Tray temperature control

On many columns tray temperature control offers a method which, although not as accurate as an analyser, provides a degree of composition control and overcomes these problems. It is not necessarily a replacement for a higher level of composition control if both are feasible then, as we shall show later, they can operate in conjunction. Tray temperature control works on the principle that hquid on the trays is at its bubble point. Bubble point is related to composition and so fixing the bubble point provides some level of composition control. As we have seen cut can also be expressed as temperature and so controlling tray temperature helps maintain cut. 

In theory we should control the temperature at the point at which the product is withdrawn from the column. However there are several reasons why this may not be practical. Firstly, the liquid may not be homogeneously at its bubble point. This is likely to be the situation close to the top tray since reflux is often subcooled. It can also occur at the base of the column if the vapour from the reboiler is superheated. Secondly, in pseudo-binary columns, the relationship between composition and bubble point also depends on the proportion of non-key components in the product The greatest proportion of LLK will occur at the top of the column, and of HHK at the bottom. Thus, if non-key composition varies, these are the regions most prone to inaccuracy. Finally, particularly on high purity columns, the temperature may not be sensitive to changes in composition. 

Similarly the relationship between temperature and its MV may be nonlinear which will give mning problems for the temperature controller. We will address this later when completing the exercise of checking the suitability of a chosen tray temperature. 

The process of selecting tray temperature(s) for composition control is usually completed using a simulation of the column, rather the real column. Properly it should be done before the column design is frozen. On columns, already bmlt, that have multiple tray temperatures installed, it is possible to execute the plant tests required but they will be very disruptive to the operation. 

The column at this stage of the control design wdl have two MVs remainmg for use to control the composition of both products. Which these variables are depends on the choice of level control configuration. If the material balance scheme is in place then, usually, distillate flow and reboil are available or, less usually, bottoms flow and reflux. If the energy balance scheme is in place then reboil and reflux are available. 

Let us first discuss using a temperature controller to maintain a tray temperature in the column. Looking at the temperature profile in Aspen Plus, we see that Stage 9 displays a fairly steep slope. Its temperature is 337.36 K. 

The program is run to make sure everything works okay without a lag or a deadtime in the loop. Now we back up and insert a deadtime element on the flowsheet between the column and the temperature controller. The reason for installing the controller initially without the deadtime element is to avoid initialization problems that sometime crop up if you attempt to install the deadtime and the controller all in one shot. 

Start test Fim8hteBt Cancel test] Help  

CONVERTING FROM STEADY-STATE TO DYNAMIC SIMULATION 

To start the test, click the Run button at the top of the screen and click the Start test button on the Tune window. To be able to see the dynamic responses, click the Plot button at the top of the controller faceplate. 

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Figure 11.7 gives results for a 5-minute shutdown of the distillation column feed pump. The flowrate of bottoms from the column drops quickly for about 20 minutes. The column tray temperature controller recovers in about 1 hour. Distillate and bottoms compositions are still changing after 5 hours. 

Tray temperatures correlate very well with product compositions for many distillation columns therefore, inferential control of distillation product composition is a widely used form of inferential control. Figure 15.58 shows the arrangement for inferential temperature control of the bottoms product composition for this column. Note that the tray temperature controller is cascaded to a flow controller. 

We want to compare tray temperature control with two types of composition control. In both, the composition of the distillate propane product is measured directly and controlled at 2 mol% isobutane impurity. The first type is direct composition control in which a single PI controller is used with reboiler heat input manipulated. The second type uses a cascade composition-to-temperature control structure. 

The tray temperature controller is the secondary (slave) controller. It is set up in exactly the same way as we did in the previous section. It looks at tray temperature (Stage 9) and manipulates reboiler heat input. However, its SP is not fixed. The SP signal is the output signal of the composition controller, which is the primary (master) controller. 

The control of partial condenser columns is more complex than total condenser columns because of the interaction among the pressure, reflux-drum level, and tray-temperature control loops. Both pressure and level in the reflux dmm need to be controlled, and there are several manipulated variables available. The obvious are reflux flow, distillate flow, and condenser heat removal, but even reboiler heat input can be used. In this section, we explore three alternative control structures for this type of system, under two different design conditions (1) a large vapor distillate flow rate (moderate RR) and (2) a very small vapor distillate flow rate (high RR). 

The tray temperature controllers are tuned by inserting a 1 min deadtime in the loop and using the relay-feedback test to determine the ultimate gain and ultimate frequency. Then, the Tyreus-Luyben settings are used. Table 8.2 gives the tuning constants. 

In the last three chapters, we have developed a number of conventional control structures dual-composition, single-end with RR, single-end with rellux-to-feed, tray temperature control, and so on. Structures with steam-to-feed ratios have also been demonstrated to reduce transient disturbances. Four out of the six control degrees of freedom (six available valves) are used to control the four variables of throughput, pressure, reflux-drum level, and base level. Throughput is normally controlled by the feed valve. In on-demand control structures, throughput is set by the flow rate of one of the product streams. Pressure is typically controlled by condenser heat removal. Base liquid level is normally controlled by bottoms flow rate. 

The other two loops, a tray temperature controller (TC) and a pressure controller, are different in the two cases. 

Tray temperature control is used in most distillation columns to infer product composition, but changes in pressure on the control tray can adversely affect the estimation of composition. Pressure is typically controlled in the condenser, not on the control tray, so changes in vapor flow rates will change tray pressure due to changes in tray pressure drops. Pressure-compensated temperature control was proposed over four decades ago to solve this problem. Measurements of both temperature and pressure on the control tray are used to estimate composition. The method has been qualitatively described in many practical distillation control books, but the author is not aware of any quantitative evaluation of its effectiveness that has appeared in the open literature. 

In the following, we will investigate the proper overall control strategy for the proposed design with the lowest TAC. Only a tray temperature control loop(s) will be used in the overall control strategy for wider industrial applications. 

The remaining manipulated variables for the heterogeneous azeotropic column are the OR flow and the reboiler duty, and the remaining manipulated variables for the preconcentrator/recovery column are the reflux flow and the reboiler duty. We will investigate the simplest overall control strategy first with only one tray temperature control loop in each column. 

Open-loop sensitivity analysis is used to determine the tray temperature control point for both the heterogeneous azeotropic column and the preconcentrator/recovery column. 

The other two manipulated variables, not used in the two tray temperature control loops, are the OR flow for the heterogeneous azeotropic column and the reflux flow for the preconcentrator/recovery column. Ratio control schemes are used so that these two manipulated variables can be adjusted in the face of disturbances. The OR flow is set to maintain a constant ratio to the feed flow of the heterogeneous azeotropic column, and the reflux ratio of the preconcentrator/recovery column is also maintained. The overall proposed control strategy is shown in Figure 8.26. 

The heterogeneous azeotropic column system using isobutyl acetate as the entrainer will be studied in detail in this section. The overall control strategy of this system will be developed in order to maintain bottom and top product specifications in spite of feed flowrate and feed composition changes. In the control strategy development, we will assume no online composition measurement is available. The composition control loops will be inferred by some tray temperature control strategy. This type of control strategy can easily be implemented in industry for wider applications. 

An optimum overall control strategy is also proposed for this column system to hold both bottom and top product specifications in spite of +10% feed rate and +10% feed water composition load disturbances. Several alternative control structures are compared using dynamic simulation. The proposed overall control strategy is very simple, requiring only one tray temperature control loop in the column. This simple overall control strategy can easily be implemented in industry. 

The above control strategy leaves two reboiler duties that can be used in a tray temperature control loop in each column. Figures 10.22 and 10.23 display the results of open-loop... 

The most important control loop in this operation is a tray temperature controller at the middle stage of the column. The temperature is controlled at the average temperature of the minimum-boUing azeotropic temperature of the entrainer-water mixture and the normal boiling-point temperature of pure acetic acid. This will ensure that the top vapor of this colurrm will approach the binary azeotrope and the bottoms will approach pure acetic acid. The manipulated variable of this important temperature loop is the organic reflux valve. Thus the accumulation of the organic phase will be automatically adjusted throughout the batch mn. 

Step 2 Start the heat up of the column and put aU controllers in automatic mode (middle tray temperature control, top column pressure control, and decanter temperature control). Step 3 When the bottom composition reaches 0.991 (mole fraction of acetic acid product), stop heating and collect bottom, aqueous, and organic products. The reason for using a stop criterion of 0.991 instead of the specification of 0.990 is to compensate for the small amounts of material on the column trays that drain into the bottom. 

Figure 13.23 shows the dynamic response of the middle tray temperature control loop. Controller performance is satisfactory until time = 1.4 h. After that, the temperature of... 

Besides using the two alternative entiainers, the batch distillation of directly separating acetic acid and water without an entrainer was also simulated for comparison purposes. The operation and control strategy as in Skogestad et al. is used in the simulation. The concept of this batch distillation is to collect water in the reflux drum and acetic acid in the bottom. A similar middle tray temperature control loop is used with the setpoint value set at 109.02°C (the average of 118.01 and 100.02°Q. The batch was stopped when the product composition at the column bottoms is the same as the case with entrainer (0.990) for direct comparison. 

Vinyl acetate is the best entrainer for acetic acid dehydration using heteroazeotropic batch distillation system. There are only several operating variables needed to be set for this simple batch operating sequence to work. They are the setpoint of the middle tray temperature control loop at 91.73°C, the entrainer preloading amounts of 1.0 kmol, and the reboiler duty fixed at 0.16 GJ/h. In this section, the settings of these three operating variables will be altered to see if the proposed batch operation is robust enough. 

We first examine the effect of changing the setpoint of the middle tray temperature control loop. Because the desirable top vapor of this column is the vinyl acetate-water two component azeotrope (at 65.45°C) and the bottom is the pure acetic acid product (n.b.p. at 118.01°C), the setpoint value of this middle temperature control loop was set to be the average of the two temperatures at 91.73°C. Table 13.7 shows five simulation runs with this temperature setpoint altered firom 71.73 to 111.73°C in increments of 10°C. This is a very wide range of values for setting this setpoint. However, even with this wide range of... 

TABLE 13.7 Effect of Varying the Setpoint Value of the Middle Tray Temperature Control. 

When the reboiler duty is fixed at a high value (0.19 GJ/h), the component acetic acid is starting to boil-up and escape from the top of the column causing the middle tray temperature control loop to fail, even with total organic reflux. This series of dynamic simulation runs show that the reboiler duty is also not too sensitive. The reboiler duty with a range of 0.14 to 0.18 GJ/h aU gave acceptable separation performance with trade-off in total batch time and energy consumption. 

We can adopt a similar approach to tray temperature controllers on distillation columns. They provide some control of product composition because this correlates with the bubble point of the liquid. However, changing pressure changes this relationship. Figure 5.11 shows the effect pressure has on bubble point, in this case water, but all liquids show a similar behaviour. 

If a distillation tray temperature controller keeps the temperature constant as the pressure changes, the composition wiU move away from target. We can resolve this by using the pressure to condition the temperature measurement. The subject of pressure compensated temperatures is covered in full in Chapter 12. 

We have seen that we can define cut both in terms of product yield but also in terms of temperature. A tray temperature controller can be configured to manipulate whatever variable remains after the level controllers are configured. And so we can control cut whether we have selected either the energy balance or the material balance scheme. However, there are occasions when tray temperature is insensitive to changes in product composition -for example when separating components with very similar bubble points. Under these circumstances the cut control provided by the material balance scheme is advantageous. 

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