Mass transfer tray column

Example 8 Calculation of Rate-Based Distillation The separation of 655 lb mol/h of a bubble-point mixture of 16 mol % toluene, 9.5 mol % methanol, 53.3 mol % styrene, and 21.2 mol % ethylbenzene is to be earned out in a 9.84-ft diameter sieve-tray column having 40 sieve trays with 2-inch high weirs and on 24-inch tray spacing. The column is equipped with a total condenser and a partial reboiler. The feed wiU enter the column on the 21st tray from the top, where the column pressure will be 93 kPa, The bottom-tray pressure is 101 kPa and the top-tray pressure is 86 kPa. The distillate rate wiU be set at 167 lb mol/h in an attempt to obtain a sharp separation between toluene-methanol, which will tend to accumulate in the distillate, and styrene and ethylbenzene. A reflux ratio of 4.8 wiU be used. Plug flow of vapor and complete mixing of liquid wiU be assumed on each tray. K values will be computed from the UNIFAC activity-coefficient method and the Chan-Fair correlation will be used to estimate mass-transfer coefficients. Predict, with a rate-based model, the separation that will be achieved and back-calciilate from the computed tray compositions, the component vapor-phase Miirphree-tray efficiencies. 

As shown in Fig. 13-92, methods of providing column reflux include (a) conventional top-tray reflux, (b) pump-back reflux from side-cut strippers, and (c) pump-around reflux. The latter two methods essentially function as intercondenser schemes that reduce the top-tray-refliix requirement. As shown in Fig. 13-93 for the example being considered, the internal-reflux flow rate decreases rapidly from the top tray to the feed-flash zone for case a. The other two cases, particularly case c, result in better balancing of the column-refliix traffic. Because of this and the opportunity provided to recover energy at a moderate- to high-temperature level, pump-around reflirx is the most commonly used technique. However, not indicated in Fig. 13-93 is the fact that in cases h and c the smaller quantity of reflux present in the upper portion of the column increases the tray requirements. Furthermore, the pump-around circuits, which extend over three trays each, are believed to be equivalent for mass-transfer purposes to only one tray each. Bepresentative tray requirements for the three cases are included in Fig. 13-92. In case c heat-transfer rates associated with the two pump-around circuits account for approximately 40 percent of the total heat removed in the overhead condenser and from the two pump-around circuits combined. 

Mass-Transfer Contact Section Where there is a strong possi-bihty that not all of the incoming vapors will be condensed in the pool, a direct-contact mass-transfer section is superimposed on the quench tank. This can be a baffle-tray section (as shown in Fig. 26-21) or a packed column sec tiou. 

Distillation design is based on the theoretical consideration that heat and mass transfer from stage to stage (theoretical) are in equilibrium [225-229]. Actual columns with actual trays are designed by establishing column tray efficiencies, and applying these to the theoretical trays or stages determined by the calculation methods to be presented in later sections. 

Example 8-42 Mass Transfer Efficiency Calculation for Baffle Tray Column (used by permission [211])... 

Fair reports that the data for mass transfer in spray, packed, and tray columns can be used for heat-transfer calculations for these columns. The pressure drop in these types of columns is usually quite low. 

Example 11.8 With highly reactive absorbents, the mass transfer resistance in the gas phase can be controlling. Determine the number of trays needed to reduce the CO2 concentration in a methane stream from 5% to 100 ppm (by volume), assuming the liquid mass transfer and reaction steps are fast. A 0.9-m diameter column is to be operated at 8 atm and 50°C with a gas feed rate of 0.2m /s. The trays are bubble caps operated with a 0.1-m liquid level. Literature correlations suggest = 0.002 m/s and A, = 20m per square meter of tray area. 

In order to develop a method for the design of distillation units to give the desired fractionation, it is necessary, in the first instance, to develop an analytical approach which enables the necessary number of trays to be calculated. First the heat and material flows over the trays, the condenser, and the reboiler must be established. Thermodynamic data are required to establish how much mass transfer is needed to establish equilibrium between the streams leaving each tray. The required diameter of the column will be dictated by the necessity to accommodate the desired flowrates, to operate within the available drop in pressure, while at the same time effecting the desired degree of mixing of the streams on each tray. 

HETP is another quantity that is used to express the efficiency of a device for carrying out a separation, particularly in which mass is transferred by a stage-wise action rather than a differential contact. For example, in a tray column, the HETP value is the tray spacing divided by the fractional overall tray efficiency. 

As might be expected, the vapour phase may offer the controlling resistance to mass transfer in high pressure distillations. Values for tray efficiencies at elevated pressure are scarce [23, 24]. The prediction of tray efficiency may be approached in several ways. One way is to utilize field performance data taken for the same system in very similar equipment. Unfortunately such data are seldom available. When they are available, and can be judged as accurate and representative, they should be used as a basis for efficiency specification [25], Another way is to utilize laboratory-or pilot-plant efficiency data. For example a small laboratory-Oldershaw tray-column can be used with the same system. Of course, the results must be corrected for vapour-and liquid mixing effects to obtain overall tray efficiencies for large-scale design [26], Another approach is the use of empirical or fundamental mass-transfer models [27-30],... 

The absorption column design represents a compromise between mass transfer factors and economic considerations. The final design specification is for a column of 1.8 m diameter, approximately 32 m high, and containing 59 sieve trays. 

The absorption column is sized according to two key parameters, these are to design for optimum mass transfer and optimum unit cost. A column internal diameter can be estimated according to the liquid and gas flowrates by utilizing graphs and nomographs such as those contained in Ref. A3. These recommendations have been refined using a computer-based mathematical model. The model predicts the required number of trays for a specified column internal diameter. These results enable a compromise to be achieved between tower cost and tower performance. 

Gas-liquid reactions form an integral part of the production of many bulk and specialty chemicals, such as the dissolution of gases for oxidations, chlorin-ations, sulfonations, nitrations, and hydrogenations. When the gaseous reactant must be transferred to the liquid phase, mass transfer can become the rate-limiting step. In this case, the use of high-intensity mixers (motionless mixers or ejectors) can increase the reaction rate. Conversely, for slow reactions a coarse dispersion of gas, as produced by a bubble column, will suffice. Because a large variety of equipment is available (bubble columns, sieve trays, stirred tanks, motionless mixers, ejectors, loop reactors, etc.), a criterion for equipment selection can be established and is dictated by the required rate of mass transfer between the phases. 

Fractionation, by definition, is simply the mass transfer between a liquid phase and a gas phase in contact with each other. A fractionation column is simply a tall, vertical, cylindrical pressure vessel that contains numerous flat internal metal plates called trays. Each tray allows liquid to flow over it, so the liquid flows from tray to tray by the force of gravity. The liquid thus enters the top tray. The liquid portion not vaporized in the column s trays is taken out in the column s bottom liquid accumulation. Gas enters the column s bottom section and flows through each tray to the top section. Entering vapor pressure is its driving force. Gas not absorbed by the liquid exits the column s top section. 

Often a part of the condensate is returned (reflux) back to the still and is mixed with the outgoing vapour. This allows further transfer of lighter components to the vapour phase from the liquid phase and transfer of heavier components to the liquid phase from the vapour phase. Consequently, the vapour stream becomes richer in light components and the liquid stream becomes richer in heavy components. Different types of devices called plates, trays or packing are used to bring the vapour and liquid phases into intimate contact to enhance the mass transfer. Depending on the relative volatility and the separation task (i.e. purity of the separated components) more trays (or more packing materials) are stacked one above the other in a cylindrical shell to form a column. 

A stagnant film model is used for two-phase boundaries (1-2), which in effect, isolates the mass transfer process to a thin region at the interface stagnant film. Once the expressions for entropy production in terms of pressure, temperature, and composition are available a transformation is made to process variables such as reflux ratio, column height, packing or tray geometry, column diameter and column efficiency. Results of this design optimization model are compared with the results obtained via traditional methods. 

In nonequilibrium models, as in the other models, the subscript j is for the stage. In a trayed column, it is the actual tray. In a packed column, j is a section of packing. By convention, transfer is to be from the liquid to the vapor with the mass transfer rate to the vapor, Nf, taken as positive. 

The number of equations, M5C + 1), for a large number of trays and components, can be excessive. The global Newton method will suffer from the same problem of requiring initial values near the answer. This problem is aggravated with nonequilibrium models because of difficulties due to nonideal if-values and enthalpies then compounded by the addition of mass transfer coefficients to the thermodynamic properties and by the large number of equations. Taylor et al. (80) found that the number of sections of packing does not have to be great to properly model the column, and so the number of equations can be reduced. Also, since a system is seldom mass-transfer-limited in the vapor phase, the rate equations for the vapor can be eliminated. To force a solution, a combination of this technique with a homotopy method may be required. 

Entrainment (Figure 6.14) is liquid transported fry the gas to the tray above. This liquid contains more of the less-volatile material than the tray above, and therefore it counteracts the mass transfer process and reduces tray efficiency. Other undesirable effects of entrainment are carryover of nonvolatile impurities upward to contaminate the overhead product and the possibility of damage to rotating machinery located in the path of tbe column overhead vapor. 

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