Plant-wide Control
T&S Chapter 5, SS&L Chapter 20
Terry A. Ring
University of Utah
PROCESS DESIGN STAGES AND TOOLS
- 10Design and Control
DESIGN AND ANALYSIS II - (c) Daniel R. Lewin
2
Instrumentation and Control Objectives
Control Valves
• Safety Aspects
• Fail Open, Fail Closed
• Types of Control Valves
• Linear
• Non-linear
• Restricted open range (gain scheduled controller)
Pressure Control Loop
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Pressure Control Valve – fail closed
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Pressure Transducer – PT
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Is fail closed best?
Pieziolelectric, diaphragm displacement
Pressure Indicator Controller – PIC
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PID, maybe gain scheduling if valve is non-linear
Gain scheduling for pressure relief above a given pressure.
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Alarm High
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Alarm Low
Level Control Loop
pump
• Level Control Valve – fail open
• Level Transducer – LT
• DP cell, guided radar, conductivity
• Level Indicator Controller – LIC
• PID, maybe gain scheduling if valve is non-linear
• Alarm High
• Alarm Low
• Pump may dead head
• Interlock or gain schedule valve to never go to zero
Flow Control
• Centrifugal Pump
• Flow Control Valve – Fail ??
• Flow Transducer
• Orifice, venture, v-cone, many others
• Flow Indicator Controller – FIC
• PID, maybe gain scheduling if valve is non-linear
• Note, check valve to prevent back flow and by-pass loop to prevent deadheading pump
• VFD can be costly
• Missing Alarms on Pressure Indicator (PI)
Flow
Ratio
Ratio Control
• Ratio Control Valve – FFV – fail Closed
• May have two control valves – one for each stream
• 2 Flow Transducers
• Orifice, venture, v-cone, many others
• Flow Ratio Controller
• PID with calculator for total flow(FT=FT1+FT2) to be held constant
• CV1 SP=Ratio*SPLC
• CV2 SP=(1-Ratio)*SPLC
Add
Pressure Relief
Temperature Control w Utility
• Utility Control Valve – TV – Fail ??
• Temperature Sensing Element (TE) and Temperature Transducer (TT) or Temperature
Indicator (TI)
• Temperature Indicator Controller
• PID, maybe gain scheduling if valve is non-linear
Vaporizer Control
• Flow Control on Liquid and Level Control in HX
• Pressure relief valve???
CSTR Contol System
• Utility Temperature Control
• TT and FT feed signals to Flow control –
Cascade Control
• Valve Fail???
• Pressure Control
• PID, gain scheduling = on/off
• Valve Fail ???
• Level Control
• PID, gain scheduling if non-linear Valve
• Valve Fail??
• Feed Flow Rate Control
• This needs ratio control!!
• Valves Fail ??
PID Tuning – Lamda Tuning
• P only P=1
• Step Test – CV +/- 10%
• Observe increase/decrease in PV
• Τd and τ (63.2%) and Kp=%PV/%CV(output)
• Kc=(Kp)-1*(Tr)/(λ+Td)
• Tr=τ, λ=0.5 to 4 (typically 3) times Max(Td, Tr)
• Cascade Control inner loop 5x to 10x faster than outer loop
Lamda Tuning
Results
Max(Td, Tr) = 4 s
These are all simple 1 input, 1 output
controllers with alarms for safety
• Plant Wide Control is Different
• Active Control System Objectives
• Meet production rate and product quality targets
• Keep system in Safe Operating Range for equipment, catalysts and materials of
construction
• Minimize Energy Utilized
• Minimize Process Variability
• Meet environmental regulations
• Air quality
• Water discharge quality
• Profit Optimizer
• Provide inputs for Active Control System
• Separate Safety Control System
• separate sensors, valves, controllers, power system
• Identify Unsafe conditions
• Take Action to Safely Bring the Plant/System Down to a Safe Point of
Operation or Shut Down
Procedure for Plant Wide Control
I. Top-down analysis to identify degrees of freedom
and primary controlled variables (look for selfregulating variables)
II. Bottom-up analysis to determine secondary
controlled variables and structure of control system
e.g. pairings (CVs with Sensors).
Degrees of Freedom (DoF) Analysis
• DoF (= Control Valves)
• DoF=Nvariables-Nexternally_defined-Nequations
• Practice DoF on Smaller Problems
Heat Exchanger Network
Flow Rate Set
Floating Flow Rates for C1 & C2
Heat Exchanger
Network Control
Hot Stream
• DoF = Nv-NDef-Neq=15-4-9=2 (# Control Valves)
• Nv=15 (9Ts,3 flows, 3 Hx Qs)
• NDef=4 (F1,To, Θo, Θ1)
• Equations for Each Exchanger Neq=3x3=9
• Q1=F1Cp1(To-T1)
• Q1=F3Cp3(Θ4- Θ3)
• Q1= U1A1ΔTLM
• With By-pass, DoF = Nv-NDef-Neq=17-4-10=3
• Θ3 =(1-φ) Θ0+ Θ’3
•
Neq=10, Nv=17(& φ,Θ’3)
Streams 2 and 3 Cold
Distillation Control with
Total Condenser
• Control Valves ????
• Degrees of Freedom Analysis DoF= Nv-NDef-Neq
• Variables, Nv= 4 NT+13, NDef=2 (Feed Flow and Composition)
Additional Control Valve
on Feed is Possible which
is used for Production
Rate Control
• Equations = 4 NT+6,
DoF=5 (# Control Valves)
Distillation Control
• What measurements?
• What should be the pairings of
measurements to control valves?
• PD controlled with CV QC
• LR controlled by CV B
• LD controlled by
• CV L if so what does CV D control??
• CV D if so what does CV L control??
• What does CV QR and CV B control?
Distillation Control
• Material Balance Control
• 4 Control Schemes
• Inferred Composition Analysis = Temp. of Stage
• Top CA or Bottom CA
With Composition
Analysis
• LV control
• DV control
• L/F V control
• D/F V control
Key Concerns for Plant-wide Control System
• Establish Control Objectives
• Safe Operation – Process Within its Constraints
• Meets Environmental Constraints
• Control Production Rate
• Feed Flow Controls
• Product (on-demand) Flow Controls
• Control Product Quality
• Determine Degrees of Freedom for Control System
• Position Control Valves
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Establish Energy Management System
Recycle Loop Flows Fixed – watch out for snowball effects
Vapor and Liquid Inventories Fixed
Improve Dynamic Controllability
Acrylic Process
AB
• Objectives
• Production Rate of B high &
constant
• Conversion in reactor highest
possible
• Constant Composition for B
Where to set the production rate? Should it be at
the inlet or outlet?
• Very important!
• Determines structure of remaining inventory (level) control system
• Set production rate at (dynamic) bottleneck
• Link between Top-down and Bottom-up parts of Plant-wide Control
System
On-demand Product
or Feed Flow Control
System
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•
•
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20.13
FC on Product B
Feed of A controlled as needed by reactor
20.14
FC on Feed A Flow
Level Controls Product B Flow Rate
• Reactor Temp is controlled with Cascade
Controller
• To meet Composition Requirements
• Note:
• Trim steam heater for Feed A for accurate
temperature control at reactor
Note: Flash Vessel will have a spray head for
cold B to condense all possible Vapors
Vapor Typically is fed above Liquid level
Demister on outlet of Vessel
Plantwide Control Design
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Determine all manipulated variable. Number of manipulated variables MAY be equal to the number of control valves.
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Determine how production rate is set: (a) upstream process; (b) downstream process; (c) free to vary. Determine the best way to control the production rate: (a) valve selection; (b)
setpoint on temperature, recycle, etc. can sometimes be used to control production rate.
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Decide how to control product quality.
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The closed loop system for product quality control should have adequately small time constant, time delays, and sufficiently large gains.
Consider interaction between different loops and decide if multivariable control is needed.
Make sure that flows are not too small to achieve the objective (i.e. it is difficult to control condenser level using small flow rate of distillate in the column with high internal flows).
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Determine and stabilize unstable units.
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Design inventory control loops.
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Develop component balance control loops (control of makeup streams of reactants, makeup gaseous streams to maintain pressure, liquid makeup to control level, etc.).
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Be careful with recycle loops: They introduce feedback (positive or negative), which makes it difficult to analyze the consequences of the particular control design on the overall
performance. For instance, it was found that unless one flow somewhere in the recycle loop is fixed, the recycle flow might grow to very high rate when disturbance occurs or when
throughput is increased (snowball effect).
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Heat integration saves cost but can make plantwide control more difficult because integrating loops also “spread” disturbances.
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Use cascade control to reduce the effect of disturbances.
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Using steady state simulations to study controllability of the designed system in face of disturbances. For example, perturb the feed flow rate or composition and study how flows
and other process variables change to compensate for the disturbances. If small disturbance requires large changes in flow rate or other manipulated variables to compensate for
its effect, the control system must be redesigned.
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Using steady state simulations, determine the ranges of acceptable disturbances, which can be taken care of while maintaining the production goals.
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Use dynamic simulation to determine dynamic response of the designed system. Redesign as needed. For instance, redesign to eliminate inverse response, if at all possible.
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Use dynamic simulation to find the range of disturbances, which can be compensated for using the designed control system.
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Use dynamic model to determine the range of stability of the closed-loop system when different model parameters change.
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Use the remaining degrees of freedom (manipulated variables) for steady state optimization (maximize the profit, etc.) and/or to improve controllability of the plant (disturbance
rejection, flexibility in changing operation point or product mix, etc).
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Ensure safe and environmentally sound operation.
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Startup shutdown, safety and emergency handling control.
Vinyl Chloride Process
• Reactor 1 FeCl3 solid Catalyst
• Conversion is >90%
• C2H4 + Cl2  C2H2Cl2
• Fired Heater for Pyrolysis Reaction
• Conversion is 60%
• C2H2Cl2  C2H3Cl + HCl
• Product
•
Vinyl Chloride Process
Vinyl Chloride Process w Control System
Approach to Optimizer
• DoF Analysis for Optimizer
• Levels are not important in Optimization
• Constraint Analysis
• Optimization Itself
Example of self-optimizing structure
Recycle process: J = V (min. boil-up => min. energy)
5
4
1
2
Given feed rate F0 and
column pressure:
Nm = 5
N0y = 2 Levels
DoF=Nss = 5 - 2 = 3
3
Constraints:
Maximize reactor volume (residence
time and thus conversion)
Product spec: xB > 0.98
Recycle process:
Selection of controlled variables
• Step 3.1 J=V (minimize energy with given feed)
• Step 3.1 DOFs for optimization: Nss = 3
• Step 3.3 Most important disturbance: Feedrate F0
• Step 3.4 Optimization: Constraints on max. Mr and
xB always active
• Step 3.5 1 DOF left, candidate controlled variables:
F, D, L, xD, ...
• Step 3.6 Loss with constant setpoint:
Recycle process: Loss with constant setpoint, cs
Large loss with c = F
Negligible loss with c = L/F
Proposed control structure
for case with J = V (minimize energy)
Active constraint
Mr = Mrmax
Active constraint
xB = xBmin= 98%
2 Active Constraints for Optimization
Reactor Full
Min Composition Required
Recycle process:
Selection of controlled variables
• Step 3.1 J=V (minimize energy with given feed)
• Step 3.1 DOFs for optimization: Nss = 3
• Step 3.3 Most important disturbance: Feedrate F0
• Step 3.4 Optimization: Constraints on max. Mr and
xB always active
• Step 3.5 1 DOF left, candidate controlled variables:
F, D, L, xD, ...
• Step 3.6 Loss with constant setpoint: Good: xD, L/F.
Poor: F, D, L
Given feedrate, the production rate set at inlet of the column
Reconfiguration is required when the bottleneck (max. vapor rate in
column) is reached (rate set at the outlet of the column)
MAX
HW 5 Heat Integration
Optimization of minimum temperature approach in a heat exchanger:
• Set up a ΔT for E1 =(TS1-TS4)
• VP = (1-t)(S-C)-i(Cequipment)
• S=0,
·
• C= Utilities for E2 and E3
• Qi = mi Cpi ΔTi
• Ci =Σ Qi * Price of Utilityi
• Cequipment
• Cost of HX
• E1+E2+E3
• Cost α f(HX Area)
• Area = Q/(U ΔTLM)
Capital and Operating Cost Optimization
ΔTthres
Problem 2 - Pinch Analysis
• To exchange heat between four streams
with ΔTmin= 20°C, the HEN in the figure
below is proposed. Determine if the
network has the minimum utility
requirements. If not, design a network
with the minimum utility requirements.
• As an alternative design a network with
the minimum number of heat
exchangers.