- Distillation Principles
- Distillation Modeling
- Distillation Operating Equations
- Distillation Calculations
- Distillation Enthalpy Balances
- Enthalpy-Concentration Method
- Equipment & Column Sizing

Before beginning most distillation calculations, a decision must be reached: does equimolal overflow apply? If so, the operating equations are lines and you have one set of options -- notably the McCabe-Thiele method. If not, energy balances must be explicitly considered.

There are several ways of incorporating the energy effects. The Ponchon-Savarit method is a graphical approach that does not require an assumption of equimolal overflow. Graphical construction is done on the enthalpy composition diagram. Your text does not include this method, but it may be useful.

In all cases, one can use a "stepping" approach. Starting from one end of the column, the component, material, and energy balances can be solved simultaneously. After a stage is determined, you step up (or down) to the next and calculate that stage. Depending on what information is known, the form of the equilibrium relation, etc., this approach may require an iterative solution.

The graphical *McCabe-Thiele Method* can be used to determine the number
of ideal stages and feed tray location. To do this, you make a plot showing
the equilibrium curve, feed line, and operating lines for the rectifying and
stripping sections (all on the same axes), and then find answers by graphical
construction.

A standard (beginner) distillation problem provides you with x_{F}, q,
x_{D}, x_{B}, and R_{D}; although not necessarily
directly. You may need to use the overall material balances to find some of the
compositions; to calculate the q factor from temperature and composition data;
and/or determine the reflux ratio based on the limiting minimum value. Once you
have these values, the solution procedure is:

- Plot the equilibrium curve.
- Calculate the slope
*q/(q-1)*of the feed line. - Plot the feed line using x=x
_{F}and the slope. - Calculate the y-intercept of the rectifying line
*x*._{D}/(R_{D}+1) - Plot the rectifying line using (xd,xd) and the intercept.
- Draw the stripping line connecting the intersection of the rectifying and feed lines and the point (xb,xb).
- Correct for stage efficiency by drawing the "effective" equilibrium curve between the equilibrium curve and the operating lines.
- "Step off" equilbrium stages.

If you don't want to draw, you can do the same thing, iteratively solving for the various equation intersections. You'll need to be pretty careful with your "bookkeeping" if you try this.

To apply Murphree tray efficiencies, construct an effective equilibrium curve between the equilibrium and operating curves, and step using the effective curve to determine actual separation stages. Remember that a partial reboiler or partial condenser is by definition an ideal stage, so you use the ideal equilibrium curve (not the effective) for these stages.

The optimum feed tray is the triangle with one corner on the rectifying line and one on the stripping line. Putting the feed anywhere else increases the number of stages needed to make the separation. To visualize this, notice that the closer the operating line is to the equilibrium curve, the smaller the stepping triangles become. Introducing the feed at the intersection of the rectifying and stripping lines maximizes the size of the triangles and so leads to the fewest steps.

When analyzing existing distillation systems, the actual feed entry point may not be at the optimum (the tray where the operating lines intersect). In this case, the tray stepping should switch from the rectifying line to the operating line at the actual feed tray location.

An example of McCabe-Thiele analysis has been prepared. It is available for download as a Mathcad 5.0+ file.

Frequently, when analyzing or designing a process, it is useful to look at limiting cases to assess the possible values of process parameters. In distillation analysis, separation of a pair of components can be improved by increasing the number of stages while holding reflux constant, or by increasing the reflux flow for a given number of stages. This tradeoff sets up two limiting cases:

- Total Reflux (minimum ideal stages)
- Minimum Reflux (infinite ideal stages)

The total reflux condition represents operation with no product removal. All the overhead vapor is condensed and returned as reflux. Consequently, the reflux ratio (L/D) is infinite. This, in turn, makes the operating lines the 45 degree line (prove it to yourself by setting D=0, and noticing that consequently L=V). With the operating lines on the diagonal, they are as far as they can get from the equilibrium curve, so if the number of plates are stepped off using the diagonal and the equilibrium curve, the number of theoretical stages will be a minimum.

Often, columns are operated at total reflux during their initial startup, and product is not withdrawn until a separation close to that desired is achieved.

The *Fenske Equation* is another method for determining the minimum
number of trays required for a given separation. It is an example of a
"shortcut" distillation method. There are a number of these approximate
methods available to get initial estimates of distillation requirements.

The Fenske equation applies to distillation systems with constant
relative volatility. Note that the form of the Fenske equation
shown calculates the minimum number of *plates*; it does not include the
reboiler (hence the -1 on the right hand side). Other texts may use a form for
the minimum number of *stages* and not subtract the reboiler.

If the relative volatility varies through a column because of temperature effects, it is possible to use a geometric mean value of the relative volatility (as is done for multicomponent distillation) and the Fenske equation to get an approximate value for the number of stages.

The intersection of an operating line and the equilibrium curve is called a
*pinch point*. A simple column will have two pinch points (because
there are two operating lines). The points change when the operating lines
do. An existing column can "pinch" if its operating line is too close to its
equilibrium curve. This means that there are several stages doing very little
separation and wasting resources.

To cure a pinch, the most direct solution is to move the feed entry point. This is often an expensive proposition. In such cases, the reflux and boilup ratios can be increased to change the operating lines. This will increase operating costs and energy consumption, but may be the only realistic option.

A pinch at the intersection of the feed line and the equilibrium curve indicates
that the column is operating at *minimum reflux*.

The minimum reflux condition represents the theoretical opposite of total reflux -- an infinite number of ideal separation stages. In this case, the intersection of the operating lines lies on the equilibrium curve itself. Thus, the distance between the equilibrium curve and the operating lines is at its minimum, the stepping triangles become very small, there is no gap between the equilibrium curve and the intersection point, so you cannot step past the feed point.

The minimum reflux rate can be determined mathematically from the endpoints of
the rectifying line at minimum reflux -- the overhead product composition point
(x_{D},x_{D}) and the point of intersection of the feed line
and equilibrium curve (x', y').

If the equilibrium curve has an inflection point, it may not be possible to
construct a line between the overhead product point and the feed/equilibrium
intersection without passing outside the equilibrium envelope. Operating curves
must always intersect within the equilibrium envelope and cannot cross outside
(in either half of the column). In this case,
minimum reflux occurs at a *tangent pinch* and the operating line is
tangent to the equilibrium curve. Calculations are based upon the intersection
of the tangent operating line and the feed point.

When designing columns, it is common to define the design reflux ratio as some multiple of the theoretical minimum reflux. The cost optimum reflux ratio is typically in the 1.1 to 1.5 range depending on energy costs, condenser coolant, and materials of construction. The rule of thumb reported most often suggests that a reflux ratio of about 1.2 times the minimum is a good design value.

It may also be necessary to distinguish between *returned reflux* (the
reflux stream flowing from the accumulator to the column) and the *effective
reflux* flowing down the column. This is a concern if the reflux is
subcooled. In this case, the effective reflux will consist of the returned
reflux plus whatever additional liquid is condensed when the cold liquid
contacts vapor on the reflux tray. It is probably best to use effective reflux
in minumum reflux calculations.

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Minimum reflux corresponds to a pinch at the intersection of the feed line and
the equilibrium curve. From its formula, the rectifying line has slope
and will connect the intersection point (x_{int}, y_{int}) and
(x_{D}, x_{D}). Consequently, we can express the slope in
terms of "rise over run", or

Algebraic rearrangement gives the desired formula:

Clearly, this formula doesn't apply if there are more than two operating regions. In these cases, it is probably smarter to calculate the reflux ratio from the ratio of the liquid and vapor flow rates.

The heating and cooling loads in the condenser and reboiler can be calculated from straighforward energy balances. The chief difficulty is in getting good values for the heats of vaporization -- an enthalpy-concentration diagram is very useful!

For the reboiler, energy must be added equal to the sum of the sensible heat needed to raise the liquid to its boiling point and the latent heat of vaporization. The steady-state energy balance on the process side of the reboiler is then:

The heating medium requirements can be calculated from an energy balance on the heating side of the reboiler. If saturated steam is the heating medium, then

A similar analysis provides the condenser load as

It is possible to imagine a case where ; that is, when . The case when the vapor rates at the top and bottom of the column are most likely to be the same occurs when the feed is at its bubble point, and . In this situation, the condenser and reboiler loads will be approximately equal.

Mass transfer limitations prevent the vapor leaving a tray from truly being in precise equilibrium with the liquid on the tray; consequently, the assumption of ideal stages is only an approximation.

An efficiency is used to represent the deviation from equilibrium. There are three types of efficiencies we will consider:

- Overall efficiency
- Local efficiency
- Murphree efficiency

An *overall efficiency* is the simplest choice. It is the ratio of the number of
ideal stages to the number of actual stages.

The *local efficiency* is the most accurate option, but also the most difficult to
use. It is defined at only a single point on a specific tray.

A *Murphree efficiency* is probably the most common choice, since
it represents a workable compromise between accuracy and ease of use. It
has the same form as a local efficiency but is based on tray average
compositions.

The overall efficiency and the Murphree efficiency are not directly related.
You **cannot** use an average Murphree efficiency in place of an
overall value.

To use a local or Murphree efficiency with a graphical method, the true equilibrium curve is
replaced with an *effective equilibrium curve* located between the true
curve and the operating curves. The effective curve is used to count stages.
Note, however, that the efficiency doesn't apply to the reboiler, so the true
equilibrium curve should be used for the last stage of the stripping section.

To construct an effective equilibrium curve is not difficult. The effective curve is given by:

- Seader, J.D. and E.J. Henley, Separation Process Principles, John Wiley, 1998, pp. 292-4, 299.

R.M. Price

Original: 2/7/97

Revised: 2/28/97, 1/26/98, 1/25/99, 2/14/2003

Copyright 1997, 1998, 1999, 2003 by R.M. Price -- All Rights Reserved