The concept of cross-flow microfiltration, described by Bertera, Steven and Metcalfe(4), is shown in Figure 8.4 which represents a cross-section through a rectangular or tubular membrane module. The particle-containing fluid to be filtered is pumped at a velocity in the range 1-8 m/s parallel to the face of the membrane and with a pressure difference of 0.1-0.5 MN/m2 (MPa) across the membrane. The liquid permeates through the membrane and the feed emerges in a more concentrated form at the exit of the module. 

FIGURE 1.35 SLM process using O-9-(l-adamantylcarbamoyl)-10,ll-dihydro-ll-octadecylsulfinylquinme and corresponding quinidine derivative as chiral carriers for the preparative separation of enantiomers of Al-derivatized amino acids (e.g., DNB-Leu). (a) ftinciple of the carrier SLM process with carrier-mediated transport (top) and (nonstereoselective) nonspecific transport processes (bottom), (b) General experimental setup of the SLM production unit with two membrane modules, (c) Multistage SLM purification process. P, permeate QD/QN, membrane modnles snpported with quinidine-derived and quinine-derived chiral carriers. R, S, D, L refers to the respective enantiomers of the selectand (DNB-Leu). (Reproduced from A. Maximini et al., J. Membr. ScL, 276 221 (2006). With permission.)... 

Cabral and coworkers [253] have investigated the batch mode synthesis of a dipeptide acetyl phenylalanine leucinamide (AcPhe-Leu-NH2) catalyzed by a-chymotrypsin in a ceramic ultrafiltration membrane reactor using a TTAB/oc-tanol/heptane reverse micellar system. Separation of the dipeptide was achieved by selective precipitation. Later on the same group successfully synthesized the same dipeptide in the same reactor system in a continuous mode [254] with high yields (70-80%) and recovery (75-90%). The volumetric production was as high as 4.3 mmol peptide/l/day with a purity of 92%. The reactor was operated for seven days continuously without any loss of enzyme activity. Hakoda et al. [255] proposed an electro-ultrafiltration bioreactor for separation of RMs containing enzyme from the product stream. A ceramic membrane module was used to separate AOT-RMs containing lipase from isooctane. Application of an electric field enhanced the ultrafiltration efficiency (flux) and it further improved when the anode and cathode were placed in the permeate and the reten-tate side respectively. 

Polymer-Assisted Ultrafiltration of Boric Acid. The Quickstand (AGT, Needham, MA) filtration apparatus is pictured schematically in Figure 3. The hollow fiber membrane module contained approximately 30 fibers with 0.5 mm internal diameter and had a nominal molecular weight cut-off of 10,000 and a surface area of 0.015 m2. A pinch clamp in the retentate recycle line was used to supply back pressure to the system. In a typical run, the transmembrane pressure was maintained at 25 psig and the retentate and permeate flow rates were 25 ml/min and 3 ml/min, respectively. Permeate flux remained constant throughout the experiments. 

Figure 19.4. The spiral wound membrane module for reverse osmosis, (a) Cutaway view of a spiral wound membrane permeator, consisting of two membranes sealed at the edges and enclosing a porous structure that serves as a passage for the permeate flow, and with mesh spacers outside each membrane for passage of feed solution, then wound into a spiral. A spiral 4 in. dia by 3 ft long has about 60 sqft of membrane surface, (b) Detail, showing particularly the sealing of the permeate flow channel, (c) Thickness of membranes and depths of channels for flows of permeate and feed solutions.

Several factors contribute to the successful fabrication of a high-performance membrane module. First, membrane materials with the appropriate chemical, mechanical and permeation properties must be selected this choice is very process-specific. However, once the membrane material has been selected, the technology required to fabricate this material into a robust, thin, defect-free membrane and then to package the membrane into an efficient, economical, high-surface-area module is similar for all membrane processes. Therefore, this chapter focuses on methods of forming membranes and membrane modules. The criteria used to select membrane materials for specific processes are described in the chapters covering each application. 

Spiral-wound modules were used in a number of early artificial kidney designs, but were fully developed for industrial membrane separations by Gulf General Atomic (a predecessor of Fluid Systems, Inc.). This work, directed at reverse osmosis membrane modules, was carried out under the sponsorship of the Office of Saline Water [112-114], The design shown in Figure 3.42 is the simplest, consisting of a membrane envelope of spacers and membrane wound around a perforated central collection tube the module is placed inside a tubular pressure vessel. Feed passes axially down the module across the membrane envelope. A portion of the feed permeates into the membrane envelope, where it spirals towards the center and exits through the collection tube. 

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]. 

Figure 4.19(a) shows the concentration of water vapor on the feed and permeate sides of the membrane module in the case of a simple counter-flow module. On the high-pressure side of the module, the water vapor concentration in the feed gas drops from 1000 ppm to about 310 ppm halfway through the module and to 100 ppm at the residue end. The graph directly below the module drawing shows the theoretical maximum concentration of water vapor on the permeate side of the membrane. This maximum is determined by the feed-to-permeate pressure ratio of 20 as described in the footnote to page 186. The actual calculated permeate-side concentration is also shown. The difference between these two lines is a measure of the driving force for water vapor transport across the membrane. At the feed end of the module, this difference is about 1000 ppm, but at the permeate end the difference is only about 100 ppm. 

Backflushing is another way of cleaning heavily fouled membranes. The method is widely used to clean capillary and ceramic membrane modules that can withstand a flow of solution from permeate to feed without damaging the... 

Figure 6.14 Backflushing of membrane modules by closing the permeate port. This technique is particularly apphcable to capillary fiber modules...

Both Mitsui [26] and Sulzer [27] have commercialized these membranes for dehydration of alcohols by pervaporation or vapor/vapor permeation. The membranes are made in tubular form. Extraordinarily high selectivities have been reported for these membranes, and their ceramic nature allows operation at high temperatures, so fluxes are high. These advantages are, however, offset by the costs of the membrane modules, currently in excess of US 3000/m2 of membrane. 

Very small systems (less than 5 MMscfd). At this flow rate, membrane units are very attractive. Often the permeate is flared or used as fuel, so the system is a simple bank of membrane modules. 

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