8.19
Distillation: Basic Controls
H. L. HOFFMAN, D. E. LUPFER
L. A. KANE
(1985)
B. A. JENSEN
B. A. JENSEN, B. G. LIPTÁK
V
(1970)
(1995)
TR
D
R or L
(2005)
Ri or Li
F
Q
V
LF
V
Q
LF
B
Flow sheet symbol
INTRODUCTION
Distillation is the most common class of separation processes
and one of the better understood unit operations. It is an
energy-separating-agent equilibrium process that uses the difference in relative volatility, or differences in boiling points,
of the components to be separated. It is the most widely used
method of separation in the process industries. The distillation
process will most often be the choice of separation unless the
following conditions exist:
•
•
•
•
Thermal damage can occur to the product.
A separation factor is too close to unity.
Extreme conditions of temperature or pressure are
needed.
Economic value of products is low relative to energy
costs.
Control involves the manipulation of the material and
energy balances in the distillation equipment to affect product
composition and purity. Difficulties arise because of the multitude of potential variable interactions and disturbances that
can exist in single-column fractionators and in the process
that the column is a part of.
Even seemingly identical columns will exhibit great
diversity of operation in the field. Therefore, this section will
not attempt to provide control strategies that can be applied
1820
© 2006 by Béla Lipták
to columns in a “cookbook” fashion. Instead, discussion will
begin with a basic description of the distillation process and
equipment, followed by techniques used to derive a mathematical column model.
The presentation in this section will then describe methods to evaluate interactions and alternative control strategies;
control models used for some product quality, pressure, and
feed flow control strategies; and finally some common feedforward advanced regulatory control strategies commonly
used in the regulation of fractionators.
The goal of this section is to provide the process control
engineer with the tools necessary to design unique control
strategies that will match the specific requirements of distillation columns.
General Considerations
Distillation separates a mixture by taking advantage of the
difference in the composition of a liquid and that of the vapor
formed from that liquid. In the processing industries, distillation is widely used to isolate and purify volatile materials.
Thus, good process control of the distillation process is vital
to maximize the production of satisfactory purity end products.
Although engineers often speak of controlling a distillation tower, many of the instruments actually are used to
control the auxiliary equipment associated with the tower.
For this reason, the equipment used in distillation will be
discussed.
8.19 Distillation: Basic Controls
DISTILLATION EQUIPMENT
Condenser
There are some basic variations to the distillation process.
One such basic difference is between continuous and batch
distillation. The main difference between these processes is
that in continuous distillation the feed concentration is relatively constant, while in batch distillation it is rich in light
components at the beginning and lean in light components
at the end. While batch distillation is also described in this
section, the emphasis is on the continuous processes.
Another basic difference is in the way the condenser heat
is handled. The more common approach is to reject that heat
into the cooling water and thereby waste it. This necessitates
the use of “pay heat” at the reboiler, which usually is a large
part of the total operating cost of the column. An alternate
approach, also discussed in this section, is “vapor recompression” (Figure 8.19a), in which the heat taken out by the
condenser is reused at the reboiler after a heat pump (compressor) elevates its temperature. While vapor recompression
controls are also discussed in this section, the emphasis is on
the traditional air- or water-cooled condenser designs.
The Column
The primary piece of distillation equipment is the main tower.
Other terms for this piece of equipment are column and fractionator, and all three terms are used interchangeably. The
tower, column, or fractionator has two purposes: First, it separates a feed into a vapor portion that ascends the column and
a liquid portion that descends; second, it achieves intimate
D
Removed
heat
wasted
D
M
F
F
B
Pay heat
added
Compr.
Work
B
Recovered
heat
FIG. 8.19a
In contrast with conventional distillation, the vapor recompression
system uses recovered heat.
© 2006 by Béla Lipták
1821
Column
Accumulator
Feed
pump
Reflux
pump
Preheater
Reboiler
FIG. 8.19b
Distillation equipment.
mixing between the two countercurrent flowing phases. The
purpose of the mixing is to get an effective transfer of the
more volatile components into the ascending vapor and corresponding transfer of the less volatile components into the
descending liquid. The other equipment associated with the
column is shown schematically in Figure 8.19b.
In continuous distillation, the feed is introduced continuously into the side of the distillation column. If the feed is
all liquid, the temperature at which it first starts to boil is
called the bubble point. If the feed is all vapor, the temperature at which it first starts to condense is called the dew
point. The feed entering the column is normally operated in
a temperature range that is intermediate to the two extremes
of dew point and bubble point. However, some optimization
strategies may call for designs where the feed is either superheated or subcooled. For effective separation of the feed, it
is important that both vapor and liquid phases exist throughout the column.
The separation of phases is accomplished by differences
in vapor pressure, with the lighter vapor rising to the top of
the column and the heavier liquid flowing to the bottom. The
portion of the column above the feed is called the rectifying
section and below the feed is called the stripping section.
Packing and Trays The intimate mixing is obtained by one
or more of several methods. A simple method is to fill the
column with lumps of an inert material, or packing, that will
provide surface for the contacting of vapor and liquid.
Another effective way is to use a number of horizontal plates,
or trays, which cause the ascending vapor to be bubbled
through the descending liquid (Figure 8.19c).
1
Tray designs are numerous and varied. Tray designs
include bubble cap plate unit, valve, sieve plate, tunnel,
dual-flow, chimney, disc-and-donut, turbogrid trays, v-grid,
Perform-Kontakt, Haselden baffle tray, Kittel trays, and other
specialty-type units. Dualflo® trays, Flexitray®, Varioflex®,
Bi-Frac®, Max-Frac®, NYE Trays®, Superfrac® trays,
Super-Flux® trays, and Ultra-Frac® trays are specialty registered tray designs from different manufacturers that are
variations of the aforementioned tray designs. Bubble caps
1822
Control and Optimization of Unit Operations
•
L
•
•
L
•
L
FIG. 8.19c
Intimate contact and therefore equilibrium is obtained as the vapor
bubbles ascend through the liquid held up on each tray, as the liquid
descends down the column.
and sieve trays are the most common designs used in distillation applications.
2
Many different types of packings are available. They are
normally classified as random or stacked. Random packings
are those that are dumped into the containing shell. Raschig
rings, Berl saddles, Intalox saddles, and Pall rings are the
most common random packings and come in various sizes
from 1/2 to 31/2 in (1.25 to 9 cm).
Stacked packings, also known as grid or stacked packing,
include large-sized Raschig rings and Lessing rings. Packings
generally give lower pressure drops at the cost of higher
installation costs. They are made of ceramic, plastic, or metal,
depending upon the type of packing and the intended application. Other packings such as Maspac®, HyPak®, Tellerette®, IMTP® FLEXIPAC® KATAMAX®: FLEXIGRID®-2,
-3, and -4, and KOCH-GLITSCH GRID® EF-25A are specialty registered packings from different manufacturers that
are just variations of the aforementioned packings.
When deciding between the use of trays and packing, the
3
following factors should be considered:
•
•
•
•
•
•
•
Because of liquid dispersion difficulties in packed towers, the design of plate towers is considerably more
reliable and requires less safety factor when the ratio
of liquid mass velocity to gas mass velocity is low.
Towers using trays can be designed to handle wider
ranges of liquid rates without flooding.
Towers using trays are more accessible for cleaning.
Towers using trays are preferred if interstage cooling
or heating is needed because of lower installation costs
of delivery piping.
Towers using trays have a lower total dry weight,
though total weight with liquid hold-up is probably
equal.
Towers using trays are preferred when large temperature changes are expected because of thermal expansion or when contraction may crush packing.
Design information for towers using trays is generally
more readily available and more reliable.
© 2006 by Béla Lipták
•
Packed towers are cheaper and easier to construct than
plate towers if highly corrosive fluid must be handled.
Diameters of packed towers are generally designed to
be less than 4 ft, while plate tower diameters are
designed to be more than 2ft.
Packed towers are preferred if the liquids have a large
tendency to foam.
The amount of liquid hold-up is considerably less in
packed towers.
The pressure drop through packed towers may be less
than for plate towers performing the same service, making packed towers desirable for vacuum distillation.
Thus, generally, trays work better in applications requiring high flow, such as those encountered in high-pressure
distillation columns, such as depropanizers, debutanizers,
xylene purification columns, and the like. Packing works best
at lower flow parameters, as the low-pressure drop of structured packing makes it very attractive for use in vacuum
columns or ethylbenzene recycle columns of styrene plants.
The contacting between the vapor and liquid in a singlestage contacting device will not produce total equilibrium.
The relationship between ideal and actual performance is the
efficiency that translates the number of ideal separation stages
into actual finite stages that must be used to accomplish the
desired final separation. Efficiency varies, not only with the
type of mixing method used (e.g., packing or trays), but also
with fluid rates, fluid properties, column diameter, and operating pressure.
The influence of plate efficiency in the operation of
the distillation tower becomes important in the control of the
overhead composition. Because plate efficiencies increase
with increased vapor velocities, the influence of the refluxto-feed ratio on overhead composition becomes a nonlinear
relationship.
Dynamics Dynamic considerations due to liquid hold-up
on the trays comes into play when discussing distillation
control. Because the liquid on each tray must overflow its
weir and work its way down the column due to tray or
packing hydraulics, this change will not be seen at the bottoms of the tower until some time has passed. The exact
dynamics depend on column size, type of tray, number of
trays, and tray spacing. The hold-up at each tray as shown
in Figure 8.19c can be modeled by the LaPlace transform of
the form
KG ( s ) =
where
KG(s)
K
T1
S
=
=
=
=
K
(T1 s + 1)
transfer function
system gain
time constant
LaPlace transfer operator
8.19(1)
8.19 Distillation: Basic Controls
X(t)
Y1
Y
4
Y2
Y 1(t)
Y1
Y40
0
X(t)
Y 2(t)
Y4(t)
Y10(t)
Y40(t)
Y1
Y2
Y4
Y10
Y40
=
=
=
=
=
1 Lag
2 Lags
4 Lags
10 Lags
40 Lags
Sum of the time constants are equal.
Time
These lags are cumulative as the liquid passes each tray
on its way down the column. Thus, a 30-tray column could
be approximated by 30 first-order exponential lags in series
of approximately the same time constant.
K
(T1s + 1)n
In that case, the condensers are called partial condensers. In
this instance, a vapor product is normally withdrawn as well
as a liquid product.
A total condenser is usually designed for accumulator
pressures up to 215 psia (1.48 MPa) at an operating temper4
ature of 120°F (49°C). A partial condenser is used from
215 psia to 365 psia (1.48 to 2.52 MPa), and a refrigerant
coolant is used for the overhead condenser if the pressure is
greater than 365 psia (2.52 MPa).
Common condensers include fin fans and water coolers.
However, in order to improve efficiency of heat recovery, heat
exchange with another process stream is often performed.
Propane is the most common refrigerant used. A pressure
drop of 5 psia (34.4 KPa) across the condenser is often
assumed if no measurements are available. The condenser
and accumulator are the key pieces of equipment with respect
to controlling pressure in the column.
Reboilers
FIG. 8.19d
Response of nth-order lags to unit step change.
KG (s) =
8.19(2)
The liquid leaving the bottom of the column is reheated in a
reboiler. A reboiler is a special heat exchanger that provides
the heat necessary for distillation. Part of the column bottoms
liquid is vaporized and the vapors are injected back into the
column as boil-up. The remaining liquid is withdrawn as a
bottom product or as residue.
As shown in Figure 8.19e, reboilers come in widely varying designs. They can be internal, but most are external to
the column. They can use natural or forced circulation.
where
n = 30 for a 30-tray column
Figure 8.19d shows the response of nth order lags to a
unit step change. The effect of increasing the number of lags
in series is to increase the apparent dead time and increase
the response curve slope. Thus, the liquid traffic within the
distillation process is often approximated by using a secondorder lag plus dead time as modeled by the LaPlace transform:
KG (s) =
Ke − t s
(T1s + 1)(T2s + 1)
e = e of log to the base e
φ = dead time
T1, T2 = time constants
Condensers
The overhead vapor leaving the column is sent to a condenser
and is collected as a liquid in a receiver, or accumulator. A part
of the accumulated liquid is returned to the column as reflux.
The remainder is withdrawn as overhead product or distillate.
In many cases, complete condensation is not accomplished.
V
Q
Q
L
B
B
Internal
8.19(3)
where
© 2006 by Béla Lipták
1823
External kettle
V
V
Q
Q
B
Vertical
thermosyphon
B
Horizontal
thermosyphon
FIG. 8.19e
Reboiler design variations. External kettle reboilers often use forced
circulation (pump), while the thermosyphon designs depend on natural circulation. The horizontal thermosyphon reboiler takes its
liquid from the bottom tray, while the others take it from the column
bottoms.
1824
Control and Optimization of Unit Operations
The kettle reboiler is the most common external forced circulation design.
Vertical and horizontal thermosiphon reboilers operate
by natural circulation. In these, flow is induced by the hydrostatic pressure imbalance between the liquid inside the tower
and the two-phase mixture in the reboiler tubes. In forced
circulation reboilers, a pump is used to ensure circulation of
the liquid past the heat transfer surface. Reboilers may be
designed so that boiling occurs inside vertical tubes, inside
horizontal tubes, or on the shell side.
A newer development in reboiler design is the concept
self-cleaning shell-and-tube heat exchangers for applications
where heat exchange surfaces are prone to fouling by the
process fluid. Common heat sources include hot oil, steam,
or fuel gas (fired reboilers). Cases where simple heat exchange
with another process stream is used for efficiency of heat
recovery are common. Thus, the choice of instrumentation to
control heat addition to the tower depends upon the type of
reboiler used.
Interheaters/Intercoolers
In some cases, additional vapor or liquid is withdrawn from
the column at points above or below the point at which the
feed enters. All or a portion of this sidestream can be used
as intermediate product. Sometimes, economical column
design dictates that the sidestream be cooled and returned to
the column to furnish localized reflux. The equipment that
does this is called a sidestream cooler, or intercooler. Multiproduct fractionators often have these intercoolers in a pumparound stream.
At other times, localized heat is required. Then, some of
the liquid in the column is removed and passed through a
sidestream reboiler, or interheater, before being returned to
the column. Interheaters are usually utilized in cryogenic
demethanizers.
Often the feed is preheated before entering the column.
Common preheat mediums include the bottoms product or
low-pressure steam. Preheating is often a convenient method
to recover heat that would otherwise be wasted.
Column Variables
Controlling a fractionator requires the identifying of the controlled, manipulated, and load variables (Figure 8.19f). Controlled variables are those variables that must be maintained
at a precise value to satisfy column objectives. These normally include product compositions, column temperatures,
column pressure, and tower and accumulator levels.
Manipulated variables are those variables that can be
changed in order to maintain the controlled variables at their
desired values. Common examples include reflux flow, coolant flow, heating medium flow, and product flows. Load variables are those variables that provide disturbances to the
column. Common examples include feed flow rate and feed
© 2006 by Béla Lipták
L
Overhead product
(D)
Feed
Steam
(V)
Bottom product
(B)
Apparent
variables:
C1
C2
C3
C4
u1
u2
u3
u4
u5
u6
u7
u8
u9
m
=
=
=
=
=
=
=
=
=
=
=
=
=
=
Independent
variables
overhead temperature
overhead pressure
overhead composition
overhead flow rate
bottom temperature
bottom pressure
bottom composition
bottom flow rate
feed temperature
feed pressure
feed composition
feed per cent vapor
feed flow rate
steam flow rate (heat input)
2
1
2
1
2
1
1
1
11
FIG. 8.19f
In a binary distillation process the number of independent variables
is eleven (11) and the number of defining equations is two (2).
Therefore, the number of degrees of freedom is nine (9), which is
the maximum number of automatic controllers that can be used on
such a process.
composition. Other common disturbances are steam header
pressure, feed enthalpy, environmental conditions (e.g., rain,
barometric pressure, and ambient temperature), and coolant
temperature.
To handle these disturbances, column controls can be so
designed as to make the column insensitive to these disturbances, or secondary controls can be designed to eliminate
the disturbances. It is also important to evaluate the expected
magnitude and duration of the likely disturbances, so that
proper control system scaling and tuning can be achieved.
Feedforward controls are designed to compensate for
these disturbance variables and are discussed later in this
section. There are other advanced control or optimization
methods that can be designed to compensate for these disturbance variables. They are discussed in Section 8.21.
Pairing of Variables The variables that should be controlled
are usually obvious. They are normally identified when process objectives are defined and understood. Load variables
are also easily identified. But identification of the manipulated variables can be more difficult. The general guidelines
for identifying which manipulated variables to associate with
which controlled variables are
8.19 Distillation: Basic Controls
•
•
•
•
•
Manipulate the stream that has the greatest influence
on the associated controlled variable.
Manipulate the smaller stream if two streams have the
same effect on the controlled variable.
Manipulate the stream that has the most nearly linear
correlation with the controlled variable.
Manipulate the stream that is least sensitive to ambient
conditions.
Manipulate the stream least likely to cause interaction
problems.
Unfortunately, the decision on pairing controlled and
manipulated variables is complicated by the fact that the
above rules may sometimes result in conflicting recommendations. Section 8.20 provides information on relative gain
calculations, which can help to optimize the pairing of controlled and manipulated variables. Once the pairings are completed, the equations are then solved for the manipulated
variables in terms of the controlled and load variables. In that
form, the equations are the mathematical representations of
the control systems.
MODELING AND CONTROL EQUATIONS
The primary application of instruments in distillation is to
control the product purity, and secondarily, to minimize
upsets to the unit caused by a change in process inputs. The
instruments calculate the effects of the input changes and
determine the corrective action needed to counteract them.
The control actions are implemented by direct manipulation
of the final control elements or by alteration of the set points
of lower level controllers.
A careful analysis of limits and operating constraints is
essential to the successful control of distillation columns. If
the system is not designed to provide limit checks and overrides to handle operating limits, frequent operator intervention will be required during upsets. This is likely to result in
a lack of confidence in the control system and will cause the
operators to remove the column from automatic control more
often than necessary, thereby not only reducing the effectiveness of the system, but also reducing safety.
The first step in the design of a good control system is
the derivation of a process model. Knowing the defining
equations, the manipulated variables can be selected, and the
operating equations for the control system can be developed.
The instrumentation is then selected for the correct solution
of these equations.
The final control system can be relatively simple or can
be a complex, interacting, multicomponent, computer-based
system. In the discussion that follows, the procedures for
designing distillation controls is followed by examples of the
more common applications in distillation column control. A
more detailed discussion of alternative strategies and
advanced distillation column controls will be presented in
Section 8.21.
© 2006 by Béla Lipták
1825
Steady-State Model
The first step in the design of a control system must be the
development of a process model. Frequently omitted in simple distillation columns, this step is essential to minimize the
need for field reconfiguration of control strategies. Even with
easily reconfigurable process automation systems (PASs), the
development of the model is essential to fully understanding
the process.
The model defines the process with equations developed
from the material and energy balances of the unit. A common
simplifying assumption is that all components of the feed
have equal heats of vaporization, which leads to the assumption of equimolal overflow. Most shortcut fractionation calculations are based upon this underlying assumption.
The model is kept simple by the use of one basic rule:
The degrees of freedom limit the number of controlled variables (product compositions) specified in the equations, as
was illustrated in connection with Figure 8.19f. Some of the
variables that can be manipulated to control a column are
shown in Figure 8.19g.
Material Balance For example, for a given feed rate only
one degree of freedom is available for material balance control. If overhead product (distillate) is a manipulated variable
(controlled directly to maintain composition), then the bottom product cannot be independent but must be manipulated
to close the overall material balance according to the following equations:
F=D+B
Accumulation = Inflow − Outflow
Accumulation = F − (D + B)
8.19(4)
8.19(5)
8.19(6)
Because accumulation is zero at steady state, B is dependent upon F and D, as expressed by Equation 8.19(4):
B=F−D
8.19(7)
V
Heat removed
Pressure
Reflux
rate (L)
Feed
temperature,
Distillate rate (D)
Composition (Y)
Composition (Z)
and rate (F)
Heat added (boilup)
Bottom rate (B)
Composition (X)
FIG. 8.19g
Variables that fix the distillation operation.
1826
Control and Optimization of Unit Operations
or if the bottoms product is the manipulated variable:
D=F−B
Di = V − Li
8.19(8)
Li = L[1 + (Cp/∆H) × (To − Tr)]
If the compositions of the feed, distillate product, and
bottoms product are known, then the component material
balance can be solved:
8.19(9)
8.19(10)
8.19(11)
V = VB + VF × F
F (z)
where:
%LLKF = lighter than light key in the feed (mol%)
%LKF = light key in the feed (mol%)
%LLKD = lighter than light key in the distillate product
VB = QB/∆H
(mol%)
%LKD = light key in the distillate product (mol%)
%HKD = heavy key in the distillate product (mol%)
%LKB = light key in the bottoms product (mol%)
In the most general case, the feed might have four components, having the concentrations of LLKF , LKF , HKF , and
HHKF . Three of these components appear in each of the
bottom and overhead products. The separation of the column
is fixed by specifying the heavy key component in the overhead product HKD and the concentration of the light key
component in the bottom product LKB.
Equations 8.19(9) to 8.19(11) assume no heavier than
heavy key is found in the distillate and that no lighter than light
key is found in the bottoms. Rearranging Equation 8.19(11)
gives
%LKD = (F • %LKF − B • %LKB)/D
8.19(12)
Substituting Equation 8.19(8) into Equations 8.19(10)
and 8.19(12) gives
%LLKD = (F • %LLKF)/(F − Β)
8.19(13)
%LKD = (F • %LKF − B • %LKB)/(F − B) 8.19(14)
Substituting Equations 8.19(13) and 8.19(14) into
Equation 8.19(9) to eliminate %LLKD and %LKD:
B/F =
(100 − %HK D − %LLK F − %LK F )
(100 − %HK D − %LK B )
8.19(15)
For a given feed composition and desired product compositions, only one bottoms-to-feed ratio, B/F (product split),
will satisfy the overall and component material balances. By
fixing the bottoms flow, the distillate flow will be fixed.
© 2006 by Béla Lipták
D(y)
L @ Tr
where:
F = feed rate (the inflow)
D = overhead rate (an outflow)
B = bottoms rate (an outflow)
100 = %LLKD + %LKD + %HKD
D × %LLKD = F × %LLKF
F × %LKF = D × %LKD + B × %LKB
QT
To
B
=
D
=
L
−
Lf −
Li −
QB −
QT −
VB −
VF −
∆H −
∆HD −
∆HL −
∆HLi −
Bi if no accumulation occurs in
the column bottoms.
Di if no accumulation occurs in
the accumulator
External reflux
Liquid flow below feed tray
Internal reflux
Heat addition at bottom
Heat removal at top
Vapor boilup rate
Vapor fraction in feed
Heat of vaporization in reboiler
Heat of condensation of distillate
Heat of vaporization of reflux
Heat of condensation of
internal reflux
Lf = Li + (1 − VF) × F
QB
Bi = Lf − VB
B(x)
Material balance: F = D + B
separation is the energy/feed
ratio of a column. For binary
y(1 − x)
process: S =
x(1 − y)
Separation should be controlled
by the more pure product.
FIG. 8.19h
Energy balance equations can be used to describe the steady-state
heat flow model of a distillation column.
However, fixing a value of product split does not fix either
the distillate or bottoms composition because many combinations of %LLKF , %LKF , %LKB, and %HKD could yield the
same value of B/F.
Energy Balance The energy balance and the separation
obtained are closely related. Conceptually, product composition
control can be thought of as a problem of the rate of heat
addition QB at the bottom of the fractionator and the rate of heat
removal QT at the top of the column. A series of energy balances
produces additional equations. Figure 8.19h shows a steady5
state internal model of these equations.
The vapor boil-up rate VB equals the heat QB added by
the reboiler divided by the heat of vaporization (∆H) of the
bottoms product:
VB = QB /∆H
8.19(16)
The vapor rate V above the feed tray equals the vapor
boil-up rate plus the vapor entering with the feed (feed rate
8.19 Distillation: Basic Controls
F times vapor fraction VF , provided the feed is neither subcooled nor superheated):
V = VB + F × VF
8.19(17)
The internal reflux rate, that is, the liquid at the top tray
of the column is derived by a heat balance around the top of
the tower. Assuming a steady-state heat balance where the
heat into the tower equals the heat out:
D × ( ∆ H D + C pD × Tt ) + LI × ( ∆ H LI + C pR × Tt )
being used by the control equation. Also, C pL and ∆HL should
be calculated near the existing pressure and temperature of
the external reflux.
The liquid rate, LF , below the feed tray equals the internal
reflux plus the liquid in the feed:
LF = LI + (1 − VF) × F
D = V − LI
8.19(18)
+ L × ( ∆ H L + C pL × To ) + LI × (C pL × Tt )
I
where
Cp = specific heat
To = overhead vapor temperature (vapor at its dew point)
L = external reflux
Tr = external reflux temperature
LI = internal reflux
Tt = top tray temperature (liquid at its bubble point)
Equation 8.19(18) reduces to:
D × C pD × (Tt − To ) + LI × ∆H LI − L × ∆H L
+ L × C pL × (Tr − To ) = 0
The bottoms rate, B, equals the liquid rate, L, minus the
boil-up, VB :
B = L − VB
Because the tower doesn’t always operate at steady state, it
is essential to also account for the dynamics of the process.
This necessitates extending the steady-state internal flow
model and requires additional considerations. Figure 8.19i
6
shows the internal flow model that includes dynamics.
Di = V − Li
DA= Di − D
QT
or
To
or
zI = L × [1 + K1 × ∆T]
V = VB + VF × F
8.19(22)
GT
F
GT & GB are second order lags
GB
8.19(23)
VB = QB/∆H
8.19(24)
Note: This equation is valid for whatever units are used
for C pL or ∆HL. Because specific heat and heat of vaporization are nearly always in mass units, care must be taken to
account for density differences whenever volume units are
© 2006 by Béla Lipták
D
Li = L[1 + (Cp/∆H) × (To − TR)]
If a total condenser is employed, the composition of the
internal reflux and external reflux are the same, i.e.,
∆H LI = ∆H L , so the constant K2 = 1.0. Thus,
LI
[1 + K1 (TO − Tr )]
L @ Tr
8.19(21)
resulting in the equation
L=
8.19(27)
The criterion for separation is the ratio of reflux (L) to
distillate (D) flows vs. the ratio of boil-up (V) to bottoms (B)
flow rates. Manipulating reflux affects separation equally as
well as manipulating boil-up, albeit in opposite directions.
Consequently, only one degree of freedom exists to control
separation. Thus, for a two-product tower, two equations
define the process. One is an equation describing separation,
and the other is an equation for material balance.
LI × ∆H LI = L × ∆H L + L × C pL × (To − Tr ) 8.19(20)
LpI /L = K2 × [1 + K1 × (Tpo − Tpr)]
8.19(26)
Dynamic Model
8.19(19)
Making a simplifying assumption that the tray temperature equals overhead vapor temperature (i.e., the dew point
of the vapor equals the bubble point of the liquid; Tt = To)
produces:

C pL
Li ∆H L 
=
⋅  1.0 +
⋅ (To − Tr ) 
L ∆H Li 
∆H L
 

8.19(25)
The distillate rate, D, equals the vapor rate, V, above the
feed tray minus the internal reflux:
I
+ L × (C pL × Tr ) = D × ( ∆ H D + C pD × To )
1827
QB
FIG. 8.19i
Dynamic internal flow model.
DA & BA represent accumulations in
the accumulator and the column
bottoms respectively.
Li = GB[GT Li + (1 − VF) × F)
Bi = Li − VB
BA = Bi − B
B
1828
Control and Optimization of Unit Operations
Because a change in the reflux rate must work its way
down the column due to tray or packing hydraulics, this
change will not be seen at the reboiler until some time has
passed. The holdup at each tray has previously been modeled
by the LaPlace transform of Equation 8.19(1). This Laplace
transform can be converted to a simple first-order exponential
lag equation of the form, which describes the response to a
step change in input:
−t
Llag = L (1 − e )
8.19(29)
where
− t1
−t 2
GB = φ1 (1 − e ) (1 − e )
−t3
−t 4
GT = φ2 (1 − e ) (1 − e )
φ1 and φ2 are the dead times
while the orientation of separation for a given degree of
separation is defined as
Orientation of Separation =
%HK D
%LK B
S=
y(1 − x )
x (1 − y)
8.19(33)
8.19(34)
where
x = mole fraction of the key light component the distillate
(%LKD)
y = mole fraction of the key light component in the bottoms, (LKB)
The relationship between separation (S) and the ratio of
boil-up to feed (V/F) over a reasonable operating range is
V/F = a + bS
8.19(35)
where a and b are functions of the relative volatility, the number of trays, the feed composition, and the minimum V/F. The
control system therefore computes V based on the equation:
GB and GT are the solution to the LaPlace transform of
Equation 8.19(3).
Changes in boil-up rates are observed at the condenser
in a matter of seconds. Normally, no dynamic terms are
necessary for vapor streams, as the value of use of computing resources to that of the benefits by compensating for the
dynamics is negligible.
The liquid inventory in the condenser or associated accumulator will change during unsteady-state actions. In the
unsteady state, the difference DI − D is the rate of accumulation of material in the accumulator. Similarly for the liquid
inventory at the bottom of the tower (the kettle), the difference BI − B is the rate of accumulation:
8.19(30)
8.19(31)
where
DA is the accumulation in the overhead accumulator
BA is the accumulation in the tower bottoms
Separation Equations
The control of product compositions for a fractionator is primarily a matter of control of the internal flows. In considering
© 2006 by Béla Lipták
(%LK D × %HK B )
8.19(32)
KB )
(%HK D × %LK
The relationship between x (the light key component) and
7
the energy balance was developed by Shinskey as a function
of separation S:
These lags are cumulative as the liquid passes each tray
on its way down the column. However, implementation of
multiple first-order lags is impractical. Fortunately, it can be
shown that multiple lags in series can be approximated by a
dead time and a second-order exponential lag as shown by
the LaPlace transform of Equation 8.19(3). For this reason,
two dynamic terms (GT and GB) are included in Figure 8.19i.
Equation 8.19(25) is then rewritten as
DA = DII − D
BA = BI − B
Degree of Separation = ln e
8.19(28)
where
L is the liquid incoming to the tray
Llag is the liquid leaving the tray
t is the time constant
L = GB[GT LI + (1 − VF) × F]
product separation, the degree of separation and the orientation of separation are important. The degree of separation is

 y(1 − x )  
V = F a + b 

 x (1 − y)  

8.19(36)
Because y is held constant, the bottom composition controller adjusts the value of the parenthetical expression if an
error should appear in x. Let V/F = y(1 − x)/(1 − y), and the
control equation becomes:
V = F(a + b[V/F])/x
8.19(37)
where [V/F] = the desired ratio of boil-up to feed.
Figure 8.19j illustrates four of the most common basic
controls for the flows and levels of a two-product fractionator,
where it is assumed that feed flow and tower pressure are
held constant. A different set of the above control equations
for controlling internal product flow rates will apply, depending upon the configuration of instrumentation used.
Scaling
The form of the control system equations influences the computing functions required. Boolean operands, such as high
and low selectors, and dynamic functions, such as dead times,
lead, and lag function, are also used. Most process automation
systems have these basic computing function blocks. Implementation in a distributed control system (DCS), programmable
8.19 Distillation: Basic Controls
PSP
PC
PSP
PC
LC
LC
D
L
FC
D
L
FC
FC
FC
DSP
F
1829
DSP
F
LSP
LSP
QSP
FC
FC
QSP
BSP
BSP
LC
LC
FC
FC
Q
Q
B
B
Case 1
Case 2
PSP
PC
LC
LC
D
L
FC
F
D
L
FC
FC
FC
LSP
DSP
F
LSP
DSP
QSP
QSP
FC
FC
LC
BSP
LC
FC
BSP
FC
Q
Q
B
Case 3
FIG. 8.19j
Four cases of conventional distillation control configurations.
© 2006 by Béla Lipták
PSP
PC
B
Case 4
Control and Optimization of Unit Operations
Process values
Normalized
values
Analog signals
1
2
3
4
Volts
50
mA dc
20
mA dc
10
20
30
40
4
8
12
16
1
2
3
4
3
6
9
12
mA dc
15 PSIG
0.2
0.4
0.6
0.8
1.0 bar
0
0.25
0.50
0.75
1.0
0
25%
50%
75%
100%
93
116
138
160
182°C
200
360°F
0
240
280
320
Temperature transmitter
0.85
1.7
2.93
0
225
5
100%
L′ = 100%
75
L′ = 75%
50
L′ = 50%
25
L′ = 25%
3.4 m3/h
450
775
Linear flow transmitter
1.7
2.4
2.95
900 GPH
0
450
636
779
Differential pressure flow transmitter
900 GPH
0
2.50
3.75
1.25
Chromatograph output
0
1.143 L′
5
Multiplier output, normalized
manipulated variable (M′)
1830
3/h
3.4 m
0%
50
100%
25
75
Multiplier input, normalized ratio (R′)
FIG. 8.19l
Multiplier output for the solution of Equation 8.19 (39).
%
5.0
FIG. 8.19k
Common analog signals and their relationship to process variables.
logic controller (PLC), or multivariable digital controllers is
vendor-specific.
The terms of the equations are sometimes scaled because
most analog instruments and some PAS systems act on normalized numbers (0–100%) rather than on actual process
values. With digital instrumentation and today’s process automation systems, those occurrences are rare. The calculations
become easier for those systems operating in engineering
units.
Analog, and many digital, transmitters also operate on
normalized values of the process variables. That is, the measurement signal will vary from 0 to 100% as the process
variable shifts from 0 to its maximum value. Figure 8.19k
illustrates the relationship among the various forms of analog
signals and some typical process measurements.
The actual value of a process measurement is found by
multiplying the analog signal by the calibrated full-scale
value (meter factor) of the process variable. In the examples
of Figure 8.19k, the temperature, represented by a 75% analog signal, is 320°F (160°C), the linear flow is 775 gph
3
(2.93 m /h), the output of the differential pressure transmitter
3
(flow squared) is 779 gph (2.95 m /h), and the composition
is 3.75%.
Example As an example, let us review a flow ratio system
in which the load stream, L, has the range of 0 to 1000 gpm
3
(0 to 3.79 m /h); the manipulated stream, M, has a range of
© 2006 by Béla Lipták
3
0 to 700 gpm (0 to 2.65 m /h); and the ratio range, R, is 0
to 0.8 (R = M/L).
700M′ = (1000L′)(0.80R′)
8.19(38)
Reducing to the lowest form,
M′ = 1.143(L′)(R′)
8.19(39)
The number 1.143 is the scaling factor. M′ is plotted as
a function of L′ and R′ in Figure 8.19l.
In applications such as the constant separation system,
exact scaling is not critical. Exact scaling is when scaling
constants must be used as calculated from instrument spans.
The alternative is flexible scaling, where exact ranges are not
needed but some arbitrary range is used to allow internal
calculations to remain within range.
The flexible scaling cannot be used (1) when compensation for feed composition is part of the model, (2) when
narrow spans must be used for reasons of stability, and
(3) when transmitter calibrations are inconsistent with material balance ratios. Exact scaling techniques must be used for
these cases.
MULTIPLE COMPONENT DISTILLATION
With binary mixtures, only two products are removed in the
distillation column. However, most separations involve multiple components. Even then, most distillations remove only
two liquid products. In other applications a vapor product is
removed, or multiple liquid products are drawn from the
tower. Sometimes only one product is withdrawn at a time.
8.19 Distillation: Basic Controls
1831
Columns with Sidedraw
Having a sidestream product in addition to the overhead and
bottom products adds a degree of freedom to a control system. The source of this extra degree of freedom can be seen
from the overall material balance equation:
F=D+C+B
8.19(40)
where C is the sidestream flow rate. Two of the product
streams can be manipulated for control purposes, and the
material balance can still be closed by the third product
stream.
The presence of this added degree of freedom makes the
careful analysis of the process even more essential to avoid
mismatching of the manipulated and controlled variables. As
in the case of the previously discussed columns, the development of a control system for sidedraw applications also
involves developing the process model and determining the
relationship among the several controlled and manipulated
variables.
In this case, for a constant feed rate and column pressure,
five degrees of freedom exist: three composition specifications
and two levels that can manipulate three product flows, and
two heat balances (V and L). Several possible combinations
of variables are available and should be explored.
The possible combinations of manipulated variables for
the column in which the bottom composition and the sidestream composition must be controlled are
Distillate and sidestream flows
Distillate and bottom flows
Distillate flow and heat input
Sidestream and bottom flows
Sidestream flow and heat input
Bottom flow and heat input
Similarly, the possible combinations of manipulated variables for the column in which the distillate composition and
the sidestream composition must be controlled are:
Distillate and sidestream flows
Distillate and bottom flows
Distillate flow and heat input
Sidestream and bottom flows
Sidestream flow and reflux
Bottom flow and reflux
LT
X
FY
FIC
V
Dynamics FY
FT
FT
FIC
L
Dynamics FY
F, z1, z2
D, y1, y2
FT
ARC
AT
C, c1, c2
Ratio
RIC controller
FT
LT
FT
FIC
LIC
FY
X
ARC
AT
B, x1, x2
FIG. 8.19m
Control of composition in two product streams with a sidedraw.
The symbols z1, y1, and c1 refer to the concentrations in
the feed, distillate, and sidestream of the component under
control in the sidestream. The concentrations of the key component in the bottom are respectively expressed by z2, x2, and
c2 for the feed, the bottoms, and the sidestream.
The resulting control system is shown in Figure 8.19m.
Note that in this configuration the ratio of heat input to feed
(and, therefore, boil-up to feed) is held constant. Separate
dynamic elements are used for the distillate loop and for the
heat input and sidestream loops.
Multiproduct Fractionators
The equations are
© 2006 by Béla Lipták
LIC
z −c 
D= F 1 1
 y1 − c1 
8.19(41)
 z − x2 
C = F 2

 c2 − x2 
8.19(42)
Multiproduct fractionators are most common in the refining
industry where multicomponent streams are separated into
many fractions. Examples of multiproduct fractionators are
crude towers, vacuum towers, and fluidized catalytic cracking
unit (FCCU) main fractionators.
Product quality controls are used to adjust local column
temperatures and sidedraw flow rates to control distillate
properties related to the product specifications. An example
is true boiling point (TBP) cut points. TBP cut points approximate the composition of a hydrocarbon mixture and are
numerically similar to the American Society for Testing and
1832
Control and Optimization of Unit Operations
SP
FT
FRC
D
PRC
PT
Accumulator
FRC
FT
D
L
Q
Q
FRC
FRC
FT
FT
FT
Main fractionation
Heat
balance
logic
FRC
Boiling
point
calculation
C
FT
FRC
C
F
B
FIG. 8.19n
Control of product flows and pump-around refluxes.
Materials’ (ASTM’s) 95%. The ASTM laboratory distillate
evaluation method is the standard used in the petroleum refining industry for determining the value (composition) of the
distillation products.
A computer is required to calculate the product boiling
point specification, such as 95% boiling point or TBP cut
point on the basis of local temperature, pressure, steam flow,
and reflux data. Local reflux is derived from internal liquid
and vapor flows, as discussed previously, and the remaining
variables are measured.
Boiling point analyzers can be used to provide the measurement signals. If there is no analyzer, the calculated boiling points can be used by themselves, or if there is one, they
can be used as a fast inner loop with analyzer trim. Because
of the volume of liquid/vapor loads within most multiproduct
fractionators, the manipulated variables that provide the
greatest sensitivity and the quickest response are generally
the product flows.
Adjustment of reflux flows, as shown in Figure 8.19n,
is an example of a heat balance control. The goal is to
8
maximize heat exchange to feed, subject to certain limits
(limits and constraints are discussed as part of the subject
of the optimization of distillation towers in Section 8.21).
The task of maximizing the heating of the feed often simplifies to recovering heat at the highest possible temperature, which means recovering it as low as possible in the
column.
© 2006 by Béla Lipták
Superfractionators
The term superfractionator is applied to towers that are physically large. These distillation units separate streams having
their light and heavy key relative volatilities quite close to
each other. Included in this classification are deisobutanizers,
which separate isobutane from normal butane; propylene
splitters, which separate propane from propylene; ethylbenzene towers, which separate ethylbenzene from xylene; and
xylene splitters, which separate para- and ortho-xylene from
meta-xylene.
Sometimes, the number of trays and subsequent height
make it necessary to physically divide these towers into two
or even three sections. Superfractionators have tremendous
internal vapor-liquid rates in order to achieve the separation.
Reflux-to-distillate ratios are very high, as are vapor-to-bottoms
ratios.
A large pressure drop through the tower also exists. Long
dead times and lag times are experienced before any response
is seen to feed rate or reflux changes. Generally, distillate
compositions of superfractionators have to be controlled with
material balance equations due to the lack of sensitivity of
response.
Batch Distillation
In batch distillation (see Figure 8.19o), an initial charge of
liquid is fed to a vessel, and the distillation process is initiated
8.19 Distillation: Basic Controls
TC
HIC
V + LqL
TT
FT
Yi
FY X & Σ mD + Yi
SP
FC
ARC
1833
Y
AT
D
FY
Distillate
D(y)
Reflux, L
FRC
D2
FT
Distillate
(D)
Receiver
Wi - Initial batch quantity
yi - Initial product concentration
v
FC
FT
Steam, Q
Wi( y2)
x
Figure 8.19p shows the control system that will accomplish this when the vapor rate from the batch column is maintained constant. The equation describing this operation is:
L − LqL
FIG. 8.19o
Batch distillation.
Y = mD + yi
by turning on the heating and cooling systems. During the
distillation process, the initial charge in the vessel continually
depletes while building up the overhead product in the distillate receiver.
Batch distillations are more common in smaller, multiproduct plants where the various products can only be manufactured at different times, and where a number of different
mixtures may be handled in the same equipment.
Equation 8.19(43) is the basic equation that describes this
operation:
W = Wi − Dt
8.19(43)
where
W = amount remaining in the bottoms
Wi = the initial charge
D = distillate rate
t = time period of operation
The basic objective of the control system of this type of
separation is to keep the composition of the distillate constant. Other goals include keeping the distillate flow constant
or maximizing the total distillate production. The main goal
of a batch distillation is to produce a product of specified
composition at minimum cost. This often means that operating time must be reduced to some minimum while product
purity or recovery is maintained within acceptable limits.
If product removal is too fast, separation and the quantity
of the product are reduced. Conversely, if the product is
withdrawn to maintain separation, its withdrawal rate is
reduced and operating time is increased. However, the set
point to a composition controller can be programmed so that
the average composition of the product will still be within
9
specifications while withdrawal rate is maximized.
© 2006 by Béla Lipták
FIG. 8.19p
Control system for batch distillation.
8.19(44)
where
y = the fraction of key component in the product
m = the rate of change of y with respect to the distillate (D)
yi = the initial concentration of the product
The only adjustment required is the correct setting of m.
The higher its value, the faster y will change and the smaller
will be the quantity of material recovered.
CONTROL OBJECTIVES AND STRATEGIES
Operating objectives include the composition specifications
for the top and bottom product streams. Other objectives can
include increasing throughput, enhancing column stability,
and operating against equipment constraints. Yet other considerations include what product composition is considered
most important to maintain during disturbances, what are
acceptable variations in product specifications, and what are
relative economic values of the product streams and cost of
10
energy used in the separation.
The column operating objectives are ultimately governed
by economic benefits that are measurable, significant, and
11
achievable. Economics of individual fractionators may continually change throughout the life of the plant. Prices and
costs may determine that energy savings are important at one
particular time but that recovery is more important at some
other time.
The economic benefits of fractionator control include
shifting of less profitable components into more profitable
products, energy conservation, and increased throughput.
Other benefits arise, including minimum disturbances propagated to downstream units, minimum rework or recycle of
off-spec products, and more consistent product quality. Thus,
a given column’s operating economics and, therefore, its
objectives may change with time.
1834
Control and Optimization of Unit Operations
When minimization of fractionator utilities is an objective, the following guidelines are recommended:
•
•
•
•
Implement control to achieve composition control on
all products of the fractionator
Operate the fractionator to produce minimum overseparation
Ascertain that the reduction in energy usage is reflected
in the energy inflow to the production complex
Minimize energy waste from blending of overseparated products
Alternative Control Strategies
Many choices confront the design engineer when selecting
the control variables for a column. The first decision involves
configuration of the top or bottom control loops, which
directly determines product compositions. Once these strategies are tentatively determined, the control strategies for the
remaining variables (e.g., pressure or levels) become easier
to select.
Pairings of controlled and manipulated variables are normally made according to the single-input single-output
(SISO) method. Multivariable control, where multiple-input
and multiple-output (MIMO) variables are paired, are discussed in Section 8.21. In these multivariable strategies,
although a controlled variable can be affected by several
manipulated variables, only one manipulated variable is used
to directly affect the controlled variable. The minimum number of controlled variables for a fractionator tower is four.
These include:
Controlled Variables
Manipulated Variables
Overhead composition
Reflux flow
Bottoms composition
Reboiler heating media flow
Accumulator level
Distillate flow
Bottoms level
Bottoms flow
This allows for 24 possible configurations (4 factorial).
Of course, most towers include pressure as a controlled variable, with condenser flow or vapor bypass as a manipulated
variable. Additional manipulated variables can include feed
flow and enthalpy. If a tower includes a sidedraw stream,
another control pair is added to the possible combinations.
In fact, additional control variables increase the number
of possible control configurations factorially (e.g., six variables produce 720 possible configurations).
The pairing of controlled and manipulated variables can
follow three general control structures: energy balance con12
trol, material balance control, and ratio control. Energy
balance control uses reflux and reboiler heating media flow
to control compositions, thus fixing the energy inputs.
Material balance control uses the distillate and bottoms
product flows to control compositions, thus fixing the overall
© 2006 by Béla Lipták
material balance. Ratio control utilizes a ratio of any two flow
rates at each end of the column. The two common examples
of ratio control are the control of reflux-to-distillate ratio and
the boil-up-to-bottoms ratio. These control configurations perform quite differently depending upon the fractionator characteristics.
CONTROL LOOP INTERACTION
The selection of which product composition to control (or
both, if control of both can be controlled) and the decision
on which variables will give better control can be aided by
calculation of a relative gain array. The concept of relative
13,14
provides a measure of the interaction that can be
gain
expected between control loops. This subject is covered in
more detail in Chapter 2 in Section 2.12 and in Section 8.20.
The concept may be used to find the control configurations
that will have the least amount of interaction. Therefore,
relative gain analysis should be considered the first step in
evaluating alternative composition control strategies.
In addition, some pairings can be made heuristically from
operating experience and on the basis of a general understanding of column dynamics (Table 8.19q).
The following are general rules used to reject some pos15,16
sible control pairings:
1. Overhead composition and bottoms composition should
not both be controlled with material balance equations
if the objective is to control product specifications at
both ends of the fractionator.
Because of lack of dynamic response the following loops
should not be paired:
1. Accumulator level should not be controlled with
reboiler heat if the reboiler is a furnace.
2. Bottoms level should not be controlled with reboiler
heat if the reboiler is a furnace.
3. Bottoms level should not be controlled with distillate
flow.
4. Accumulator level should not be controlled with bottoms product flow.
5. Overhead composition should not be controlled with
bottoms product flow.
6. Bottoms composition should not be controlled with
distillate flow.
7. Bottoms level should not be controlled with reflux
flow.
8. Bottoms composition should not be controlled with
reflux flow if the number of trays is greater than a
minimum limit (approximately 20).
9. Bottoms level should not be controlled with reboiler
heat if the diameter of the column is greater than a
minimum limit (approximately 15–20 ft (4.5–6 m),
indicating a high volume of liquid in the bottoms).
8.19 Distillation: Basic Controls
1835
TABLE 8.19q
4
Dynamic Response and Sensitivity Limitations on the Pairing of Distillation Control Variables
(Both compositions should not be controlled by material balance (B,D) if both specifications are important)
Manipulated
Variable
Controlled
Variable
Composition of Overhead Product (ACy)
Distillate Flow
(D)
OK if L/D 6
Note 3
Composition of Bottoms Product (ACx)
Accumulator Level (LCa)
Note 3
OK if L/D 6
Bottoms Level (LCb)
Notes: 1.
2.
3.
4.
Bottoms Product Flow
(B)
OK if V/B 3
Vaporization Rate (V) or
Heat Input at Reboiler (O)
Reflux Flow Rate
(L)
Notes 1 and 2
Note 2
Notes 1 and 2
OK if trays 20
Not good with furnace
OK if V/B 3
OK if L/D 0.5
Not good if furnace is used
OK if diameter at bottom
20 ft
Control that concentration (x or y) which has the shorter residence time by throttling vapor flow (v).
More pure product should control separation (energy).
Less pure product should control material balance.
When controlling both x and y, the only choices for possible pairings are:
a. Control y by D and x by V.
b. Control y by D and x by L.
c. Control y by L and x by V.
d. Control y by B and x by L.
Of these, choice d is not recommended because a y/B combination is not responsive dynamically.
10. Accumulator level should not be controlled with
reboiler heat if the control objective is to maintain
overhead product specification and the V/B ratio is
less than a minimum limit (approximately 3).
Because of lack of sensitivity, these loops should not be
paired:
1. Overhead composition should not be controlled with
reflux flow if the reflux ratio (L/D) is less than a minimum value (approximately 6).
2. Accumulator level should not be controlled with distillate flow if the reflux ratio (L/D) is less than a maximum value (approximately 6).
3. Accumulator level should not be controlled with reflux
flow if the reflux ratio (L/D) is less than a maximum
value (approximately 0.5).
4. Bottoms composition should not be controlled with
sidedraw flow if the sidedraw is a vapor phase.
5. Overhead composition should not be controlled with
sidedraw flow if the sidedraw is a liquid phase.
6. Bottoms composition should not be controlled with
sidedraw flow if the sidedraw is a liquid phase and the
sidedraw tray number is greater than a minimum number (approximately 20).
7. Sidedraw composition should not be controlled with
reflux or distillate flow if the difference between the total
number of trays and the number of the sidestream tray
is greater than a minimum value (approximately 20).
8. Bottoms level should not be controlled with sidedraw
flow if the difference between the bottoms and the
© 2006 by Béla Lipták
number of the sidestream tray is greater than a minimum value (approximately 100).
9. Bottoms level should not be controlled with bottoms
flow if the V/B ratio is greater than a minimum limit
(approximately 3).
Choices for controlling product compositions include
(1) controlling top or bottom composition only (generally
suitable for constant separation conditions, where specifications for one product are loose or where effective feedforward/
feedback systems can be designed to compensate for load
changes) and (2) controlling of both product compositions
(minimizes energy use and provides tight specification top and
bottom products for columns in which the problems of interaction are small).
These choices can be broken down further into considerations such as manipulation of distillate-boil-up, DV configuration (generally suitable for high reflux columns) or
manipulation of reflux-boil-up, LV configuration (generally
suitable for low reflux columns), and so forth.
Further considerations include the use of decoupling control schemes (can present practical problems, such as insensitive control, operating problems, and high sensitivity to errors)
and the use of temperature measurements to infer composition
or analyzers to measure composition directly (generally an
economic decision based on how well a temperature-sensitive
control point can be determined and the costs of analyzer
hardware and maintenance). These choices are based on operating objectives of the column, expected disturbance variables,
and the degree of control loop interaction.
1836
Control and Optimization of Unit Operations
PRODUCT QUALITY CONTROL
Water
Conceptually, product control is a problem of making precise
adjustments to the rate of heat addition and the rate of heat
removal from the tower. Heat removal determines the internal
reflux flow rate, and the internal reflux as measured on the
top tray is a direct reflection of the composition of the distillate. Heat added determines the internal vapor rate. These
internal vapor and liquid flow rates determine the circulation
rate, which in turn determines the degree of separation
between two key components.
Once interaction of the various variable pairings has been
established, and the column’s operating objectives and disturbance variables are considered, the primary composition
control loops of the column can be selected. Measurement
of these control variables can be either direct or inferred.
PRC
PT
TRC
Set
TT
LT
FRC
LRC
Set
FT
FRC
FT
FIG. 8.19s
If overhead composition is to be controlled, the reflux flow to the
column is throttled by a temperature controller.
Inferring Composition from Temperature
If the cost of on-line analyzer hardware and maintenance is
prohibitive, or if backup is desired in case of analyzer failure
or maintenance, and because the results of laboratory analysis
take too long to be usable for effective control, temperature
measurement often can be used to infer composition.
Because distillation separates materials according to their
difference in vapor pressures, and because vapor pressure is a
temperature-controlled function, temperature measurement has
historically been used to indicate composition. This presumes
that the column pressure remains constant, or that the temperature measurement is compensated for pressure changes, and
that feed composition is constant. Then, any change in composition within a column will be detected as a temperature change.
The best point to locate the temperature sensor cannot be
established from generalizations. The important consideration
is to measure the temperature on a tray that strongly reflects
the changes in composition. When composition of the bottom
product is important, it is desirable to maintain a constant temperature in the lower section. This can be done by letting the
temperature measurement manipulate the reboiler steam supply
by resetting the steam flow controller set point (Figure 8.19r).
When composition of the distillate is more important, it
is desirable to maintain a constant temperature in the upper
section, as in Figure 8.19s. In this configuration the sensing
point for column pressure control should be located near the
temperature control point. Keeping the sensor locations close to
each other helps to fix the relation between temperature and
composition at this particular point.
If column temperature profiles caused by small positive
and negative changes in manipulated variables, such as a ±1%
change in distillate flow (Figure 8.19t), can be generated, the
Stage
Sensor location for
maximum sensitivity
18
16
14
1% decrease in D
1% increase in D
12
10
TRC
TT
8
Set
FRC
6
FT
4
Steam
LT
LIC
2
−0.3
FIG. 8.19r
In this configuration the reboiler heat input is throttled by a
temperature controller to keep the bottoms product composition
constant.
© 2006 by Béla Lipták
−0.2 −0.1
0
0.1
0.2
0.3
Stage temperature change in °C
FIG. 8.19t
Example of column temperature profiles resulting from a 1%
increase and from a 1% decrease in distillate flow.
8.19 Distillation: Basic Controls
14
TT
6
TT
TDRC
FRC
FT
SP
1
1837
FRC
FT
Steam
AT
ARC
FIG. 8.19u
Heat input controlled by temperature difference.
FIG. 8.119v
Distillate withdrawal controlled by chromatograph.
following criteria may be helpful in selecting sensor loca17
tions: (1) The sensitivity of the temperature-manipulated
variable pairing should be in the range of 0.1 to 0.5°C/% and
(2) equal temperature changes should result when increasing
and when decreasing the manipulated variable.
For a two-product fractionator, distillation temperature
is an indication of composition only when column pressure
remains constant or if the temperature measurement is pressurecompensated. When separation by distillation is sought between
two compounds having relatively close vapor pressures, temperature measurement, as an indication of composition, is
not satisfactory.
Fixing two temperatures in a column is equivalent to
fixing one temperature and the pressure. Thus, by controlling
two temperatures, or a temperature difference, the effect of
pressure variations can be eliminated. The assumption used
here is that the vapor pressure curves for the two components
have constant slopes.
Controlling two temperatures is not equivalent to controlling a temperature difference. A plot of temperature difference vs. bottom product composition exhibits a maximum.
Thus, for some temperature differences below the maximum
it is possible to get two different product compositions.
Separation of normal butane and isobutane (in the
absence of other components, such as pentanes and heavier
substances) can be accomplished very well by using temperature difference control. Figure 8.19u illustrates how the heat
input to such a column can be controlled by a temperature
difference controller.
used. (For details, refer to Chapter 8 of Volume 1 of this
handbook.)
Once, the time required for a chromatographic analysis
(several minutes) was a great barrier to its use for automatic
control. Since then, the equipment has been enhanced so that
analyses can now be made in less than 5 min, and in many
cases for low-volatility hydrocarbons, the analysis can be
made continuous.
With careful handling, the under 5 min sampling rate will
permit closed-loop distillate control. In fact, fractionators are
successfully controlled with cycle times as long as 7–10 min
by applying dead time compensation algorithms.
Light ends fractionators have been satisfactorily controlled by the use of chromatography. Figure 8.19v illustrates
the controls of a superfractionator designed to separate isobutane and normal butane. In this case, the chromatograph
continuously analyzes a sample from one of the intermediate
trays, and this measurement is used by the analyzer controller
to modulate the product draw-off valve.
Overhead and bottoms analyzers typically measure the
loss of a valuable product or the presence of impurities.
Impurity components are chosen because small concentration
variations can be measured more precisely and with better
repeatability, and can provide a more sensitive measure of
separation. For example, the change of an impurity from 1.0
to 1.1% can be measured with greater precision than a change
of the major component from 99 to 98.9%.
When composition analyzers are used in feedback control, several configurations can be considered. These include
1) direct control of a manipulated variable, 2) cascade control
adjusting the set point of a slave temperature controller, and
3) analysis control in parallel with temperature control in a
selective control configuration. The configuration used
depends on the control objective, sensitivity of control, and
analysis dead time.
Control by Analyzers
Analytical or composition control is a way to sidestep the
problems of temperature control. Although additional investment is needed for the analytical equipment, savings from
improved operation usually results. Several types of instruments are available for composition analysis. Of these, the
gas chromatograph is the most versatile and most widely
© 2006 by Béla Lipták
Direct Control by Analyzers Analyzer controllers in a feedback configuration can be considered when the dead time of
each analysis update is less than the response time of the
1838
Control and Optimization of Unit Operations
SP
FT
SP
FRC
FT
Σ
ARC
ARC
PRC
AT
PRC
X
TT
SP
FRC
FT
Accumulator
Fractionation
FIG. 8.19w
Direct control of overhead product composition by an analyzer
controller (ARC) throttling the set point of a reflux flow controller
(FRC).
process. Because it is the control of the composition of the
product, which is often the objective, direct control by an
analyzer controller would seem to be better than indirect
control by temperature.
The composition controller provides feedback correction
in response to feed composition changes, pressure variations,
and variations in tower efficiencies. Figure 8.19w shows the
configuration of a control system, in which a chromatograph
analyzes a liquid sample from the condenser rundown line.
A sample probe gathers the liquid sample and the
sampling system conditions and vaporizes the liquid sample
to provide a representative vapor sample to the chromatograph. The analyzer controller (ARC) uses the chromatographic measurement to manipulate the reflux flow by adjusting the set point to the reflux flow controller (FRC).
Smith Predictor Often the analyzer is so slow that it introduces a significant delay time that degrades the controllability
of the process. In that case, some type of dead time compensation is used (see Section 2.19 in Chapter 2). A Smith
predictor compensator can serve to model the process to
predict what the analyzer measurement should be between
analysis updates. When the actual measurement is completed,
the model’s prediction is compared to the actual measurement
and the input to the controller is biased by the difference.
SP
FRC
AY
FT
+
AY
+
−
AY
AY
Lag
Dead
time
AT
Accumulator
D
L
F
B
© 2006 by Béla Lipták
TT
D
L
F
PT
+∆
Fractionation
PT
AY
FRC
B
FIG. 8.19x
Analyzer controller with dead-time compensation cascaded to reflux
flow control.
Figure 8.19x shows the same configuration as did
Figure 8.19w except that the analyzer controller is equipped
with a first-order Smith predictor that provides dead time
compensation.
In Figure 8.19x, the multiplier, lag, and dead time calculations (AY) provide the predicted analysis. (The lag represents
the first-order process.) This predicted response is subtracted
from the actual measurement to give a differential of the
actual process from its own model. This delta is added to the
model without dead time to provide a modified pseudomeasurement to the analyzer controller. Thus, the analyzer measurement, which has a significant dead time due to sampling
and cycle times, provides a trim to the predicted measurement
of the model.
Triple Cascade and Selective Control Analyzer control cascaded to temperature control can be used when stable temperature on a particular tray is desired and the tower operates
at a constant, maintainable, and controllable pressure. An
example is cascading the analyzer controller to the overhead
temperature of a tower, which in turn is cascaded to the reflux
flow rate. Because temperature is an indicator of composition
at this pressure, the analyzer controller only serves as a trim
correcting for variations in feed composition. Figure 8.19y
shows this triple cascade configuration of an analyzer controller setting the temperature controller setting the reflux
flow controller.
8.19 Distillation: Basic Controls
1839
SP
FT
FRC
PRC
PT
PRC
PT
TT
ARC
SP
FRC
FT
AT
Accumulator
D
L
FY
<
SP
FRC
TRC
Fractionation
F
Absorber stripper
SP
TRC
FT
Reboiler
ARC
Lean oil to
absorber
B
FIG. 8.19y
Triple cascade configuration of overhead composition control.
Analyzer controls can be used in a high or low select
configuration in combination with temperature when a high
or low limit based on temperature is important. The temperature controller is a constraint controller (see Section 2.28 in
Chapter 2 for details) serving to prevent the temperature from
exceeding a limit.
An example of using this control configuration is the
control of the bottoms of an absorber stripper. Here, the
temperature should not exceed a certain value, as no additional stripping of the light component in the bottoms of the
column could be accomplished. Even though an analyzer
controller may call for more heat, this heat would only
increase the bottoms temperature of the recycled oil to the
absorber without removing the impurity, thereby reducing
the absorption capability at the absorber.
Figure 8.19z depicts an analyzer controller in a low select
configuration with the temperature constraint controller.
Note that both cascade and selective control configurations require external feedback to protect them from reset
windup. Figure 8.19aa illustrates how the external feedback
(EF) is applied to the master controller in a cascade configuration (TIC) and to both controllers (FIC and PIC) in a
selective control configuration. For more details on external
feedback, refer to Section 2.28 in Chapter 2.
Fractionator Trains A controller may use as its measurement the analysis of a single component, or may use the ratio
of two components. A ratio (e.g., ethane-to-propane, C2/C3)
© 2006 by Béla Lipták
FIG. 8.19z
Analyzer control in a low select configuration with a temperature
constraint controller.
is often used when the fractionator is not the final step in the
18
separation sequence. This often occurs in a natural gas
liquids separation train where a de-ethanizer, a depropanizer,
a debutanizer, and a deisobutanizer (butane splitter) produce
the products ethane, propane, butane, isobutane, and gasoline
as shown in Figure 8.19bb.
Specifications for the primary overhead products may
include limitations on the amount of both light and heavy
impurities. For example, the propane product from the overhead of the depropanizer would have limitations on ethane
as well as isobutane. The problem is that the light impurity
(lighter than light key) cannot be controlled in the tower that
produces that product. Rather, it must be controlled in an
upstream tower.
Set
point
TIC
EF
FIC
Master
Override
<
FY
PIC
PT
Slave
TT
FT
EF
EF
A/O
FIG. 8.19aa
In a cascade configuration the external feedback signal (EF) is the
slave measurement, while in selective control configurations, it is
the signal that is throttling the control valve.
1840
Control and Optimization of Unit Operations
Isobutane
product
AX C3/IC4
AX C5+
Deisobutanizer
Propane
product
Debutanizer
AX C2/C3
AX NC4
AX IC4
Depropanizer
NGL
Deethanizer
AX C3
Ethane
product
N-butane
product
AX IC4
Straight run
gasoline
AX NC4
FIG. 8.19bb
Analyzer placement in a fractionator train.
The lighter than light key specification in the distillate
of the downstream tower can be controlled more easily by
controlling the ratio in the bottoms of the upstream tower.
That is, the ethane content in the propane product (depropanizer distillate) is maintained by controlling the C2/C3 ratio
in the bottoms of the de-ethanizer. Measuring the C2/C3 ratio
in the bottoms requires an additional analyzer but eliminates
the dead time of obtaining the concentration in the overhead
of a downstream tower.
A feed analyzer is sometimes included as a part of feedforward control. The feed analysis is used in predicting internal reflux/overhead flow and bottoms/heat input. However,
when feed composition changes slowly or when results from
the analyzer cannot be obtained faster than the dynamics of
the tower, this analyzer is omitted and the burden is placed
on feedback control from the product analyzers.
In practice, a feed analyzer is the exception rather than
the rule. Its use is mainly when the analyzer is already in
place, because it is controlling an upstream tower. For example, the NGL separation train in Figure 8.19bb has a deethanizer bottoms analyzer that could also be considered the
depropanizer feed analyzer.
Analyzer Selection The choice of analyzer control depends
upon the analytical equipment available and on the type of
separation desired. Each type of separation requires a compromise between the controllability and the delay of the
control system. For example, the NGL train (Figure 8.19bb)
was studied to determine the best analyzer system. In the
© 2006 by Béla Lipták
depropanizer (where isobutane was to be measured in the
presence of ethane, propane, and normal butane) and in
the deisobutanizer (where isobutane was to be measured in the
presence of normal butane and isopentane), an infrared analysis
was to be preferred.
However, in the debutanizer the goal was to measure the
combined isopentane plus normal pentane concentrations in
the presence of isobutane and normal butane to control the
butane-pentane separation. Here, investigation revealed that
gas chromatography provides the best solution.
Some boiling point analyzers are reliable enough to be
used for on-line control (see Section 8.50 in Chapter 8 in
Volume 1 of this handbook). Normally, cut points between
overhead products and side-cuts are maintained by temperature controllers. These controllers generally influence reflux
rate or product draws to achieve the desired results. Laboratory distillation results are used to adjust the set points to the
temperature controllers. This method of control, however, is
cyclical because of the time lags involved in temperature
control.
To avoid exceeding the target cut points and to meet
required product specifications, the cut point is set below
specification. This results in downgrading the more valuable
product to the stream of lesser value. This downgrading can
be minimized through the use of on-line boiling point analyzers. Justification of a boiling point analyzer depends upon
the value of the products, how much downgrading is occurring, and the cost of analyzer maintenance. Figure 8.19cc
illustrates an end-point analyzer.
Viscosity is another property that can be measured continuously to give faster control corrections. In vacuum distillation, the viscosimeter monitors each of the streams for
which viscosity is a specification. Any deviation from the
desired viscosity is corrected by a change in the set point of
the control loop involved.
Deviation from the desired viscosity and the subsequent
downgrading of the product can occur because of frequent
variations in tower operating conditions and feed composition. In addition to normal operation, the use of a viscosity
analyzer minimizes downgrading during major upsets and
large feed compositions changes. With such an arrangement,
low viscosity vacuum bottoms can be detected quickly and
diverted to recoverable feed for profitable reprocessing.
Once again, profitability determination requires a thorough analysis of column operation and an assessment of the
engineering, operating, and maintenance capabilities at the
location where this type of control is to be implemented.
Many other analytical instruments are being moved out
of the laboratory and into the processing area. Mobile units
containing several different kinds of analyzers can be used
to learn the best place to locate on-stream analyzers. In cases
in which permanent analyzers cannot be justified, the mobile
unit is connected to the process long enough to find the best
operating conditions. Then, the mobile unit can be moved
elsewhere.
8.19 Distillation: Basic Controls
Conditioned
sample
Ok
(fast but
not repeatable)
PSV
PC
Overhead
condenser
Percent
evaporated
temperature
controller
TT
Ok
(slow)
NG
Packed
column
PI
1841
Distillate
product
NG (D)
TRC
Feed
(F)
To
distillation
column
controls
Ok
(fast)
FO
TC
Boiler pot
Sample
outlet
valve
Float
Drain
valve
Analyzer
effluent
FIG. 8.19cc
The end-point distillation analyzer is a miniature version of the
process column.
Sampling Proper sampling of material in a column is necessary if analyzers are to control effectively. A poor sampling
system often is responsible for the unsatisfactory performance of plant analyzers. For details on sampling system
design, refer to Chapter 8, Section 8.2, in Volume 1 of this
handbook.
The sampling points for composition analysis should be
at, or very near, the column terminals for the following reasons: (1) freedom from ambiguity in the correlation of sample
composition with terminal composition, and (2) improved
control loop behavior as a result of reduction of transport lag
(dead time) and of the time constants (lags) describing the
sampling point’s compositional behavior. This assumes that
the controller applies its manipulation at the same terminal
(steam or reflux) where the controlled variable is measured.
The factors favoring moving the point of sampling nearer
to the feed entry point are (1) improved terminal composition
behavior as a result of earlier recognition of composition
transients as they proceed from the feed entry toward the
column terminals, and (2) less stringent analytical requirements as a result of (a) analyzing the control component at
a higher concentration and over a wider range, and (b) simplifying the multicomponent mixture, because nonkey components tend to exhibit constant composition zones in the
column.
© 2006 by Béla Lipták
Bottoms
product
(B)
Heater
Ok
(slow & low
pressure)
Ok
(slow)
FIG. 8.19dd
The choices of sampling point locations for analyzers used in distillation column control.
Figure 8.19dd shows some of the typical sample locations
of a distillation tower. Most analyzers are designed to accept
a clean, dry, noncorrosive sample at low temperatures, pressures, and flow rates. Such conditions seldom exist in the
process, so the sampling system must be designed and operated to overcome the difference between the conditions in
the process and the conditions required by the analyzer.
The sample system must provide a current and representative sample of the stream being analyzed. It must transport
the sample from the sample point to the analyzer with a
minimum of transport lag (preferably less than 30 sec and
definitely not greater than 1 min). Transportation times are
minimized using high flow rate bypass streams taken from
the process sample point and returned to the process at a
lower pressure. The sample system must condition the sample
to remove traces of foreign materials through filtering, maintain pressure and temperature, and maintain or change phase
for introduction into the analyzer.
For chromatographs, liquid sample points are generally
preferred (Figure 8.19ee). This is because vapor streams have
historically not provided representative samples. Vapor samples do not tend to produce repeatable values as consistently
19
and reliably because of condensation at the sample probe
and in the sample lines when hydrocarbons of high boiling
1842
Control and Optimization of Unit Operations
Analyzer
Cooling
in
Sample
Out
In
Water
out
Control
Self
valve
cleaning
filter
FI
Sample
cooler
Shut-off
valves
Flow
FI indicator
Check
valve
To
drain
Pressure
gauge
Temperature
gauge
PI
T1
Flow
indicator
FI
with
needle
valve
Flow
indicator
Pressure
regulator
Calibration
sample
Shut-off
valves
Lab sample
take off
Coalescer
Flow
indicator
FI
with
needle
Check
valve
valve
Pressure
relief valve
Pressure
gauge
PI
Shut-off
valve
FIG. 8.19ee
Sampling system for a liquid product in a refinery application.
points are present in the sample. When the sample lines are
long, some separation between components can also occur.
However, vapor samples can be used when warranted and if
the proper care can be taken.
A satisfactory point for measuring bottoms product composition is at the point of highest pressure. This approach
will ensure a representative sample and will provide the pressure drop to return the sample bypass. The point of highest
pressure is generally immediately after the product pump.
However, if liquid holdup in the reboiler and kettle is large, a
long lag is introduced, which slows the transient response of
the measurement and control system. Alternative sample
points such as a bottoms tray or seal pan may be used, but
may require extra expense for the sample system.
A satisfactory sample-point location for measuring the
distillate is the outlet liquid of the overhead vapor condenser.
Sampling the overhead accumulator liquid after the reflux or
distillate pump should be avoided because of the tremendous
process lag it introduces. Sampling the overhead vapor
reduces the process lag of sampling after the condenser if a
repeatable, representative sample can be obtained.
PRESSURE CONTROL
Most distillation columns are operated under constant pressure. However, floating-pressure operation can have advantages in many processes. One reason for the resistance to the
use of floating-pressure control is based on the fact that
temperature is sensitive to pressure changes, and therefore,
© 2006 by Béla Lipták
it requires pressure compensation if the pressure varies. As
analyzers are increasingly replacing temperature-based controls, the argument favoring constant pressure operation is
also lessening. However, even when temperature control is
used, the temperature measurements can be compensated for
pressure variations.
The primary advantage of floating-pressure control is the
ability to operate at the minimum column pressure within the
constraints of the system. Lower pressure reduces the volatility of distillation components, thereby reducing the heat
input required to effect a given separation. Other advantages
include increased reboiler capacity and reduced reboiler fouling
due to lower tower temperatures.
In the following paragraphs, floating-pressure control strategies will be described for the following conditions: (1) liquid
distillate withdrawn when noncondensables are present,
(2) vapor distillate withdrawn when noncondensables are
present, and (3) liquid distillate withdrawn when the amount
of noncondensables is negligible.
Liquid Distillate and Inerts
In some separation processes the problem of pressure control
is complicated by the presence of large percentages of inert
gases. The noncondensables must be removed, or they will
accumulate and blanket off the condensing surface, thereby
causing loss of column pressure control.
The simplest method of handling this problem is to bleed
off a fixed amount of gases and vapors to a lower pressure unit,
such as to an absorption tower, if one is present in the system.
8.19 Distillation: Basic Controls
Vent
used to start opening the purge control valve (VPCV-2), when
the opening of PCV-1 reaches some preset limit. This can be
done by means of a calibrated valve positioner or by using
a valve position controller (VPC-2) in Figure 8.19ff.
Partially
flooded
condenser
VPCV
2
VPC
2
PCV
1
Vapor Distillate and Inerts
PRC
1
PT
1843
LT
LIC
FIG. 8.19ff
Column pressure control with inerts present.
If an absorber is not present, it is possible to install a vent
condenser to recover the condensable vapors from this purge
stream. Often in refinery applications such noncondensables go
to the fuel gas system or to flare.
It is recommended that the fixed continuous purge be
used whenever economically possible; however, when this is
not permitted, it is possible to modulate the purge stream.
This might be desirable when the amount of inerts is subject
to wide variations over time.
As the noncondensables build up in the condenser, the
pressure controller will tend to open the control valve (PCV-1
in Figure 8.19ff) to maintain the proper rate of condensation.
The controller signal that is throttling PCV-1 could also be
In the case where the distillate is in the vapor phase and inerts
are present, the overhead product is removed under pressure
control as shown in Part A of Figure 8.19gg. In this configuration the system pressure will quickly respond to changes
in the distillate vapor flow. In this control system a level
controller is installed on the overhead receiver to regulate the
cooling water to the condenser, so that it will condense only
enough condensate to provide the column with reflux.
This control system will operate properly only if the
condenser is designed to provide a short residence time for
the coolant, which will minimize the level control time lag.
If this is not the case, the cooling water flow should be
maintained at a constant rate.
In this case (Part B in Figure 8.19gg), the level controller
can regulate the flow of condensate through a small vaporizer
and mix it with the vapor from the pressure control valve.
If the cooling water has fouling tendencies, it is preferable to use the control system shown in Figure 8.19hh, where
a pressure controller regulates a vapor bypass around the
condenser.
Liquid Distillate with Negligible Inerts
In distillation processes where the distillate is in the liquid
phase and the amount of inerts is negligible, the column
pressure is usually controlled by modulating the rate of condensation in the condenser. The method of controlling the
Dry
condenser
Dry
condenser
Water
Vaporizer
Water
PRC
PRC
PT
PT
Vapor
(D)
Vapor (D)
LRC
LT
LT
Receiver
A. Variable water flow
LRC
B. Constant water flow
FIG. 8.19gg
Column pressure control when the distillate is in the vapor phase and contains inerts for variable (A) and constant condenser water flow
configurations (B).
© 2006 by Béla Lipták
1844
Control and Optimization of Unit Operations
Partially
flooded
condenser
∆P
Water
PRC
∆P
PT
PRC
Vapor
(D)
∆P
PT
LT
∆P
LIC
LT
LRC
Liquid
(D)
B. Liquid distillate
A. Vapor distillate
FIG. 8.19hh
Column pressure controlled by hot gas bypass throttling in case of vapor (A) and liquid (B) distillate processes.
rate of condensation depends upon the mechanical construction of the condensing equipment.
Controlling the Cooling Water Flow Figure 8.19ii describes
a control configuration where the column pressure is controlled by throttling the cooling water flow from the condenser. This method of control is recommended only when
the cooling water is treated with chemicals that prevent the
fouling of the tubes in the event of high temperature rise
across the condenser tubes. In such configuration, the maintenance costs are low, because the control valve is on the
water side and the control performance is acceptable, provided the condenser is properly designed.
The best condenser for this service is a bundle-type unit
with the cooling water flowing through the tubes. This water
Water
PT
Dry
condenser
PRC
LT
LIC
FIG. 8.19ii
Column pressure control by throttling condenser water.
© 2006 by Béla Lipták
should be flowing at a rate of more than 4.5 ft/s (1.35 m/s),
and the water should have a residence time of less than 45 sec.
The shorter the residence time of the water, the better will
be the quality of control obtained, owing to the decrease in
dead time or lag in the system.
With a properly designed condenser, the pressure controller needs only proportional control, because a narrow
throttling range is sufficient. However, as the residence time
of the water increases, the time lag of the system will
increase, and consequently the controller will require a wider
throttling range and will need automatic reset to compensate
for the load changes.
The control obtained by using a wide proportional band
is not satisfactory for precision distillation columns because
of the length of time required for the system to recover from
an upset. Also, the dead time varies with load, and therefore
the integral setting of the PRC should be set to match the
variation in residence time.
Therefore, it is unacceptable to use this control system
on a condenser box with submerged tube sections, because
there would be a large time lag in the system due to the large
volume of water in the box. It would require the passage of
a significant amount of time before a change in water flow
rate would change the temperature of the water in the box
and finally would affect the rate of condensation.
Controlling the Condensate Flow To reduce such unfavorable time lags, it becomes necessary to use a different type
of control system, one that permits the water flow rate to
remain constant and controls the amount of surface exposed
to the condensing vapors. This is done by modulating the
flow of condensate from the condenser.
When the column pressure is dropping, this condensate
throttling valve reduces the condensate flow, causing it to
8.19 Distillation: Basic Controls
Partially
flooded
condenser
PCV
2
PRC
1
PRC-1>PRC-2
∆P
PRC
2
PCV
1
PRC
PT
PT
1845
PT
∆P
LT
PRC
Inerts
PT
LIC
LT
LRC
Liquid
(D)
Liquid (D)
B. With inerts
A. No inerts
FIG. 8.19jj
High-speed column pressure control.
build up and flood more tube surface and, consequently, to
reduce the condensing surface exposed to the vapors.
Thereby, the condensing rate is reduced, and the pressure in
the column rises. In such designs, a vent valve should be
installed to purge the noncondensables from the top of the
condenser if it is expected that noncondensables could build
up and blanket the condensing surface.
Control of Hot Vapor Bypass A third possible control configuration for applications with liquid distillate containing
negligible inerts is used when the condenser is located below
the receiver. This is frequently done to make the condenser
available for servicing and to save on steel work. It is usual
practice to elevate the bottom of the accumulator 10–15 ft
(3–4.5 m) above the suction of the pump in order to provide
a positive suction head on the pump.
In this type of installation the control valve is placed in
a bypass of the vapor line to the accumulator (see
Figure 8.19hh, Part B). When this valve is open, it equalizes
the pressure between the vapor line and the receiver, causing
the condensing surface to become flooded with condensate
because of the 10–15 ft of head in the condensate line from
the condenser back to the receiver.
The flooding of the condensing surface causes the pressure to build up because of the decrease in the rate of condensation. Under normal operating conditions, the subcooling
that the condensate receives in the condenser is sufficient to
reduce the vapor pressure in the receiver. The difference in
pressure permits the condensate to flow up the 10–15 ft of
pipe between the condenser and the accumulator.
A modification of this latter system controls the pressure
in the accumulator by throttling the condenser bypass flow
© 2006 by Béla Lipták
(Figure 8.19jj, Part A). The column pressure is maintained
by throttling the flow of vapor through the condenser. Controlling the rate of flow through the condenser provides faster
pressure regulation for the column.
Part A of Figure 8.19jj shows the operation if there are
no inerts, as follows: If column pressure rises, PRC-1 opens
PCV-1. This increases the vapor pressure in the condenser,
which pushes some of the condensate out of it and increases
the condensing surface area exposed to the vapors. Therefore,
the rate of condensation is increased, and thereby the column
pressure is lowered back to the set point of PRC-1.
At this higher rate of condensation, the pressure drop
(∆P) across PCV-2 is also reduced (the valve opens). If the
column pressure drops, the opposite sequence occurs: PCV-1
closes and the flooding of the condenser increases, reducing
the rate of condensation and increasing the pressure drop
(∆P) across PCV-2 by slightly closing it. The setting of PRC-1
must always be above that of PRC-2.
The most common pressure control configuration is
shown in Part B of Figure 8.19jj. Here, the column pressure
controller is throttling the hot vapor bypass, as was the case in
Part A of Figure 8.19hh, but in addition a second pressure
controller is utilized on the accumulator. This PRC is set at
about 5 PSIG below the required tower pressure and is used
to vent the inert gases that may build up in the system.
Vacuum Systems
For some liquid mixtures, the temperature required to vaporize the feed would need to be so high that decomposition
would result. To avoid this, it is necessary to operate the
column at pressures below atmospheric. Steam jet ejectors
are often used to create vacuum in distillation systems. These
1846
Control and Optimization of Unit Operations
PRC
Cold/cold
exchanger
F
Steam
Air
C3
Propane
chiller
PIC
PT
Turboexpander
PT
PRC
Surge
bypass
SP
LIC
SC
Demethanizer
LT
FIG. 8.19kk
Vacuum column pressure control.
can be used singly or in stages, when a wide range of vacuum
conditions are required. The acceptance of steam jet ejectors
is due to their having no moving parts and requiring very
little maintenance.
Most ejectors are designed for a fixed capacity and work
best at one steam condition. Increasing the steam pressure
above the design point will not usually increase the capacity
of the ejector; instead, it will sometimes decrease the capacity
because of the choking effect of the excess steam in the
diffuser throat.
Steam pressure below a critical value for a jet will cause
the ejector operation to be unstable. Therefore, it is recommended that a pressure controller be installed on the steam
to keep it at the optimum pressure required by the ejector.
The recommended control system for pressure control in
vacuum distillation applications is shown in Figure 8.19kk.
Here, a controlled rate of air or gas is bled into the vacuum
line just ahead of the ejector. Closing this bleed valve makes
the maximum capacity of the ejector available to handle any
surges or upsets in the process load. A control valve regulates
the amount of bleed air used to maintain the pressure on the
reflux accumulator. Using the pressure of the accumulator
for control involves less time lag than if the column pressure
were used as the control variable.
Because ejectors are fixed capacity, the variable load is
met by air bleed into the system. At low loads this represents
a substantial waste of steam. Therefore, if substantial load
variations are expected, operating costs can be lowered by
installing a larger and a smaller ejector. This makes it possible
to automatically switch to the small unit when the load drops
off, thereby reducing the steam demand.
Vapor Recompression
Vapor recompression is another means of improving energy
efficiency of the operation. The overhead vapor from the
© 2006 by Béla Lipták
Turbine
Compressor
TT
TRC
Reboiler
Residue
gas
NGL
FIG. 8.19ll
Vapor recompression pressure control.
distillation column is compressed to a pressure where its
condensation temperature is greater than the boiling point is
at the pressure of the tower bottoms. The heat of condensation
of the overhead can then be used as the source of heat for
reboiling the bottoms.
This scheme is known as vapor recompression. It is used
fairly often when the distillation involves a relatively closeboiling mixture, and the boiling points of the top and bottom
products are similar. In cryogenic demethanization processes,
illustrated in Figure 8.18ll, the column pressure is controlled by the throttling of the speed of the recompression
compressors.
The heat of condensation of the overhead is also used
as the heat for the reboilers in propylene fractionators.
Figure 8.19mm shows the pressure controls needed for the
operation of this vapor recompression via the heat pump on
this particular tower.
FCCU main fractionators and crude towers make use of
compressors to “draw” vapors from the tower because operation is essentially at atmospheric pressures. The pressure
control system used in this case is shown in Figure 8.19nn.
In this configuration, the tower pressure is maintained by
controlling the speed of the compressor. This is accomplished
by the manipulation of the steam used to drive the compressor
turbine.
8.19 Distillation: Basic Controls
1847
Feed Flow Control
A flow controller in the feed line can maintain a constant flow
rate. In some instances, the feed pump of a distillation unit is
a steam-driven pump instead of an electrically driven one. In
this case, the controller modulates the steam to the driver.
Feed composition has a great influence upon the operation of a distillation unit. Unfortunately, feed composition is
seldom subject to adjustment. For this reason, it is necessary
to make changes elsewhere in the operation of the column
in order to compensate for the variations in feed composition.
The corrective steps are discussed later. The discussion below
assumes a constant feed composition.
PT
F
Heat pump
Propylene tower
PRC
CW
SP
TRC
FRC
Accumulator
Variable Column Feed
FT
Propylene
(C3−)
Reboiler
Propane
(C3)
FIG. 8.19mm
Propylene tower vapor recompression pressure control.
FEED CONTROLS
One of the best means of stabilizing the operation of almost
any continuous-flow-process, including distillation, is to hold
the flow rates and operating temperatures constant. Therefore, whenever possible, a flow controller should be used on
the feed to maintain a constant rate of flow.
SP
FT
Steam
PRC CW
PT
FRC
Accumulator
Having constant feed conditions simplifies the amount of
control required to achieve stable operation. However, the
distillate product is often fed to a second column. Then, any
changes that occur in the first column are reflected in the
quantity and composition of the feed to the second. If the
feed flow to the column is controlled by a liquid level controller of the previous column, that controller can be tuned
with a low gain, so that the level can swing over a wide range
without drastically upsetting the flow of product.
Nevertheless, the second column will receive a varying
flow of feed if it is linked to the first column. One way to
iron out temporary variations caused by liquid level changes
is to cascade the level to a flow controller in the product lines.
Flow controllers also serve to smooth out the pressure fluctuations caused by the distillate/reflux pump.
With variable feed rates and variable feed compositions,
cascade controls are justified. If the feed rate and composition
are relatively constant, resetting the major control loop manually
is sometimes adequate. In other cases the flow controller is
arranged as the cascade slave of the level controller
(Figure 8.19oo). The control algorithm for the level controller
in Figure 8.19oo is usually selected to be nonlinear to allow
the level to float in the surge tank without changing the FRC
set point, which would upset the feed to the next column.
Therefore, the nonlinear controller is so configured that
as long as the level in the surge tank is between 25 and 75%,
Main fractionator
Turbine
Compressor
LT
Nonlinear
LRC
Set
FRC
FT
Reboiler
FIG. 8.19nn
Main fractionator pressure control.
© 2006 by Béla Lipták
FIG. 8.19oo
Feed flow to the next column is kept relatively constant by the use
of a nonlinear level controller (LRC) on the surge tank acting as
the cascade master of the slave flow controller (FRC).
1848
Control and Optimization of Unit Operations
To column
TRC
Set
FRC
TT
FT
Steam
Distillate
FT
Feed
FY
X
Feed
FC
m
FIG. 8.19pp
Feedforward control minimizes feed rate disturbances.
the set point to the FRC remains constant. This will allow
the surge tank to fulfill its purpose and smooth out the load
variations between the related processes. If the level drops
below 25% or rises above 75%, the FRC set point is reduced
or increased respectively to protect it from draining or flooding the tank.
If feed rate disturbances must be accepted by the column,
a feedforward control system as shown in Figure 8.19pp can
10
be used to minimize the impact of these disturbances. The
ratio, m, is selected by performance of a simple material
balance around the column. Changing the product flow in
proportion to the feed flow minimizes internal column transients and, thus, the quantity of off-spec material during
recovery.
The value of m, however, is accurate only for one feed
composition and will have to be readjusted either manually
or automatically for different feed compositions. Dynamic
compensation, which will be discussed in more detail in
Section 8.21, is also recommended here.
Another method of minimizing feed rate disturbances is
to use adaptive tuning or other nonlinear level control techniques (see Section 2.36 in Chapter 2) on the level controller.
The key is to allow the accumulator, or the tower kettle, to
utilize its capacity to accommodate transient material balance
accumulations and act as a “surge drum” to minimize feed
flow changes to the next unit.
Feed Temperature Control
The thermal condition of the feed determines how much
additional heat must be added to the column by the reboiler.
For efficient separation, it usually is desirable to have the
feed preheated to its bubble point when it enters the column.
Unless the feed comes directly from some preceding distillation step, an outside source of heat is required to achieve
that.
Steam may be used to heat the feed, and a thermocouple
inside a thermowell can detect the temperature inside the feed
© 2006 by Béla Lipták
FIG. 8.19qq
The use of a temperature-flow cascade loop improves the column
feed temperature controls provided by a preheater.
line. In this configuration, the temperature of the feed leaving
this preheater controls the steam flow into the preheater. In
such a cascade configuration, the temperature master is usually a three-mode controller. On start-up, the initially large
correction provided by rate action of this three-mode controller helps to get the unit lined out faster. A full discussion
of the advantages of cascade loops is provided in Section 2.6
in Chapter 2. Figure 8.19qq describes such a control system
on a preheater application.
An alternate feed preheating configuration is to use an
economizer on the feed stream. An economizer is a heat
exchanger designed to take advantage of the waste heat to
preheat the feed. Often, if the bottoms product is just sent to
storage, it must be cooled anyway. Therefore, exchanging its
heat content with the feed stream accomplishes both the
objective of feed preheating and that of product cooling.
Two control configurations are common. If heat from the
bottoms product stream is not sufficient, a second exchanger
using steam is also installed to augment the heating of the
feed, and on this second exchanger, a temperature control
system like that previously shown in Figure 8.19qq is used.
If the available bottoms product heat is more than sufficient,
temperature control is achieved by manipulation of a bypass
valve around the economizer, as shown in Figure 8.19rr.
Constant temperature feed does not necessarily mean
constant feed quality. If feed composition varies, its bubble
point also varies. It is common practice to set the temperature
control at a point that is equivalent to the bubble point of
the heaviest feed. As the feed becomes lighter, some of it will
vaporize, but this variation can be handled by subsequent
controls.
FEEDFORWARD CONTROLS
This section discusses the basic single-input–single-output
feedback control loops, serving to control the product qualities, feed rate and temperature, and tower pressure. Such
feedback systems are capable of compensating for deviations
and disturbances only after they have occurred and have been
8.19 Distillation: Basic Controls
B
Economizer
TF
F
TT
TRC
Reboiler
FIG. 8.19rr
The heat content of the bottoms products can be utilized to preheat
the feed to the column in an economizer preheater.
detected. When using these simple control schemes, the operators are required to manually adjust the set point of these
SISO loops in response to changing plant conditions as they
occur. This approach is usually sufficient to keep the distillation column in operation, but it is not sufficient to achieve
optimal performance.
As will be discussed in Section 8.21 in more detail,
feedforward strategies attempt to compensate for process disturbances in the shortest time possible by accounting for
process dynamics, dead times, time delays, and loop interac1
11
tions. The benefits of better control are:
•
•
•
•
•
•
•
Increased throughput
Increased product recovery
Energy conservation
Reduced disturbances to other processing units
Minimum rework or recycle of off-spec products
Reduced operating personnel
Increased plant flexibility
It has been reported that feedforward-based product composition control of distillation can give energy savings of
5
5–15%.
While feedback-based basic distillation controls only aim
at running the processing unit at current conditions, the objective of feedforward control is not only to do that, but also to
account for conditions that can be anticipated. The challenge
is to utilize the technique, the tools, and available resources
to design unique feedforward control strategies that will
match the specific objectives for the distillation columns. The
choice between any of these control techniques depends upon
factors such as preference and familiarity, complexity, degree
© 2006 by Béla Lipták
1849
of compensation, hardware for application, and number of
variables monitored and controlled by a single strategy.
Often, additional instrumentation is not necessary when
building upon basic feedback control designs to implement
feedforward control. However, in many cases, if key measurements are not available and are needed for the feedforward
calculation or compensation, the installation of new in-line
sensors is also required.
Unlike basic feedback control, where much of the control
could be implemented by simple analog control devices,
feedforward control strategies generally require more sophisticated level computing systems. In this section, the common
applications of feedback-based distillation column controls
are discussed. Section 8.21 will discuss feedforward,
model-based, and other advanced control systems, including optimization.
CONCLUSIONS
This section deals with some of the more basic control configurations for distillation towers. Section 8.20 describes the
calculation of relative gains and Section 8.21 is devoted to
the more advanced and optimized control strategies. The
separation between these three areas is not very sharp, and
some overlap does exist.
In this section, the control strategies for some of the more
common distillation problems have been described. Although
many other system configurations can exist, they usually are
combinations of those presented. Control strategies today are
no longer hardware dependent. Most modern microprocessorbased systems are designed with control function modules to
execute a variety of the basic strategies that were discussed
in this section. Multivariable unit operations controllers of
both the model-predictive and the model-free variety are also
slowly becoming available and will be discussed in
Section 8.21.
It is important to emphasize that control by feedback
methods alone cannot approach the quality of control possible by predictive (feedforward) techniques. This is true even
though it is likely that the predictive control equations may
need to be updated by feedback. In effect, predictive control
tends to substantially reduce the size of the errors that are
left to be handled by feedback. Further discussion of feedforward strategies and of other techniques for optimization
are provided in more detail in Section 8.21.
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8.19 Distillation: Basic Controls
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O’Conner, J., and Illing, H., “The Turbine Control Valve. A New Approach
to High-Loss Applications,” Instrumentation Technology, December
1973.
1974
Ackley, W. R., “Feedforward Control Strategy for Distillation Towers,”
Instrumentation Technology, February 1974.
Adiutori, E. F., The New Heat Transfer, Cincinnati, OH: Ventuno Press, 1974,
Chapter 7.
Kline, P. E., “Technical Task Force Approach to Energy Conservation,”
Chemical Engineering Progress, February 1974.
Shinskey, F. G., “Values of Process Control,” Oil and Gas Journal, February
18, 1974.
Ellerbe, R. W., “Steam Distillation Basics,” Chemical Engineering (NY),
March 4, 1974.
Lupter, D. E., “Distillation Column Control for Utility Economy,” presented
at 53rd Annual GPA Convention, Denver, March 25–27, 1974.
Waller, K. V. T., “Decoupling in Distillation,” AIChE Journal, May 1974.
Farwell, W., “More Cooling at Less Cost,” Plant Engineering, August 1974.
Gallier, P. W., and McCune, L. C., “Simple Internal Reflux Control,” Chemical Engineering Progress, September 1974.
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Distillation Systems with Energy Integration,” AIChE Journal,
September 1974.
© 2006 by Béla Lipták
1851
Buckley, P. S., “Material Balance Control in Distillation Columns,” presented at AIChE Workshop on Industrial Process Control, Tampa, FL,
November 11–13, 1974.
1975
Bojnowski, J. H., Crandall, J. W., and Hoffman, R. M., “Modernized Separation Systems Saves More than Energy,” Chemical Engineering
Progress, October 1975.
Boyd, D. M., “Fractionation Column Control,” Chemical Engineering, June
1975.
Buckley, P. S., Cox, R. K., and Rollins, D. L., “Inverse Response in a
Distillation Column,” Chemical Engineering Progress, June 1975.
Franzke, A., “Save Energy with Hydraulic Power Recovery Turbines,”
Hydrocarbon Process, March 1975.
Luyben, W. L., “Steady-State Energy Conservation Aspects of Distillation
Column Control Design,” I&EC Fundam., Vol. 14, No. 4, pp. 321–325,
1975.
Monroe, E. S., “Vacuum Pumps Can Conserve Energy,” Oil and Gas Journal, February 3, 1975.
Rademaker, O., Rijnsdorp, J. E., and Maarleveld, A., Dynamics and Controls
of Continuous Distillation Units, Amsterdam: Elsevier, 1975.
Shinskey, F. G., “Controlling Unstable Processes, Part II: A Heat Exchange,”
Instrument and Control Systems, January 1975.
Sternlicht, B., “Low-Level Heat Recovery Takes on Added Meaning as Fuel
Costs Justify Investment,” Power, April 1975.
Tyreus, B., and Luyben, W. L., “Control of a Binary Distillation Column
with Sidestream Drawoff,” I&EC Process Design and Development,
Vol. 14, No. 4, p. 391, 1975.
Tyreus, B., and Luyben, W. L., “Two Towers Cheaper than One?” Hydrocarbon Processing, Jun., 1975.
Wolf, C. W., Weiler, D. W., and Ragi, E. G., “Energy Costs Prompt Improved
Distillation,” Oil and Gas Journal, September 1, 1975.
1976
ISA Standard S51.1, “Process Instrumentation Terminology,” Research Triangle Park, NC: Instrument Society of America, 1993.
Shinskey, F. G., “Energy-Conserving Control for Distillation Units,” Chemical Engineering Progress, May 1976.
Wright, R. M., “A Better Approach to Distillation Control,” Instruments and
Control Systems, June 1976.
Van Horn, L. D., and Latour, P. R., “Computer Control of a Crude Still,”
Instrumentation Technology, November 1976, pp. 33–39.
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1977
Hadley, K., “Control Objectives Analysis,” National Petroleum Refiners
Association Computer Conference, New Orleans, 1977.
Hobbs, J. W., “Unifying Concepts for Fractionator Product Composition
Control,” Proceedings of the Automatic Control Conference, San
Francisco, 1977.
Evans, L. B., “Impact of the Electronics Revolution on Industrial Process
Control,” Science, Vol. 195, No. 4283, March 18, 1977.
Hughart, C. L., “Designing Distillation Units for Controllability,” Instrumentation Technology, May 1977.
Lieberman, N., “Instrumenting a Plant to Run Smoothly,” Chemical Engineering, September 12, 1977, pp. 140–154.
Shinskey, F. G., “The Stability of Interacting Loops with and without Decoupling,” presented at the IFAC Symposium on Multivariable Control,
Fredericton, New Brunswick, July 4–8, 1977.
1852
Control and Optimization of Unit Operations
1978
Buckley, P. S., Cox, R. K., and Luyben, W. L., “How to Use a Small
Calculator in Distillation Column Design,” Chemical Engineering
Progress, June 1978, pp. 49–55.
Buckley, P. S., “Distillation Column Design Using Multivariable Control,”
Instrumentation Technology, September/October 1978.
Buford, B. N., Bush, B. A., and Staten, H. W., “Computers Conserve Energy
in NGL Fractionators,” Oil & Gas Journal, December 1978.
Doukas, N., “Control of Sidestream Columns Separating Ternary Mixtures,”
Instrumentation Technology, June 1978.
Doukas, N. and Luyben, W. L., “Economics of Alternative Distillation Configurations for the Separation of Ternary Mixtures,” I&CE Process
Design and Development, Vol. 17, No. 272, 1978.
Griffin, D. E., Parsons, J. R., and Smith, D. E., “The Use of Process Analyzers for Composition Control of Fractionators,” Proceedings of the
ISA Spring Joint Conference, Houston, TX, May 1978.
Latour, P. R., “Composition Control of Distillation Columns,” Instrumentation Technology, July 1978.
Lamb, M. Y., “Computer Control of a Propylene Upgrading Unit,” 85th
National AIChE Meeting, Philadelphia, PA, 1978.
McCoy, R. D., “Adding Capabilities to Process Chromatography with
Microprocessor-Based Programmers,” Proceedings of the ISA Joint
Spring Conference, Houston, TX, May 1978.
Mix, T. J., Dweck, J. S., Weinberg, M., and Armstrong, R. C., “Energy
Conservation in Distillation,” Chemical Engineering Progress, April
1978.
Painter, J. W., and Gonnella, J. L., “Improved Control of a Distillation
Column Using a Minicomputer and an On-Line Gas Chromatograph,”
Texas A&M Instrumentation Symposium, January 1978.
Tolliver, T. L., and McCune, L. C., “Distillation Column Control Design
Based on Steady State Simulation,” ISA Transactions, Vol. 17, No. 3,
pp. 3–10, 1978.
1979
Watkins, R. N., Petroleum Refinery Distillation, 2nd ed., Houston, TX: Gulf
Publishing Company, 1979.
Douglas, J. M., Jafarey, A., and McAvoy, T. J., “Shortcut Techniques for
Distillation Column Design and Control, Part 1: Column Design,” I&EC
Process Design and Development, Vol. 18, pp. 197–202, April 1979.
Van Horn, L. D., “Computer Control. How to Get Started,” Hydrocarbon
Processing, September 1979.
Chin, T. G., “Guide to Distillation Pressure Control Methods,” Hydrocarbon
Processing, October 1979, 145–153.
Black, J. W., “Model Estimates Analyzer Payouts,” Hydrocarbon Processing, September 1981.
Shinskey, F. G., “Controlling Distillation Processes for Fuel-Grade Alcohol,”
Instruments and Control Systems, December 1981.
Thurston, C. W., “Computer-Aided Design of Distillation Column Controls,”
Hydrocarbon Processing, Part 2, August 1981, p. 5.
1982
Rinne, R., Sunnel, H., Latour, P. R., and Payntex, K. K., “Experience with
Distillation Unit Computer Control,” Hydrocarbon Processing, March
1982.
Roffel, B., and Rijnsdorp, J. E., Process Dynamics, Control, and Protection,
Ann Arbor, MI: Ann Arbor Science, 1982.
DiBano, R. J., “Advanced Control: General Purpose Design vs. Working It
Out in the Field,” Hydrocarbon Processing, August 1982.
Dweck, J. S. and Mix, T. W., Conserving Energy in Distillation, MIT Press,
1982.
1984
Mix, P. E., The Design and Application of Process Analyzer Systems, New
York: John Wiley & Sons, Inc., 1984.
Stephanopoulos, G., Chemical Process Control, An Introduction to Theory
and Practice, Englewood Cliffs, NJ: Prentice Hall, 1984.
Barduhn, A. J., “Setting the Pressure at which to Conduct a Distillation,”
Chemical Engineering Education, Vol. 18, No. 1, pp. 38–40, 1984.
1985
Deshpande, P. B., Distillation Dynamics and Control, Research Triangle
Park, NC: Instrument Society of America, 1985.
Buckley, P. S., Luyben, W. L., and Shunta, J. P., Design of Distillation
Column Control System, Research Triangle Park, NC: Instrument Society of America, 1985.
Gani, R., Romagnoli, J. A., and Stephanopoulos, G., “Control Studies in an
Extractive Distillation Process: Simulation and Measurement Structure,” Chem. Engng. Commun., 40, pp. 281–302, 1985.
1986
Tsai, T. H., Lane, J. W., and Lin, C. S., Modern Control Techniques for the
Process Industries, Vol. 23, Marcel Dekker, 1986.
Hansen, T. T., and Jørgensen, S. B., “Optimal Open Loop Control of Binary
Batch Distillation,” Chem. Eng. J., 33 , pp. 151–155, 1986.
1980
1987
Treybal, R. E., Mass-Transfer Operations, 3rd ed., New York: McGraw-Hill
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Van Horn, L. D., “Crude Unit Computer Control. How Good Is It?” Hydrocarbon Processing, April 1980.
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Black, C., “Distillation Modeling of Ethanol Recovery and Dehydration
Process for Ethanol and Gasohol,” Chemical Engineering Progress,
September 1980.
Andrews, A. J., and Griffin, D. E., “Performance Audits Evaluate Two
Distillation Control Projects,” Oil & Gas Journal, December 8, 1980.
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Binary Batch Distillation with Recycled Waste Cut,” Chem. Eng. J., 34,
pp. 57–64, 1987.
1981
Griffin, D. E., “Tighten Distillation Column Control and Save Energy,”
Instruments and Control Systems, March 1981.
© 2006 by Béla Lipták
1988
Balchen, J. G., Process Control, Structures, and Applications, New York:
Van Nostrand Reinhold, 1988.
Birky, G. J., McAvoy, T. J., and Tyreus, B. D., “Expert System for Design
of Distillation Controls,” Proceedings of the ISA/88 International Conference and Exhibit, Houston, TX, October 1988.
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Triangle Park, NC: Instrument Society of America, 1988.
Nichols, G. D., On-Line Process Analyzers, New York: John Wiley & Sons,
Inc., 1988.
Roat, S. D., Moore, C. F., and Downs, J. J., “A Steady-State Distillation
Column Control System Sensitivity Analysis Technique,” Proceedings
IEEE Southeast Con., 1988, pp. 296–300.
8.19 Distillation: Basic Controls
1989
Finco, M. V., Luyben, W. L., and Polleck, R. G., ‘‘Control of Distillation
Columns with Low Relative Volatilities,” Ind. Eng. Chem. Res., Vol. 28,
January 1989, pp. 75–83.
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for a Full-Scale Industrial Distillation Column,” Control Systems Magazine, January 1989, pp. 91–96.
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and Exhibit, Part 1, Philadelphia, PA, October 1989.
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Projects,” Control, October 1989.
Luyben, W. L., Process Modeling, Simulation, and Control for Chemical
Engineers, 2nd ed., New York: McGraw-Hill Book Company, 1989.
Kister, H. Z., Distillation Operation, New York: McGraw-Hill Publishing
Company, 1989.
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Process Model-Based Control,” ISA89, Philadelphia, PA, October 1989.
Riggs, J. B., “Nonlinear Process Model-Based Control of a Distillation
Column with a Sidestream Draw-Off,” presented at the annual AIChE
meeting, San Francisco, CA, November 1989.
Riggs, J. B., Sinha, R., and McDaniel, R., “Comparison of Control Techniques for High Purity Distillation Columns,” AIChE Spring National
Meeting, Houston, TX, April 1989.
1853
Riggs, J. B., Watts, J., and Beauford, M., “Advanced Model-Based Control
for Distillation,” National Petroleum Refinery Association Meeting,
Seattle, WA, October 1990.
Riggs, J. B., Watts, J., and Beauford, M., “Industrial Experience with Applying Nonlinear Process Model-Based Control to Distillation Columns,”
ISA90, New Orleans, LA, October 1990.
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1991
Coughanowr, D. R., Process Systems Analysis and Control, 2nd ed., New
York: McGraw-Hill Book Company, 1991.
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Processing, May 1991.
Papastathopoulou, H. S., and Luyben, W. L., “Control of a Binary Sidestream
Distillation Column,” Industrial & Engineering Chemistry Research,
April 1991, pp. 705–713.
Rijnsdorp, J. E., Integrated Process Control and Automation, Amsterdam:
Elsevier, 1991.
Sandelin, P. M., Haeggblom, K. E., and Waller, K. V., “Disturbance Rejection
Properties of Control Structures at One-Point Control of a Two-Product
Distillation Column,” Industrial & Engineering Chemistry Research,
June 1991, pp. 1182–1186.
Sandelin, P. M., Haeggblom, K. E., and Waller, K. V., “Disturbance Sensitivity Parameter and its Application to Distillation Control,” Industrial
& Engineering Chemistry Research, June 1991, pp. 1187–1193.
1990
1992
Papastathopoulou, H. S., and Luyben, W. L., “Turning Controllers on Distillation Columns with the Distillated Bottoms Structure,” Industrial and
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Chien, I.-L., and Fruehauf P. S., “Consider IMC Tuning to Improve Controller Performance,” Chemical Engineering Progress, Vol. 86, No. 10,
pp. 33–41, October 1990.
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Dynamic Model of Binary Distillation Column,” Chemical Engineering Science, Vol. 45, No. 12, pp. 3585–3592, 1990.
Kister, H. Z., Distillation Operation, New York: McGraw-Hill Book Company, 1990.
Li, R., Olson, J. H., and Chester, D. L., “Dynamic Fault Detection and
Diagnosis Using Neural Networks,” Proceedings of the Fifth IEEE
International Symposium on Intelligent Control, Philadelphia, PA,
September 1990, pp. 1169–1174.
McGreavy, C., Dynamics and Control of Chemical Reactors, Distillation
Columns, and Batch Processes, Pergamon Press, 1990.
Papastathopoulou, H. S. and Luyben, W. L., “Potential Pitfalls in Ratio
Control Schemes,” Industrial & Engineering Chemistry Research,
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Pitt, M. J., Instrumentation and Automation in Process Control, New York:
Horwood, 1990.
Skogestad, S., Lundstrom, P., and Jacobsen, E. W., “Selecting the Best
Distillation Control Configuration,” AIChE Journal, Vol. 36,
pp. 753–764, 1990.
Skogestad, S., Jacobsen, E. W., and Morari, M., “Inadequacy of SteadyState Analysis for Feedback Control. Distillate. Bottom Control of
Distillation Columns,” Industrial & Engineering Chemistry Research,
December 1990, pp. 2339–2346.
Pandit, H. G., and Rhinehart, R. R., “Process Model-Based Control of a
Nonideal Binary Distillation Column,” Proceedings of the Annual
AIChE Meeting, Chicago, IL, November 1990.
Pedersen, N. H., and Jørgensen, S. B., “A GC Subsystem for Fast On-Line
Concentration Profile Measurement for Advanced Distillation Control,” Analytica Chemica Acta, 238, pp. 139–148, 1990.
Fruehauf, P. S., and Mahoney, D. P. “Distillation Column Control Design
using Steady-State Models, Usefulness and Limitations,” in Advances
in Instrumentation and Control, Vol. 47, Part 1, Research Triangle
Park, NC: Instrument Society of America, 1992, pp. 92–120.
Skogestad, S., “Dynamics and Control of Distillation Columns: A Critical
Survey,” Preprints IFAC Symposium, DYCORD +92, College Park,
MD, pp. 1–25.
Luyben, W. L., Practical Distillation Control,” New York: Van Nostrand
Reinhold, October 1992.
Koggersbøl, A., and Jørgensen, S. B., “Dynamics and Control of a Distillation Column with a Sidestream,” IChemE Symposium Series, 128,
pp. A429–A449, 1992.
© 2006 by Béla Lipták
1993
Fruehauf, P. S., and Mahoney, D. P., “Distillation Column Control Design
Using Steady State Models: Usefulness and Limitations,” ISA Transactions, Research Triangle Park, NC. 1993.
Balchen, J. G., Dynamics and Control of Chemical Reactors, Distillation
Columns, and Batch Processes, Pergamon Press, April 1993.
Ganguly, S., “Model Predictive Control of Distillation,” ISA/93 Technical
Conference, Chicago, IL, September 1993.
1994
Gokhale, V., Shukla, N., and Munsif, H., “Analysis of Advanced Distillation
Control on a C3 Splitter and a Depropanizer,” 1994 AIChE National
Annual Meeting, San Francisco, CA, November 1994.
Fruehauf, P. S., and Mahoney, D. P., “Improve Distillation Column Control
Design,” Chemical Engineering Progress, March 1994.
1995
Fleming, B., and Sloley, A. W., “Feeding and Drawing Products: The Forgotten Part of Distillation,” Proceedings of the ChemShow and Exposition, New York, December 1995.
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Musch, H. E., and Steiner, M., “Robust PID Control for an Industrial Distillation Column,” Control System Magazine, Vol. 15, No. 4, 46–55,
1995.
Hurowitz, S. E., and Gokhale, V., “A Dynamic Model of a Superfractionator: A Test Case for Comparing Distillation Control Techniques,”
DYCORD ‘95, 4th IFAC Symposium, Helsingor, Denmark, June
1995.
Lundstrom, P., and Skogestad, S., “Opportunities and Difficulties with 5 × 5
Distillation Control,” J. Process Control, Vol. 5, 249–261, 1995.
Rawlings, J. B., “Dynamics and Control of Chemical Reactors, Distillation
Columns, and Batch Processes (Dycord+‘95),” a Postprint Volume
from the 4th IFAC Symposium on Dynamics and Control of Chemical
Reactors, Distillation Columns and Batch Processes (DYCORD ‘95),
Helsingor, Denmark, June 1995.
Diwekar, U. M., “Batch Distillation: Simulation, Optimal Design, and Control (Series in Chemical and Mechanical Engineering),” Taylor & Francis, September 1995.
Banerjee, A., and Arkun, Y., “Control Configuration Design Applied to the
Tennessee Eastman Plantwide Control Problem,” Computers & Chem.
Eng., 19(4), 453–480, 1995.
1996
Koggersbøl, A., Andersen, R., Nielsen, J. S., Jørgensen, S., “Control Configuration for Energy Integrated Distillation,” Computers & Chem.
Eng., 20 (supplement), pp. S853–S858, 1996.
1997
Linsley, J., “New, Simpler Equations Calculate Pressure-Compensated Temperatures,” Oil & Gas Journal, May 24, 1997, 58–64.
Ming T. Tham, “Distillation,” Base Document URL: http://lorien.ncl.ac.uk/
ming/distil/distil0.htm, October 1997.
Anderson, N. A., Instrumentation for Process Measurement and Control,
3rd edition, Boca Raton, FL: CRC Press, October 1997
Hurowitz, S. E. and Anderson, J. J., “Distillation Configuration Selection
for Dual Composition Control,” AIChE Spring National Meeting,
Houston, TX, April 1997.
Hurowitz, S. E. and Anderson, J. J., “Control of High Purity Distillation
Columns,” Control 97 Conference, Sydney, Australia, October 1997.
Mahoney, D. P. and Fruehauf, P. S., “An Integrated Approach for Distillation
Column Control Design Using Steady-State and Dynamic Simulation,”
Aspentech technical articles, March 1997. www.aspeutech.com/
corporate/press/publications.
Skogestad, S., “Dynamics and Control of Distillation Columns: A Tutorial
Introduction,” Trans. IChemE., Vol. 75, Part A, pp. 539–562, 1997.
Riggs, J. R., “Improve Distillation Column Control,” Chemical Engineering
Progress, October 1998, 31–47.
Stichlmair, J. G. and Fair, J. R., Distillation: Principles and Practices, New
York: John Wiley & Sons, 1998.
1999
Eker, I. and Sakthivel, K., “Automation & Lube Oil Additives Blending Plant
Using an S88.01 Consistent Batch Software: A Case Study,” Proceedings of the World Batch Forum, San Diego, CA, April 1999.
Hurowitz, S., Anderson, J., Duvall, M., and Riggs, J.B., “An Analysis of
Controllability Statistics for Distillation Configuration Selection,” presented at the AIChE Annual Meeting, Dallas, TX, November 1999.
2000
Andrew W. Sloley, “Steady Under Pressure: Distillation Pressure Control,”
presented at the American Institute of Chemical Engineers Spring
Meeting, March 6–9, 2000.
Betlem, B.H.L. “Batch Distillation Column Low-Order Models for Quality
Control Program,” Chemical Engineering Science, 55, pp. 3187–3194,
2000.
Roffel, B., Betlem, B. H. L., and De Ruijter, J. A., “Modeling and Control
of a Cryogenic Distillation Column,” Computers and Chemical Engineering, 24, pp. 111–123, 2000.
Roffel, B., “Distillation: Instrumentation and Control Systems,” in Encyclopedia of Separation Science, London: Academic Press, 2000.
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guide),” 2000. http://lorien.ud.cc.uk/ming/control/g
2002
Florez, M., “Batch Distillation: Practical Aspects of Design and Control,”
Proceedings of the World Batch Forum, Woodbridge Lake, NJ, April,
2002
Cook, B., Engel, M., Landis, C., Tedeschi, S., and Zehnder, A., “Synthesis
of Optimal Batch Distillation Sequences,” Proceedings of the World
Batch Forum, Woodbridge Lake, NJ, April 2002.
2003
Kralj, F., “Application of the S88 Model in the Control of Continuous
Distillation Facilities,” Proceedings of the World Batch Forum, Woodbridge Lake, NJ, April 2003.
Jones, M., and Kilian, A., ‘‘Tricky Pressure Control in Distillation Column,”
July, 2003. http://instrumentation.co.za/regulator
1998
2004
Betlem, B. H. L., Krijnsen, H. C., and Huijnen, H., “Optimal Batch Distillation Control Band on Specific Measures,” Chemical Engineering
Journal, 71, pp. 111–126, 1998.
© 2006 by Béla Lipták
Hurowitz, S., Anderson, J., Duvall, M., and Riggs, J. B., “Distillation Control
Configuration Selection,” submitted to J. Process Control, March 2004.