Control of Continuous Distillation Columns

GBH Enterprises, Ltd.

Process Engineering Guide:
GBHE-PEG-MAS-608

Control of Continuous Distillation
Columns
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Process Engineering Guide:

Control of Continuous
Distillation Columns

CONTENTS

SECTION

0

INTRODUCTION/PURPOSE

4

1

SCOPE

4

2

FIELD OF APPLICATION

4

3

DEFINITIONS

4

4

GENERAL DESCRIPTION OF A DISTILLATION COLUMN 5

5

REGULATORY CONTROL

7

5.1
5.2
5.3
5.4
5.5

Composition Control
Mass Balance Control
Design of Feedback Control Systems
Pressure and Condensation Control
Reboiler Control

7
12
15
17
25

6

DISTURBANCE COMPENSATION

32

6.1
6.2
6.3

Feed-forward Control
Cascade Control
Internal Reflux Control

32
36
37

7

CONSTRAINT CONTROL

38

7.1
7.2
7.3

Override Controls
Flooding
Limiting Control

39
39
40


8

MORE ADVANCED TOPICS

43

8.1
8.2
8.1
8.2
8.1
8.2

Temperature Position Control
Inferential Measurement
Floating Pressure Control
Model Based Predictive Control
Control of Side-streams
Extractive/Azeotropic Systems

43
44
45
46
47
50

9

REFERENCES

51

TABLES
1

SYMPTOMS OF IMBALANCE AND THE REGULATORY
VARIABLES

12

PRACTICAL LINKAGES BETWEEN CONTROL
(P, R, B, C) AND REGULATION VARIABLES
(h, r, d, b, c, v)

15

3

COMPOSITION REGULATION

16

4

COMPOSITION REGULATION - VERY SMALL FLOWS

17

2


FIGURES

1

DISTILLATION COLUMN SHOWING A SELECTION OF
POSSIBLE CONTROL AND REGULATION POINTS

4

2

SCHEMATIC OF TYPICAL SIEVE TRAY

6

3

COLUMN TOP AND BOTTOM TEMPERATURE CONTROL

8

4

TYPICAL TEMPERATURE PROFILE

9

5

PROFILE FOR A COLUMN WITH VERY PURE TOP
PRODUCT BUT MIXED COMPOSITION BOTTOM PRODUCT

9

PROFILE FOR A COLUMN WITH VERY PURE BOTTOM
PRODUCT BUT MIXED COMPOSITION TOP PRODUCT

10

TEMPERATURE PROFILE SHOWING MULTIPLE SENSITIVE
REGIONS

10

8

MASS BALANCE CONTROL

14

9(a)

REGULATION OF VAPOR PURGE

18

9(b)

REGULATION OF COOLANT

18

9(c)

FLOODED CONDENSER

18

10

A HOT-GAS BYPASS IS THE MOST COMMON MEANS OF
CONTROLLING AIR-COOLED CONDENSERS

21

CONDENSATION CONTROL SYSTEMS FOR OPERATION
UNDER VACUUM

22

CONDENSATION USING A BOILING REFRIGERANT IN
CONDENSER SHELL

23

CONDENSATION USING A BOILING REFRIGERANT IN
CONDENSER TUBES

24

REFLUX DRUM MOUNTED LEVEL WITH OR ABOVE
CONDENSER

24

6

7

11

12

13

14


15

REFLUX DRUM MOUNTED LEVEL WITH OR ABOVE
CONDENSER BUT WITH REDUCTION OF PRESSURE
DROP ROUND REFLUX LOOP

25

INSERTING A TUBE BUNDLE EITHER DIRECTLY IN THE
COLUMN BASE OR IN AN EXTERNALLY MOUNTED KETTLE

26

ALTERNATIVE ARRANGEMENTS OF THERMOSYPHON
REBOILERS

27

18

HEAT INPUT CONTROL

29

19

DIAGRAM SHOWING THE VALVE-POSITION CONTROLLER
(VPC) ADJUSTING THE TEMPERATURE SO THAT ONE
VALVE IS ALMOST FULLY OPEN

30

HEATER PASS CONTROL- HOW THE VPC ADJUSTS FLOWS
TO MAINTAIN ONE VALVE FULLY OPEN

31

16

17

20

21

HOW ENERGY INTEGRATION FORCES THE BOIL-UP OF ONE
COLUMN TO BE DEPENDENT UPON ANOTHER
32

22

CONVENTIONALLY CONTROLLED COLUMN

33

23

AN EXAMPLE OF MORE COMPREHENSIVE FEED-FORWARD

34

24

FEED-FORWARD CONTROL WITH COMPOSITION
CORRECTION

35

25

A SUDDEN FEED CHANGE COULD CAUSE A REBOIL INVERSE
RESPONSE
36

26

TEMPERATURE/FLOW CASCADE

37

27

TYPICAL SYSTEM FOR MAINTAINING INTERNAL REFLUX
CONSTANT WHEN CONDENSER SUBCOOLING VARIES

38


28

THE OVERRIDE CONTROL OF REBOIL TO PROTECT
BASE LEVEL DRYING OUT

39

EP OVERRIDE OF REBOIL TO PREVENT COLUMN
FLOODING

40

30

LOW FEED FORWARD REFLUX FLOW LIMIT

41

31

COMPUTER SET BOTTOM OFF-TAKE WITH LEVEL
REGULATING FEED VALVE

42

FLOW CONTROLLER TUNED TO ENSURE THAT LONG TERM
DESIRED BOTTOM TAKE OFF-TAKE IS OBTAINED

42

33

SHARP TEMPERATURE PROFILE

43

34

POSITIONAL TEMPERATURE CONTROLLER

44

35

DIAGRAM SHOWING HOW THE VALVE-POSITION
CONTROLLER SLOWLY ADJUSTS THE PRESSURE SET
POINT TO KEEP THE CONDENSER FULLY LOADED IN
THE LONG TERM

46

DIAGRAM SHOWING HOW ALL THE HEAT USED IN CRUDE
OIL DISTILLATION ENTERS WITH THE FEED

48

DIAGRAM SHOWING HOW VIRTUALLY ALL THE METHANE
FROM THE BOTTOM OF THE DEMETHANISER LEAVES
WITH THE LOW GRADE ETHYLENE

49

29

32

36

37

DOCUMENTS REFERRED TO IN THIS PROCESS
ENGINEERING GUIDE

52


0

INTRODUCTION/PURPOSE

As an industrial separation process, distillation has been in worldwide use since
the 1920s. It is one of the leading unit operations in process engineering.
A simple distillation column (see Figure 1) with only feedback control has over
700 possible control arrangements from simple pairing of variables alone. When
other factors are taken into account such as choice of temperature measurement
point, direct composition control, different possible methods of condensation and
pressure control, different reboil arrangements, feed forward control and the
question of the need for more advanced control techniques, the number of
possible control systems for continuous distillation columns can be measured in
the thousands.
The choice of any control system should ideally be based on a quantitative
understanding of the process and the comparative performance of different
control structures. Mathematical modeling, simulation and controllability
assessment are often used to design control systems where the design is not
obvious a priori. In the case of distillation, modeling is difficult and time
consuming. In addition, there are many different possible control arrangements.
Fortunately, by a logical approach and argument it is possible in most cases to
reduce the realistic possibilities to very few. This is the purpose of this Guide
and the problem is tackled in the following way.
(a)

After a general description of a column, regulatory control for steady state
operation is addressed by considering the related questions of
composition and material balance.

(b)

Next, enhancements to cope with constraints and for superior control
during transients are described.

(c)

Finally, the relevance and application of more advanced techniques is
discussed. The Guide concludes with a number of references for further
study.


1

SCOPE

This Process Engineering Guide deals with the control of continuous distillation
columns which are used widely in petrochemical and chemical manufacture and
refining. It covers regulatory control, disturbance compensation, constraint
control and more advanced topics.
The purpose of the Guide is to display control options; it does not deal with
control hardware.

2

FIELD OF APPLICATION

This Guide applies to the process and control engineering communities in GBH
Enterprises worldwide.

3

DEFINITIONS

For the purposes of this Guide, the following definitions apply:

VPC

Valve Position Controller

PIC

Proportional, Integral and Derivative Control

MBPC


DMC

Dynamic Matrix Control


FIGURE 1

DISTILLATION COLUMN SHOWING A SELECTION OF
POSSIBLE CONTROL AND REGULATION POINTS


4

GENERAL DESCRIPTION OF A DISTILLATION COLUMN

The main part of the column consists of a series of plates or a packed bed in
which fractionation takes place (see Figure 2). The feed enters the column at
some point between its ends chosen to minimize the mismatch between the feed
and the material being processed at that point. Vapor from the reboiler enters at
the lower end and passes upward through the column where it mixes intimately
with the descending liquid. This liquid is that part of the condensed vapor
returned as reflux from the condenser at the top of the column. In this way
the excess enthalpy of the vapor is given up to boil the liquid on the plates of the
column.
FIGURE 2

SCHEMATIC OF TYPICAL SIEVE TRAY


Vapor thus produced gives up its latent heat to boil liquid in the plate above and
so on. The rising vapor becomes richer in the more volatile component of the
feed while the residual liquid becomes richer in the heavier component and is
returned to the plate below. The overhead vapor is said to strip the more volatile
component from the liquid while the reflux scrubs the less volatile component
from the vapor. The action of a packed column is analogous with the
distinction that the process is continuous in space.
A pattern is thus set up such that the proportion of more volatile component
increases with column height while the proportion of the heavier component
decreases. If the column equilibrium conditions are upset the pattern is distorted
and a measure of this distortion can be provided by composition analyses or
temperature measurements within the column.
In general, there are only two independent quantities which may be held constant
within a column. The reflux and reboil rates are the two independent variables
which have a significant effect on the separation and may be externally adjusted
to compensate for disturbances in feed conditions.
That part of the overhead vapor which is not returned in condensed form as the
reflux is removed as top product or distillate to maintain a mass balance in the
condensation system. A measure of a lack of mass balance may be detected
from the pressure or the liquid level, depending on the particular form of
condensation system used. If the top product is removed entirely in the liquid
phase, the system is said to use a total condenser but if a substantial part
of the top product is removed in the vapor phase it uses a partial condenser.
Similarly, at the bottom of the column the down-flow liquid enters a reboil system
and that part which is not returned to the column as reboiled vapor is removed as
bottom product.
The measurement which provides an indication of any lack of mass balance in
the reboil system is usually the liquid level in the base of the column or in the
reboiler itself.
Some distillation columns feature multiple feeds and side-streams. Control of
side-streams is covered in 8.5.


5

REGULATORY CONTROL

5.1

Composition Control

Disturbances to column operation are most usually due to changes in feed-rate,
composition or enthalpy but atmospheric conditions can also affect the column particularly some condensation systems.
There are normally only two quantities which can be independently adjusted by a
control system to compensate for disturbances. Reference to Figure 1 should
make this clear. Five independent adjustments (control valves) are shown. There
are 3 mass balances to be maintained, viz: base liquid, tops liquid and vapor.
This leaves two for composition control. This means that it is not possible to keep
the composition profile in the column constant – only two quantities related to
that profile. In practice, we are normally interested in the composition at the top
and/or bottom of the column.
The composition of an n-component mixture is determined by n-1 quantities and
so, if n > 3 the column can never be controlled to give a constant top and bottom
composition as feed conditions change. A degree of over-purification is always
necessary in the steady state to ensure acceptable performance during transient
conditions.
5.1.1 Temperature Control
Temperature is most often used to infer composition. However, it should be
remembered that while at constant pressure the composition of a multicomponent mixture determines the boiling point uniquely, the converse is not
true.
As we are usually interested in the composition of products at the top and bottom
of the column, it is sometimes suggested that the bottom temperature could be
used to regulate reboil and the top to regulate reflux (see Figure 3). This can give
rise to problems of interaction.
If an analysis is carried out of the short term effects on temperature of changes in
feed-flow, composition and enthalpy then it can be shown (see Ref.1) that:
(a)

the relationships between the temperatures and the flows they might
regulate change, dependent on the disturbance and the condition of the
feed, and


(b)

some short term effects are often in the opposite direction to the
long term steady state requirements.

Note: When the words Ref.1 etc are shown in the text this refers to that number
document in the References in Clause 9.
This suggests that dual temperature control regulating both ends of the column is
fraught with difficulties and is likely to be highly interactive. This form of control is
not recommended, therefore, for columns with simple feedback control systems.
Dual composition control is feasible under some circumstances with mixed
temperature and direct analysis (see 6.1) and with model based predictive
control (see 8.4) where the column interactions and dynamics are built into the
control system.


FIGURE 3

COLUMN TOP AND BOTTOM TEMPERATURE CONTROL


5.1.1.1

Temperature Profiles

The temperature profile (Figure 4) shows a ''sensitive region'' (X) for temperature
control where the rate of change of temperature is relatively high. It is in this
region where temperature should be controlled to provide the maximum
sensitivity for composition control. In the region of maximum sensitivity pressure
variations are also likely to have a correspondingly small effect on the measured
temperature. Generally, for good temperature control, the effect on temperature
of normal pressure changes should be at least 5 times less than the effect of
normal composition changes.
In some instances there is a very large temperature change over only a few trays
in the column and under these circumstances simple temperature control may be
too sensitive. This problem is addressed in 8.1.
In practice the column temperature profile is likely to be very different from that
shown in Figure 4. It is necessary to distinguish carefully between the concept of
high product purity and a product containing a constant proportion of a particular
component. In the situation where a column is producing a very pure top product
but a bottom product with different components the profile may well be as shown
in Figure 5 while the opposite case is shown in Figure 6. In some instances the
column temperature control may show multiple sensitive regions (see Figure 7)
which are characteristic of particular internal separations.
In all these cases it is important to be clear about the control objectives. Which of
them need to be met under all circumstances and which might be relaxed and by
how much? What are the key heavy and light components? What are the design
feed rates? What is the normal turndown? Furnished with this information it is
usually possible:
(a)

to make a reasonable decision on where to site the temperature
measurement point;

(b)

to consider whether on-line direct analysis is feasible for composition
control or as a supplement to temperature control, and

(c)

to decide whether the composition controller should regulate the reflux or
the reboil rate to the column.


FIGURE 4

TYPICAL TEMPERATURE

FIGURE 5 PROFILE FOR
A COLUMN WITH VERY PURE
TOP PRODUCT BUT MIXED
COMPOSITION BOTTOM
PRODUCT

.


5.1.1.2

Choice of Reboiler or Reflux Regulation

If temperature control is to be used and if the temperature profile has multiple
sensitive regions, it is important to determine which corresponds to the key
separation in the column and this is where the control temperature should be
measured. In more complex azeotropic or extractive distillations (see 8.6), the
presence of such a region can be dependent on maintaining other
conditions.
The choice of reboil or reflux regulation is usually decided on the basis of
dynamic response. It is then usual to fix the other variable on flow control at a
value to allow the column to cope with disturbances. Feed-forward compensation
is dealt with later. In general, reboil changes affect the column temperature
(composition) much more quickly than corresponding changes in reflux.
It follows then that reflux regulation should be used only when the control
measurement can be made very close to the top of the column. This is likely to
be the case where one is interested in obtaining a top product containing a
constant proportion of a given component, not a pure product. It is difficult to be

definitive about where the breakpoint should be between choice of reboil or
reflux. F G Shinskey (Ref.4), argues that reflux should only be used if the control
temperature is measured on the top tray. This is probably too restrictive but if the
temperature has to be measured lower than, say, 20% from the top of the column
then regulation by reflux is likely to be unsuitable. Because of uncertainty in the
expected shape of the profile, and to provide flexibility for later changes in
operation, there is a need in the design phase to provide an adequate number of
optional temperature measurement points around the expected region of
sensitivity.
If analytical measuring instruments of sufficient reliability and speed are
available, direct composition analysis is to be preferred. The control loop
measurement delay and analyzer time constant should be short compared to the
column dynamics. There is then no objection to controlling top product
composition by varying the reflux rate when one is particularly interested
in the composition of the top product or the bottom product composition by
varying the reboil rate when one is interested in the composition of the bottom
product. The control loop is very short in either case. The poor dynamics of reflux
control are not then important.

5.1.1.3

Effect of Pressure

As described above, temperature measurement used to infer composition is
sensitive to pressure variations and this is why it is best to measure the
temperature in a region of high sensitivity to composition changes. On occasions
it is not possible completely to mask the pressure effect. A number of solutions
are possible.
(a)

Column pressure may be controlled at the point of temperature
measurement. This assumes that it is desirable to hold column pressure
constant (standard practice) but modern thinking suggests that a control
system that allows pressure to float can be more economical (see 8.3).

(b)

The temperature can be mathematically compensated for variations of
pressure so that the compensated temperature is seen by the temperature
controller. The objective is to reference a temperature made at a variable
pressure to one at some base pressure, e.g. atmospheric. The
compensated temperature Tc is related to the measured temperature
Tm, pressure p and base pressure P b as:


(c)

Difference or double difference temperature measurements are
sometimes made which compensate for pressure variations. They are all
affected by pressure changes to approximately the same extent. However
if these arrangements are to be successful they need to be carefully
matched to column characteristics (Ref. 2 & 3).

5.1.2 Direct Composition Control
Modern analytical devices are much more reliable than hitherto and have
the capability to supplement temperature control in many applications. The
advantage is that they can be sited to measure directly top or bottom
compositions and to regulate reflux or reboil in fast loops as appropriate.
In the case of azeotropic or extractive distillation, where it is necessary to
maintain a minimum component composition at some point in the column,
this concentration can be measured directly. Analyzers are regularly
provided to monitor product quality in run down lines. There is an
opportunity to use these analyzers to control quality directly, by siting them
in the column, rather than to monitor the performance of a less effective
inferential system. Unfortunately this option is rarely taken. The most
common form of on-line analyzer is the chromatograph which has the
advantage of being able to measure multiple components. It only performs
an analysis at discrete intervals, however, and can be very slow. If it is
possible to use IR or UV analyzers then their continuous output makes
them much more suitable for control.


All analyzers are more expensive to install and maintain than temperature
measurements. For this reason analyzer use is likely to be restricted to
applications where there is a strong economic incentive to reduce costs, to
minimize environmental pollution or to maximize recovery of a particularly high
value product. Generally, but not exclusively, they will be used in conjunction with
temperature based composition control. A particularly successful mode of
operation is, for example, to trim reflux/feed ratio when the reboil is regulated by
a conventional temperature control loop (see 6.1). In this way, a measure of dual
composition control is possible.
Care should be taken in the installation of all direct composition control loops to
minimize the analyzer sample delay which is a pure dead time. The analyzer
should be sited as close as possible to the column and a vapor sample is
generally to be preferred since the higher velocity reduces the dead-time in the
sample line.
Chromatographs present additional problems since their analysis is produced at
discrete intervals known as the "sample interval". When a new analysis is
produced, the controller should apply immediate control action and then wait until
a new analysis measurement is made. This sampled/data control has the benefit
of introducing an integral action effect into the control loop without the normal
disadvantages of increasing period. If a computer is used, performance will be
improved if its sample period is keyed to the sample period of the analyzer so
that the control can take immediate action on new information.

5.2

Mass Balance Control

The choice of regulation variable to control composition has been discussed
above. This leaves many other variables to be regulated from other
measurements. The correct values of these can be obtained from mass balance
considerations. Sufficient symptoms of mass imbalance are available to enable a
conventional feedback control system to be set up. To reduce the options to a
manageable level it is helpful to consider the various mass balances that it is
necessary to achieve and those regulatory variables which can affect them
directly or indirectly.
The balances are:
(a)

Vapor balance in the column.

(b)

Liquid balance in the condensation system.


(c)

Liquid balance in the reboil system.

Table 1 shows the symptoms of imbalance and the regulatory variables. The
situation is shown diagrammatically in Figure 8.
TABLE 1:

SYMPTOMS OF IMBALANCE AND THE REGULATORY
VARIABLES

Note: Not all of these measurements and regulatory variables may be available
on any given column.
(a)

Vapor balance in the column
Vapor is added to the system by the reboiler, by evaporation of liquid from
the plates and possibly from the feed. It is removed by condensation in the
condenser, by condensation on the plates and possibly as top product
taken off in vapor form (non-condensables in the feed or air leakage in the
case of a vacuum system may be additional contributors to the vapor
balance). Any lack of balance between the sum of the rates of production
and the sum of the rates of removal of vapor causes a change in the
pressure in the system so pressure measurement at some point provides
a symptom of lack of vapor balance. Inspection of Figure 8 shows that of
the quantities mentioned above, the only ones available for pressure
control are the rate of generation of vapor in the reboiler (h), the rate of
condensation in the condenser (c) and the rate of removal of top product
in the form of vapor (v). At least one of them should be left available to be
adjusted either directly or indirectly by the column pressure controller.


The details of how the condensation rate can be adjusted in various types
of condenser and how the reboil rate can be adjusted in various types of
reboiler are covered in later sections of this Guide. The present indications
are symbolic.
(b)

Mass balance of the liquid in the condensation system.
The condensation system is that part of Figure 8 enclosed by the dotted
lines and includes the condenser and reflux drum. Liquid enters this
system by condensation of vapor in the condenser (C). It leaves as
distillate in condensed form (d) and as reflux returned to the column (r). To
maintain a liquid balance it is necessary, in the steady state, that (c)
= (r) + (d). Any lack of balance is indicated by a rise in or fall in the reflux
drum level (R). It is not necessary to hold a reflux drum level precisely.
Indeed, it is usually best to use the reflux drum and the base of the column
as surge capacities to smooth disturbances to downstream units (see
5.2.1). All three of the quantities (c), (d) and (r) are available for direct
control of reflux drum level.


FIGURE 8

(c)

MASS BALANCE CONTROL

Mass balance on the liquid in the reboil system.
Liquid enters this system from the column itself and leaves it as bottom
product or by evaporation in the reboiler. The bottom product flow (b) and
the rate of vaporization (h) are available to control the level directly.


From the preceding discussion it will have been apparent that certain
variables affect more than one mass balance or composition. This is
shown in the matrix in Table 2. Added to the matrix is an additional row
showing possible linkages for composition control.
Direct linkages between regulation variables and the symptoms of
imbalance are shown by a cross in the appropriate location. An (x) shows
an indirect, feasible but less desirable potential linkage.
At least one, or preferably both, of bottoms off take (b) and distillate (d)
should be controlled by either a symptom of imbalance or composition.
The composition link should only be considered if a sloppy split is
adequate.
TABLE 2

PRACTICAL LINKAGES BETWEEN CONTROL (P, R, B, C) AND
REGULATION VARIABLES (h, r, d, b, c, v)

Sometimes adjustment of a column feed is used to control base level. This
generally provides poor control due to excessive dead time and process delay
before changes in flow can affect the level. All other flows in the column should
also be adjusted in the same ratio to maintain separation conditions. This form of
control should be avoided if possible.
An example of how to design a feedback control system for a particular
arrangement is given in 5.3.


5.2.1 Surge Capacity and Level Control
It is a general rule that most capacities on a distillation column and its peripheral
equipment should be used to the maximum extent to smooth out flow fluctuations
and thereby help protect downstream equipment from disturbances. An
exception to this general rule may be when, for example, reflux drum level is
regulating reflux and it is important to increase reflux to match any increase in
reboil rate. The reasons for setting up such a form of control is discussed later.
In all other cases the level control system should be tuned to maximize the use of
surge capacity. This means that a proportional only controller should be used
tuned so that the capacity never empties or overfills. Sometimes proportional
(error)2 control is suggested for such applications so that little control action
takes place at the mid measurement point but more drastic action takes place as
the measurement reaches its limits. In the experience of the author properly
tuned proportional only control is normally acceptable. It is simpler and is better
understood.

5.3

Design of Feedback Control Systems

(a)

Design Example
Consider a distillation column with the following specification:
Liquid feed (entering about half way up the column) - 10 te/hr
Distillate - 4 te/hr
Reflux ratio - 2 : 1, i.e. r/d
The column has a total condensation system and operates at 20 bar g.
The sensitive region for temperature control is 20% of the height above
the base.
It can easily be deduced that the column reflux is 2 x 4 = 8 te/hr and the
bottom off-take is 6 te/hr. The flow into the base will be approximately the
feed rate + reflux = 18 te/hr. The bottom off take is 6 te/hr and so the
reboil rate will be approximately 12 te/hr.
There are no critical flows, so, from Table 2 first make a choice of
composition regulation. This is fairly evidently reboil (h) and the other
controls can be built up for the reasons given in Table 3.


TABLE 3

COMPOSITION REGULATION

This leaves reflux uncontrolled. In these circumstances it is normal to flow
control reflux. This is discussed later.
We have not considered the type of condensation system provided. If, for
example, it were a water cooled condenser it is quite possible that no
regulation would be available because the necessary valves are likely to
be very large and cooling water is relatively inexpensive. How should
pressure be controlled?
In this circumstance there might be a temptation to control pressure by
regulating reflux flow. But then there would be no independent way of
setting the energy input to the column and, if the pressure and
temperatures are mismatched (as they always will be), the reboil and
reflux would ramp up or down until a balance is found. This might be
maximum (or minimum) energy input.
In these circumstances the correct action is probably to let the pressure
float to achieve the lowest possible value at any time. The problems of
temperature control with the variable pressure are covered in 5.1.1.3. The
reflux should be flow controlled or ratioed to feed. Then the reflux flow and
its temperature sets the heat input to the column to balance the reflux flow
to maintain steady compositions.
5.3.1 Very Small Flows
On some occasions a top or bottom off-take will be very small since only trace
impurities are being removed. Consider the following case, identical to the above
example except that the top off-take is 9 te/hr with the bottom off-take 1 te/hr.
Carrying out a similar calculation to previously, it can be shown that the ratio of
reboil to bottoms off-take is 27:1. In this circumstance, it is very difficult to
maintain a bottoms mass balance by regulating bottoms off-take; reboil has a
much greater effect.

In cases like this it is not uncommon to regulate the bottoms level by adjusting
reboil and to use the composition measurement, particularly if it is close to the
bottom of the column, to regulate the bottoms off-take. Similar arguments apply
to the control of distillate and reflux from a temperature control near the top of the
column and the reflux drum level.
Consider the following table. In this case the bottoms mass balance is the critical
parameter to be measured so consider it first.
TABLE 4

5.4

COMPOSITION REGULATION - VERY SMALL FLOWS

Pressure and Condensation Control

5.4.1 Principles of Condensation Control
The overhead vapor leaving the top of the column consists of a mixture of
condensable vapors containing all the components of the feed in various
proportions together with a certain amount of non-condensables which
may be either true inerts or material of too high a vapor pressure to
condense to any appreciable extent under the conditions prevailing in the
condensation system.
Overhead condensers may be either total condensers where the
condensable part of overhead vapor should ideally be completely
condensed leaving only the inerts to be purged from a suitable point in the
condensation system. Alternatively, they may be partial condensers where
only a fraction of the condensable components are condensed, the
remainder passing out with the inerts as a vapor top product. In any
condenser the total rate of condensation depends on:


(a)

The partial pressure of the condensable fraction of the
vapor in contact with the cooling surface.

(b)

The temperature on the cold side of the condenser.

(c)

The overall heat transfer coefficient.

(d)

The available surface area for condensation.

(e) The effective latent heat of the mixture.
Condensation control methods usually vary one or more of the first four of these
quantities. They are considered in turn.
Typical condensation arrangements are shown in Figures 9(a) to 9(c). In Figures
9(b) and 9(c) the vapor top off-take may not be present if condensation is total
but there is always likely to be some means of manually venting inerts which
have built up in the system.
(1)

Consider Figure 9(a). If it is required to reduce the rate of condensation,
the control valve closes and the fraction of non-condensables in the
condenser will then build up.
The particular system shown can only be used with columns operating
above atmospheric pressure but alternatives using the same principle can
be used when operating at below atmospheric pressure. This is
considered later. Strictly speaking, adjustment of the vapor purge valve
also varies the overall heat transfer coefficient as well as the partial
pressure of the condensables since the vapor side heat transfer coefficient
is dependent on the fraction of the non-condensables in the vapor.

(2)

Figure 9(b) shows a system where the flow of cooling medium is varied. At
high flow rates the average temperature of the coolant is reduced so the
rate of condensation is increased. Sometimes boiling refrigerant is used
as the cooling medium, absorbing the latent heat of condensation as its
own latent heat of vaporization.


FIGURE 9 (a)

REGULATION OF VAPOR PURGE

FIGURE 9 (b)

REGULATION OF COOLANT


FIGURE 9 (c)

FLOODED CONDENSER

In this case the refrigerant is in equilibrium with its own vapor so its temperature
can be changed simply by changing the pressure in the vapor space above it.
This can be used for condensation control (see 5.4.6).
Manipulation of the coolant flow is rarely used to adjust heat transfer in water
cooled condensers on distillation columns. Response of the condensing rate to
water-flow variations is non-linear and slow. The speed of response also varies
with coolant flow which can cause loop stability problems. Low velocity and also
leads to high fouling and metal corrosion. For more details see (Ref.2).
(3)

The surface area available for condensation can be varied by operating
the condenser partly flooded with liquid as shown in Figure 9(c) where a
control valve is placed in the line carrying the liquid away from the
condenser. In this case only the part of the surface which is not
submerged is available for condensation of vapor and the liquid will leave
the condenser sub-cooled.
The basic condensation control systems differ in dynamic performance.
For example, the speed of operation of the system shown in Figure 9(a)
will depend entirely upon the proportion of inerts present in the overhead
vapor from the column. If the proportion is very small then, even if the vent
valve is closed, it could take a long time to build up the sufficient
proportion of inerts in the condensation system to affect heat transfer.
Sometimes a split range control system is provided where nitrogen can be
added to the system to speed up the response.


The speed of operation of the system shown in Figure 9(b) depends very
much on the thermal capacity of the tube wall separating the hot and cold
sides of the condenser and on the residence time of the cooling medium
when this is a cold liquid.
The system in Figure 9(c) has slow dynamics. Movement of the control
valve affects the rate of change of the flooding depth in the condenser and
this in turn affects the rate of change of condensation and hence pressure

5.4.2 Air Condensers
In this system the air is directed upward by a fan through a horizontal bundle of
finned tubes. The temperature of the condensing vapor tends to be constant if
the tubes are not flooded with liquid and the following equation applies:

The rate of heat transfer is limited by the air film and so the 0.2 power in the
exponent is justified. As a consequence, heat flow should be reasonably linear
with air flow.
Controlling air flow is another matter. Variable speed fans are rarely used
because of the high cost of drive units. Most air coolers use multiple fans which
can be energized in stages but this gives only incremental control. Some fans are
equipped with variable pitch blades or adjustable louvers but they have a
tendency not to work satisfactorily.
A further problem with air condensers is the effect of ambient conditions. Rain
tends to convert the dry condenser into a wet condenser which can have a
marked affect on the temperature of the condensate and hence the internal reflux
in the column. In critical conditions this may require the provision of internal reflux
control computation. This is discussed later.

Adjustment of cooling in the air condenser can be effected by:
(a)

flooding with condensate, or,

(b)

bypassing.

Most air cooled condensers are horizontal and hence flooding does not give
smooth control. Sometimes the condenser is mounted at an angle and this helps
but since the tubes are of large capacity and can contain a substantial amount of
liquid, response can be slow.
Bypassing hot vapors around the condenser is shown in Figure 10. Column or
reflux drum pressure can be controlled by adjustment of the bypass valve. The
sizing of the bypass valve is difficult but critical to the effective operation of the
condensation system. In particular, it is necessary to decide what proportion of
the vapor shall be bypassed under different operational circumstances,
throughput and ambient conditions.
For more information on air cooled heat exchangers see (Ref.3).

5.4.3 Internal Condensers
In some columns a condenser is mounted inside the column above the top plate.
There is then no reflux drum and the condensed liquid reflux falls directly from
the condenser onto the top plate through some distributor system. Uncondensed
vapor can be taken off directly as top product or an internal weir may be fitted,
allowing the distillate to be withdrawn. This limits the options for control. Ideally,
all condensate should be trapped and withdrawn for metering and control. If this
is not done the system behaves as if the reflux flow were regulated via a very
high gain reflux drum level controller. (R <-> r) should be assumed when
examining Table 1. Large and sudden changes in top off-take can cause
corresponding changes in reflux and, in extreme situations, reflux flow may be
lost for a period.


5.4.4 Operation at Atmospheric Pressure with a Total Condenser
Many columns are designed to operate at atmospheric pressure by linking the
condenser to atmosphere through a vent and scrubbing system with a small
pressure drop. In principle, this balances the rate of condensation to the
overhead vapor flow by altering the proportion of inerts (air) present in the
condensing vapor. Pressure control is rapid but varies with atmospheric
pressure.
From the point of designing an overall control system it is as if the pressure is
linked to the condensation rate and (P <-> c) has to be assumed in Table 1.
When a column is one of a number in a distillation train and a partial condenser
is used, then the vapor top product may form the vapor feed to the next column
in the train. If some method of pressure control is used on the second column,
the first may be maintained at very nearly the same pressure as the second
simply by ensuring that the line carrying the vapor offers little resistance to flow.
Referring to the Table 1, this effectively means that the pressure is related to the
vent rate (v) although physically there is no controller or control valve. Assume (P
<-> v).


FIGURE 10 A HOT>GAS BYPASS IS THE MOST COMMON MEANS OF
CONTROLLING AIR>COOLED CONDENSERS


5.4.5 Operation at Sub-Atmospheric Pressure
When the column operates at below atmospheric pressure, the proportion
of inerts in the vapor entering the condenser is likely to be significant
because of leaks of air into the system. The purge line is often connected
to the suction side of some form of vacuum pump or steam ejector. With
this arrangement it is possible to use any of the basic condensation
control systems but, since the proportion of inerts in the overhead vapor
may well be high, a modification of the system shown in Figure 9(a) is
normally used. Three common arrangements are shown in Figures 11(a)
to 11(c).
FIGURE 11 CONDENSATION CONTROL SYSTEMS FOR OPERATION
UNDER VACUUM


5.4.6

System Using a Refrigerant as Cooling Medium
When distillation is carried out at low temperature, a boiling refrigerant is
frequently used as the cooling medium in the condenser.
The temperature of the boiling refrigerant depends on the pressure
maintained in the vapor space above it which may very easily be regulated
by manipulating a control valve in the refrigerant exit line. The refrigerant
is often contained in the shell of a condenser as shown in Figure 12 and
condensation of the process vapor takes place in the tubes. A constant
refrigerant level is maintained by admitting a controlled flow of liquid
refrigerant to the shell. The vapor condensation rate can be varied to a
certain extent by altering the refrigerant level but the temperature of
refrigerant boiling is usually used as the principal method of condensation
control.


FIGURE 12 CONDENSATION USING A BOILING REFRIGERANT IN
CONDENSER SHELL

When partial condensation is required, and the vapor velocity in the
condenser tube is high, the condensing liquid is carried out of the tubes by
the vapor flow. If the vapor velocity is low, as will be the case with a total
condenser, it may not be easy to keep the tubes well drained. In this case,
better heat transfer may be obtained by condensing the vapor in the
shell with the refrigerant vaporizing in the tubes. A system of this type is
shown in Figure 13. The refrigerant is held in equilibrium with its vapor in a
separate accumulator to which a control flow of refrigerant liquid is added
as again, the temperature of this liquid is determined by the pressure in
the vapor space above it.

5.4.7 Systems Where the Reflux Drum is Mounted Level to or Above the
Condenser
Sometimes, with free standing distillation columns, both condensers and
reflux drums are mounted at ground level to be easily accessible for
maintenance and to avoid the need for expensive supporting steelwork.
The reflux drum is then sited at the same level as the condenser or even
slightly above it. Liquid formed by condensation cannot fall into the reflux
drum under gravity. In practice, to avoid the need to pump liquid into the
reflux drum, a pressure differential is maintained between condenser and
reflux drum.

FIGURE 13 CONDENSATION USING A BOILING REFRIGERANT IN
CONDENSER TUBES

The basic method of operation is shown in Figure 14. The condenser works at
almost the same pressure as the top of the column and the rate of condensation
is varied by flooding part of the cooled surface with condensed liquid.


FIGURE 14 REFLUX DRUM MOUNTED LEVEL WITH OR ABOVE
CONDENSER

The liquid entering the reflux drum is therefore cooled below its condensation
temperature from contact with the submerged tubes. Its vapor pressure, which is
the reflux drum pressure, is therefore lower than the pressure in the condenser.
This pressure difference drives the condensed liquid into the reflux drum against
the hydrostatic head due to the difference in level and the frictional losses in the
piping.
Pressure control is effected as shown in Figure 15 by bypassing some vapor
from the column to the reflux drum. This alters the EP between condenser and
drum and so alters the flow of condensate and thus the level in the condenser.
This alters the condensation rate and so the column pressure.


FIGURE 15 REFLUX DRUM MOUNTED LEVEL WITH OR ABOVE
CONDENSER BUT WITH REDUCTION OF PRESSURE DROP
ROUND REFLUX LOOP

Experience suggests that there can be operational problems with these
systems and great care needs to be taken with the hydraulic design. If the
bypass valve is adjusted to control the column pressure there is an
inverse response which can cause instability. Controlling the reflux
drum pressure by adjusting the bypass appears to be more satisfactory.

5.5

Reboiler Control
Reboilers can be classified into two groups depending on whether the
mechanism for transfer of heat to the column is natural or forced
circulation. The different physical construction or heat sources can affect
control.
Figure 16 shows an example of natural circulation where the reboiler tube
bundle is immersed in liquid either in a kettle exterior to the column or


internally. In either case the heating medium may be steam or liquid within
the tubes. Both achieve 100% vaporization.
FIGURE 16 INSERTING A TUBE BUNDLE EITHER (a) DIRECTLY IN THE
COLUMN BASE OR (b) IN AN EXTERNALLY MOUNTED
KETTLE


Natural circulation or "thermosyphon" reboilers may be mounted either
horizontally or vertically. Vertical thermosyphon reboilers are used
primarily in the chemical industry and are typically steam heated (see
Figure 17(a). Horizontal thermosyphon reboilers are more common in
petroleum refining or similar operations in the chemical industry. They are
usually heated with circulating oil, see Figure 17(b).
Forced circulation is used with vacuum distillation or when the heat input
is obtained from oil or gas furnaces.
FIGURE 17 ALTERNATIVE ARRANGEMENTS OF THERMOSYPHON
REBOILERS


5.5.1 Steam Reboilers
Steam is the most common vapor used for heating. In some low
temperature applications a refrigerant is used. Superheated steam is a
relatively poor heat transfer medium, as are most gases. Saturated steam
is an excellent medium because of good heat transfer and a high latent
heat of vaporization. The boil-up rate may be controlled either by steam
flow or shell pressure. The steam flow measurement is usually made
upstream of the control valve where the pressure is normally constant.
The steam rate to the reboiler may be controlled either by a valve in the steam
line or in the condensate off-take. With the control valve on the inlet to the
reboiler, the saturation pressure in the shell varies with heat load. Since the heat
is being transferred between a condensing and a boiling fluid, neither changes
temperature greatly in the process. A steam trap or similar condensate seal is
necessary to drain condensate without releasing steam.
If pressure rather than flow is controlled, the boil-up automatically increases if the
process fluid boiling point falls. This action compensates correctly for changes in
bottom product composition but not for changes in column pressure. However,
the rate of boil-up is not linear with steam pressure nor is steam pressure zero at
zero boil-up. For these reasons, steam flow control is preferred to pressure
control to establish and maintain boil-up rate, prevent flooding
and control product quality.
There are advantages and disadvantages in placing the control valve in the
condensate line. In the first place the valve can be smaller - typically one third of
the line size of a steam valve used for the same service. In addition, the steam
reaching the reboiler is at a higher pressure than with a valve in the steam line
and hence the maximum heat transfer rate is higher.
The dynamic response characteristics of the two approaches differ significantly.
Manipulating the control valve in the steam line causes steam flow to change
immediately. Shell pressure, and hence rate of heat transfer, will lag behind for
only a few seconds.
By contrast, the condensate valve has no direct effect on steam flow and as the
condensate level determines steam flow this level takes time to change. The
slow response of a system with a valve in the condensate line usually means that
it is not suitable if a boil-up is being used for base level control.


5.5.2 Hot Oil Reboilers
When a liquid stream is used to boil-up a column there arises the question of
how to measure and control the heat input. With steam heating, the rate of
condensation and therefore heat input is directly proportional to the steam flow.
With liquid media, however, the relationship is very non-linear since, as the flow
of hot oil increases, the temperature difference between inlet and exit also
changes because the heat transfer coefficient is dependent on flow rate. Heat
input can be calculated from the equation:

The calculation of actual heat transfer can be implemented either with digital or
analogue components as shown in Figure 18. This heat transfer equation
assumes a steady state. In practice, the indicated heat flow always leads the true
heat flow by the residence time of the liquid in the reboiler.
Hot oil reboilers are often provided with heat from a fired heater (see Figure 19).
A low oil flow requires higher reboiler inlet temperatures to transfer given flows of
heat to the columns. Additionally, higher inlet temperatures require a higher flue
gas temperature in the heater which causes higher stack losses. Maximum
efficiency is realized when oil flow is at maximum and oil temperature is at its
minimum. Conventional heater controls always include a bypass, recirculating
hot oil back to the cold oil line, to protect against loss in flow through the heater.
The bypass valve is usually manipulated to control differential pressure between
hot and cold oil lines and, as less heat is required by the reboilers, the bypass
valve opens to maintain a constant flow through the heater. A better arrangement
is to ensure that the reboiler demanding the most heat, receives full oil flow at all
times, with the oil temperature set to deliver that heat.


FIGURE 18 HEAT INPUT CONTROL

Flow to the other reboilers may then be throttled to match their requirements.
This can be achieved by use of a valve position controller (VPC) as shown in
Figure 19. Each column will have its own heat input controls manipulating hot oil
flow. The valve position signals are compared in a high selector and the highest
is sent to the VPC. This device then adjusts oil temperature until the highest
valve signal is at or near full opening. The oil temperature will then be at its
minimum acceptable value as will the hydraulic power loss through the control
valves. The valves are free to be manipulated by the individual column controls
for fast response in the short term while the slower acting VPC minimizes energy
loss in the long term. If such as scheme is implemented bypassing is not
normally required since one of the load control valves is always nearly full open.
However, a bypass valve that fails open on high ΔP should still be used to
protect the heater against controller failure.
For more information on reboilers see (Ref.5).


5.5.3 Fired Heaters
A fired heater may also be used as a reboiler. This is often done for the
fractionation of crude oil and for other products having boiling points in excess of
150°C. The flow of bottoms product to the heater is nearly always controlled to
maintain efficient heat transfer.
In addition, the flow is usually conducted through several parallel passes in the
furnace and some means of equalizing flow through these passes should be
provided. This can easily be done by use of a VPC.
FIGURE 19 DIAGRAM SHOWING THE VALVE POSITION CONTROLLER
(VPC) ADJUSTING THE TEMPERATURE SO THAT ONE VALVE
IS ALMOST FULLY OPEN


The highest valve signal is selected as the input position controller which adjusts
the set points of all flow controllers to hold that valve exactly at the fully open
position (see Figure 20). This arrangement results in the minimum attainable
pressure loss through the valves.

5.5.4 Inverse Response
Inverse response can sometimes be a problem when heat input is used to control
column base level. The percentage vaporization increases with heat input such
that the volume of vapor bubbles in the reboiler liquid mass increases. This
moves liquid from the reboiler to the column base or, in the case of an internal
tube bundle, increases the "voidage" in the base. In either case this can cause
the measured base level to rise. This is only a short term effect because the
increased rate of boiling will drive off mass in the long term causing the level to
fall.
In these circumstances the base level controller should be tuned for the long
term response where the level falls with increased heat input. The short term
transient response in the opposite direction can cause severe stability problems.
The situation is analyzed on page 165 of (Ref. 4).


FIGURE 20 HEATER PASS CONTROL> HOW THE VPC ADJUSTS FLOWS
TO MAINTAIN ONE VALVE FULLY OPEN


5.5.5 Heat Integration
It is attractive, for energy efficiency reasons, to integrate heat sources and sinks
in distillation trains. This can be simply the use of a cooler on a bottom product to
heat the feed to the column itself or the use of the top or bottom streams of some
other column to provide feed heating. Such apparently innocent systems can
sometimes transmit disturbances which may exacerbate an already difficult
control problem. The reader should be aware of this possibility.
On some distillation trains the condensers of one column provide the majority, if
not all, of the reboil load for another column (see Figure 21). Such systems are
difficult to operate efficiently since there can be a number of competing operating
objectives but a distinct shortage of effective degrees of freedom, even when
heater bypasses are provided. In practice, the control system should be set up to
rank the objectives and where there is a limitation in, say, heat input, it should be
remembered that some of the objectives may not be met fully.
FIGURE 21 HOW ENERGY INTEGRATION FORCES THE BOIL-UP OF ONE
COLUMN TO BE DEPENDENT UPON ANOTHER


Ratio control and simple computation can go some way towards providing an
efficient system but control based on conventional PID controllers is unlikely to
be entirely effective. Model based predictive control, which takes account of
measured process interactions, may well provide a better form of control in these
cases and should be considered (see 8.4).

6

DISTURBANCE COMPENSATION

6.1

Feed-forward Control

Up to now in this Guide we have discussed only feedback control systems.
These are systems which feedback a correction as a result of a deviation from
set point - an error. In other words, the process has to be disturbed before a
feedback system can begin to make a correction. Feed-forward control, on the
other hand, takes account of disturbances which can upset quality and, if
properly designed, can apply corrective action to minimize deviations of the
control variable. Feed-forward is theoretically capable of perfect control but in
practice errors will always remain. Nevertheless a feed-forward calculation
accurate to only ± 10% will reduce the sensitivity of product quality to these
measured disturbances by a factor of 10; a good return for modest accuracy. It is
difficult to understand why feed-forward is so little used; it is usually only
considered when feedback systems fail. To a first approximation, to achieve
constant separation, all flows in a distillation column should increase or decrease
in sympathy with the feed rate. Feed rate changes are a common disturbance in
distillation columns and so it is logical to feed-forward changes in feed in such a
way as to adjust other column inputs to compensate for the feed change. Figure
22 shows a conventionally controlled column. Without feed-forward, the
temperature controller will make the appropriate adjustments to the reboil but this
will take time and the column will be subject to a larger disturbance than is
necessary. Where reflux is flow controlled it would be set at the highest rate
required for normal operation which could be costly in energy consumption.


FIGURE 22 CONVENTIONALLY CONTROLLED COLUMN

The most basic feed-forward system would ratio reflux to feed. A more
comprehensive system would also ratio reboil to feed and compensate this ratio
from the composition measurement (see Figure 23). As the feed rate changes,
reflux and reboil will also change. This would speed up the column response to
disturbances and minimize off specification product.


FIGURE 23 AN EXAMPLE OF MORE COMPREHENSIVE FEED>FORWARD


Feed composition changes are normally only compensated for by the standard
feedback control system but, if they are very severe, it may be possible to
analyze the feed and to make the appropriate changes to column operation. This
is rare.
Feed-forward allows a measure of dual composition control on a column because
it helps to decouple feedback loops. For example, if the main separation is
effected by a temperature controller near the base of the column, but it is
important to maintain a fixed purity level in a top product, a system as shown in
Figure 24 may be an option where an analyzer placed near the top of the column
and monitoring a key tops impurity is used to modify the reflux ratio.


FIGURE 24 FEED-FORWARD CONTROL WITH COMPOSITION
CORRECTION


6.1.1 Dynamic Compensation
It is often not sufficient that the slave flows be ratioed to the feed in a
steady state relationship. The correct dynamic relationship should be
established between the variables. Feed, reflux and reboil are physically
located at different points in the column and their effect on product quality
will differ in speed of response. When differences exist between the
responses of the manipulated inputs, feed-forward correction by steady
state calculation alone is likely to cause transient errors. If this is the case,
particularly for rapid disturbances, a dynamic lag should be placed in the
feed-forward loop to match as closely as possible the response of the feed
forward loop to that of the column response to feed changes.
The need for dynamic compensation is clearly illustrated by examining the
response of a column to feed and boil up variations. Consider the scheme
shown in Figure 25. Here the bottom product flow is ratioed to the liquid
feed through a feed-forward loop and the level in the column base is
controlled by regulating heat input.
FIGURE 25 A SUDDEN FEED CHANGE COULD CAUSE A REBOIL
INVERSE RESPONSE


This is a common form of control where variation in the bottoms off-take is not
sufficient to control the base level. If no dynamic compensation is applied, a step
increase in feed will cause a step increase in bottoms flow causing a reduction in
bottoms mass hold-up and therefore a decrease in boil-up. In the long run the
boil-up will ultimately increase but the lack of dynamic compensation has caused
a transient response in the wrong direction.
Compensation with a simple first order lag in the feed-forward loop would
moderate the inverse response but not remove it. Exact compensation in the
requires a dead time equal to the difference in dead times of the effect of the
feed and bottom flow on the bottom level. In addition, a lag equal to the lag of the
feed on bottom level and a lead equal to the lag of bottom flow on the level is
required. Since most responses can be characterized by a dead time followed by
a first order lag, dynamic compensators provide dead time plus lead-lag.

6.2

Cascade Control

Cascade control is commonly applied on distillation columns - often where it is
not needed!
A typical cascade loop is shown in Figure 26. Here the "master" loop,
temperature, is regulating reboil via a steam reboiler. The flow control "slave"
flow loop is provided to compensate for steam main pressure fluctuations. Any
change in pressure is picked up by the flow loop and compensated before the
reboil is affected. This implies two things:
(a)

The dynamics of the slave loop should be faster than the master loop.

(b)

There are disturbances which can be usefully compensated by the slave
loop.

Another example of a good application of cascade control is with hot oil reboilers
where a combination of flow and temperature to calculate heat input can be
compensated for by the slave (see 5.5.2).


FIGURE 26 TEMPERATURE/FLOW CASCADE

Examples of often inappropriate applications of cascade are level/flow
combinations where it may not matter if the level is affected by some disturbance
in the flow line and the system can be simplified by removal of the slave loop.
There are other cases where cascade can be useful. For example, if the valve
has a very nonlinear installed characteristic the slave flow loop can, in effect,
"linearized" the characteristic and in those special cases a level flow cascade
may be appropriate. This is likely to be a way of fixing a poor installation rather
than a way of designing an effective control system.
The message is to be aware of the benefits of cascade control when it properly
applied but to avoid the added complexity when it is not necessary.


6.3

Internal Reflux Control

If the external reflux is sub-cooled compared to the column top temperature, it
will cause additional condensation in the column and the internal reflux will be
greater than the externally measured flow. So long as the sub-cooling is
constant, this does not matter. If, however, a flooded condenser or air condenser
is used the external reflux temperature may vary. If this is an important factor, a
simple on-line calculation may be carried out to compensate for the effect,
viz:

The calculated internal reflux is fed to a flow controller which regulates the
external reflux flow valve (see Figure 27).


FIGURE 27 TYPICAL SYSTEM FOR MAINTAINING INTERNAL REFLUX
CONSTANT WHEN CONDENSER SUB>COOLING VARIES

7 CONSTRAINT CONTROL
In this Guide simple regulatory control for essentially steady state operation was
first described. Then we discussed, under disturbance compensation, controls
designed to provide better dynamic behavior.
A more extreme situation is the need for constraint control. Control for
constrained operation is not common but is essential if maximum performance is
to be achieved since, if at least one variable is not being held at its limit,
separation efficiency or productivity could still be improved.


Constraint control is particularly useful where, for whatever reason, column
conditions have been pushed to an extreme or where a more important
parameter should be safeguarded at the expense of a less important. An obvious
example is override control of reboil from a measure of column flooding. Here it is
more important to prevent flooding than to try to maintain column composition.
Model based predictive control systems with constraint avoidance fall into this
category. In all cases detection of an override is necessary to warn the operator
that the control structure has, in effect, been changed.
Examples of a number of different common override situations are given below.
A comprehensive discussion of controlling within constraints can be found in
(Ref.4).
FIGURE 28 THE OVERRIDE CONTROL OF REBOIL TO PROTECT BASE
LEVEL DRYING OUT


7.1

Override Controls
A simple example of an override control would be that of a base level
controller, normally regulating bottoms off-take overriding the composition
control of the reboil in an extreme situation (Figure 28). It is necessary to
provide the temperature controller with external reset from the output of
the selector relay to prevent integral saturation of the controller when the
level controller takes over regulation of the reboil heat input. A reboil
deviation alarm is also well worth providing to warn the operators of the
potential loss of composition control.
A similar arrangement could be provided for reflux inventory control but in
both cases it is better to set up the mass balance controls to be selfsufficient in their operation. A more common example of override control is
that to prevent flooding when a column is being pushed to its limits. This is
covered in the next section.

7.2

Flooding
Flooding is the term used to describe various conditions which cause a
loss of tray efficiency at high column throughput. These can be
entrainment, foaming or downcomer flooding. The first two are most
common and are due to poor vapor/liquid disengagement.
Downcomer flooding occurs where vapor velocity and hence pressure
drop in the column is so high that the liquid is prevented from flowing
down the column. It can be caused by an excessive reboil rate. To monitor
this, overhead vapor flow can be measured directly but it is usually an
expensive option on a large vapor line. An alternative, and more common
arrangement, is to measure the differential pressure across the column
using the trays themselves as orifices.
The differential pressure responds rapidly to heat input and can be used
as an override control on the column reboiler (Figure 29). Here the
temperature controller normally regulates the reboiler but if the ΔP rises
too high the ΔP controller takes over. The temperature controller is
fitted with external reset to prevent integral saturation when it is
overridden, the ΔPC would normally be a narrow band proportional action
only controller and so this is not appropriate.


FIGURE 29 EP OVERRIDE OF REBOIL TO PREVENT COLUMN FLOODING

7.3

Limiting Control
In 6.1 feed-forward control of reflux from feed was recommended to cope
with feed rate changes. This works well. A linear dependence on total feed
is usually good enough. Dynamic compensation may well be provided
and, in some circumstances, feedback adjustment of the ratio from a
composition measurement. A problem arises when the plant rates are
reduced. It is important to maintain some minimum liquid (and vapor) load
on the column to ensure separation. This can be done by providing a low
limit on the reflux (or reboil) flow. This is easily provided by a computing
element or a ratio relay and a low selector as shown in Figure 30.


Good descriptions of override limiting for reboilers, start-up, shutdown and
standby operation are given in (Ref.4). There are many other examples of
constraint control systems but most need to be tailor-made for the specific
problem.
FIGURE 30 LOW FEED FORWARD REFLUX FLOW LIMIT


For example, consider the following. Part of a column control system was set up
as shown in Figure 31. Here the bottoms off-take was set from a computer
optimization and the mass balance was normally achieved by reflecting back
changes in level to feed. This worked adequately under steady operation but
when the column was disturbed control became chaotic. There was a significant
dead time before changes in feed affected the level. In addition changes in feed
upset the general operation of the column and the preceding column. The fact
that there was additional heat integration between the two columns added to the
difficulty of the situation.
FIGURE 31 COMPUTER SET BOTTOM OFF>TAKE WITH LEVEL
REGULATING FEED VALVE

A better form of control, bearing in mind that the optimization requirements for a
particular bottom off-take were long term steady state objectives, would be to
implement the system shown in Figure 32.


FIGURE 32 FLOW CONTROLLER TUNED TO ENSURE THAT LONG TERM
DESIRED BOTTOM TAKE OFF-TAKE IS OBTAINED

Here, short term changes in mass balance are compensated by changes in the
bottoms off-take rate but, in the longer term (typically, -1 hr), the off-take is
brought back to its desired value by slow adjustment of the column feed. If it
were necessary to ensure a total quantity of bottoms off-take rather than a
specific rate in the steady state, some form of flow integration could be
provided.


8

MORE ADVANCED TOPICS

8.1

Temperature Position Control
Most distillation columns have a region of temperature sensitivity as
shown in Figures 4, 5 and 6 In such cases, a carefully chosen temperature
point located in the region of maximum rate of change of temperature is
usually the best place for temperature control if direct composition
measurement is not feasible.
Occasionally the temperature control profile is much "sharper" and
virtually the full column temperature change can occur over only a few
trays (see Figure 33). In this circumstance, control using a single
temperature point can be very difficult because of the very great system
gain change from the sensitive to insensitive regions. The consequence is
often a sustained limit cycle oscillation with frequent complete loss of
control.

FIGURE 33 'SHARP' TEMPERATURE PROFILE TEMPERATURE
HEIGHT


The answer is to invert the problem and, in effect, switch from controlling a
temperature on a particular tray to controlling a tray at a particular temperature.
Figure 34 illustrates the idea. A number of temperature points are installed
around the region of sensitivity and linear interpolation is used to locate the
position of the temperature of interest. This "position" is then
fed to a conventional controller where the set point is the desired position.
Clearly the technique will only work where there is an adequate provision of
temperature points; at least one point should always be measuring within the
region of profile sensitivity.
An example application and a detailed discussion of the technique is given in
(Ref.7), together with a software listing.
FIGURE 34 POSITIONAL TEMPERATURE CONTROLLER


8.2

Inferential Measurement

In this Guide, use of temperature as a substitute for direct composition control of
columns has been discussed. It has been argued that where it is possible to use
direct analysis with adequate safeguards, reliability, etc., this is to be preferred
when it is possible to measure the specific quality control variable directly.
Inferential measurement is a technique still in its infancy but gathering in
applications and popularity. It means using easily made measurements to infer
the value of a difficult-to-make (usually quality) measurement
It is fundamentally different from using temperature as a substitute for
composition measurement. Rather, a set of "secondary measurements", e.g.
temperature, feed-flow, reflux are used to infer (estimate) the value of the desired
or "primary" measurement. This might, for example, be a top composition. The
situation usually is that the top composition is available by other means on an
intermittent basis - perhaps from a chromatograph with a long sample delay
time or by labs analysis. The method involves the derivation of a mathematical
model which relates the infrequent primary measurement to the frequent
secondary measurements. This model is then used to infer the primary
measurement at the time frequency of the secondary measurements. The result
can either be used open loop to inform the operator of likely trends
in the primary measurement or, better, in a closed loop control system.
This technique is potentially useful for the determination of difficult-to-make
laboratory type quality measurements. A full description of the technique is given
in (Ref.8).

8.3

Floating Pressure Control

Precise control of column pressure is not necessary but it is common practice.
While there needs to be an upper limit to column pressure because of vessel
constraints, there is no similar lower limit and reboil heat transfer is enhanced. If
it were possible to reduce the pressure excessively there might be a limitation on
condenser heat transfer.
Within these limits it is advantageous to allow column pressure to float to achieve
the minimum possible under any circumstances. This maximizes the relative
volatility of most components allowing an increase in recovery or reduction in
energy in any given situation. Other advantages reported are reduced boiler
fouling as well as increased boiler capacity.

Resistance to floating pressure operation is traditional, probably based on the
necessity of holding pressure constant to achieve reliable temperature
measurement for control purposes. The increasing use of analyzers for direct
composition control makes this argument obsolete. Even where it is necessary to
use temperature control, temperature can be compensated for pressure
variations as discussed in 5.1.1.3.
While it would be possible to operate a total condensing distillation column with
no pressure control at all and with all condensation control valves removed, this
is not recommended. Sudden changes in pressure should be avoided. For
example, if a column with an air condenser is suddenly exposed to a rainstorm,
the wetted surfaces of the air condenser transfer more heat and there can be a
sudden drop in column pressure as a result. This fall in pressure may result in a
marked transient increase in boil-up and vapor rate as sensible heat is
converted into latent heat. In the extreme this can be enough to flood a column.
Certainly high boiling components will be moved higher up the column at the
expense of the tops specification.
What is required is a combination of short term protection against rapid pressure
changes combined with a control system which, in the long term, achieves the
lowest possible pressure operation. One way is to instruct the operators at all
times to reduce the set point of the pressure controller to the lowest possible
achievable limit with the condensation system working at a maximum. There are
two objections to this approach:
(a)

operators make small step changes which tend to upset unit operations;
and

(b)

they always operate cautiously and will soon stop making continuous
adjustments if not regularly exhorted to do so!

The correct solution is to provide a simple VPC to adjust the pressure set point to
ensure that it operates at minimum at all times. This can be done with the system
shown in Figure 35. Here, the VPC has, as its set point, a position of, say 90% or
10%, whichever corresponds to maximum condensation. The output of the VPC
sets the pressure controller set point which, in turn, adjusts the condenser valve.
The operators can easily revert to normal operation by breaking the cascade and
setting the pressure control set point manually.


If this happens it may be difficult to restore fully automatic operation because the
VPC would then integrate to an extreme. This is solved by feeding back, as
external reset, the column pressure. The short and long term responses of the
system are provided by the different settings of the PIC and VPC. The PIC is set
up conventionally to control the pressure tightly while the VPC, which should be
integral only, is set with the longer integral time constant to adjust the pressure
set point over a period of, say, half an hour.
FIGURE 35 DIAGRAM SHOWING HOW THE VALVE POSITION
CONTROLLER SLOWLY ADJUSTS THE PRESSURE SET
POINT TO KEEP THE CONDENSER FULLY LOADED IN THE
LONG TERM

8.4

Under the headings of Regulatory Control, Disturbance Control and
Constraint Control, methods have been described to circumvent problems
with the control of highly interactive distillation columns. This has been
necessary because of the use of single loop feedback controllers as
the basic unit of control.


These controllers are usually very effective at maintaining particular
variables at their desired values but they are operated in isolation one
from the other. They do not take account of the effect on other loops of
changes made to hold the variable at its set-point. Any system so
controlled is always subject to continuous disturbances because of this
interaction and the poor dynamic compensation provided by conventional
feedback controllers. Feed-forward, ratio and dynamic compensation
controls can be fitted to compensate for these interactions as described
earlier. In most cases the normal less-than-perfect control is accepted as
adequate and the ubiquitous single-loop PID controller is well established
in the industry.
In an increasing number of cases, particularly with more modern
intensively designed plants or where tight operation against constraints is
important, this form of control is definitely not adequate. Then, recognizing
and compensating for interactions and abnormal dynamics is necessary
and Model Based Predictive Control (MBPC) may be the answer.
The concept of MBPC is well described in (Ref.9). A particular version of
MBPC, Dynamic Matrix Control (DMC) is probably the leading technology
worldwide and has been very successfully applied in the refrigeration
complex within of a European plant. This is described in (Ref.10). An
alternative MBPC technology, An application to a reactor system is
described in (Ref.11).
The fundamental concept of MBPC is not difficult to understand. A
dynamic model of the process itself is derived from plant tests. This model
links all the control variables (outputs) to all of the important regulation
(controlled inputs) and disturbances (uncontrolled inputs). Thus it
is possible, using a plant computer and knowing the past history of
disturbances and the actions taken by the plant control system to
compensate for them, to predict into the future the inevitable effect on all
of the outputs (controlled variables). Since the future desired values of
all of the set-points can also be known it is possible to calculate an error
trajectory into the future. The purpose of the controller is then to minimize
this error by making a series of future moves which will compensate, so
far as is possible, this future error due to the historical actions. Clearly, in
addition to this feed-forward effect, there needs to be some feedback
correction to allow for model errors and unmeasured disturbances.


A well designed MBPC works well and some packages deal very
effectively with constraints so that plants with very high cash flows can be
operated very close to their real limitations. Interaction is also dealt with
very well since the controller "knows" about the interaction in advance and
makes the
appropriate set of control actions to meet the control objectives.
There are many applications of DMC applied to distillation control. It has a
particular value in multiple column applications where there are difficult
interaction problems due, for example, to heat integration making PID
control rather ineffective. Another significant success has been in
dual column arrangements where, for example, quality control of the
downstream column. The very long dead-times, process and analytical
delays inherent in such a form of control make automatic control by
conventional means virtually impossible. A successful application in this
area is reported in (Ref 12).
8.5

Control of Side-Streams
Multi-component mixtures can be separated into multiple products in one
column. This is economic in that only one column and a single heat source
and sink are required, but the product specification which is achievable will
be limited. Common applications of the multiple side-stream concept are
crude towers in refineries (see Figure 36) and pasteurizing columns, as
typically found on an olefin unit ethylene fractionator. A typical set-up is
shown in Figure 37, in this case there are two final products - low and high
grade ethylene. The top product contains virtually all the methane in the
splitter feed, with the high grade ethylene being removed as a side-stream
lower down the column. More methane could have been removed earlier
in the chain in the demethanizer but at the expense of a higher boil-up and
ethylene loss. In practice, the rate of tops off-take will be regulated to
control the purity of the side-stream product - almost certainly an
application for direct composition control.


FIGURE 36 DIAGRAM SHOWING HOW ALL THE HEAT USED IN CRUDE
OIL DISTILLATION ENTERS WITH THE FEED


FIGURE 37 DIAGRAM SHOWING HOW VIRTUALLY ALL THE METHANE
FROM THE BOTTOM OF THE DEMETHANISER LEAVES WITH
THE LOW>GRADE ETHYLENE

Side-streams can be vapor or liquid. Changes in the flow of a vapor sidestream can have a marked effect on the vapor loading above the off-take
point while changes in a liquid off-take can significantly affect the internal
reflux below the off-take. The consequent effects should be considered
when designing overall control systems. In practice, for vapor side-stream
off takes a form of achieving constant vapor flow above the side-stream
such as ΔP control of reboil may be indicated.

For liquid off-takes constant internal reflux below the side-stream may
be desirable and the control system can be designed to achieve this.
There can be severe interactions in multiple side-stream columns. Where
this is experienced or expected a model based predictive control system
may be well worth considering (see 8.4).

8.6

Extractive/Azeotropic Systems
The subject of extractive and azeotropic distillation is complex. It is
discussed in some detail in (Ref.4).

8.6.1 Extractive Distillation
Extractive distillation uses a solvent to improve the relative volatility of the
components of interest. The solvent is selected to have a particular affinity
for a class of components and in this way acts very much like a selective
absorbent. At least two columns are required. In the first column, in the
case of a binary mixture, the extractant attaches itself to one component
while the other is distilled. In the second column the captured component
is stripped from the extractant which is then recycled. In many respects
the process is similar to conventional absorber/stripper combinations.
In both columns, because of the presence of the extractant, temperature
profiles can be confusing and care shall be taken to select the appropriate
sensitive region if temperature is used as the means of composition
control.
8.6.2 Azeotropic Distillation
An Azeotrope is a mixture of two or more volatile components of identical
vapor and liquid compositions and equilibrium. The azeotropic mixture will
boil at a higher or lower temperature than its components depending on
the nature of the system. If one attempts to distil a mixture of two
components forming an azeotrope only one of the components can be
removed, the azeotrope becoming the other product. Since the azeotrope
cannot be separated by conventional distillation, other methods such as
extraction, need to be combined with distillation.


Azeotropes may also be either homogeneous or heterogeneous. The
latter separates into two liquid phases when condensed from the vapor.
This makes their separation easier. One method of separating
homogeneous azeotropes is extractive distillation.
The heterogeneous azeotrope can be easier to separate than many ideal
mixtures. The immiscibility of the liquid phases condensed from the vapor
effectively breaks the azeotrope. This allows complete separation of the
components. This property can be used to advantage to assist in this
separation of close boiling mixtures by deliberately forming an azeotrope.
The approach is to add a third component, form a low boiling
heterogeneous azeotrope with one of the other difficult to separate
components.
This often means introducing into the column the additional azeotroping
stream, usually of a component already present, water perhaps. The
additional control problem is then that of regulating this extra "feed" to
maintain a sufficiently high, but not excessive, concentration of
the azeotroping agent in the appropriate region of the column. If this is to
be done properly, it is usually necessary to use direct composition control
to regulate the azeotroping agent flow. This will be in addition to any other
composition control system. Interaction is a real probability and a
poorly set up or maintained control system can lead to very poor and
inefficient operation. An example of an effective control system for an
azeotropic separation is found in (Ref.13).


9

REFERENCES

(1)

Control of Distillation Columns by R Jackson & J S Anderson,
Billingham Division Central File Report B.125,196, 22 October
1962.

(2)

PEG.HEA.103 "Selection and Design of Condensers" P.D.Hills,
1992.

(3)

GBHE-PEG-HEA-513"Air Cooled Heat Exchangers" P.D.Hills,
1991.

(4)

Distillation Control, F G Shinskey, McGraw Hill, 1977.

(5)

PEG.HEA.102 "Selection of Reboilers for Distillation Columns"
P.D.Hills, 1992.

(6)

Design of Distillation Column Control Systems, Buckley, Luyben,
Shunt, ISA, 1985.

(7)

Control of the Cumene Column Temperature Profile by T H C
Ankcorn, EDN7015, June 1989.

(8)

Inferential Measurement by M J Oglesby, Report No IC05895,
January 1991.

(9)

Model Based Predictive Control by M J Oglesby, IC05923, July
1991.

(10)

ICEE Award 1990-1991, "The Application of DMC to Olefines 6" by
J S Anderson, M J Oglesby and N J Cordingley, March 1991.

(11)

Application of Model Based Predictive Control to a Hydrocarbon
Processing Plant by H v Spreckelsen et al, IChemE Symposium,
Advances in Process Control III, York, 23-24 September 1992.

(12)

Dynamic Matrix Control on Benzene and Toluene Towers by D
Tran and C R Cutler, Instrument Society of America, Philadelphia,
22-27 October 1989.

(13)

Problems in the Control of Distillation Columns by J S Anderson
and J McMillan, International Symposium on Distillation I.Chem.E.,
Brighton, September 1969.


DOCUMENTS REFERRED TO IN THIS PROCESS ENGINEERING GUIDE
This Process Engineering Guide makes reference to the following documents:
ENGINEERING GUIDES
GBHE-PEG-HEA-513

Air Cooled Heat Exchangers
(referred to in Clause 9)

GBHE-PEG-HEA-507

Selection of Reboilers for Distillation Columns

GBHE-PEG-HEA-508

Selection and Design of Condensers


Control of Continuous Distillation Columns

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