1

Part 3: Regulatory («stabilizing») control

Inventory (level) control structure– Location of throughput manipulator

– Consistency and radiating rule

Structure of regulatory control layer (PID)– Selection of controlled variables (CV2) and pairing with manipulated

variables (MV2)

– Main rule: Control drifting variables and "pair close"

Summary: Sigurd’s rules for plantwide control

 

Wed AM

2

Outline

• Skogestad procedure for control structure designI Top Down

• Step S1: Define operational objective (cost) and constraints

• Step S2: Identify degrees of freedom and optimize operation for disturbances

• Step S3: Implementation of optimal operation

– What to control ? (primary CV’s)

– Control Active constraints + self-optimizing variables

• Step S4: Where set the production rate? (Inventory control)

II Bottom Up • Step S5: Regulatory control: What more to control (secondary CV’s) ?

• Step S6: Supervisory control

• Step S7: Real-time optimization

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Step S4. Where set production rate?

• Very important decision that determines the structure of the rest of the inventory control system!

• May also have important economic implications

• Link between Top-down (economics) and Bottom-up (stabilization) parts– Inventory control is the most important part of stabilizing control

• “Throughput manipulator” (TPM) = MV for controlling throughput (production rate, network flow)

• Where set the production rate = Where locate the TPM? –Traditionally: At the feed

–For maximum production (with small backoff): At the bottleneck

4

TPM (Throughput manipulator)

• Definition 1. TPM = MV used to control throughput (CV)

• Definition 2 (Aske and Skogestad, 2009). A TPM is a degree of freedom that affects the network flow and which is not directly or indirectly determined by the control of the individual units, including their inventory control.

– The TPM is the “gas pedal” of the process

– Value of TPM: Usually set by the operator (manual control)• Operators are skeptical of giving up this MV to the control system (e.g. MPC)

– The TPM is usually a flow (or closely related to a flow), e.g. main feedrate, but not always.• It can be a setpoint to another control loop

• Reactor temperature can be a TPM, because it changes the reactor conversion,

• Pressure of gas product can be a TPM, because it changes the gas product flowrate

• Fuel cell: Load (current I) is usually the TPM

– Usually, only one TPM for a plant, but there can be more if there are• parallel units or splits into alternative processing routes

• multiple feeds that do not need to be set in a fixed ratio

– If we consider only part of the plant then the TPM may be outside our control. • throughput is then a disturbance

5

TPM and link to inventory control

• Liquid inventory: Level control (LC)– Sometimes pressure control (PC)

• Gas inventory: Pressure control (PC)

• Component inventory: Composition control (CC, XC, AC)

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Production rate set at inlet :Inventory control in direction of flow*

* Required to get “local-consistent” inventory control

TPM

7

Production rate set at outlet:Inventory control opposite flow*

TPM

* Required to get “local-consistent” inventory control

8

Production rate set inside process*

TPM

* Required to get “local-consistent” inventory control

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General: “Need radiating inventory control around TPM” (Georgakis)

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Consistency of inventory control

• Consistency (required property):

An inventory control system is said to be consistent if the steady-state mass balances (total, components and phases) are satisfied for any part of the process, including the individual units and the overall plant.

11

CONSISTENT?QUIZ 1

12

Local-consistency rule

Rule 1. Local-consistency requires that

1. The total inventory (mass) of any part of the process must be locally regulated by its in- or outflows, which implies that at least one flow in or out of any part of the process must depend on the inventory inside that part of the process.

2. For systems with several components, the inventory of each component of any part of the process must be locally regulated by its in- or outflows or by chemical reaction.

3. For systems with several phases, the inventory of each phase of any part of the process must be locally regulated by its in- or outflows or by phase transition.

Proof: Mass balancesNote: Without the word “local” one gets the more general consistency rule

13

CONSISTENT?QUIZ 1

14

Local concistency requirement -> “Radiation rule “(Georgakis)

15

Flow split: May give extra DOF

TPM

TPM

Split: Extra DOF (FC) Flash: No extra DOF

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Consistent?Local-consistent?

Note: Local-consistent ismore strict as it impliesconsistent

QUIZ 2

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Closed system: Must leave one inventory uncontrolled

QUIZ 3

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OK? (Where is production set?

NO. Two TPMs (consider overall liquid balance).Solution: Interchange LC and FC on last tank

QUIZ 4

TPM1

TPM2

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Example: Separator controlAlternative TPM locations

Compressor could be replaced by valve if p1>pG

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Alt.1 Alt.2

Alt.3 Alt.4

Similar to original but NOT CONSISTENT (PC not direction of flow)

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Example: Solid oxide fuel cell

Solid oxide electrolyte TPM = current I [A] = disturbanceO2- e-

CH4

H2O

Air

PC

PC

TC

CC

xCH4,s

(in ratio with CH4 feedto reduce C and CO formation)

CH4 + H2O = CO + 3H2

CO + H2O = CO2 + H2

2H2 + O2- → 2H2O + 2e-

(excess O2)

O2 + 4e- → 2O2-

Ts = 1070 K(active constraint)

xH2??

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LOCATION OF SENSORS

• Location flow sensor (before or after valve or pump): Does not matter from consistency point of view– Locate to get best flow measurement

• Before pump: Beware of cavitation

• After pump: Beware of noisy measurement

• Location of pressure sensor (before or after valve, pump or compressor): Important from consistency point of view

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Where should we place TPM?• TPM = MV used to control throughput

• Traditionally: TPM = Main feed valve (or pump/compressor)– Gives inventory control “in direction of flow”

Consider moving TPM if:

1. There is an important CV that could otherwise not be well controlled– Dynamic reasons

– Special case: Max. production important: Locate TPM at process bottleneck* !• TPM can then be used to achieve tight bottleneck control (= achieve max. production)

• Economics: Max. production is very favorable in “sellers marked”

2. If placing it at the feed may yield infeasible operation (“overfeeding”)– If “snowballing” is a problem (accumulation in recycle loop), then consider placing TPM

inside recycle loop

BUT: Avoid a variable that may (optimally) saturate as TPM (unless it is at bottleneck)– Reason: To keep controlling CV=throughput, we would need to reconfigure (move TPM)**

**Sigurd’s general pairing rule (to reduce need for reassigning loops): “Pair MV that may (optimally) saturate with CV that may be given up”

*Bottleneck: Last constraint to become active as we increase throughput -> TPM must be used for bottleneck control

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QUIZ. Distillation. OK?

LV-configuration

TPM

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DB-configuration OK???

TPM

26

cc

*But DB-configuration is not recommended!

DB-configuration: Level control NOT consistent by itself, but can still work* if we add one (or preferably two) composition/temperature loops

TPM

cc

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QUIZ

* Keep p ¸ pmin

** Keep valve fully open

*

**

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29

QUIZ

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LOCATION OF SENSORS

• Location flow sensor (before or after valve or pump): Does not matter from consistency point of view– Locate to get best flow measurement

• Before pump: Beware of cavitation

• After pump: Beware of noise

• Etc.

• Location of pressure sensor (before or after valve, pump or compressor): Important from consistency point of view

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Often optimal: Locate TPM at bottleneck!

• "A bottleneck is a unit where we reach a constraints which makes further increase in throughput infeasible"

• If feed is cheap and available: Located TPM at bottleneck (dynamic reasons)

• If the flow for some time is not at its maximum through the bottleneck, then this loss can never be recovered.

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Single-loop alternatives for bottleneck controlBottleneck.Want max flow here

Alt.1. Feedrate controls bottleneck flow (“long loop”…):

FCFmax

Alt. 2: Feedrate controls lost task (another “long loop”…):Fmax

Alt. 3: Reconfigure all upstream inventory loops:Fmax

Traditional: Manual control of feed rate

TPM

TPM

TPM

TPM

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May move TPM to inside recycle loop to avoid snowballing

Example: Eastman esterification processAlcohol recycle

Alcohol + water + extractive agent (e)

Ester product

Reach max mass transferrate: R increases sharply(“snowballing”)

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First improvement: Located closer to bottleneck

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Final improvement: Located “at” bottleneck + TPM is inside “snowballing” loop

Follows Luyben’s law 1 to avoid snowballing(modified): “Avoid having all streams in a recycle system on inventory control”

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Where should we place TPM?• TPM = MV used to control throughput

• Traditionally: TPM = Main feed valve (or pump/compressor)– Operators like it. Gives inventory control “in direction of flow”

Consider moving TPM if:

1. There is an important CV that could otherwise not be well controlled• Dynamic reasons

– Special case: Max. production important: Locate TPM at process bottleneck* !• Because max. production is very favorable in “sellers marked”

• TPM can then be used to achieve tight bottleneck control (= achieve max. flow)

2. If placing it at the feed may yield infeasible operation (“overfeeding”)– If “snowballing” is a problem (accumulation in recycle loop), then consider placing TPM

inside recycle loop

BUT: Avoid a variable that may (optimally) saturate as TPM (unless it is at bottleneck)– Reason: To keep controlling CV=throughput, we would need to reconfigure (move TPM)**

**Sigurd’s general pairing rule (to reduce need for reassigning loops): “Pair MV that may (optimally) saturate with CV that may be given up”

*Bottleneck: Last constraint to become active as we increase throughput -> TPM must be used for bottleneck control

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A purely top-down approach: Start by controlling all active constaints at max. throughput (may give moving TPM)

Economic Plantwide Control Over a Wide Throughput Range: A Systematic Design Procedure

Rahul Jagtap, Nitin Kaisthaand Sigurd Skogestad

 

Step 0: Obtain active constraint regions for the wide throughput range

Step 1: Pair loops for tight control of economic CVs at maximum throughput – Most important point economically

– Most active constraints

Step 2: Design the inventory (regulatory) control system

Step 3: Design loops for ‘taking up’ additional economic CV control at lower throughputs along with appropriate throughput manipulation strategy

Moving TPM

Warning: May get complicated, but good economically because of tight control of active constraints

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Figure 5. Plantwide control structure for maximum throughput operation of recycle process (Case Study I)

A + B CB + C D

Strippe

r

Column

A, B RecycleB

A

C

D

FC

PC

LCFC

TC

X

LC

FC

TC

TC

RC

LC

FC

FC

CCD

CCB

CCB

PC

LC

V1MAX V2

MAX

0.98%

0.02%

LVLRxrMAX

TPM for < max throughput

165 °C

xBRxr Opt HS3

TRxrOpt HS2

LS1

SP1

SP3

LC2

SP2

SP1 > SP2 > SP3

LC1

LC3

LS2

TRxrMAX

39

Conclusion TPM (production rate manipulator)

• Think carefully about where to place it!

• Difficult to undo later

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Outline

• Skogestad procedure for control structure designI Top Down

• Step S1: Define operational objective (cost) and constraints

• Step S2: Identify degrees of freedom and optimize operation for disturbances

• Step S3: Implementation of optimal operation

– What to control ? (primary CV’s) (self-optimizing control)

• Step S4: Where set the production rate? (Inventory control)

II Bottom Up • Step S5: Regulatory control: What more to control (secondary CV’s) ?

– Distillation example

• Step S6: Supervisory control

• Step S7: Real-time optimization

41

Structure of regulatory control layer (PID)

Structure of regulatory control layer (PID)– Selection of controlled variables (CV2) and pairing with manipulated

variables (MV2)

– Main rule: Control drifting variables and "pair close"

Summary: Sigurd’s rules for plantwide control

 

42

II. Bottom-up

• Determine secondary controlled variables (CV2) and structure (configuration) of control system (pairing, CV2-MV2)

• A good control configuration is insensitive to parameter changes

Regulatory layer

43

Step 5. Regulatory control layer

• Purpose: “Stabilize” the plant using a simple control configuration (usually: local SISO PID controllers + simple cascades)

• Enable manual operation (by operators)

• Main structural decisions:• What more should we control?

(secondary cv’s, CV2, use of extra measurements)

• Pairing with manipulated variables (mv’s u2)

CV1

CV2 = ?

Regulatory layer

44

Objectives regulatory control layer1. Allow for manual operation

2. Simple decentralized (local) PID controllers that can be tuned on-line

3. Take care of “fast” control

4. Track setpoint changes from the layer above

5. Local disturbance rejection

6. Stabilization (mathematical sense)

7. Avoid “drift” (due to disturbances) so system stays in “linear region”– “stabilization” (practical sense)

8. Allow for “slow” control in layer above (supervisory control)

9. Make control problem easy as seen from layer above

10. Use “easy” and “robust” measurements (pressure, temperature)

11. Simple structure

12. Contribute to overall economic objective (“indirect” control)

13. Should not need to be changed during operation

Regulatory layer

45

Stabilizing control: Use inputs MV2=u2 to control “drifting” variables CV2

GKCV2s u2

CV2

CV1

Key decision: Choice of CV2 (controlled variable)

Also important: Choice of MV2=u2 (“pairing”)

Primary CV

Secondary CV(control fordynamic reasons)

Process control: Typical «drifting» variables (CV2) are•Liquid inventories (level)•Vapor inventories (pressure)•Some temperatures (reactor, distillation column profile)

Regulatory layer

46

Degrees of freedom unchanged

• No degrees of freedom lost as setpoints y2s replace inputs u2 as new degrees of freedom for control of y1

GKCV2s u2

CV2

CV1

Original DOFNew DOF

Cascade control:

Regulatory layer

47

Example: Exothermic reactor (unstable)

• u = cooling flow (q)

• CV1 = composition (c)

• CV2 = temperature (T)

u

TCCV2=T

CV2s

CCCV1=c

CV1s

feed

product

cooling

LC

Ls=max

Active constraints (economics):Product composition c + level (max)

Regulatory layer

48

49

Degrees of freedom for optimization (usually steady-state DOFs), MVopt = CV1sDegrees of freedom for supervisory control, MV1=CV2s + unused valvesPhysical degrees of freedom for stabilizing control, MV2 = valves (dynamic process inputs)

Optimizer (RTO)

PROCESS

Supervisory controller (MPC)

Regulatory controller (PID) H2

H

y1=CV1

CV2s

y

ny

d

Stabilized process

CV1s

y2=CV2

Physicalinputs (valves)

Optimally constant valves

Always active constraints

u2u1

50

Details Step 5 (Structure regulatory control layer)

(a) What to control (CV2)?

1. Control CV2 that “stabilizes the plant” (stops drifting)

2. Select CV2 which is easy to control (favorable dynamics)

• In addition, active constraints (CV1) that require tight control (small backoff) may be assigned to the regulatory layer.*

*Comment: usually not necessary with tight control of unconstrained CVs because optimum is usually relatively flat

Regulatory layer

51

A. “Mathematical stabilization” (e.g. reactor): • Unstable mode is “quickly” detected (state observability) in the

measurement (y2) and is easily affected (state controllability) by the input (u2).

• Tool for selecting input/output: Pole vectors – y2: Want large element in output pole vector: Instability easily

detected relative to noise – u2: Want large element in input pole vector: Small input usage

required for stabilization

B. “Practical extended stabilization” (avoid “drift” due to disturbance sensitivity):

• Intuitive: y2 located close to important disturbance• Maximum gain rule: Controllable range for y2 is large compared to

sum of optimal variation and measurement+control error • More exact tool: Partial control analysis

“Control CV2s that stabilizes the plant (stops drifting)”

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“Control CV2 that stabilizes the plant (stops drifting)” In practice, control:

1. Levels (inventory liquid)

2. Pressures (inventory gas/vapor) (note: some pressures may be left floating)

3. Inventories of components that may accumulate/deplete inside plant• E.g., amine/water depletes in recycle loop in CO2 capture plant

• E.g., butanol accumulates in methanol-water distillation column

• E.g., inert N2 accumulates in ammonia reactor recycle

4. Reactor temperature

5. Distillation column profile (one temperature inside column)

• Stripper/absorber profile does not generally need to be stabilized

Regulatory layer

53

(b) Identify pairings =Identify MVs (u2) to control CV2, taking into account:

• Want “local consistency” for the inventory control– Implies radiating inventory control around given flow

• Avoid selecting as MVs in the regulatory layer, variables that may optimally saturate at steady-state (active constraint on some region), because this would require either– reassigning the regulatory loop (complication penalty), or – requiring back-off for the MV variable (economic penalty)

• Want tight control of important active constraints (to avoid back-off)

• General rule: ”pair close” (see next slide)

Details Step 5b….Regulatory layer

54

Step 5b…. Main rule: “Pair close”

The response (from input to output) should be fast, large and in one direction. Avoid dead time and inverse responses!

Regulatory layer

55

Sigurd’s pairing rule: “Pair MV that may (optimally) saturate with CV that may be given up”

• Main reason: Minimizes need for reassigning loops

• Important: Always feasible (and optimal) to give up a CV when optimal MV saturation occurs. – Proof (DOF analysis): When one MV disappears (saturates), then we have one less optimal CV.

• Failing to follow this rule: Need some “fix” when MV saturates to remain optimal, like – reconfiguration (logic) – backoff (loss of optimality)

• BUT: Rule may be in conflict with other criteria– Dynamics (“pair close” rule)– Interactions (“avoid negative steady-state RGA” rule)– If conflict: Use reconfiguration (logic) or go for multivariable constraint control (MPC which may provide “built-in” logic)

Regulatory layer

LV

TCTs

. loop

LV

TCTs

LV

TCTs

TCTC TS

(a) Normal: Control T using V (b) If V may saturate: Use L

56

Why simplified configurations?Why control layers?Why not one “big” multivariable controller?• Fundamental: Save on modelling effort

• Other: – easy to understand

– easy to tune and retune

– insensitive to model uncertainty

– possible to design for failure tolerance

– fewer links

– reduced computation load

Regulatory layer

57

Hierarchical/cascade control: Time scale separation

• With a “reasonable” time scale separation between the layers(typically by a factor 5 or more in terms of closed-loop response time)

we have the following advantages:

1. The stability and performance of the lower (faster) layer (involving y2) is not much influenced by the presence of the upper (slow) layers (involving y1)

Reason: The frequency of the “disturbance” from the upper layer is well inside the bandwidth of the lower layers

2. With the lower (faster) layer in place, the stability and performance of the upper (slower) layers do not depend much on the specific controller settings used in the lower layers

Reason: The lower layers only effect frequencies outside the bandwidth of the upper layers

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XC

TC

FC

ys

y

Ls

Ts

L

T

z

XC

Cascade control

distillation

With flow loop +T-loop in top

59

QUIZ: What are the benefits of adding a flow controller (inner cascade)?

q z

qs

1. Counteracts nonlinearity in valve, f(z)• With fast flow control we can assume q = qs

2. Eliminates effect of disturbances in p1 and p2

Extra measurement y2 = q

60

Summary: Sigurd’s plantwide control rules

Rules for CV-selection:

1. Control active constraints

– Purity constraint on expensive product is always active (overpurification gives loss):  

2. Unconstrained degrees of freedom (if any): Control “self-optimizing” variables (c).

 

•The ideal variable is the gradient of J with respect to the inputs (Ju = dJ/du), which always should be zero, independent of disturbances d, but this variable is rarely available

– Exception (is available!): Parallel systems (stream split, multiple feed streams/manifold) with given throughput (or given total gas flow, etc.)

– Should have equal marginal costs Jiu = dJi/du, so Ju = J1u - J2u, etc.

– Heat exchanger splits: equal Jächke temperatures, JT1 = (T1 – Th1)^2/(T1-T0)

•In practice, one prefers to control single variables, c=Hy (where y are all available measurements and H is a selection matrix), which are easy to measure and control, and which have the following properties:

• Optimal value for c is almost constant (independent of disturbances): Want small magnitude of dcopt(d)/dd.

• Variable c is sensitive to changes in input: Want large magnitude of gain=dc/du (this is to reduce effect of measurement error and noise).

– If the economic loss with single variables is too large, then one may use measurement combinations, c=Hy (where H is a “full” matrix).

 

3. Unconstrained degrees of freedom: NEVER try to control a variable that reaches max or min at the optimum (in particular, never control J)

– Surprisingly, this is a very common mistake, even (especially?) with control experts

 

 

Rules for inventory control:

1. Use Radiation rule (PC, LC, FC ++)

2. Avoid having all flows in a recycle system on inventory control (this is a restatement of Luyben’s rule of “fixing a flow inside a recycle system” to avoid snowballing)

– A special case is a closed system

 

Rules for pairing:

1. General: “Pair close” (large gain and small effective time delay)

2. CV1: Sigurd’s pairing rule: “Pair MV that may (optimally) saturate with CV that may be given up”

3. CV2 (stabilizing loop): Avoid MV that may saturate

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Part 1 (3h): Plantwide control

Introduction to plantwide control (what should we really control?) Part 1.1 Introduction.

– Objective: Put controllers on flow sheet (make P&ID)– Two main objectives for control: Longer-term economics (CV1) and shorter-term stability (CV2)– Regulatory (basic) and supervisory (advanced) control layer

Part 1.2 Optimal operation (economics)– Active constraints– Selection of economic controlled variables (CV1). Self-optimizing variables.

Part 1.3 -Inventory (level) control structure– Location of throughput manipulator– Consistency and radiating rule

Part 1.4 Structure of regulatory control layer (PID)– Selection of controlled variables (CV2) and pairing with manipulated variables (MV2) – Main rule: Control drifting variables and "pair close"

Summary: Sigurd’s rules for plantwide control 

62

Plantwide control. Main references

• The following paper summarizes the procedure: – S. Skogestad, ``Control structure design for complete chemical plants'',

Computers and Chemical Engineering, 28 (1-2), 219-234 (2004).

• There are many approaches to plantwide control as discussed in the following review paper: – T. Larsson and S. Skogestad, ``Plantwide control: A review and a new

design procedure'' Modeling, Identification and Control, 21, 209-240 (2000).

• The following paper updates the procedure: – S. Skogestad, ``Economic plantwide control’’, Book chapter in V.

Kariwala and V.P. Rangaiah (Eds), Plant-Wide Control: Recent Developments and Applications”, Wiley (2012).

• More information:

All papers available at: http://www.nt.ntnu.no/users/skoge/

http://www.nt.ntnu.no/users/skoge/plantwide

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PC

TC