Distillation columns component balance

Distillation Columns. Distillation is by far the most common separation technique in the chemical process industries. Tray and packed columns are employed as strippers, absorbers, and their combinations in a wide range of diverse appHcations. Although the components to be separated and distillation equipment may be different, the mathematical model of the material and energy balances and of the vapor—Hquid equiUbria are similar and equally appHcable to all distillation operations. Computation of multicomponent systems are extremely complex. Computers, right from their eadiest avadabihties, have been used for making plate-to-plate calculations. 

First, the old standby methods of checking the overall individual component balances and checking dew and bubble points will help verify distillate and bottoms concentrations. The total overhead (distillate plus reflux) calculated dew point is compared to the column overhead observed temperature and the bottoms calculated bubble point is compared to the column bottom observed temperature. If the analyses are not felt to be grossly in eiTor. the following method wfill also prove very helpful. 

For a simple distillation column separating a ternary system, once the feed composition has been fixed, three-product component compositions can be specified, with at least one for each product. The remaining compositions will be determined by colinearity in the ternary diagram. For a binary distillation only two product compositions can be specified independently, one in each product. Once the mass balance has been specified, the column pressure, reflux (or reboil ratio) and feed condition must also be specified. 

The component balance equations for reflux drum, column enriching section, feed plate, column stripping section and the column base and reboiler are very similar to those derived previously for the binary distillation example CONSTILL, but are now expressed in multicomponent terms as described in Section 3.3.3.4. 

Interaction is unavoidable between the material and energy balances in a distillation column. The severity of this interaction is a function of feed composition, product specification, and the pairing of the selected manipulated and controlled variables. It has been found that the composition controller for the component with the shorter residence time should adjust vapor flow, and the composition controller for the component with the longer residence time should adjust the liquid-to-vapor ratio, because severe interaction is likely to occur when the composition controllers of both products are configured to manipulate the energy balance of the column and thereby "fight" each other. 

In alternative (a) pure products are obtained in each column. Since the relative volatility diminishes with the pressure, higher reflux is needed in the H P column. The balance of duties can be obtained by adjusting the split of the feed. Roughly speaking, by double-effect distillation the energy consumption is divided by two. In alternative (b) there is a large temperature difference between top and bottom that may be exploited by a sloppy split in the HP column with the heavy component, while in alternative (c) this is done with the light component. Alternative (c) is the best for the present case study since it allows a lower temperature of the hot utility. 

Summing up, if the inventory of the main components can be handled by local control loops, the inventory of impurities has essentially a plantwide character. The rates of generation, mainly in chemical reactors, and of depletion (exit streams and chemical conversion), as well as the accumulation (liquid-phase reactors, distillation columns and reservoirs) can be balanced by the effect of recycles in order to achieve an acceptable equilibrium state. Interactions through recycles can be exploited to create plantwide control structures that are not possible from a standalone unit viewpoint. 

The converged mass and heat balances and the exergy loss profiles produced by the Aspen Plus simulator can help in assessing the thermodynamic performance of distillation columns. The exergy values are estimated from the enthalpy and entropy of the streams generated by the simulator. In the following examples, the assessment studies illustrate the use of exergy in the separation sections of a methanol production plant, a 15-component two-column... 

In the second control structure (Fig. 2.11b), which does work, the fresh feed makeup of the limiting reactant (.F0B) is flow-controlled. The other fresh feed makeup stream (FCvl) is brought into the system to control the liquid level in the reflux drum of the distillation column. The inventory in this drum reflects the amount of A inside the system. If more A is being consumed by reaction than is being fed into the process, the level in the reflux drum will go down. Thus this control structure employs knowledge about the amount of component A in the system to regulate this fresh reactant feed makeup to balance exactly the amount of B fed into the process. 

This effect is best explained by a simple illustration. Suppose we feed a column with 50 mol/h of A and 50 moVh of B, and A is the more volatile component. Suppose the distillate contains 49 mol/h of A and 1 mol of B, and the bottoms contains 1 mol/h of A and 49 mol/h of B, Thus the distillate flowrate is D = 50 mol/h and the purity of the distillate is xDA = 0.98. Now we attempt to fix the distillate flowrate at 50 mol/h and also hold the distillate composition at 0.98 mole fraction A. Suppose the feed composition changes to 40 mol/h of A and 60 mol/ h of B. The distillate will now contain almost all of the A in the feed (40 mol/h), but the rest of it (10 mol/h) must be components. Therefore the purity of the distillate can never be greater than xD A = 40/50 = 0.80 mole fraction A. The overall component balance makes it impossible to maintain the desired distillate composition of 0.98. We can go to infinite reflux ratio and add an infinite number of trays, and distillate composition will never be better than 0.80. 

If we use distillate flow from the recycle column to control overhead receiver level, then we see that all of the flows around the liquid recycle loop are set on the basis of level. This violates our original statement to fix the toluene recycle flow. We are then left with the question how to control the overhead receiver level in the recycle column. We can use the fresh makeup toluene feed to control this level since it represents the toluene inventory in the process. Such a scheme limits large flowrate changes to the refining section and automatically ensures the component balance for toluene. 

If one or more unit operations have been given infeasible specifications, then the flowsheet will never converge. This problem also occurs with multicomponent distillation columns, particularly when purity specifications or flow rate specifications are used, or when nonadjacent key components are chosen. A quick manual mass balance around the column can usually determine whether the specifications are feasible. Remember that all the components in the feed must exit the column somewhere. The use of recovery specifications is usually more robust, but care is still needed to make sure that the reflux ratio and number of trays are greater than the minimum required. A similar problem is encountered in recycle loops if a component accumulates because of the separation specifications that have been set. Adding a purge stream usually solves this problem. 

The model of the distillation column used throughout the paper is developed by [10], composed by the mass, component mass and enthalpy balance equations used as basis to implement a SIMULINK model (figure 1) which describes the nonlinear column dynamics as a 2 inputs Q, Lvo/) and 2 output (Ad, Ab ). Implementations details for the overall column dynamics are given in [11]. 

Steps 5 and 6 Applied to Units 2 and 3 Combined The subsystem boundary for analysis encompasses the distillation column and the condenser. Two components exist, and two values of the flow streams are unknown, D and B hence a unique solution exists (if the mass balances are independent, as they are). 

Steps 5 and 6 All the compositions are known and three stream flows, D, IV, and R, are unknown. No tie components are evident in this problem. Two component material balances can be made for the still and two for the condenser. Presumably three of these are independent hence the problem has a unique solution. We can check as we proceed. A balance around either the distillation column or the condenser would involve the stream R. An overall balance would involve D and W but not R. 

The diagram in Fig. P2.58 represents a typical but simplified distillation column. Streams 3 and 6 consist of steam and water, and do not come in contact with the fluids in the column which contains two components. Write the total and component material balances for the three sections of the column. How many independent equations would these balances represent ... 

In some processes, such as distillation columns or reactors, heat transfer and enthalpy changes are the important energy components in the energy balance. Work, potential energy, and kinetic energy are zero or quite minor. However, in other... 

Given F or D and the composition of the three streams (v/j, Zp and Xp, we can determine the two unknown flowrates just by doing a total mole balance and a component mole balance over the entire distillation column (the blue system in and figure at right). 

A distillation column, for instance, would be modeled with a column section for the stripping section, a column section for the rectifying section, and single equilibrium stages for the feed tray, the condenser, and the reboiler. In order to solve the distillation column separation equations as one unit, two sets of Equation 12.33, each with the appropriate stripping factors for the corresponding section, would have to be solved simultaneously along with a component balance around the feed tray, the condenser, and reboiler equations. Such a solution does exist for conventional distillation and for certain extraction problems (Smith and Brinkley, 1960). 

The following equations represent material balances on the liquid flow rates of one of the components at the live stages of a distillation column. Use the Thomas algorithm to solve for L L, L, L, and L, the liquid flow rates of the component on trays 1, 2, 3, 4, and 5. These flow rates could constitute part of the initial estimates required for starting the rigorous column solution. The feed rate is 45 kmol/h, sent to stage 3. 

In addition, a component mass balance can be performed which describes the net flowrate of component A in the upper section for any tray. For clarification, trays in distillation columns are typically numbered from the top to the bottom, such that the top tray is stage 1 and the bottom is stage m. Thus, liquid flows down the column from tray n to tray n + 1... 

The calculation procedures [the 0 method, Kb method, and constant composition method] developed in Chap. 2 for conventional distillation columns are applied to complex distillation columns in Sec. 3-1. For solving problems involving systems of columns interconnected by recycle streams, a variation of the theta method, called the capital 0 method of convergence is presented in Secs. 3-2 and 3-3. For the case where the terminal flow rates are specified, the capital 0 method is used to pick a set of corrected component-flow rates which satisfy the component-material balances enclosing each column and the specified values of the terminal rates simultaneously. For the case where other specifications are made in lieu of the terminal rates, sets of corrected terminal rates which satisfy the material and energy balances enclosing each column as well as the equilibrium relationships of the terminal streams are found by use of the capital 0 method of convergence as described in Chap. 7. 

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