To view the recording of Deminar #7, click on the above picture. If you want to just view the slides click on Deminar #7 - PID Control of True Integrating Processes

Time is money. If you can get to optimum setpoints faster during fed-batch operations and for startup and product transitions of continuous operations, the increase in production revenue can be significant. For continuous operations there may also be an appreciable decrease in the processing, recycle, and waste treatment costs of off-spec material.

For cascade control, the speed of the secondary PID setpoint response largely determines the ability of the primary PID to get to its setpoint quickly and reject disturbances in the primary loop. A slow secondary PID setpoint response may require detuning of the primary PID to prevent interactions between the secondary and primary loops.

In Deminar #7 we explored how we could use PID structure options, setpoint feedforward, and bang-bang control to improve the setpoint response for integrating (e.g. batch) processes. The concepts are also applicable to the continuous process startup and transitions. The demos showed a big reduction in rise time (time to reach setpoint) by the use of "PID on Error" instead of "I on Error, PD on PV." The benefit of the additional bump from derivative action on error is rather marginal for the small rate setting used. In other words most of the speedup in the setpoint response could be achieved by "PI on Error, D on PV" unless there is a large secondary lag and hence a large rate time setting. The use of setpoint feedforward helped reduce overshoot, rise time, and settling time by about 25%. For deadtime dominant self-regulating processes, the improvement would have been more impressive. The most dramatic improvement occurred for full throttle bang-bang control. With some adjustment of logic and resting value as noted on slide 6, the bang-bang logic can also be effectively used for self-regulating processes. You can try out setpoint feedforward and bang-bang control on the virtual plant website starting August 20.

If you are intrested in reducing batch cycle time and startup time, check out Deminar #7 at 10:00 am CDT on Aug 11.

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Higher value added products are generally produced by batch operations. Often these products are sold out and extra batches translate to significant increases in revenue. Prime opportunities are specialty chemicals and drugs, especially new biopharmaceuticals where optimization took a backseat to time to market in the initial plant and automation system design.

I looked over my past experience with Monsanto, Solutia, and Emerson and have come up with myriad of methods to reduce batch cycle time. I have divided them up into opportunities to help feedback loops to get to setpoint faster that are important for Fed-Batch operations and for startups and transitions of continuous operations (Part 1) and opportunities to shorten phases and holds that are important for Pure-Batch operations (Part 2). These techniques like all new configurations and strategies should be thoroughly tested by simulation and closely monitored and adjusted for safe and efficient operation. Today's blog is a preview of Deminar #7 on July 14.

Fed-Batch Opportunities

1) PID on Error Structure - This structure maximizes the kick of the controller output for a setpoint change. The overdrive (driving of output past resting point) is essential for getting slow loops, such as temperature, to the optimum setpoint as fast as possible.

2) SP Track PV - With this control option the setpoint is changed to its optimum with the controller in automatic providing the kick from the PID structure (1). For batch operations this option is commonly used. For continuous operations with few setpoint changes (no grade transitions) and extremely long run time (e.g. years), the setpoint is held at its last value. However, even here loops with slow reset action (large reset times), such as level, the use of the SP Track PV option can prove useful when putting these loops back in service after maintenance.

3) SP Feedforward - For low controller gains (controller gain less than inverse of process gain), a setpoint feedforward is useful. The setpoint feedforward gain is the inverse of the dimensionless process gain minus the controller gain on a percent basis. If the setpoint and controller output are in engineering units the feedforward gain must be adjusted accordingly. The feedforward action is the process action, which is the opposite of the control action, taking into account valve action. In other words for a reverse control action, the feedforward action is direct provided the valve action is inc-open or the analog output block, I/P, or positioner reverses the signal for a inc-close (fail open) valve.

4) Output Lead-Lag - A lead-lag on the controller output or in the digital positioner can kick the signal though the valve deadband and sticktion, get past split range points, and make faster transitions from heating to cooling and vice versa. When combined with the enhanced PID algorithm described in Deminar #1, the lead-lag can potentially provide faster control when online analyzers are used for closed loop control of the integrating response associated with batch operations.

5) Deadtime Compensation - The simple addition of a delay block with the deadtime set equal to the total loop deadtime to the external reset signal for the positive feedback implementation of integral action (see Deminar #3). The controller reset time can be significantly reduced and the controller gain increased if the delay block deadtime is equal or slightly less than the process deadtime (see Advanced Application Note 3 entry March 25, 2009 on this website).

6) Full Throttle Batch - The controller output is put at its output limit to maximize the rate of approach to setpoint. When the projected PV equals the setpoint less a bias, the controller output is repositioned to the final resting value captured from the last batch. The output is held at the resting value for one deadtime. For more details, check out the Control magazine article "Full Throttle Batch and Startup Response."

7) Feed Maximization - Valve position control, Model Predictive Control (see Advanced Application Notes 1 and 2 entries March 25, 2009 on this website), or override control is used to maximize feeds to limits of operating constraints (e.g. maximum vent, overhead condenser, or jacket valve position with sufficient sensitivity). Alternatively, the limiting valve can be set wide open and the feeds throttled for temperature or pressure control. For pressure control of gaseous reactants, this strategy can be quite effective. For temperature control of liquid reactants, the user needs to confirm that the inverse response from the addition of cold reactants to an exothermic reactor and the lag from the concentration response does not cause temperature control problems. All of these methods require tuning and may not be particularly adept at dealing with fast disturbances unless some feedforward is added. Fortunately the prevalent disturbance is a feed concentration change that is often slow enough due to raw material storage volume to be corrected by feedback control.

8) Profile Control - If you have a have batch measurement that should increase to a maximum at the batch end point (e.g. maximum reaction temperature or product concentration), the slope of the batch profile of this measurement can be maximized to reduce batch cycle time. For application examples checkout "Direct Temperature Rate of Change Control Improves Reactor Yield" in a Funny Thing Happened on the Way to the Control Room E-book April 3, 2009 entry on this website and the Control magazine article "Unlocking the Secret Profiles of Batch Reactors"

This blog was kind of fun to write "With A Little Help from My Friends" (beer and music). By the way, the album "Joe Cocker" was apparently only produced on vinyl and 8-track tape. I fondly remember riding in my roadster with the top down listening to "Saint James Infirmary Blues" on the way to Monsanto's New Orleans plant. There are great songs on this album that carried me through startups that never made it onto a CD.

At my recent presentation to the ISA Saint Louis section meeting on "pH measurement", I had several people around my age say how nice it was to see me still involved in advancing our profession. Maybe it was the beer and the top ten lists but just maybe it was also that I represent a generation of expertise rapidly disappearing via retirement. The ability to still learn and share keeps me going but I realize time is running out so I intend to take this blog to the next level by coupling it with a web lab series to provide an interactive self-learning experience for exploring process control improvements (PCI). I intend to start the web lab series on April 7. Recordings of the PCI topics and demos along with instructions on using the associated labs will be viewable anywhere anytime.

In the meantime, we need to finish up this series so let's see what we can do as automation engineers to minimize loop errors.

The first thing is to make sure the measurement is fast and precise enough. So far as loop performance is concerned, precision is more critical than accuracy. The bias or offset in a measurement and control valve position can be corrected by feedback control. The offset in valve position is eliminated by the process loop. Similarly, the offset in a process loop is eliminated when the loop is in cascade or remote cascade modes. For loops operated in the auto mode, operations have often compensated for the measurement offset by tweaking the set point. This is not to say that measurement accuracy is not important.

Improving the loop's speed of response often comes down to keeping sensors clean (e.g. electrodes and thermowells), minimizing signal damping and filtering, selecting sensor locations that eliminate transportation and stagnation delays, using boosters for big valves, maximizing positioner sensitivity, minimizing deadband, and maximizing the controller gain (last week's blog).

Control loops have a difficult time dealing with the poor precision experienced as excessive stick-slip and backlash (deadband) in control valves and insufficient resolution, repeatability, and sensitivity in measurements (older measurements technologies, such as floats and rotameters can also exhibit stick-slip and backlash). Fortunately, an increase in A/D input card bits have greatly improved the resolution of transmitted signals so that sensitivity and repeatability is the remaining focus. This is unfortunately not the case for variable frequency drives manufacturers whose standard input cards have only 8 bits. A resolution limit is more degrading than a sensitivity limit. For example for a 1% resolution and 1% sensitivity and a change in the true process variable of 1.5%, the changes in measurement would be 1% and 1.5%, respectively. The deadband setting in wireless transmitters is really a sensitivity setting. When the change in a wireless transmitter measurement exceeds this setting regardless of the direction, the full change in the process variable is communicated.

Pages 12 through 15 of EffectsLoopTuning&Dynamics-KPI.pdf show the relative effect of measurement accuracy and resolution on variability. For control valves, process variability is introduced when excessive slip-stick and deadband causes an appreciable limit cycle in loops that have single and two or more integrators, respectively (pages 19 and 20).

The total loop deadtime can be approximated as the sum of all the delays and small lags in the loop whether they are in the DCS, valve, process, or measurement. For flow, pressure, level, and inline temperature and pH loops, most of the loop deadtime comes from the automation system. If you consider that the remaining loops that have significant process deadtime, such as vessel or column temperature, have seriously detuned controllers that create an effective deadtime per Advanced Application Note 5, you realize you have the opportunity as a process control engineer to make big reductions in loop deadtime that are also low cost and quick compared to changing process piping or equipment to reduce transportation or mixing delays.

Fast disturbance originate from manual operation, on-off actions, sequences, or setpoint changes. The elimination of operator actions, on-off control (e.g. level switches), and the use of set point rate of change limits and fed-batch rather than pure sequential batch, can dramatically slow down disturbances since throttling control by intention is smooth. If we keep all loops in their highest design mode and limit on-off valves to SIS actions and isolation, we could eliminate step disturbances. Page 22 shows how slowing down the disturbance dramatically reduces the peak and integrated errors for an integrating process. Not shown here is the fact that slowing down disturbances can also reduce interaction between loops. This phenomenon explains why it is difficult to get pharmaceutical companies excited about doing a better job of bioreactor control after reaching setpoint. The disturbances from cells are incredibly slow (e.g. process time constants of days).

Maybe we should not slow down disturbances because all of our control texts are based on step disturbances. Slowing down the upsets relegates us to improving the set point response in the startup of a continuous process or for changes in phase in a batch process. Whoops, even here we could use strategies such as "Full Throttle Batch and Setpoint Response" to eliminate most of the job of the loop.

There are always opportunities to make us more appreciated even when we are not improving loops. Since spouses were at the ISA Section Saint Louis Meeting, I interjected the following list. The spouses laughed although a second opinion was suggested for some of the items. See what your spouse or significant other thinks.

Top Ten Reasons Why an Automation Engineer Makes a Great Spouse or at Least a Wedding Gift

(10) Reliable from day one
(9) Always on the job
(8) Low maintenance - minimal grooming, clothing, and entertainment costs
(7) Many programmable features
(6) Stable
(5) Short settling time
(4) No frills or extraneous features
(3) Relies on feedback
(2) Good response to commands and amenable to real time optimization
(1) Readily tuned

Startups, grade transitions, and abnormal conditions are the most difficult, operator intensive, hazardous, and inefficient periods of plant operation. Operators often believe these conditions require operator evaluation and action. The conditions are thought to be too special and the response too situation dependent to automate. The operators are right in saying these periods of operation require the best in operator expertise. However, case histories show that the power of the PID can be used to automate the best operator responses and build on them to provide faster, safer, and more efficient plant operation during these difficult process conditions. For some specific examples dealing with compressors and reactors check out the two chapters "Wally and the Beave Automate Reactor Startups" and "Wally and the Beave Return to Automate Another Reactor Startup" in my E-book on this website A Funny Thing Happened on the Way to the Control Room. For impressive examples for chemical, mining, and pulp and paper operations, check out the Control Talk columns "Show Me the Money - Part 1" (November 2009) and "Show Me the Money - Part 2" (December 2009) in Control magazine.

An extensive interview of the operators and process engineers is necessary to capture the best responses for a preliminary functional description of the control system. There are often a lot of surprises hidden by the diversity of actions that are inevitable from human responses. Free will implies these decisions are basically unpredictable. The operator actions consistent with first principles and process knowledge offer a good starting point but not the final strategy. During the commissioning of the control system, the plant response must be carefully observed and the best operator actions verified and improved by the use of the many options built into a PID loop to deal with rampant problems as the plant goes from zero to full rate, or vice versa. For example, output tracking, dynamic reset limiting, set point ramping, PID structure, gain scheduling, adaptive control, and override control can be used to deal with the problems at low rates such as noisy or inaccurate flow signals, excessive valve stick-slip near the closed position, larger transportation delays, and unrepresentative measurements. One of the common solutions is to head start (initialize) the controller output via output tracking to the best valve position for startup, transition, or abnormal situation. The initial position can be a "Full Throttle" position for fastest set point response. When the set point approaches the set point, the controller output can be momentarily set to a resting value based on experience or average position captured from a representative operating point from the last run. For fast loops such as flow and pressure, the resting value can be used as the "head start". One of the common mistakes is for process engineers to get carried away with trying to sequence the PID controller output too much or hold the controller output in the track mode for too long. For shutdown, the output must normally be held but otherwise the PID controller should be returned to automatic as soon as possible to deal with disturbances, unknown process effects, raw material variability, and nonlinearities. The process is not known or measured well enough to sequence flows without feedback control. It is particularly important to return pressure loops to automatic as fast as possible. Smart techniques for startup, transitions, and abnormal situations that take full advantage of the flexibility of the PID controller have been the source of the most impressive benefits in process control improvement. In general, these were also "quick hits" in that they were implemented in a matter of a couple of weeks by just configuration changes and controller tuning.