Extractive distillation column control

Modeling, Simulation and Control of an Extractive Distillation Column... 

The digital simulation of an extractive distillation column was performed in order to understand the dynamic behaviour of the system. Based on this results a considerably simplified dynamic model of sufficient accuracy could be developed. This model was employed in the design of a state observer and of an optimal control. After implementation in the large scale plant this new control system has proved to be highly efficient and reliable. 

Separation constraints The separation in a column can be expressed as the impurity levels of the key components in the two products xg.LK in the bottoms and xD Hx in the distillate. Separation is limited by the minimum reflux ratio and the minimum number of trays. We must always have more trays than the minimum and a higher reflux ratio than the minimum. If the number of trays in the column is not large enough for the desired separation, no amount of reflux will be able to attain it and no control system will work. In extractive distillation columns, there is also a maximum reflux ratio limitation, above which the overhead stream becomes less pure as the reflux increases. 

Write the trajectory of an extractive distillation column at separation of a four-component mixture at minimum reflux, bottom feeding being the control one. 

In this section we use an extractive distillation column as an example to demonstrate how to build a steady-state simulation. This is one column of an overall two-column system for separating isopropanol and water. The detailed design and control of the overall distUlation system will be given in Chapter 10. 

Some simple regulatory control loops are determined first. The levels of the reflux drums for both columns are controlled by manipulating the distillate. Top pressures of both columns are controlled by the condenser duty. The bottom level of the extractive distillation column is controlled by manipulating the bottoms flow. The entrainer feed temperature is controlled at 72°C by manipulating the cooler duty. The initial control structure tested is one in which the reflux ratios in the two columns are controlled by manipulating the reflux. This was used in the overall control structure in Luyben. ... 

The conversion of methanol and ammonia to methylamines is achieved over dehydration catalysts operated in the temperature range 300450°C and 0.12 MPa pressure. The reactions are exothermic, and excess ammonia is used to control the product distribution. The dehydration catalysts are generally promoted Si-Al composites. The promoters include molybdenum sulfide and silver phosphate [68]. In the commercial Leonard process, a gas-phase downflow catalytic reactor operating at about 350°C and 0.62 MPa is used [69]. Recovery of the desired product is achieved throu a series of four distillation and extractive distillation columns. Unwanted product is recycled, suppressing further formation of the undesired component in the reactor. A very small amount of methanol is lost to CO and H2, and yields from the commercial process based on methanol and anunonia are >97% [70]. 

In Chapter 8 we explored the steady-state design of the TAME reactive distiUalion system. The reactive column is part of a multiunit process that includes other columns for recovery of the methanol reactant. The recovery is necessary because the presence of methanol/C5 azeotropes unavoidably removes methanol from the reactive column in the distillate stream. The economics of two alternative methanol recovery systems were evaluated in Chapter 8. In this chapter the dynamic control of the process is studied, and an effective plantwide control structure is developed. The process has three distillation columns one reactive column, one extractive distillation column, and one methanol/water separation column from which methanol and water are recycled. 

The control structure for the prereactor and reactive distillation column Cl is shown in Figure 14.7, The control structure for the extractive distillation column C2 and the methanol recovery column C3 is shown in Figure 14.8. 

Whereas there is extensive Hterature on design methods for azeotropic and extractive distillation, much less has been pubUshed on operabiUty and control. It is, however, widely recognized that azeotropic distillation columns are difficult to operate and control because these columns exhibit complex dynamic behavior and parametric sensitivity (2—11). In contrast, extractive distillations do not exhibit such complex behavior and even highly optimized columns are no more difficult to control than ordinary distillation columns producing high purity products (12). 

Nonlinear programming Staged-Distillation column (12.1) < Liquid extraction column (12.2) Gas transmission network (13.4) Ammonia reactor (14.2) Alkylation reactor (14.3) CVD reactor (14.5) Refrigeration process (15.2) Extractive distillation (15.3) Operating margin (15.4) Reactor control (16.3)... 

Extractive Distillation Recovery of Isoprene. A typical flow-sketch and material balance of distillation and solvent recovery towers for extracting isoprene from a mixture of cracked products with aqueous acetonitrile appears in Figure 13.25. A description of the flowsheet of a complete plant is given in Example 2.9. In spite of the fact that several trays for washing by reflux are provided, some volatilization of solvent still occurs so that the complete plant also has water wash columns on both hydrocarbon product streams. A further complication is that acetonitrile and water form an azeotrope containing about 69 mol % solvent. Excess water enters the process in the form of a solution to control poly-... 

In the extractive-stripping (ES) mode the extractor and stripper are combined together. The extractor section is contacted with the lean gas and the rich gas flows down into the stripper column coming in contact with countercurrent stripped hydrocarbons from the reboiler at the bottom of the ES column. The recovery of the desired compounds is controlled by the lean gas flow rate, stripper bottom temperatures and operating pressures. The rich solvent leaving from the ES column is expanded to the pressure of operation in the product column, which is essentially a distillation column. Here the... 

The disturbances F, and T, of the stirred tank heater (Figure 1.1) are easily measured thus they are considered measured disturbances. On the other hand, the feed composition for a distillation column, an extraction unit, reactors, and the like, is not normally measured and consequently is considered an unmeasured disturbance. As we will see later, unmeasured disturbances generate more difficult control problems. 

In spite of its wide use, there are still three major problems with the Wittig reaction. 1) The stereochemistry often cannot be controlled. 2) Ketones and hindered aldehydes fail to react with phosphoranes that are hindered or are stabilized by strongly electron withdrawing substituents. 3) The by-product triphenylphosphine oxide can be difficult to separate from the product alkene. Often the alkene and the triphenylphosphine oxide cannot be separated by extraction, distillation, or crystallization, and column chromatography is required. 

The steady-state design of a two-column extractive distillation system was developed in Chapter 5. Now, we want to design an effective control structure for this system. The process has two distillation columns, and a plantwide control structure must be developed that accounts for the interaction between the columns and for the solvent recycle. 

Different physical modes are sometimes available for the same unit operation. A distillation column can, for example, be modeled on the basis of theoretical stages or using a rate-based model, taking into account the mass transfer on the column internals. A simulation of this kind can be used to extract the data for the design of the process equipment or to optimize the process itself During recent years, dynamic simulation has become more and more important. In this context, dynamic means that the particular input data can be varied with time so that the time-dependent behavior of the plant can be modeled and the efficiency of the process control can be evaluated. 

The general algorithm of calculation of the minimum reflux mode for columns of extractive distillation with two feeds requires the check-up of the conditions of trajectories joining for the cases of bottom and top control feed and requires the determination of the values of (E/E) bigger of these two... 

In this chapter, design and control of the IPA dehydration process via extractive distillation will be studied. Since, in this distillation system, entrainer selection is an important step before working on the optimal design of the column sequence, we will start by comparing two alternative entrainers for this separation system in the following section. 

The Aspen Plus file of this extractive distillation system is exported to Aspen Dynamics after dynamic parameters are specified (equipment sizes). Figure 11.8 shows the control stmcture developed for this system, which is based on the extractive distillation control structure proposed by Grassi. Relay-feedback testing and Tyreus-Luyben tuning of the temperature loops give the controller parameters given in Table 11.2. The temperature controllers have 1 min deadtimes in the loops. Reflux ratios are held constant in each column (3.44 in the extractive column and 1.61 in the methanol column). 

A two-column extractive distillation process is used with dimethyl sulfoxide (DMSO) as the solvent. The two components are separated into 99.5 mol% pure products leaving in the distillate streams from two distillation columns. The solvent flowrate that minimizes total energy consumption is determined. A control stmeture that is capable of handling very large dismrbances in throughput and feed composition is developed. The control of two tray temperatures in the extractive column is found to be necessary to handle feed composition dismrbances. 

The design and control of a maximum-boiling azeotropic system has been studied in this chapter. Extractive distillation is shown to be capable of producing quite pure products. A conventional control structure is developed that provides effective disturbance rejection for both production rate and feed composition changes. Dual temperature control is required in the extractive column in order to handle feed composition disturbances. 

In addition. Cooper et al. [91] discussed the use of a Raman analyzer to provide feedback and feed-forward data on a number of chemical manufacturing processes originating from crude oil, where several of the production steps involved a distillation separation again, in these examples, the Raman analyzer was positioned at the outlet to the distillation tower. They claim from their work that a Raman analyzer would be useful for monitoring and controlling aromatic extraction, liquid paraffin aromatization, and the production of cumene, cyclohexane from benzene, ethylbenzene, xylene isomers, dimethyl terphthalate, and styrene. It should also be noted that in all these processes, at least one and in several cases multiple distillation columns are involved. 

One must take care in determining the number of steady-state and dynamic degrees of freedom for more complex columns. Tyreus [4] describes the determination of the degrees of freedom for an extractive distillation system and for an azeotropic column with an entrainer. In the case of an extractive distillation system, recycle streams reduce the dynamic degrees of freedom through an increase in the steady-state degrees of freedom if the recycle contains a component that neither enters nor leaves the process. Also, if it is important to control the inventory of a trapped component, such as an entrainer for azeotropic distillation, then it is necessary to provide extra control valves to account for the loss of degrees of freedom. The loss comes from the addition of a side stream. 

The bottoms water stream is combined with a water makeup stream (46.2 kmol/h) and cooled before being fed as the extraction water feed to C2. The distillate methanol stream is combined with the fresh feed of methanol (232 kmol/h), and the total is split between the prereactor and the reactive distillation column. Figure 14.6 displays the temperature profile. There are two areas where the temperature changes from tray to tray are fairly large, which suggests either one-end temperamre control or two-temperature control may be possible. 

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