Reactor Temperature Control Using Feed Manipulation

4 REACTOR TEMPERATURE CONTROL USING FEED MANIPULATION 

When reactor capacity is limited by heat removal, an often-recommended control structure is to run with maximum coolant flow and manipulate feed flowrate to control reactor temperature (Tr F0 control). This control scheme has the potential to achieve the highest possible production rate. However, if the feed temperature is lower than the reactor temperature, the transfer function between temperature and feed flowrate contains a positive zero, which degrades dynamic performance, as we demonstrate quantitatively in this section. The choice of a control structure for this process presents an example of the often encountered conflict between steady-state economics and dynamic controllability. 

The reactor is the jacket-cooled CSTR with an irreversible, exothermic, liquid-phase reaction A — B, which was considered in Section 3.1. In that section the flowrate of the cooling water Fj to the jacket was the manipulated variable for the reactor temperature controller (TR — Fj control). In this section we explore the use of the flowrate of the fresh feed F() to control reactor temperature (TR — F0 control). 

The flowrate of the cooling/heating medium is usually the manipulated variable that is changed by a reactor temperature controller, either directly or through a 

Greg Shinskey1 presents a practical discussion of some of the advantages and problems of this control structure. The use of a valve position control (VPC) structure is recommended. A reactor temperature controller sets the coolant valve position. Then a VPC looks at the position of the coolant valve and adjusts the flowrate of the reactor feed to keep the coolant valve near its wide open position. 

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An exothermic reaction involving two reactants is run in a semi-continuous reactor. The heat evolution can be controlled by varying the feed rate of one component This is done via feedback control with reactor temperature measurement used to manipulate the feed rate. The reactor is cooled by a water jacket, for which the heat transfer area varies with volume. Additional control could involve the manipulation of the cooling-water flow rate. 

Two reactor temperature controllers are now used. The first manipulates the hot and cold streams used to heat or cool the reactor. The second controller manipulates the feed flowrate. More details about this control structure are presented in Section 4.2. 

Two temperature controllers are used. The first manipulates the flowrate of the A feed. A 45 min ramp in this reactor temperature controller is used with Kci = 0.5 and Tj2 = 20 min. Two 30-s lags are included in the loop. The span of the temperature transmitter is 50 K, and the maximum flowrate FA0 of the reactant A is 0.004 m3/s. The second temperature controller setting the flowrates of the hot and cold streams to the jacket is proportional-only with a 330 K setpoint and a gain of 0.05. The maximum cold water and hot water flowrates are 0.005 and 0.002 m3/s, respectively. 

As stated in the previous section, the major reactant feed was chosen as the manipulated variable. In the trial this feed was subjected to a pseudo-random binary sequence (PRBS) signal in an open loop operation of the process. The results of the trial, plotted in Fig. 2, show a strong -- but delayed -- cross-correlation between the manipulated feed rate and the reactor temperature. Using techniques described by Box and Jenkins (2), a transfer function relating the manipulated variable to the control variable of interest can be developed. The general form of this transfer function is... 

The most common continuous emulsion polymerization systems require isothermal reaction conditions and provide for conversion control through manipulation of initiator feed rates. Typically, as shown in Figure 1, flow rates of monomer, water, and emulsifier solutions into the first reactor of the series are controlled at levels prescribed by the particular recipe being made and reaction temperature is controlled by changing the temperature of the coolant in the reactor jacket. Manipulation of the initiator feed rate to the reactor is then used to control reaction rate and, subsequently, exit conversion. An aspect of this control strategy which has not been considered in the literature is the complication presented by the apparent dead-time which exists between the point of addition of initiator and the point where conversion is measured. In many systems this dead-time is of the order of several hours, presenting a problem which conventional control systems are incapable of solving. This apparent dead-time often encountered in initiation of polymerization. 

All of these systems have some common control loops. The system pressure is controlled by manipulating the fresh feed of A (F0A). The concentration controller with ratio control is used to control reactor inlet gas composition by manipulating the fresh feed of B (F0B). Bypassing (Fhy) around the FEHE is used to control gas mixture temperature Tmix. Reactor inlet temperature (Tin or T ) is controlled by manipulating the furnace heat input QF. The setpoints of these two temperature controllers are the same, and the controller output signals are split-ranged so that bypassing and furnace heat input cannot occur simultaneously. 

One of the three options considered in Chapter 5 was a cooled tubular reactor with a coolant temperature that is the same down the length of the bed. With this type of system, the temperature of the coolant can be used as the manipulated variable to control some variable in the reactor. We will illustrate the control of the peak temperature by using several temperature measurements at different locations and selecting the highest to feed to a temperature controller as the process variable PV signal. The output signal OP of this controller will be the coolant temperature. 

Note that the changes in production rate occur more quickly when the toluene recycle flowrate handle is used, compared to the reactor inlet temperature handle. Fresh feed rates of toluene and hydrogen change more quickly, as does benzene product flowrate. So if rapid transitions in production rate are important, toluene recycle flowrate manipulation is better than reactor inlet temperature manipulation. If tight product quality control is more important, the opposite is true. 

Monomer conversion can be adjusted by manipulating the feed rate of initiator. If on-line MMD is available, initiator flow rate or reactor temperature can be used to adjust MM [59]. A more complex approach involves adjusting initiator feed rate to control monomer conversicm, while by-passing part of the water and monomer around the first reactor in a train to control the PSD [60,61]. Direct control of the surfactant feed rate, based on surface tension measurements also can be used. Optimal policies for changing product grades without shutting down the reactor train are desirable, but what little is done in this area is based on operating experience rather than on optimization approaches. 

Due to the shift reaction features, for a given reactor feed stream composition and flow rate, the temperature of the feed stream can be manipulated in order to maximize CO conversion. Consequently, if the industrial reactor has in-line/on-line composition analyzers as well as flow rate measurement instrumentation, on-line optimization can be successfully implemented. The problem is that this kind of in-line/ on-line instrumentation is not only expensive but also may need continuous and excessive calibration and maintenance. Therefore, some question arises can other real-time measurements be used for CO conversion How does temperature influence CO conversion correlated This is an important issue as temperature measurements are usually reliable, accurate, real-time and low cost. Moreover, fixed bed reactors can have temperature instruments installed along the reactor length, providing a temperature profile. Towards this, a novel and alternative approach will be presented in order to overcome this issue, focusing on CO conversion control. 

Monomer conversion can be adjusted by manipulating the feed rate of initiator or catalyst. If on-line M WD is available, initiator flow rate or reactor temperature can be used to adjust MW [38]. In emulsion polymerization, initiator feed rate can be used to control monomer conversion, while bypassing part of the water and monomer around the first reactor in a train can be used to control PSD [39,40]. Direct control of surfactant feed rate, based on surface tension measurements also can be used. Polymer quality and end-use property control are hampered, as in batch polymerization, by infrequent, off-line measurements. In addition, on-line measurements may be severely delayed due to the constraints of the process flowsheet. For example, even if on-line viscometry (via melt index) is available every 1 to 5 minutes, the viscometer may be situated at the outlet of an extruder downstream of the polymerization reactor. The transportation delay between the reactor where the MW develops, and the viscometer where the MW is measured (or inferred) may be several hours. Thus, even with frequent sampling, the data is old. There are two approaches possible in this case. One is to do open-loop, steady-state control. In this approach, the measurement is compared to the desired output when the system is believed to be at steady state. A manual correction to the process is then made, based on the error. The corrected inputs are maintained until the process reaches a new steady state, at which time the process is repeated. This approach is especially valid if the dominant dynamics of the process are substantially faster than the sampling interval. Another approach is to connect the output to the appropriate process input(s) in a closed-loop scheme. In this case, the loop must be substantially detuned to compensate for the large measurement delay. The addition of a dead time compensator can... 

Just as we approached reactor control in Chap. 4, we will start by exploring the open-loop effects of thermal feedback. Consider Fig. 5.19, which shows an adiabatic plug-flow reactor with an FEHE system. We have also included two manipulated variables that wall later turn out to be useful to control the reactor. One of these manipulated variables is the heat load to the furnace and the other is the bypass around the preheater. It is clear that the reactor feed temperature is affected by the bypass valve position and the furnace heat load but also by the reactor exit temperature through the heat exchanger. This creates the possibility for multiple steady states. We can visualize the different... 

Since the production rate is controlled by the toluene fresh feed rate, the conversion in the reactor must be kept constant. This can be achieved by manipulating the heat duty to the furnace. As measurement we can use either the temperature of the reactor inlet stream or the temperature in the reactor. The first is preferred because it allows us to take control action earlier (Figure 25.7). 

The following manipulated variables are available toluene feed (Fi), hydrogen feed (F2), gas recycle (Fr), purge (Fp) and furnace duty (q h). Furnace duty may be used to control the reactor inlet temperature. The setpoint of this control loop may be coupled, in a cascade manner, with other variable from the previous list. 

Make sure that the overall component balances for all chemical components can be satisfied. Light, heavy, and intermediate inert components must have a way to exit the system. Reactant components must be consumed in the reaction section or leave the system as impurities in product streams. Therefore, either reaction rates (temperature, pressure, catalyst addition rate, etc.) must be changed or the flow rates of the fresh feed makeup streams must be manipulated somehow. Makeups can be used to control compositions in the reactor or in recycle streams, or to control inventories that reflect the amount of the specific components contained in the process. For example, bring in a gaseous fresh feed to hold the pressure somewhere in the system, or bring in a liquid fresh feed to hold the level in a reflux drum or column base where the component is in fairly high concentration (typically in a recycle stream). 

An optimal predictive controller was developed and implemented to allow for maximization of monomer conversion and minimization of batch times in a styrene emulsion polymerization reactor, using calorimetric measiuements for observation and manipulation of monomer feed rates for attainment of control objectives [31]. Increase of 13% in monomer conversion and reduction of 28% in batch time were reported. On-line reoptimization of the reference temperature trajectories was performed to allow for removal of heater disturbances in batch bulk MMA polymerizations [64]. Temperature trajectories were manipulated to minimize the batch time, while keeping the final conversion and molecular weight averages at desired levels. A reoptimization procediue was implemented to remove disturbances caused by the presence of unknown amounts of inhibitors in the feed charge [196]. In this case, temperatiue trajectories were manipulated to allow for attainment of specified monomer conversion and molecular weight averages in minimum time. 

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