Saturday, March 7, 2009

De Dietrich distillation

Distillation is an operation whereby the vaporization of a liquid mixture yields a vapor phase containing more than one component, and it is desired to recover one or more of these components in a pure state. This is distinct from evaporation which is not discussed in this section.

Distillation as such is a major unit operation in the chemical processing industry for the purification or separation of liquid mixtures. Other principle unit operations of separation are evaporation and extraction. Distillation and evaporation are often considered simultaneously because of their use of heat to achieve separation.

Distillation in practical operations can be effected on either a continuous or a batch mode of operation. A simple reboiler and condenser arrangement with an unpacked column provides one theoretical stage of separation. This is suitable for mixtures containing a volatile solvent with non volatile impurities. Columns which are packed or fitted with trays provide several theoretical stages of separation. Such arrangements are valuable for liquid mixtures in which both components have similar relative volatilities.

Although the effectiveness of the distillation operation is dependent upon such theoretical considerations as the relative volatility of the components. In practice, the design of a distillation unit is extremely important.

Borosilicate glass equipment has been used successfully for many years in the field of distillation operations. Many units have been installed using the distillation operation, especially for the recovery of solvents. However, details on this particular aspect are not covered in this section.

Glass distillation columns are normally filled with packing materials made of borosilicate glass, but other packing materials can also be supplied. Cooling arrangements for the distillate can use either shell and tube or coil type heat exchangers. Specific advice on both the optimum packing material and the method of operation can be given by our Chemical Engineers.

Glass columns can vary in diameter from 80 to 1000 mm and columns have been erected to heights of up to 30 meters. Glass distillation units can operate at atmospheric pressure or at high vacuum, using special low pressure loss packings.

Distillation Unit with Coil Type Condenser

The distillation arrangement with descending type coil condenser is one of the simplest types of condenser arrangements and includes a total condenser, product cooler and vent condenser.

By combining glass process plant equipment with other materials, we can offer well proven units in a a variety of sizes. Glass reaction vessels are available up to 400 liters in capacity and, where larger units are required, glass-lined and stainless steel vessels are commonly used. Coil type condensers have surface areas from 0.2m2 up to 12m2. Using cooling water at 20&degC these condensers have heat transfer coefficients of up to 250 kcal/m2/h&degC.

This up and over type distillation unit is suitable for use under vacuum conditions and is an ideal facility for reactions involving total reflux.

Distillation Unit with Shell and Tube Condenser

This distillation arrangement is virtually identical to the previous arrangement. The main difference is that shell and tube heat exchangers have been incorporated instead of coil condensers. Shell and tube type condensers have surface areas from 3m2 to 26m2. Using cooling water at 20&degC these condensers have heat transfer coefficients of up to 900 kcal/m2h&degC.

The most important feature of this arrangement is the low installation height.

Distillation Unit with Phase Separation

This distillation arrangement illustrates the use of a horizontal separation vessel in a distillation unit.

In this arrangement, the light phase flows back into the boiler as the reflux and the heavy phase flows through a condenser into the receiver as the product. It is easy to reverse the system to enable the heavy phase to become the reflux.

The arrangement can be operated under vacuum and the illustration shows suitable receivers.

Batch Distillation Column

Batch distillation columns always have a reboiler vessel sized to accommodate the entire batch of the material to be distilled.

For small batch operations, glass vessels up to 200 liters in capacity are normally used and, where larger batches are required, glass-lined or stainless steel vessels are available.

Glass distillation columns are available from 80 to 100 mm in nominal bore and are ideal for operation under both atmospheric and vacuum conditions.

Columns of this nature are operated under conditions of either fixed reflux ratio or variable reflux ratio. This illustration shows a batch distillation column which includes a total condenser, product cooler and the facility for providing reflux at the top of the column by means of the control valve. Reflux timers are also available.

Continuous Distillation Column

Continuous distillation columns always have this same basic construction. For the reboiler, a circulatory evaporator is often selected and, as pre-heaters, our HEB type boilers are ideal.

The main point to consider in continuous distillation columns is the automatic control system. Many standard systems are available and, in this illustration an electro-magnetically controlled reflux divider with timer is shown.

Glass distillation columns are available from 80 to 1000 mm in nominal bore and ideal for operation under both atmospheric and vacuum conditions.

Advanced Process Modelling for separation systems

Optimal process and control system design for an azeotropic distillation system

Based on the report "Simultaneous Design and Control of the Shell Azeotropic Distillation System using Mixed-Integer Dynamic Optimization" by Vikrant Bansal and Roderick Ross, Centre for Process Systems Engineering and PSE (2001). Please note that for confidentiality reasons, exact details of the process and the control and design alternatives mentioned are not provided.

Summary

Mixed-integer and dynamic optimisation were used to select the best of several proposed control schemes for an existing two-column coupled distillation unit operating under low and high-frequency disturbances. This led to a significant improvement in the controllability of the unit. Mixed integer Optimisation (MIO) was then used to improve the design of future units by selecting optimal feed and draw tray locations, while simultaneously optimising the column diameters. This identified significant improvements in capital and annual operating cost.

The project was performed on a Shell alcohol separation unit at Pernis in the Netherlands.

Process

The two-column distillation system forms part of the Shell alcohol separation train at Pernis in the Netherlands. Figure 1 shows the separations taking place. There is a ternary azeotrope at the top of Column 1. The unit is subject to regular low frequency disturbances resulting from changes in feed rate and high-frequency disturbances resulting from fluctuation in the feed composition.

Azeotropic distillation process

Figure 1. The two-column azeotropic distillation process

For many years the system had been very difficult to control. Good operation depends critically on the composition around the side draw tray in Column I being within a certain range. If the composition goes outside that range, Column II cannot operate correctly and the system enters into a sustained period of unsteady behaviour.

Study 1 - operational improvement on an existing unit

Several control schemes were proposed in order to improve the process behaviour under disturbance. These are shown in approximate form in Figure 2.

Modelling approach

The system was modelled in PSE's gPROMS package, using a detailed dynamic tray-by-tray model for the distillation columns. The model was tuned to existing plant operational data, using gPROMS' parameter estimation capabilities, in order accurately to quantify parameters such as tray efficiency and heat transfer coefficients.

Once a suitably accurate model of the process had been built, the existing control scheme was added and the model tested for response against various disturbances. Having established that the existing control scheme and tunings were not capable of restoring stability under these disturbances, a dynamic optimisation was performed (using the same input disturbances) in order to establish the optimal control tunings. This showed some improvements over the existing settings.

Azeotropic control scheme

Figure 2. Control scheme alternatives

Control scheme alternatives

Having proven that there was scope for improvement in the control system, several alterative proposed control schemes were added to the model. A dynamic optimisation was set up which included integer (discrete) selection between the alternative control schemes. This was configured to allow the dynamic optimisation to select the best control scheme to handle the disturbances, and to ensure that only one scheme was selected rather than a combination of the schemes.

Result

The results of the dynamic optimisation not only identified a better control scheme than the existing one, but provided optimal tunings (gain and integral action) for each of the controllers. The improved control scheme meant that for the first time it was possible to control the system properly during normal operation.

By including alternative optimisation variables in subsequent runs - for example, column diameter - it was shown that significant improvements could be made in the design of such plants that would lead to improved inherent controllability.

In addition to the results obtained here it would also be possible to use the dynamic model to, for example, optimise start-up procedure, or investigate optimal transition policy between different modes of operation.

Project 2 - process and control improvements for new design

Having built a detailed predictive model of the system, it was possible to use it to design more economic units for the future, with better controllability characteristics.

A particular feature of the project described here was the use of integer optimisation to determine the optimal locations of the Column I feed and side draw locations, while simultaneously optimising the values of other "continuous" variables such as the column diameters.

Objectives

The objective was to design the distillation system and its control system at minimum total annualised cost. It was necessary that:
  • the unit was capable of feasible operation over the whole of a given time horizon in the face of disturbances in the feed flow rate and composition
  • the solution satisfied composition specifications on the various column product streams
  • the column diameters calculated would be sufficient to avoid flooding in either column, and ensure that entrainment limits were observed.

This required the simultaneous determination of the optimal process design and the optimal control design, by calculating optimal values for the following integer (discrete) and continuous variables:

  • locations of the feed and draw-off trays in Column I (discrete decisions)
  • the column diameters, reboiler heat duties and flow rates of the draw-off stream to Column I and the return stream back to Column II (continuous decisions)
  • tuning parameters - gain and reset - of all control loops (continuous decisions).

In principle, other discrete decisions could also have been considered, such as the optimal return tray location in Column I and the numbers of trays in the two columns. However the client asked for these to be excluded.

Modelling approach

The modelling approach was essentially the same as that used before, with the optimised control scheme from the previous project (above) implemented. However it was necessary to providing a number of potential locations for feed tray and side draw locations for Column I, shown in Figure 3, from which the optimiser could choose the optimal combination.

The optimisation used an economic objective function that includes both capital and operating cost, to find the lowest annualised cost for the system.

Modelling approach

Results

A summary of the results is shown in Table 1. It can be seen that the optimiser has changed the feed and draw tray location significantly, and has altered the relative sizes of Column I and II.

The total annualised cost of the re-designed unit is 18% lower than the existing Shell design. These savings are achieved by drawing more product B from Column I, resulting in a smaller Column I, with the size of Column II increased correspondingly. The reduced capital cost and reboiler duty for Column I easily outweighs the increases in these quantities for Column II.

The resulting optimal solution is about 18% cheaper than the existing design, indicating both the benefits of considering process design and process control simultaneously and the viability of using this MIO technology for solving complex, industrial problems.

The optimal solution was found on the fourth iteration of the integer optimisation loop. In fact the MIO algorithm was able to find three structures (second, fourth and eighth solutions) that are cheaper than the original structure. If optimisation had not been used, it would have been necessary to evaluate 68 alternative combinations in order to achieve this result.

Table 1 - existing vs. optimal design

Optimal Design table

The study showed that if such a simultaneous optimisation approach had been used when the distillation system was originally designed, not only would the total annualised cost of operation been substantially lower, but the operability difficulties experienced during its operation woul

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Methylene Chloryde, Isopropyl Alcohol (IPA), Acetone, Tryethylammine Diethylammine Start-up Pharmaceutical Distillation Plants

It's very well completed the start-up of CMG's solvents distillation plants for an anti-biotic pharmaceutical company in IRAN.
The plants are the following:
  • Continuous distillation of Methylene Chloryde
  • Continuous distillation of Isopropyl Alcohol (IPA)
  • Batch/continuous distillation of Acetone (95%)
  • Continuous distillation of Acetone (99%)
  • Recovery-distillation of TEA-DEA (Try ethyl ammine, Di ethyl ammine)
  • Dehydration of Methylene Chloryde
In the following picture the distillation units



In order to design the distillation plants, CMG used the most upgrade distillation design program, and the obtained results of the laboratory tests made by the research centre Polo Tecnologico La Magona. CMG has been part of this technological research centre
The CMG supply is a turn-key comprising distillation units, complete engineering of the storage tanks farm, of the piping and of complete carpentry and pipe-racks. Following the picture of one of the 3d drawing of the whole project.



The goodness of the products of distillation units is confirmed by the client with its request for new distillation unit for other mixtures of solvents.
Following some technical data of the distillation units:
1- Methylene Chloryde Unit
Specification of Feed Flow-rate 1000 Kg/h, recovery 95-99%
Composition of FEED, Distillate, Bottom column product (%w)

FEEDDISTILLATEBOTTOM COLUMN
METHYLENE CHLORYDE87,7499,774.34
METHANOL0,060,0219,56
PIVALIC ACID8,7-----60,79
IPA1,20,053,28
DMAC1,1-----19,56
WATER0,50,160,89
OTHER0.6-----10,77

The distillation is the extractive type. The most important obtained results are: distillate Methylene Chloryde without Methanol (just traces) and the elimination of bad odour from the bottom column's product.The complete dehydration of Methylene Chloryde is done using molecular sieve and regeneration by hot nitrogen.

2- Distillation di Isopropil-alcol (IPA)
Specification :Feed Flow-rate 1100 Kg/h, recovery del 97%
Composition of FEED, Distillate, Bottom column product (%w)


FEEDDISTILLATEBOTTOM COLUMN
IPA38.2979.491.8
TEA1.470.027.43
WATER5813.6486.6
MIBK2.26.50.2
OTHER0.50.13.97

3- Distillation di Acetone
Specification : Feed Flow-rate 800 Kg/h, recovery 95-99%
Composition of FEED, Distillate, Bottom column product (%w)

FEEDDISTILLATEBOTTOM COLUMN
ACETONE53.1599.30.99
WATER44.360.4894.17
DMF1,26----3.18
METHANOL0.950.070.94
OTHER0.280.150.72

The first distillation unti is a extractive distillation, it permits to split Methanol from Acetone and other solvents. The second distillation unit is a "classical" continuous distillation plant. In order to have the maximum operative flexibility the first distiallation plant can be utilized either in batch configuration or in continuous configuration.

4- Recovery and distillation of TEA - DEA (try ethyl ammine, Diethyl ammine)
Specification: INITIAL FEED =3000-3500 KgComposition of FEED, Distillate, Bottom column product (%w)

FEEDSolvents after 1° distillationBefore the final distiallation
TEA25.410.1391.73
IPA1559.74.07
ACETONE18.75---0.17
MIBK14.2110.270.14
METHANOL0.591.09-----
DEA0.12---0.17
WATER25.3128.753.14
OTHER0.610.610.04

Following the process description of TEA, it's a batch procedure.
  • The salification of initial solution with HCl
  • Phase separation: the light (organic) phase is sent to thermal oxidation, the acqueous phase is storaged for the successive phase
  • Distillation to remove residual solvents;
  • Basification of the residue of distillation using NaOH solution;
  • Phase separation at about to limit the solubility of water in TEA: heavy (aqueous) phase is sent to thermal oxidation.
  • Storing of the organic phase, so that we can accumulate enough amount to be distilled;
  • Distillation to recover TEA.
more info see : http://www.cmgimpianti.com