Gas-liquid distributor

The compressive strength of Monoliths would be much higher than the catalyst particles generally used in packed. Therefore, the use of Monoliths would enable deep beds to be constructed without using intermediate supports and gas-liquid distributors. 

Maiti, R.N., and Nigam, K.D.P. (2007), Gas-liquid distributors for trickle-bed reactors A review, Industrial Engineering Chemistry Research, 46(19) 6164-6182. 

This last issue was tackled in a recent paper [38]. The hydrogenation of2-butyne-l,4-diol was conducted at 55 °C and 2 bar pressure in monolithic reactor containing more than 5000 channels (1 mm) with a total diameter as large as 10 cm, equipped with a cocurrent downflow gas-liquid distributor. The results reveals that in the monolith reactor the selectivity for the intermediate 2-butene-l,4-diol (Scheme 9.1) can be kept very high (99.6%) even at quantitative substrate conversion, a much better value than in traditional stirred-tank or trickle-bed reactors (<95%). Since the Pd catalysts used in the different reactors are different, with much better dispersed Pd particles in the monolith reactor, the observed better selectivity is partly due to the intrinsic properties of the catalyst. The productivity numbers of the different reactors are not discussed. 

D. Hold-Down Grid (see Figure 3). According to Strigle, Normally, the upper surface of the packed bed is at least 6 in. below a liquid distributor or redistributor. The bed limiter or hold-down plate is located on top of the packed bed in this space. It is important to provide such a space to permit gas disengagement from the packed bed. Such a space allows the gas to accelerate to the velocity neces-... 

The liquid distributor is the most important internal structure of a packed column. The distributor strongly influences packing efficiency. It must spread the liquid uniformly, resist plugging/fouling, provide free space for gas flow, and allow operating flexibility. 

Westerterp et al. (W5) measured interfacial areas in mechanically agitated gas-liquid contactors. The existence of two regions was demonstrated At agitation rates below a certain minimum value, interfacial areas are unaffected by agitation and depend only on nominal gas velocity and the type of gas distributor, whereas at higher agitation rates, the interfacial areas are... 

Much higher shear forces than in stirred vessels can arise if the particles move into the gas-liquid boundary layer. For the roughly estimation of stress in bubble columns the Eq. (29) with the compression power, Eq. (10), can be used. The constant G is dependent on the particle system. The comparison of results of bubble columns with those from stirred vessel leads to G = > 1.35 for the floccular particle systems (see Sect. 6.3.6, Fig. 17) and for a water/kerosene emulsion (see Yoshida and Yamada [73]) to G =2.3. The value for the floe system was found mainly for hole gas distributors with hole diameters of dL = 0.2-2 mm, opening area AJA = dJ DY = (0.9... 80) 10 and filled heights of H = 0.4-2.1 m (see Fig. 15). 

For the sake of developing commercial reactors with high performance for direct synthesis of DME process, a novel circulating slurry bed reactor was developed. The reactor consists of a riser, down-comer, gas-liquid separator, gas distributor and specially designed internals for mass transfer and heat removal intensification [3], Due to density difference between the riser and down-comer, the slurry phase is eirculated in the reactor. A fairly good flow structure can be obtained and the heat and mass transfer can be intensified even at a relatively low superficial gas velocity. 

Internal-loop airlift reactors (ALRs) are widely used for their self-induced circulation, improved mixing, and excellent heat transfer [1], This work reports on the design of an ALR with a novel gas-liquid separator and novel gas distributor. In this ALR, the gas was sparged into the annulus. The special designed gas-liquid separator, at the head of the reactor, can almost completely separate the gas and liquid even at high gas velocities. 

Figure 5.15 Schematic of the multiphase packed bed reactor. Gas inlet (A) liquid distributor (B) catalyst inlets (Q exit port manifold (D) [11].

The catalyst is a fixed bed. Flows of gas and liquid are cocurrent downwards. Liquid feed is at a such a low rate that it is distributed over the packing as a thin film and flows by gravity, helped along by the drag of the gas. This mode is suited to reactions that need only short reaction times, measured in seconds, short enough to forestall undesirable side reactions such as carbon formation. In the simplest arrangement the liquid distributor is a... 

The simplest choice of a liquid distributor is a perforated plate with 10 openings/dm2 (10 openings/15.5 in2), where the gas enters through several risers about 15 cm (5.9 in) high. More sophisticated distributors like caps ate also used. The thickness of the liquid film developed in trickle-bed reactors has been estimated to vary between 0.01 and 0.2 mm (Perry and Green, 1999). 

High-pressure conditions favour a smaller bubble size and narrower bubble-size distribution, and therefore lead to higher gas hold-up in BSCR, except in systems operated with porous plate distributors and at low gas velocities. For design purposes in BSCR at high pressure, where the liquids operate in the batch mode, Luo et al. [31] proposed the following formula for the calculation of the gas hold-up, based on their proper experimental data and those of many other authors [1,26,31-34] for various systems of gas, liquid and solids ... 

Previous workers have studied the influence of the ratio of the cross-section area of the downcomer to the riser [4,5], the reactor height [6,7], the gas-liquid separator configuration [8], and the distributor type and location [9]. All these affect the flow characteristics and mass transfer. Most previous works focus on global parameters, such as the liquid circulation velocity [10-13] and the average gas holdup in the riser [14-16]. Although much work has been carried out on EL-ALRs, the proper design and scale-up of an EL-ALR is still difficult because any variation in the physical properties of the gas or the liquid and the reactor structural feathers can have a considerable effect on the hydrodynamics... 

Liquid distribution may be an important parameter, as demonstrated in the HOC1 process, where different liquid distributors provided significantly different results (8). The initial contact of the liquid with the rotor influences the mass transfer performance of the RPB in gas continuous operations (15). Although the use of a packing support at the inside diameter of the rotor would be expected to impact this initial liquid contact with the rotor, experiments did not show any reduced mass transfer performance (36). 

Packed-tower efficiency and turndown are strongly dependent on the quality of initial liquid distribution. Uneven distribution may cause local variations in the liquid/gas ratio, localized pinch conditions, and reduced vapor-liquid contact. Figure 14 shows two common liquid distributor types, the ladder type (shown as the top distributor) and the orifice type (shown as the redistributor). The ladder type is a horizontal header of pipes, which are perforated on the underside. The orifice type is a flat perforated plate equipped with round or rectangular risers for gas passage. Other common types of distributors are a header equipped with spray nozzles (spray distributor) and a header of horizontal channels, with V notches cut in the vertical walls of the channels (notched-trough distributor). 

Ladder and spray distributors rely on pressure for their action. They provide a large gas flow area but a somewhat limited liquid flow area they are light and cheap but are sensitive to corrosion, erosion, and to a certain extent plugging. They are most suitable for high gas/liquid ratio applications. 

Pilot-plant experiments have been carried out at real process conditions in the coke plant August Thyssen (Duisburg, Germany). The DN 100 pilot column (Fig. 9.11) was made from stainless steel and equipped with about 4 m of structured packing (Sulzer MELLAPAK 350Y), three liquid distributors, and a digital control system. Several steady-state experiments have been compared with the simulation results and supported the design optimization of the coke gas purification process [91]. 

In fact, the concept of the quasi-homogeneous gas/liquid mixture, on which also the formulation of the target pi-number Y = (kLa/v) with intensity quantities is based, and which was fully verified in bubble columns with perforated plates as gas distributors, proves to be totally inappropriate when injectors are used as gas dispersers. The explanation for this fact is that in the case of injectors the coalescence takes place both in the free jet of the G/L dispersion and at its disintegration into a bubble swarm, while in the case of gas distribution with perforated plates this process has already been completed just above the perforated plate. 

The volumetric gas-liquid mass transfer coefficient, khaL, largely depends on power per unit volume, gas velocity (for a gassed system), and the physical properties of the fluids. For high-viscosity fluids, kLaL is a strong function of liquid viscosity, and for low-viscosity fluids (fi < 50 mPa s), kLaL depends on the coalescence nature of the bubbles. In the aeration of low-viscosity, pure liquids such as water, methanol, or acetone, a stable bubble diameter of 3-5 mm results, irrespective of the type of the gas distributor. This state is reached immediately after the tiny primary bubbles leave the area of high shear forces. The generation of fine primary gas bubbles in pure liquids is therefore uneconomical. 

Fig. 1.3. Top of H2S04-making ( absorption ) tower, courtesy Monsanto Enviro-Chem Systems, Inc. www.enviro-chem.com The tower is packed with ceramic saddles. 98.5 mass% H2S04, 1.5 mass% H20 sulfuric acid is distributed uniformly across this packed bed. Distributor headers and downcomer pipes are shown. The acid flows through slots in the downcomers down across the bed (see buried downcomers below the right distributor). It descends around the saddles while S03-rich gas ascends, giving excellent gas-liquid contact. The result is efficient H2S04 production by Reaction (1.2). A tower is 7 m diameter. Its packed bed is 4 m deep. About 25 m3 of acid descends per minute while 3000 Nm3 of gas ascends per minute.

Figure 3. Trickle-bed reactor. Key A, gas b, liquid c, liquid distributor d, thermocouple e, alumina particles f, jacket g, catalyst h, glass beads i, gas distributor j, packing supporting plate K distributor plate.

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