T&S Chapter 5, SS&L Chapter 20 Terry A. Ring University of Utah Plant-wide Control T&S Chapter 5, SS&L Chapter 20 Terry A. Ring University of Utah

PROCESS DESIGN STAGES AND TOOLS 10 - Design and Control DESIGN AND ANALYSIS II - (c) Daniel R. Lewin

Instrumentation and Control Objectives

Control Valves Safety Aspects Types of Control Valves Fail Open, Fail Closed Types of Control Valves Linear Non-linear Restricted open range (gain scheduled controller)

Pressure Control Loop Pressure Control Valve – fail closed Is fail closed best? Pressure Transducer – PT Pieziolelectric, diaphragm displacement Pressure Indicator Controller – PIC PID, maybe gain scheduling if valve is non-linear Gain scheduling for pressure relief above a given pressure. Alarm High 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 ?? Ratio 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 dead- heading pump VFD can be costly Missing Alarms on Pressure Indicator (PI)

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 Pressure 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 self- regulating 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) Equations = 4 NT+6, DoF=5 (# Control Valves) Additional Control Valve on Feed is Possible which is used for Production Rate Control

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 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 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 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

Plantwide Control Design Determine all manipulated variable. Number of manipulated variables MAY be equal to the number of control valves. 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. Decide how to control product quality. 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). Determine and stabilize unstable units. Design inventory control loops. Develop component balance control loops (control of makeup streams of reactants, makeup gaseous streams to maintain pressure, liquid makeup to control level, etc.). 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). Heat integration saves cost but can make plantwide control more difficult because integrating loops also “spread” disturbances. Use cascade control to reduce the effect of disturbances. 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. Using steady state simulations, determine the ranges of acceptable disturbances, which can be taken care of while maintaining the production goals. 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. Use dynamic simulation to find the range of disturbances, which can be compensated for using the designed control system. Use dynamic model to determine the range of stability of the closed-loop system when different model parameters change. 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). Ensure safe and environmentally sound operation. 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 Constraint Analysis Levels are not important in Optimization Constraint Analysis Optimization Itself

Example of self-optimizing structure Recycle process: J = V (min 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: 3 Constraints: Maximize reactor volume (residence time and thus conversion) Product spec: xB > 0.98 Nm = 5 N0y = 2 Levels DoF=Nss = 5 - 2 = 3

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 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 Capital cost goes down when A is less. This is caused by delta T being larger for Q to remain the same. 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.