Cut column energy consumption

With increasing energy prices and public sensitivity to energy consumption, many plants are paying more attention to optimizing distillation. Simple and low-risk operational changes often can provide substantial savings.

By Andrew Sloley, VECO U.S.A.

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Distillation units account for a significant fraction of most plants’ energy requirements. With increasing energy prices and public sensitivity to energy consumption, many plants are paying more attention to optimizing distillation. Energy improvements can result from changes in equipment, process and operations.

Of these, operational changes pay off most quickly and reliably. Usually because plants over time invariably stray from optimized operation many opportunities for improvement exist.

A number of factors can contribute to unnecessary energy consumption: improper control points; selection of the wrong control variables; unused system capabilities; excessive reflux or boil-up: poor equipment efficiency; and missing basic heat-integration steps. Nearly every plant can capture profits by paying attention to these basics. In many cases, improvements involve only simple operating changes and low investment.

So here we’ll look at simple steps and procedures to verify energy efficiency and to identify possible operating changes and minor projects and their benefits.

Is your control measurement in the right place?

First let’s consider control measurement on the bottom of a column, which commonly relies on temperature to infer bottoms composition (Figure 1).

Figure 1. The bottom tray usually is not the best place to make an inferential measurement of composition.

Figure 1. The bottom tray usually is not the best place to make an inferential measurement of composition.

While it may seem obvious to take the temperature measurement at the bottom of the column, the proper measurement location depends upon many factors including control response, process gain, system composition and behavior [1], as well as energy costs.

Simple inferential composition control most often works best with the control point a number of trays — typically from three to 20 depending upon the system — away from the column bottoms. The same holds true for inferred control on the column overhead.

In one typical case, shifting the control point from the column bottoms to six trays up yielded a $60,000 annual energy savings, and required only a new thermowell and a minor instrumentation run. Column energy consumption dropped by 4.3%, with 3% due to reduced variation in heat duty at a constant target operation and the remainder from the tighter operation allowing for a reduced operating margin above minimum specification.

Are you controlling the right thing?

Thorough process analysis [2] to ensure you are measuring the right control variable is important in energy conservation, as Figure 2 helps illustrate. It shows a compound absorber/stripper using an absorption fluid to reduce product losses and a reboiler for stripping.

Figure 2. Common control scheme is susceptible to significant errors in inferred measurement.

Figure 2. Common control scheme is susceptible to significant errors in inferred measurement.

It also depicts a common, but poorly chosen, control scheme — the temperature on a tray above the bottoms infers bottoms composition. In most stripping services small temperature changes produce large changes in bottoms composition.

Figure 3 charts a specific case of bottoms composition versus the control-point temperature. To keep the bottoms composition under the required operating point (line D) the temperature operating point needs to be set high enough to maintain purity specifications during all reasonable temperature excursions. Line F shows the resulting average composition for a 2°F offset. In comparison, using an alternative control specification, relative stripping vapor recycle flow allows operation at line E (a 2.5% flow offset).

Figure 3. Control should take into account that small temperature changes can cause large swings in composition.

Figure 3. Control should take into account that small temperature changes can cause large swings in composition.

The average energy saving is 3.6% of reboiler duty for the alternative control scheme. Additionally, that scheme cuts consumption by an extra 2.7% because of non-symmetrical response to control variations.

The control of the overhead gas rate is an excellent choice for strippers as long as feed composition is relatively constant. If feed composition varies occasional updates to the ratio of overhead gas to feed may be required.

Can you control to system limits?

Take advantage of the full capability of existing equipment. Don’t let specification-sheet design values restrict you. Make the most of the safety margin that invariably is included.

For instance, consider reducing column pressure to leverage potentially under-used capacity in the tower and the overhead system. Some authorities claim up to 25% savings in distillation energy consumption depending upon operating conditions [3].

For nearly all systems, decreasing the pressure increases the relative volatility. Higher relative volatility reduces the boilup and reflux requirements at a constant product split and so cuts energy needs. Factors favoring lower pressure are:

  1. actual number of separation stages close to the theoretical minimum;
  2. high purity requirements;
  3. low relative volatility; and
  4. relatively large changes in relative volatility with pressure.

The first three factors tend to create systems with high reflux ratios. The fourth identifies systems sensitive to pressure changes. High reflux ratios (high energy demand) coupled with a sensitive system favor pressure minimization.

Simulations can rapidly identify systems with potential for significant energy savings. If a simulation isn’t available, accurate shortcut methods [4] can quickly pinpoint likely candidates.

Systems generally are limited either by overhead condensing capability or tower capacity. If the overhead system can control pressure within the operating margin, it isn’t a limit. A quick review of control parameter ranges (for example, valve position) in the overhead system will rapidly identify if you have an overhead system limit.

In general, most towers will eventually flood as pressure drops. However this isn’t always true, so be careful. In the unusual case where flooding isn’t a concern, you can safely reduce pressure. In the more common situation where tower flooding becomes more of a problem at lower pressure, you need to know how to keep out of flood. This requires monitoring pressure drop across the tower’s limiting section, so you can operate as close as possible to tower flood (and at the lowest and cheapest pressure). Direct measurement of pressure drop is most common, but the flow of the vapor [5] or liquid by internal restrictions also can provide such data.

Are you controlling energy consumption at all?

Figure 4 shows part of a distillation scheme with heat integration to provide reboiler heat from two sources — recovery from reactor effluent and 600-psig steam.

Figure 4. In this scheme no process value has an impact on the heat input to the tower.

Figure 4. In this scheme no process value has an impact on the heat input to the tower.

Taking advantage of the reactor effluent certainly reduces unit energy costs. However, the apparent complexity of the control scheme hides one major factor: no process value affects the heat input into the tower. The total heat input is set by the hand control valve (HC) on the reactor effluent split and the flow controller on the 600-psig steam. Essentially, the tower operates at constant duty. Constant duty operation, either in the condenser or the reboiler, nearly always is a sign of wasted energy.

A survey of this large plant showed that 42% of the reboiled towers used some form of constant energy control. While such control is necessary in some situations, it’s unlikely that nearly half the towers in a typical processing plant require it.

Check your control configurations for constant energy control in either the condenser or reboiler. Odds are you can change many of them to reduce energy use.

Are you effectively using stripping steam?

Steam stripping often is chosen for light-ends removal or recovery. It reduces the partial pressure of other components in a system, allowing them to vaporize. The heat of vaporization comes from the process fluid, not the steam. So, frequently far less steam is required compared to using a reboiler. This saves energy — but only if stripping is effective.

Tray towers. The relatively low steam rate can pose difficult equipment-design problems, especially for trays. Many stripping trays provide nearly 0% efficiency because their design doesn’t suit the service requirements. Even under the best conditions stripping tray efficiency rarely rises above 35%. Still the difference between 0% and 35% is huge.

A common design miscue is not accounting for the large changes in vapor rate that often occur across the stripping section. For trays to work effectively, they must have a minimum pressure drop to prevent liquid bypass and to provide good vapor distribution. Typically, a sieve tray needs a minimum of 0.05-psi pressure drop. When vapor rate varies, the number of sieve holes required to get the minimum pressure drop varies. If the vapor rate changes a lot, different trays may have to have significantly different hole layouts.

Table 1 summarizes stripping vapor loads for a service with eight real trays at three different efficiencies. In all cases the product (overhead) yield is constant. Zero percent efficiency shows the steam required for the stripping using only a steam flash.

Table 1. Ignoring vapor rates can take a severe toll on stripping efficiency and thus energy use.

Table 1. Ignoring vapor rates can take a severe toll on stripping efficiency and thus energy use. (Click to enlarge.)

Raising this plant from essentially 0% efficiency on the stripping trays to 25% percent saved 81,500 pounds per day of steam, which translates to $200,000 per year.

Improving efficiency demands accounting for the vapor profile through the trays. Figure 5 plots the vapor rate across the eight trays (assuming a constant 25% efficiency per tray). Initially, the vapor rate changes rapidly across each tray then levels out.

Figure 5. A single tray design generally can’t effectively handle wide differences in vapor rate.

Figure 5. A single tray design generally can’t effectively handle wide differences in vapor rate.

The rate on Tray 8 is more than 1.5 times the rate on Tray 2. Using a tray design suitable for the maximum rate means that the design for the bottom tray is nearly completely ineffective. Because the bottom tray doesn’t do any stripping, the vapor rate remains the same and Tray 2 doesn’t work either. This cascades up the stripping section and none of the trays work well. Average efficiency is close to zero.

Stripping sections with rapid changes in vapor rate must use multiple tray designs. Our example calls for four designs with different hole area for Tray 1, Trays 2 and 3, Trays 4 and 5, and Trays 6 to 8. This results in 25% efficiency instead of 0% efficiency and reduces energy costs by 75%.

Packed towers. The same problems occur with high-efficiency packed stripping sections. Here the major problem is steam distribution. Natural vapor distribution across a packed bed depends upon pressure drop: the higher the pressure drop, the better the distribution. However, if you design for low vapor rates, you have low pressure drops. The low pressure drop results in poor vapor flow distribution to the packed bed. You lose the benefit that you might get from a highly efficient stripping section.

For a packed stripping section not performing as well as required, check if you need better vapor distribution to the bed. If so, consider a steam sparger. Such a unit installed below an 8-ft.-diameter packed section improved effective packing performance from nearly 0% efficiency (zero stages) to two stages of separation.

How active are your trays?

Normally we assume tray flexibility ranges of 2:1 to 3:1 for sieve trays, and 3:1 and higher for valve trays. Many people think bubble cap trays have operable ranges of up to 10:1. These flexibility ranges generally do hold for single-pass trays with normal-to-high pressure drops, that is, a minimum of 0.05 psi to 0.06 psi per tray on a 2-ft. tray spacing. However, multiple-pass trays and those with low-pressure-drop designs have much lower vapor-handling flexibility. The clearest way to explain this is to look at multiple-pass trays.

Once bubbling starts on a specific spot on a tray deck, the fluid on the deck has a lower density than still liquid with the same depth next to it. So, the next increment of vapor flow tends to go through the tray where vapor bubbling has already started. If the vapor load is high enough, the entire tray surface is active. In contrast, at low vapor loads vapor tends to channel through one area rather than spreading out across the tray. Multiple flow passes for the liquid make the problem more severe. Table 2 shows standard values for minimum area of the tray that must be active to prevent severe bypassing [6].

Table 2. A greater number of flow paths poses greater demands for tray active area.

Table 2. A greater number of flow paths poses greater demands for tray active area. (Click to enlarge). Source: Ref. 6.

Efficiency suffers if not enough liquid contacts the vapor. Small efficiency penalties might have been tolerable when energy was cheap but not now. So, even a modest gain in efficiency may justify some minor modifications to tray decks to improve activity. Consider blanking strips, blocking valves down, and even replacement tray sections to increase efficiency.

Have you checked basic heat integration?

Never presume that a plant’s current heat-integration configuration is logical, let alone optimal. Consider the quirky scheme used in a solvent recovery plant (Figure 6).

Figure 6. The scheme boasts an unusual element and the opportunity for significant improvement.

Figure 6. The scheme boasts an unusual element and the opportunity for significant improvement.

The overhead is heat integrated after the first cooling-water exchanger, which is really unusual and has never been adequately explained. Also, while the overhead is heat integrated, neither the hot vapor sidedraw nor the tower bottoms streams have heat recovery.

The fix (Figure 7) was to reconfigure the overhead system and add heat recovery.

Figure 7. Investment in reconfiguration and added heat recovery was paid back in a few months.

Figure 7. Investment in reconfiguration and added heat recovery was paid back in a few months.

The sequence of overhead exchangers was reversed to heat-integrate before going to cooling water. One existing hot-oil exchanger was converted to feed preheat versus the vapor product draw. Finally, a new heat-recovery exchanger was added to the bottoms stream (10% of feed flow). One hot-oil exchanger was left on the feed stream (E3) to make the unit easier to start-up. Although the unit has a feed rate of only 360 gal/hr the heat duty was reduced by 3 million btu/hr. The changes involved a capital expenditure of less than $100,000 and yielded a savings of $280,000 per year.

Many other plants can take advantage of simple heat-integration steps between tower feed and tower bottoms. Don’t overlook these relatively easy ways to save money.

Simple steps can save energy

Regretfully, energy savings alone, despite today’s high energy prices, rarely justify major investment at plants. However, you can improve the energy efficiency of distillation columns without spending a lot of money. The small steps outlined here don’t involve much capital or process risk. This makes them more likely to get done.


Andrew Sloley is principal engineer for VECO U.S.A. in Bellingham, Wash. He also writes the Plant InSites column in Chemical Processing. E-mail him at ASloley@ putman.net.


References:

  1. Tolliver, T.L. and L.C. McCune, “Finding the optimum temperature control trays for distillation columns,” InTech, 27 (9), p. 75 (Sept. 1980).
  2. Sloley, A.W. and G. Martin, “Process modeling for control system design and analysis,” presented at IASTED Conference on Modelling, Simulation, and Control in the Process Industry, Ottawa (May 1994).
  3. Shinskey, F.G., “Distillation control for productivity and energy conservation,” McGraw-Hill, New York (1977).
  4. Kister, H.Z. and I.D. Doig, “When would floating pressure strategy save energy?,” Chem. Eng. Prog., 77 (9), p. 55 (Sept. 1981).
  5. Sloley, A.W., “Sidestep side-draw control surprises,” Chem. Proc., 67 (7), p. 33 (July 2004).
  6. Sloley, A.W. and B. Fleming, “Successfully downsize trayed columns,” Chem. Eng. Prog., 90 (3), p. 39 (Mar. 1994).
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