Continuous steam stripping

Remaining VCM is removed in a continuous steam-stripping column. The column is specifically designed to minimize product degra-... 

Continuous steam stripping, in which the solvent-rich mixture is fed to the top of the column and steam is injected into the base, has a lower steam requirement than batch steam distillation. To reduce a toluene/involatile mixture from 0.9 to 0.1 mole fraction of solvent will theoretically require 0.315 kg steam/kg toluene, not allowing for the heating up of the mixture, compared with batch steam distillation at about 0.43 kg/kg. 

Some slurry processes use continuous stirred tank reactors and relatively heavy solvents (57) these ate employed by such companies as Hoechst, Montedison, Mitsubishi, Dow, and Nissan. In the Hoechst process (Eig. 4), hexane is used as the diluent. Reactors usually operate at 80—90°C and a total pressure of 1—3 MPa (10—30 psi). The solvent, ethylene, catalyst components, and hydrogen are all continuously fed into the reactor. The residence time of catalyst particles in the reactor is two to three hours. The polymer slurry may be transferred into a smaller reactor for post-polymerization. In most cases, molecular weight of polymer is controlled by the addition of hydrogen to both reactors. After the slurry exits the second reactor, the total charge is separated by a centrifuge into a Hquid stream and soHd polymer. The solvent is then steam-stripped from wet polymer, purified, and returned to the main reactor the wet polymer is dried and pelletized. Variations of this process are widely used throughout the world. 

The reaction is completed after 6—8 h at 95°C volatiles, water, and some free phenol are removed by vacuum stripping up to 140—170°C. For resins requiring phenol in only trace amounts, such as epoxy hardeners, steam distillation or steam stripping may be used. Both water and free phenol affect the cure and final resin properties, which are monitored in routine quaHty control testing by gc. OxaHc acid (1—2 parts per 100 parts phenol) does not require neutralization because it decomposes to CO, CO2, and water furthermore, it produces milder reactions and low color. Sulfuric and sulfonic acids are strong catalysts and require neutralization with lime 0.1 parts of sulfuric acid per 100 parts of phenol are used. A continuous process for novolak resin production has been described (31,32). An alternative process for making novolaks without acid catalysis has also been reported (33), which uses a... 

One of the key benefits of anionic PS is that it contains much lower levels of residual styrene monomer than free-radical PS (167). This is because free-radical polymerization processes only operate at 60—80% styrene conversion, whereas anionic processes operate at >99% styrene conversion. Removal of unreacted styrene monomer from free-radical PS is accompHshed using continuous devolatilization at high temperature (220—260°C) and vacuum. This process leaves about 200—800 ppm of styrene monomer in the product. Taking the styrene to a lower level requires special devolatilization procedures such as steam stripping (168). 

Solvent extraction in batch or continuous systems is used to recover most of the residual oil from the presscake. Heptane, hexane, or a mixture of these solvents is used to recover the oil. The solvent-extracted presscake is steam stripped to recover solvent and a residual meal known as castor pomace, containing 1% residual oil. The solvent extracted oil is also processed for solvent recovery (qv). The oil from the extraction procedure is darker than the mechanically pressed oil and has a higher free fatty acid content. It is sometimes referred to as a No. 3 castor oil and is used for blending with higher quaUty oils that are well above No. 1 specifications. 

Batch with Constant Reflux Ratio, 48 Batch with Variable Reflux Rate Rectification, 50 Example 8-14 Batch Distillation, Constant Reflux Following the Procedure of Block, 51 Example 8-15 Vapor Boil-up Rate for Fixed Trays, 53 Example 8-16 Binary Batch Differential Distillation, 54 Example 8-17 Multicomponent Batch Distillation, 55 Steam Distillation, 57 Example 8-18 Multicomponent Steam Flash, 59 Example 8-18 Continuous Steam Flash Separation Process — Separation of Non-Volatile Component from Organics, 61 Example 8-20 Open Steam Stripping of Heavy Absorber Rich Oil of Light Hydrocarbon Content, 62 Distillation with Heat Balance,... 

The ethylene resides in the reactor about two minutes, and as it polymerizes, it remains dissolved in the cyclohexane. To keep the concentration of polyethylene in the cyclohexane at 35-40%, a solution of the feed, solvent, and product are continuously drawn off. Downstream, the ethylene flashes off to be recycled in a flash tower. A precipitator removes the polyethylene from the cyclohexane by centrifuge. The polyethylene is steam-stripped to remove any remaining cyclohexane, then dried, extruded, pelletized, and packaged. 

In a typical fluid catalytic cracker, catalyst particles are continuously circulated from one portion of the operation to another. Figure 9 shows a schematic flow diagram of a typical unit W. Hot gas oil feed (500 -700°F) is mixed with 1250 F catalyst at the base of the riser in which the oil and catalyst residence times (from a few seconds to 1 min.) and the ratio of catalyst to the amount of oil is controlled to obtain the desired level of conversion for the product slate demand. The products are then removed from the separator while the catalyst drops back into the stripper. In the stripper adsorbed liquid hydrocarbons are steam stripped from the catalyst particles before the catalyst particles are transferred to the regenerator. 

A steam stripping section located at the bottom of the tower frees the carbon of adsorbed material, which is drawn off directly above or passes on up the tower as reflux. However, traces of heavy material remain on the surface of the carbon and these are removed by continuously withdrawing a fraction of the circulating carbon and passing it to a high-temperature steam stripping section. 

The purpose of the final cooler is to remove the heat of compression added by the exhauster and to cool the gas to its final temperature so that downstream absorbers will operate more efficiently. Final cooling is typically achieved by direct contact with the cooling medium, either water or oil. An important function of final cooling is removal of naphthalene. In final coolers using wash oil, the naphthalene dissolves in the oil, and a side stream is steam stripped to remove the naphthalene. If water cooling is used, the condensed naphthalene must be absorbed by contacting the water with tar. The tar is continuously exchanged with fresh tar to prevent naphthalene buildup. 

The Unipol process employs a fluidized bed reactor (see Section 3.1.2) for the preparation of polyethylene and polypropylene. A gas-liquid fluid solid reactor, where both liquid and gas fluidize the solids, is used for Ziegler-Natta catalyzed ethylene polymerization. Hoechst, Mitsui, Montedison, Solvay et Cie, and a number of other producers use a Ziegler-type catalyst for the manufacture of LLDPE by slurry polymerization in hexane solvent (Fig. 6.11). The system consists of a series of continuous stirred tank reactors to achieve the desired residence time. 1-Butene is used a comonomer, and hydrogen is used for controlling molecular weight. The polymer beads are separated from the liquid by centrifugation followed by steam stripping. 

K-Resin SBC was invented by Alonzo Kitchen, a research chemist at Phillips Petroleum Research and Development laboratories. With inventorship came the opportunity to name the new resin, which he called K-Resin . The first pilot plant resins were made in 1967, and commercial samples were prepared for test marketing in 1968. Commercial production started in October of 1972 at the SBC plant in Borger, Texas, on a 10 million pound per year capacity line. Initially, the solution product was steam stripped to remove the hydrocarbon solvent, but this left a significant haze in the resin. The finishing system was quickly converted to a devolatilizing extruder. Commercial production continued at this plant until 1979, ending with the opening of a new production facility at Adams Terminal (later renamed the Houston Chemical Complex) in Pasadena, Texas. The new plant had a nameplate capacity of 120 million pounds per year. Plant expansions increased the production capacity in 1988 and 1994 to a total nameplate capacity around 300 million pounds per year. 

With the sealed ampoule process used for their kinetic studies, Root, Saeman, Harris, and Neill [20] achieved fiirfiiral yields well in excess of 70 % at temperatures above 220 °C, whereas industrial furfural processes, operating at lower temperatures and featuring a continuous removal of the furfural by steam stripping, have typical yields below 60 %. By contrast, in analytical chemistry, at a proven yield of 100 % [21], the formation of furfural from xylose or pentosan is routinely used for the quantitative determination of these substances. It is of great importance to elaborate the reasons for this yield paradox . 

Delayed coking is the only main process in a modern petroleum refinery that is a batch-continuous process. The flow through the tube oven is continuous. The feed stream is switched between two drums. One drum is filling with coke while the other drum is being steam-stripped, cooled, decoked and warmed up (see chapter 6). The overhead vapors from the coke drums flow to a rectification unit. The rectification tower has a reservoir in the bottom where the fresh feed is combined with condensed product vapors (recycle) to make up the feed to the coker heater. 

Delayed coking is a well developed commercial process (6), and operates on a semi-continuous basis. Feed, usually vacuum residue, mixed with steam, is continuously pumped through tubular heaters in which it is heated to its incipient coking temperature. At this temperature the feed is injected into an insulated drum where coking takes place. The vapors produced in the drum during coking are continuously removed and fractionated. The fractions usually include coker naphtha and light and heavy coker gas oils. As a drum fills up, feed is switched to another drum. Meanwhile, the full drum is steam stripped, cooled and the coke drilled out. Whereas feed is continuously supplied to the drum, the coke is recovered intermittently. 

Another efficient method is steam stripping. A flow of steam, at 110-125 °C, is introduced continuously, in counter flow, with a descending flow of polyether, in a classical column with plates, situated under vacuum [142] (Figure 4.36). 

Solvent (hexane) extraction of soybeans is a diffusion process achieved by immersing solid in solvent or by percolating solvent through a bed of solids. Rotary (deep-bed), horizontal belt, and continuous loop extractors are used for soybeans (Woerfel 1995). Solvent is recovered from the mixture of solvent and extracted oil (miscella) by double-effect evaporator and steam stripping and from flake by a desolventizer-toaster, and is recycled. 

The odoriferous substances in oils are free fatty acids, peroxides, aldehydes, alcohols and other organic compounds. Experience has shown that the flavor and odor removal correlates well with the reduction of free fatty acids. Therefore, all commercial deodorization consists of steam stripping the oil for free fatty acid removal. Currently, batch, semi-continuous, and continuous systems of various designs are utilized to produce deodorized oils. All of the systems utilize steam stripping with four interrelated operation variables vacuum, temperature, stripping steam rate, and holding time. 

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