Dynamic Modeling Using Aspen Hysys® For Oil And Gas - Course Number Eb1105.06.04

Dynamic Modeling using Aspen HYSYS® for Oil and Gas AspenTech Customer Education Training Manual: Workbook Course Number EB11 05.06.04

----~_

...

_._.~-~=-=_._~----

r '- ~-

star*awardlff':C" hall of falTJe

Dynamic Modeling using Aspen HYSYS for Oil and Gas

Contents

Contents Lesson

Getting Started in Steady State Transitioning from Steady State to Dynamics Pressure Flow Theory Dynamic Details Expanding the Model Compressor TEG Dehydration Tower The Event Scheduler, Cause and Effect Matrix, and Spreadsheet Upstream Options - Hydraulics Appendix A - Basic Control Theory

©2007 AspenTech. All Rights Reserved.

Aspen Technology, Inc.

Contents

Dynamic Modeling using Aspen HYSYS for Oil and Gas

( ( (

(

( (

(

, \

Aspen Technology, Inc.

ii

©2007 AspenTech. All Rights Reserved.

Getting Started in Steady State

(

Getting Started in Steady State

© 2007 AspenTech - All Rights reserved. EB1105.06.04 01_GettingStartedlnSteadyState.doc

1

2

2

Getting Started in Steady State

Getting Started in Steady State

3

Workshop (

(

(

The Getting Started in Steady State module introduces you to some of the basic concepts necessary for creating simulations in Aspen HYSYS. Some ofthe things you will learn from this module are: •

Methods for moving through different environments

•

Selecting property packages and components

•

Navigating the new property views of Aspen HYSYS

•

Adding streams

•

Adding operations

("

(

You will use Aspen HYSYS to define streams and unit operations to develop a Flowsheet for an offshore platform.

Learning Objectives After you have completed this section, you will be able to: •

Define a Fluid Package (property package, components)

•

Add streams

•

Add operations

For those who are familiar with Aspen HYSYS, the steady state simulation can be built with the information listed on the following PFD. Step-by-step instructions follow the PFD.

3

4

Getting Started in Steady State

(

(

4

c Getting Started in Steady State

c

c (

(

(

5

Building a Steady State Simulation The Simulation Basis Manager Aspen HYSYS uses the concept of the Fluid Package to contain all necessary information for performing flash and physical property calculations. This approach allows you to defme all information (such as property package, components, hypothetical components, interaction parameters, reactions, tabular data) inside a single entity. There are four key advantages to this approach: •

All associated information is defined in a single location, allowing for easy creation and modification of the information.

•

Fluid Packages can be stored as a completely defined entity for use in any simulation.

•

Component lists can be stored separately from the fluid packages as completely defined entities for use in any simulation.

•

Multiple Fluid Packages can be used in the same simulation. However, they are all defmed inside the common Basis Manager.

(

The Simulation Basis Manager is a property view that allows you to create and manipulate multiple fluid packages or component lists in the simulation. The opening tab of the Simulation Basis Manager allows for the creation of component lists which are independent of- but can be associated with - the individual fluid packages in the case.

5

6

Getting Started in Steady State

The Components tab of the Simulation Basis Manager allows you to manage the component list(s) used in your case. There are a number of buttons available: •

View. Allows you to access the property view for the selected Component List.

•

Add. Allows you to create a Component List. Note: Component Lists can also be added using the Fluid Package property view.

•

Delete. Removes the selected Component List from the Simulation.

•

Copy. Makes a copy of the selected Component List.

•

Import. Allows you to import a pre-defined Component List from disk. Component Lists have the file extension .crnI.

(

•

Export. Allows you to export the selected Component List to disk. The exported Component List can be retrieved into another case by using the Import function.

(

•

(

(

Refresh. Allows you to reload and update all the pure component property data from the database. If you have a case created with an older version of Aspen HYSYS, you can update the pure component database to the latest version by using the Refresh function.

( ( !

\

6

c

(~

(~ .~.

( Getting Started in Steady State

(

7

c (

The Fluid Pkgs tab allows you access to the fluid packages/flowsheet associations list as well as the fluid package definition. As with older versions, Aspen HYSYS allows the user to use multiple fluid packages within a single simulation by associating the fluid packages with various flowsheets and linking the flowsheets together. However, beginning with Aspen HYSYS 3.0, the user no longer requires the use offlowsheets to employ multiple fluid packages within a single simulation. The user can now use the Stream Cutter operation to incorporate multiple fluid packages into a single flowsheet.

( (

(

( (

(

C

Basis Environment icon

•

View. This is only active when a Fluid Package exists in the case. It allows you to access the property view for the selected Fluid Package.

•

Add. Allows you to create and install a Fluid Package into the simulation.

•

Delete. Removes the selected Fluid Package from the simulation.

•

Copy. Makes a copy ofthe selected Fluid Package. Everything is identical in the copied version except the name. This is useful for modifying fluid packages.

•

Import. Allows you to import a pre-defined Fluid Package from disk. Fluid Packages have the file extension .:tpk.

•

Export. Allows you to export the selected Fluid Package to a disk. The exported Fluid Package can be retrieved into another case by using the Import function.

You can use the Ctrl+B hot key to re-enter the Simulation Basis Manager from any point in the simulation or click the Basis Environment icon from the tool bar. (In the Simulation Basis Manager, this is the Home View button.)

7

8

Getting Started in Steady State

Defining the Simulation Basis Add a Property Package New Case icon

1.

Start a new case by clicking the New Case icon.

2.

Go to the Fluid Pkgs tab and create a fluid package by clicking the Add button.

3.

Choose the Peng-Robinson Equation of State model.

4.

Change the Name from the default Basis-l to Oil&Gas Plant.

5.

Click the View button in the Component List Selection section of the Set Up tab. This will allow you to add components to the Component List that is now associated with the Oil&Gas Plant fluid package.

)

(

(

8

(

, (

c

(

Getting Started in Steady State

9

c

c

Add Components

(

( You can select components for your simulation using several different methods.

(

(

To Use This...

(

( (

Match Cell

Do This...

1.

You can add a range of components by highlighting the entire range and clicking the Add Pure button.

Select one of the three name formats by selecting the corresponding radio button:

• • •

Formula

Click the Match cell and enter the name of the component. As you start to type, the list will change to match what you have entered.

3.

After the desired component is highlighted, either:

/

Family Filter

Full Name/Synonym

2.

(

Component List

SimName

• •

Press Enter.

•

Double-click the component to add it to your simulation.

Click the Add Pure button.

1.

Using the scroll bar for the main component list, scroll through the list until you find the desired component.

2.

To add the component, either:

•

Press Enter.

• •

Click the Add Pure button. Double-click the component to add it to your simulation.

1.

Ensure the Match cell is empty, and click the View Filter button.

2.

Select the Use Filter checkbox to display the various family filters.

3.

Select the desired family (such as Hydrocarbons) from the list of Family Filters to display only that type of component.

4.

Use either of the two previous methods to select the desired component.

9

-------------- - - - - - - -

)

10

Getting Started in Steady State

)

1.

\ )

Select the library components N2, C02, CI, C2, C3, i-C4, n-C4, i-C5, n-C5, C6, and H20.

(

) )

(

2.

Select the Hypothetical branch in the Add Component tree to add hypothetical components to the Fluid Package.

3.

Click the Hypo Manager button to create a hypothetical group of components.

4.

Click the Add button and rename the group name Heavies.

"

( When you click the Quick Create a Hypo Component button, Aspen HYSYS adds a hydrocarbon class hypo by default. If you want to add a hypo from another class, click the Hypo Manager button and then in the view that displays, click the View Group button. This will open The Tabular Hypothetical Input, where you can add nonhydrocarbon class hypotheticals.

\.

(

(

10

c-----------------( c

Getting Started in Steady State

11

A hypothetical component can be used to model non-library components, defmed mixtures, undefined mixtures, or solids. You will be using a hypothetical component to model five components in the gas mixture.

( ( (

(

( (

5. The minimum information required for defining a hypo is the NBP or the density and MW.

Click the Add Hypo button five times in order to add five hypothetical components. Use the following table to add Name, NBP, and Liq Density information. Name

NBP

NBP200

93.3 °C (200 OF)

NBP280

138°C (280 OF)

NBP425

218°C (425 OF)

NBP630

332°C (630 OF)

NBP920

493°C (920 OF)

Liq Density

961 kg/m3 (60 Ib1ft3)

('

".

11

12

Getting Started in Steady State

6.

Select the Estimate Unknown Props button in order to estimate all of the unknown properties and close the window.

7.

On basis manager view, go to the Components tab, highlight Component List1 and click the View button. Click the Hypothetical branch in the Add Component tree and add the hypo components to the Selected Components list by selecting Heavies from the Available Hypo Groups and then clicking the Add Group button.

You can use the Sort List button to order the Component List.

You have now ftnished deftning the fluid package. You can view the Peng-Robinson binary coefftcients for your selected components by selecting the Binary Coefficients tab. 8.

12

Press the Enter Simulation Environment button.

c

f-

e' (

Getting Started in Steady State

f

13

Selecting a Unit Set

( (

When you build the simulation, you will:

C (

(

•

Select a unit set

•

Add streams

•

Attach utilities

In Aspen HYSYS, it is possible to change the unit set used to display the different variables.

C C (

1.

From the Tools menu, choose Preferences.

2.

Click the Variables tab.

3.

Select the SI unit set.

4.

Close this view to return to the simulation.

(

13

-

-

-

------~--

14

Getting Started in Steady State

Adding Streams In Aspen HYSYS, there are two types of streams, Material and Energy. Material streams have a composition and parameters such as temperature, pressure, and flowrates. They are used to represent process streams. Energy streams have only one parameter: heat flow. They are used to represent the duty supplied to or by a unit operation. There are a variety of ways to add Materials streams in Aspen HYSYS. To Use This...

Do This...

Menu Bar

Select Add Stream from the Flowsheet menu or press the F11 Hot Key. The Stream property view opens.

Workbook

Open the Workbook and go to the Material Streams tab. Type a stream name into the **New** cell.

Object Palette

From the Flowsheet menu, select Open Object Palette or press F4 to open the Object Palette. Double-click the stream icon.

In this case, you will add three streams to represent three different gas wells. Each stream will be added using a different method of installation.

14

Getting Started in Steady State

(

15

\.

(

C , \

(

Adding a Stream from the Menu Bar To add a stream using the Fll hot key: 1.

( (

Press FIt. The Stream property view displays. Ifthe Stream property view is not displayed, double-click the newly created stream (from the PFD) to bring up the property view.

2.

( (

To rename the stream, highlight the Stream Name cell and type in the new name. Rename the stream Alpha.

3.

Enter the following information and press Enter.

( \

C Temperature Pressure

6480 kPa (940 psia)

Molar Flow

10900 kgmol/hr (24000 Ibmole/hr)

Entering Stream Compositions There are two different pages for entering stream compositions: When using the...

Do This...

Conditions Page

Double-click the Molar Flow cell to enter mole fractions or Double-click the Mass Flow cell to enter mass fractions or Double-click the Std Ideal Liquid Volume Flow cell to enter volume fractions. The Input Composition for Stream view displays.

Composition Page

Click the Edit button. The Input Composition for Stream view displays.

4.

If the Input Composition for Stream view is not already open, double-click the Mass Flow cell.

5.

Click the Mole Fractions radio button in the Composition Basis group to change the basis from mass to mole fractions.

6.

Enter the following compositions.

15

16

Getting Started in Steady State

For this Component

Enter this mole fraction

Nitrogen

0.015

CO2

0.023

Methane

0.759

Ethane

0.066

) \

Aspen HYSYS automatically normalizes the compositions once you click the OK button. Ensure that you have entered the right composition values before clicking OK.

Propane

0.034

i-Butane

0.006

n-Butane

0.013

i-Pentane

0.004

If there are <empty> values either enter 0 or press the Normalize button.

n-Pentane

0.005

n-Hexane

0.006

water

0.038

NBP200

0.007

NBP280

0.014

NBP425

0.006

NBP630

0.003

NBP920

0.001

)

,

(

\

, \ J

Figure 9

(

(

7.

Click OK to close the Input Composition for Stream dialog box.

(

16

)

Getting Started in Steady State

17

Adding a Stream from the Workbook

c (

To open or display the Workbook, click the Workbook icon on the button bar.

(

1.

Enter the stream name Bravo in the **New**cell.

( (

2.

Supply the following information and compositions (double-click the Molar Flow cell to enter mole fractions).

Workbook icon

Nitrogen

0.005

CO2

0.010

Methane

0.719

Ethane

0.062

Propane

0.041

i-Butane

0.010

n-Butane

0.026

i-Pentane

0.010

n-Pentane

0.019

n-Hexane

0.010

water

0.040

NBP200

0.010

NBP280

0.021

NBP425

0.010

NBP63 0

0.005

NBP920

0.002

3.

Click OK to close the Input Composition for Stream view.

17

--------.---.--------------------------------

18

Getting Started in Steady State

Adding a Stream from the Object Palette 1.

Ifthe Object Palette is not open, press the F4 hot key to open it.

2.

Double-click the Material Stream icon on the Object Palette. The Stream property view displays.

3.

Change the name of the stream to Charlie and supply the following information and compositions:

Material Stream icon

Nitrogen

0.017

C02

0.026

Methane

0.730

Ethane

0.074

Propane

0.039

i-Butane

0.006

n-Butane

0.014

i-Pentane

0.004

n-Pentane

0.006

n-Hexane

0.007

water

0.043

NBP200

0.008

NBP280

0.016

NBP425

0.006

NBP630

0.003

NBP920

0.001

4.

18

Close the Stream property view.

c

- f~ (

(-----------------( (

( (

19

Saving Your Case You can use one of several different methods to save a case in Aspen HYSYS: •

From the File menu select Save to save your case with the same name.

•

From the File menu select Save As to save your case in a different location or with a different name.

•

Click the Save icon on the toolbar to save your case with the same name.

(

(

Getting Started in Steady State

Save your case as: Feed Stream.hsc.

(

(

(

Adding Operations As with streams, there are various ways to install Unit Operations in Aspen HYSYS. Some of these methods immediately open the operation property view and some do not. The initial flowsheet for the platform will be built in Steady State mode. Generally, all of the necessary information required to specify the unit operations for process design purposes is on the Design-Parameters tab. After all the necessary information has been entered, the status OK will display in the status bar and the color ofthe status bar will change to green. To use this...

Do this...

Menu Bar

Choose the Flowsheet menu bar, and then select Add Operation or Press the F12 hot key.

Workbook

From the Unit Ops page of the Workbook, click the Add Unit Op button.

Object Palette

Double-click the Operation icon and the Property view will display.

Choosing one ofthe three methods listed previously, add the following operations:

• • • • •

Pipe segment Mixer Separator Cooler Valve

19

20

Getting Started in Steady State

Add Pipe Segments

•••••

1.

Choose Alpha as the inlet streams, Q-Alpha as the energy stream and Alphal as the product stream.

2.

Go to rating tab sizing page and click Append Segment.

3.

Select pipe schedule 40 and click nominal diameter 609.6 (the last one), then click Specify button and dismiss the schedule window. Set length to be lOOOm.

4.

Go to Heat transfer page, Click Overall HTC, and fill in 25C for ambient temperature.

5.

Click Estimate HTC, check the 4 "Includes" check boxes. The segment should solve now.

Pipe segment icon

20

C f~

C' Getting Started in Steady State

21

c ((

( (

( (

(

(

( ( (

C 6.

Repeat the above steps to add PIPE-1 01 and PIPE-1 02 for Bravo and Charlie, respectively.

Add Valves Downstream of Pipe Segments The streams from the pipe segments are going through valves. The downstream pressure is 6480 kPa (940 psia). Valve icon

7.

Add a valve with the following information:

8.

Since pressure rises across the valve, increase pressure of Alpha to be 7000kPa to eliminate the problem.

9.

Add two more valves for the other 2 lines, but do not specify pressures.

21

---

----.-_.

--

( 22

Getting Started in Steady State

Add a Mixer 10. Choose Alpha2, Bravo2, and Charlie2 as the inlet streams and ToSep as the product stream. Mixer icon

11. Go to Parameter page and select Equalize All.

Add a Separator 12. Add a separator and enter the following information: Separator icon

Name

HPSep

Inlet

ToSep

Vapor Outlet

HPVap

Liquid Outlet

HPLiq

Note that the Separator unit operation is completely defined in Steady State without having to specify a pressure drop and a volume, but these are important parameters for dynamic simulation analysis.

22

~---~---------------~ - ~----~------~----

Getting Started in Steady State

(

23

(

Add a Cooler The vapor stream from the High Pressure Separator is cooled to 27°C.

(

c

13. Add a cooler and enter the following information:

Cooler icon

E-100·

Duty

J .g:oWe+()()7JkJi~

i=:.eed·'Tel1)p_e:r~tur~'1 p:tl)-q~~:T~'rnP~rat.u.re'.,

(

?'1.8.5,:"!-'C 1"ii,oo:!G.

Name

E-100

Inlet

HPVap

Outlet

HotVap

Energy

qcool

Add a Valve The liquid stream from the High Pressure Separator is let down through a valve. The downstream pressure is 390 psia. Valve icon

14. Add a valve with the following information:

The flowsheet should be completely solved. Save your case as EMlss.hsc.

23

24

24

Getting Started in Steady State

Transitioning from Steady State to Dynamics

Transitioning from Steady State to Dynamics

© 2007 AspenTech - All Rights reserved. EB1105.06.04 02_TransitioningFromSteadyState.doc

1

2

2

Transitioning from Steady State to Dynamics

~~~~ .. _~~---

. - _.

--~~--~~

Transitioning from Steady State to Dynamics

3

Workshop This module examines the process of changing a Steady State simulation into a Dynamic one. The process for doing this is not difficult, but it will become easier as you gain experience with dynamic simulations in Aspen HYSYS. Starting with the steady state simulation model that you prepared in Module I, you will add the necessary equipment information and Flowsheet specifications to permit dynamic simulation analyses.

Learning Objectives Once you have completed this module, you will be able to: •

Size equipment

•

Derme pressure flow specifications

•

Add strip charts and controllers

•

Run a simple dynamic simulation and observe the role of the various controllers

Prerequisites Before beginning this section you need to know how to: •

Add Streams

•

Add Unit Operations

•

Maneuver through the Aspen HYSYS interface

3

__

.-

-

4

4

Transitioning from Steady State to Dynamics

Er ( ( Transitioning from Steady State to Dynamics

5

(-

Design Dynamics " I

Aspen HYSYS Dynamics has been designed to permit a two-tiered approach to simulation. With numerous options to supply different levels of equipment design and performance information, Aspen HYSYS Dynamics provides modeling capabilities aimed at both process design and detailed design activity. For the design activity simulation, users typically enter basic design information and Aspen HYSYS Dynamics estimates reasonable defaults for the detailed equipment information. Typically these basic design parameters can be found on the Design tab of unit operations.

5

6

Transitioning from Steady State to Dynamics

The default equipment information estimated by Aspen HYSYS Dynamics is highlighted with a red color-coding. Generally, this detailed information can be found on the Rating tab ofthe Unit Operation property view.

In the next few modules, we will focus on design dynamics in order to illustrate the fundamental concepts underlying the use and configuration of Aspen HYSYS Dynamics. In later modules, we will expand the design dynamics model by incorporating detailed equipment and performance information and explore the detailed rating capabilities that Aspen HYSYS Dynamics provides. We will begin this module by loading the case that was saved in the previous module, EMlss.hsc, and prepare the model for dynamic simulation analyses.

6

( Transitioning from Steady State to Dynamics

(

7

(

(

Transitioning from Steady State to Dynamics

(

( (

(

If you switch to Dynamics from Steady State (click the Dynamics button), Aspen HYSYS will ask you if you would like to resolve the items which need attention before switching to dynamics. Answer No. You will notice that some pieces of equipment appear red.

(

Dynamics Mode icon

Aspen HYSYS Dynamics will indicate that each unit is missing some information necessary for dynamic simulation analyses. This is because the equipment is not sized. Equipment sizing is a very important step in dynamic modeling.

Flow in the plant occurs as a result of the pressureflow relationships between nodes.

If you have clicked the Dynamics Mode icon, click the Steady State Mode icon to return to Steady State, and activate the solver. If pipes do not solve after going back to steady state, you may go to their rating/heat transfer/hest loss page to delete the heat loss numbers there. Before a transition from Steady State to Dynamics occurs, the simulation Flowsheet should be set up so that a pressure drop exists across the plant. This pressure drop is necessary because the flow in Aspen HYSYS Dynamics is determined by the pressure drop throughout the plant. No pressure drop means no flow.

7

8

Transitioning from Steady State to Dynamics

Examine the following areas when setting up a simulation in Steady State and transitioning to Dynamics:

More information on these steps can be found in the Aspen HYSYS Dynamic Modeling manual, Section 1.5.2. Moving from SteadyState to Dynamics.

1.

Adding Unit Operations

2.

Equipment Sizing

3.

Adjusting Column Pressure

4.

Logical Operations

5.

Adding Control Operations

6.

Entering the Aspen HYSYS Dynamic Environment

7.

Adding Pressure-Flow Specifications

8.

Troubleshooting

We will now examine these eight areas as we convert our Steady State simulation into a Dynamic one.

Adding Unit Operations Identify material streams which are connected to two unit operations with no pressure flow relation and whose flow must be specified in Dynamic mode. Unit operations without a pressure flow relation include the Separator and the tray sections in Columns. You will need to add unit operations such as Valves, Heat Exchangers, or Pumps which define a pressure flow relation to these streams. Which stream in this case connects two unit operations with no pressure flow relation and will need an additional unit operation?

Equipment Sizing All unit operations in the simulation need to be sized using actual plant equipment or pre-defined sizing techniques. Vessels should be sized to accommodate actual plant flowrates and pressures while maintaining acceptable residence times.

Sizing the Valves For dynamic operation, it is necessary to size every valve in the simulation. Aspen HYSYS Dynamics automates valve sizing based on the: •

8

Valve Operating Characteristics group (Linear, Quick Opening, Equal Percentage, and User Table)

(-

------------------

(

Transitioning from Steady State to Dynamics

9

c_ ..

\,

(

•

Stream Conditions (Nonnal Valve Opening Position, Pressure Drop, Mass Flow Rate, Inlet Pressure, and Molecular Weight)

•

Sizing Method (Cv and Cg)

(

The Cv radio button is the default selected option. Aspen HYSYS Dynamics calculates the Cv that will allow the valve to pass 100% of the upstream Flowrate through the valve at the design valve opening position. This action is automatically done when you switch to dynamics if you did not click the Dynamics Mode icon before:

(

1.

Double-click the control valve VLV-I03.

2.

On the Rating tab of the Valve property view, select Universal Gas Sizing as the Valve Manufactures option and select Linear as the Valve Operating Characteristics option. Set the Valve Opening (%) in the Sizing Conditions group to 50%. Click the Size Valve button to complete the sizing. If you are using SI units, the screen should look like the following figure. If you are using field units, the numbers may be different, but the Cv value should be close to the one shown in the following screenshot.

(

'-..

""

-

In this instance, the Size Valve calculation determines that a Cv of 109 USGPM will pass 347 m3 /hr (52000 bblJd) when the valve is 50% open with a pressure drop of 3790kPa (550 psi). Similarly size the 3 valves at the downstream of the pipes. What Cv does Aspen HYSYS calculatefor VLV-IOO? Andfor VLV-IOl?

_

Andfor VLV-I02?

9

10

Transitioning from Steady State to Dynamics

Sizing the Separator Appropriate vessel sizing is important for dynamic simulation analyses. The vessel hold-up will affect the system's transient response during dynamic analyses as the user moves from one operating regime to the next. In addition, the vessel size affects the pressure calculations that are associated with this unit operation. On the Rating tab ofthe Separator, enter a Volume of127 m3 (4500 ft3).

If you do not know the dimensions of your process equipment, calculate a vessel size based on an approximate residence time.

•

10 minutes is typically a suitable residence time for liquid phase holdups.

•

2 minutes is typically a suitable residence time for vapour phase holdups.

Unlike separators, a number of unit operations have equipment volumes that are defaulted - for example, heat exchangers, heaters, and coolers. When adding these unit operations to your flowsheet, make sure that reasonable equipment volumes are specified.

10

c ( Transitioning from Steady State to Dynamics

(

11

(~

c

Sizing the Cooler

(

1.

On the Dynamics tab of the Cooler, enter a volume of 15 m3 (500 fe).

2.

Save your case as: EM2.hsc.

l ( (

Dynamic Specifications and Steady State Specifications: The Differences Dynamics Mode icon

Steady State icon

In Aspen HYSYS Dynamics, the simultaneous solution ofthe pressure-flow relationships within the Flowsheet requires the user to make a number of dynamic operating specifications. The possible pressure or flow type specifications for a Flowsheet include:

• • • • • •

Pressure specification on a material stream Flow specification on a material stream Fixed pressure drop across equipment PressurelFlow equation Resistance calculation (for valves) Conductance calculations (for process equipment)

11 _.

r

12

Transitioning from Steady State to Dynamics

Pressure flow specifications are made on the Specs page ofthe Dynamics tab of Unit Operations and Streams.

Aspen HYSYS Dynamics provides users with a great deal of flexibility in making pressure flow specifications to tackle challenging simulation problems. A number of different combinations of pressureflow specifications will lead to a consistent solution and solve the process Flowsheet.

Dynamic Specifications Boundary Streams

Insert a valve on all boundary streams (feed/product streams) within the Flowsheet that are not connected to conductance devices (such as heat exchangers, coolers, heaters).

Pressure Specifications

Place a pressure specification on all boundary streams (feed/product streams) within the Flowsheet.

Distillation Column

Distillation columns with condensers require an extra specification around the condenser. Make a flow specification for the reflux flow.

Valves

Use the pressure/flow relationship as the dynamic specification for a valve.

Kvalue

Use the overall K value as the dynamic specification for coolers, heaters, and heat exchangers and LNG exchangers.

Pressure Gradients

Be sure to account for pressure gradients throughout the Flowsheet and specify reasonable pressure drops/rises in the Flowsheet. Pressure differentials are the driving force for flow through the process Flowsheet.

Tray Sizing

Use the tray sizing utility to estimate the column geometry and pressure profile.

Mixers

Use the Equalize All option as the pressure specification for mixers.

Tees

Remove Use Splits as Dynamic Flow Specs on tees.

Rotating Equipment (Pumps, Compressors, Expanders)

Use Efficiency and either Head or Pressure Rise as dynamic specifications for rotating Equipment. Compressor and Pump Curves, if available, make excellent dynamic specifications.

Hold-ups

Be sure to properly size equipment with hold-ups.

Dynamic specifications can only be modified when the Integrator is stopped. Once the Integrator is started the value of the dynamic specification can be changed (its value displays in blue). but the choice of dynamic specification can not be changed.

12

E; (

C (

(.-------------------

0--: C ( (

Transitioning from Steady State to Dynamics

13

Making Pressure-Flow and Dynamic Specifications Analysis of the Process Flowsheet Based on the previous table, we can make the following observations: •

The boundary stream are Alpha, Bravo, Charlie, HotVap, and HPLiql.

•

HPLiql is connected to valve VLV-I03, so the addition ofa valve is not required.

•

HotVap is connected to a conductance device, E-100, so the addition ofa valve is not required.

•

The upstream has 3 pipes and 3 valves to control the flows, so no need for anymore.

Make the Appropriate Pressure-Flow Specifications All boundary streams in the Flowsheet must have a pressure specification. The boundary streams are: Alpha, Bravo, Charlie, HotVap, and HPLiq 1. 1.

On the Dynamics tab of the stream Alpha, select the Pressure Specification by checking the Active checkbox. Uncheck the flow spec if necessary.

13

14

Transitioning from Steady State to Dynamics

Make sure that only the pressure specification is active. If the flow specification is active, click the Active box to remove the specification.

Use Dynamic P/F spec color scheme

2.

Do the same for streams Bravo, Charlie, HotVap, and HPLiq1. Check all other streams and make sure that neither the pressure or flow specifications are active (Alpha, Bravo, and Charlie could have a flow specification active and Alpha2 could have a pressure spec). As a shortcut, instead of going to each stream's Dynamics tab, open each unit operation and go to the Worksheet-PF Specs tab to view the pressure flow specifications for each of the connected streams.

14

(.

c ( (

(-------------------c,

Transitioning from Steady State to Dynamics

15

k is the conductance-to-flow constant for the cooler. The value of k is calculated based on the current delta P, density, and Flowrate through the cooler.

(

\

(

3.

On the Dynamics-Specs tab for the cooler, click the Calculate k button.

4.

After the k value has been calculated, activate the Overall k box and deactivate the Overall Delta P box.

5.

Check that all the control valves have the Pressure Flow relation specification active.

l ( ( (

Having the "k" value as the active specification means that the pressure drops across that unit will change with the flow. This is more realistic than having a constant pressure.

r

Save your case as: EM3-Specs.hsc. The model is now ready to run in Dynamics. Click the Dynamic Mode icon and start the Integrator by clicking the Integrator Active icon.

Summary Pressure specifications have been made on all Boundary streams. No pressure or flow specifications have been made on the internal Flowsheet streams (ToSep, HP Vap, HP-Liq, Charlie#, Bravo#, Alpha#). 1.

Open the property view for anyone of these streams to verify that this is true.

Aspen HYSYS Dynamics has automatically selected resistance to flow specifications (Pressure Flow Relations) for Valves. Check if the sizing has been done handling ·multi-phase systems rigorously on the Rating-Options tab. 2.

Open the property view Dynamics-Specs tab to verify that this is correct.

3.

Conductance specifications have been made on process equipment (Cooler).

From the tool bar select Simulation I Equation Summary View (the solver must be in Dynamics mode). Click the Full Analysis button and then select to the Spec Eqns tab for a complete list of all specifications. From the PFD view, click the Color Scheme icon. Select the default schemes Dynamic P/F Specs. Colour Scheme icon

15

16

Transitioning from Steady State to Dynamics

Controllers Controllers can be added to the Flowsheet using the same methods as for other unit operations. The Pill Controller button on the palette represents this unit operation. Once the Controller has been added to the Flowsheet:

PID icon on Object Palette

1.

Make the necessary connections for the Process Variable Source and Output Target Object.

2.

Select the Minimum and Maximum values for the Process Variable. These values should bracket all possible PV values.

3.

Size the valve - controller range. This is not necessary if a valve was chosen as the Output Target Object.

4.

Select Controller Action, Reverse or Direct.

5.

Enter Controller Tuning Parameters.

6.

If desired, choose the mode of the controller: Off, Manual, or Automatic.

Add the Proposed Control Strategy 1.

Add a Flow Controller that will control the Mass Flow rate of Alpha1 to the Separator:

Action

Reverse

Range PV Minimum

o kg/hr (0 Ib/hr)

Range PV Maximum

365000 kg/hr (800000 Ib/hr)

Mode

Manual

OP

50%

Kc

0.25

TI

0.10

Td

16

---------

c;

(~

( ( ,-

(,

----------------------

Transitioning from Steady State to Dynamics

17

C (

2.

Insert a Controller Face Plate for monitoring by clicking the Face Plate icon on the property view. Figure 8

c

3.

Add another PID controller to control the mass flow of Bravo.

C

Action

Reverse

Range PV Minimum

okg/hr (0 Ib/hr)

Range PV Maximum

365000 kg/hr (800000 Ib/hr)

Mode

Manual

OP

50%

Kc

0.25

Ti

0.10

~

Td

17

18

Transitioning from Steady State to Dynamics

4.

Insert a face plate for FC- Bravo.

5.

Add another Pill controller to control the mass flow of Charlie.

Action

Reverse

Range PV Minimum

o kg/hr (0 Ib/hr)

Range PV Maximum

365000 kg/hr (800000 Ib/hr)

Mode

Manual

OP

50%

Kc

0.25

Ti

0.10

Td

6.

Insert a face plate for FC- Charlie.

7.

Add a Level Controller to Control the amount ofliquid in the HP Sep vessel.

Action

Direct

Range PV Minimum

0%

Range PV Maximum

100%

Mode

Manual

OP

50%

Kc

1.0

Ti

3.0

Td

18

Transitioning from Steady State to Dynamics

'--

19

8.

Insert a face plate for LC-HP Sep.

9.

Add another PID controller to control the temperature of the Hot Yap stream by manipulating the Feed Cooler Duty.

Action

Direct

Range PV Minimum

10°C (50 OF)

Range PV Maximum

65°C (150 OF)

Mode

Manual

OP

50%

Kc

E (

Ti

2

Td

~-

c ( ( (~

10. Insert a face plate for TC- HotVap. The control valve view will vary depending on the cooler model chosen (see the Dynamics-Spec tab for the cooler). If the Product Temp Spec radio button was selected, then there would be no need for the controller (the duty will float in order to maintain the desired temperature). If the Duty Fluid radio button was selected, then the controller would control the flow rate of the duty fluid. The user would supply a VA, duty fluid Cp, and the duty fluid inlet temperature.

(

( ( (

(

E::::-

19

--------------------------------------------------

20

Transitioning from Steady State to Dynamics

Cascade Control Cascade control is a common control technique that uses two controllers within one feedback loop. One controller is "nested" inside the other. This means that the two controllers are not independent, but linked together with the "primary" controller setting the SP for the "secondary" controller. Cascade control can improve the dynamic response and controllability of a process that has considerable dead time, or where the time response of the primary loop is very large.

Adding the Primary Controller The secondary controller for our cascade loop will be the FC-Alpha controller that already exists in our simulation. Complete the following steps to add and define the primary controller. 1.

Add another Pill controller to the simulation using the drag and drop method from the Object Palette. Name it PC-HPSep.

2.

Connect the pressure of the HP Sep vessel as the PV for the controller. Press and hold the key and move the cursor over onto the newly created controller. Note that three green boxes will display indicating the possible connection points. One of the boxes will have the label Sensor Input Object. (This label will display in a 'fly-by' text box.) Drag this box onto the logical connection point of the vessel icon (PV). A window will display asking you which variable you want; select Vessel Pressure and click OK.

3.

Connect the OP of this controller in the same way. Drag the green box over to the FC-Alpha Controller.

(

,

(

(

l,

20 \

J

Transitioning from Steady State to Dynamics

21

,--

4.

/-

Open the property view for the PC-HPSep controller and enter the following information on the Parameters page.

Action

Reverse

Range PV Minimum

5500 kPa (800 Psia)

Range PV Maximum

6900 kPa (1000 Psia)

Mode

Manual

OP

35%

Kc

3.0

Ti

2.0

The FC-Alpha controller should now have three connections. Note that this is the first time that more than two connections have been used on a controller and it is typically done in cascade control setups.

21

22

Transitioning from Steady State to Dynamics

Strip Charts While the Flowsheet is running dynamically, it is difficult to observe the simulation variables. Individual variables can be observed while in the PFD environment, or multiple variables can be seen on the workbook. All variables are updated constantly as the dynamic simulation is running. Using a strip chart allows the user to observe several variables in real time as the dynamic simulation runs.

Adding Strip Charts The Strip Chart provides a method for easily monitoring key process variables in a graphical environment. Strip Charts are installed individually using the Strip Charts page. Variables can only be connected to the Strip Chart though this page. Multiple Strip Charts are allowed, and each of these can have an unlimited number of variables charted. Perform the following steps to create a Strip Chart to monitor the flow rates of all the streams.

22

1.

Create a Strip Chart called Flows. Open the Tools menu and select Databook, or press the Ctrl+D hot key.

2.

Select the Variables tab and click the Insert button.

3.

Add the following variables:

• •

Bravo I - Mass Flow

•

Charlie I - Mass Flow

•

HotVap - Mass Flow

•

HPLiq1 - Mass Flow

Alphal - Mass Flow

4.

Select the Strip Charts tab.

5.

Click the Add button in the Available Strip Charts group.

6.

Aspen HYSYS installs the new Strip Chart and automatically names it. In this case, the name is DataLoggerl.

7.

Change the name to Flows in the Logger Narne field.

,..-.-

Transitioning from Steady State to Dynamics

Normally, all strip chart variables must be added to the Databook before they are available to the strip chart. However, there are two other approaches to adding variables to strip charts: •

•

Variables can be "drag and dropped" onto an active strip chart from the parent object. Each unit operation has a strip chart section on the Dynamics tab.

A user can "Quick Create" a strip chart using a predefined subset of variables.

8.

In the Individual Strip Chart Data Selection group box, check the Active checkbox for the mass flows ofthe previous streams.

9.

To view existing Strip Charts use one of the following methods: •

HigWight the Strip Chart name in the Available Strip Charts group and click the Strip Chart button in the View group.

•

Double-click the name of the Strip Chart.

23

10. Click the Setup button for the selected Strip Chart to configure the amount of data available on the Strip Chart and the sample interval for each data point. For all of the Strip Charts in these workshops, we will set the Logger Size at 3600 points and the Sample Interval at 20 s. 11. Create a second strip chart called HP Separator and insert the following variables: •

HPSep - Vessel Pressure

•

HPSep - Liquid Percent Level

•

HotVap - Temperature

To modify Strip Chart parameters, right-click the Strip Chart to open the Strip Chart Property view, from which you can edit the Strip Chart parameters. Save your case as: EM3-DynO.hsc.

23

24

Transitioning from Steady State to Dynamics

Face all the controller Plates and both Charts as follows: Figure 11

If you are already in Dynamics mode, start the Integrator. Otherwise, enter the Dynamics mode, and start the Integrator. Observe the Feed system Strip Chart. Let the Integrator run for a few minutes. Does the System achieve a Steady State Solution?

Place the level controller in Auto. Does the System achieve a Steady State Solution? -,;..~

Place the flow and temperatures in Auto. Finally, change the mode ofFC-Alpha from Auto to Cascade and Pressure controller in Auto. The simulation should be stable at the steady state values. Save your case as: EM3-Dynl.hsc.

24

Transitioning from Steady State to Dynamics

25

12. Reopen case EM1ss.hsc. 13. Run dynamic assistant, review the recommendations there - what do you find out?

25

26

26

Transitioning from Steady State to Dynamics

c Pressure Flow Theory

Pressure Flow Theory

© 2007AspenTech - All Rights reserved. EB1105.06.04 03_PressureFlowTheory.doc

1

2

2

Pressure Flow Theory

Pressure Flow Theory

3

Workshop The Pressure Flow Theory module introduces you to the underlying concepts necessary for developing your own dynamic simulations with Aspen HYSYS Dynamics. Some of the things you willieam from this module are: •

The underlying assumptions of dynamic simulation with Aspen HYSYS

•

How to analyze your Flowsheet to make appropriate pressure flow specifications

•

Which pressure-flow specifications make sense

•

How to troubleshoot the process Flowsheet for inconsistent pressure-flow specifications

Learning Objectives After you have completed this section, you will understand: •

The basic concepts of dynamic simulation in Aspen HYSYS

•

Dynamic pressure flow specifications

•

Process Flowsheets

3

4

Pressure Flow Theory

Theoretical Foundations The Pressure-Flow Solver: A Boundary Value Problem In terms of pressures and flows, perhaps the simplest way to view the pressure flow solver in Aspen HYSYS Dynamics is to consider the Flowsheet as a Boundary Value Problem.

If you were to make pressure or flow specifications on all the boundary streams (feeds/product streams in a Flowsheet), then all the internal pressures and flows would be solved simultaneously at each integration step by the pressure-flow solver. The internal stream pressures and flow rates are calculated from the pressure gradients in the Flowsheet. Flow rates are determined from: Since pressure gradients are the driving force for flow in Aspen HYSYS, care should be taken to ensure that the pressure profile of the f10wsheet has been properly specified.

•

Changes in vapour pressure nodes (vessels with hold-ups) within the Flowsheet system

•

Resistance across valves

•

Conductance through equipment (coolers, heaters, heat exchangers)

Pressure Nodes All unit operations (with hold-up) represent pressure nodes. Some unit operations may contribute to one or more nodes. For example: •

Heaters/Coolers with multiple zones

•

Heat Exchanger - shell side/tube side

•

Columns with multiple stages (trays)

Fundamental Principle Vessel equipment has a fixed geometry and thus a fixed volume. Mathematically, this means that: This concept is fundamental to performing dynamic simulation analyses with Aspen HYSYS.

dV

dt

o

(1)

Therefore, for a fixed volume, a pressure node (vessel pressure) is calculated as a function of the vessel temperature and the vessel hold up.

4

Pressure Flow Theory

5

In dynamic mode, the rate of change in vessel pressure is related to the rate of change of temperature (enthalpy) and the rate of change of material hold-up (level):

dP

dt

where:

=

fn(V, F, T)

(2)

V = Fixed volume F = Change in flow (hold-up)

T= Temperature (change in enthalpy) A volumetric flow balance around the vessel can be expressed as follows: (3)

where: LlVp

=

Volume change due to pressure change

LlVF = Volume change due to flow changes LlVr = Volume change due to temperature change The total volume change must always be zero.

5

6

Pressure Flow Theory

Example Consider the operation of a separator in dynamic mode that is initially at steady state with a level of 60%: Figure 1

r------- Assume fixed flow Flow in Fixed geometry

60%

.....-----Assume fixed flow

Remember:

In Steady State,

Flow into separator = Flow out of separator, no accumulation. But in Dynamics, if the separator feed flow increases with the product flow rates (vapour and liquid) remaining unchanged, the level (hold up), temperature (enthalpy) and pressure ofthe vessel must all change from the steady state condition.

Liquid Level Increases Since Liquid Flow In - Liquid Flow Out = Accumulation (hold-up), an increase in the feed liquid Flowrate with a constant liquid product Flowrate results in the liquid level (hold-up) increasing.

6

(

Pressure Flow Theory

7

r

Vessel Pressure Increases The vessel pressure would increase for two reasons. 1.

Vapour Flow In - Vapour Flow Out = Accumulation. An increase in the feed vapour Flowrate with a constant vapour product Flowrate results in the vapour (hold up) increasing. Because vapour is a compressible fluid, the accumulation of vapour occupying a smaller volume will cause the vessel pressure to rise.

2.

The increase in liquid level also causes the vapour hold-up to occupy a smaller volume within the vessel, causing the vessel pressure to rise.

Distributed and Lumped Models Most chemical engineering systems have thermal and component gradients in three dimensions (x, y, z) as well as in time. This is known as a distributed system. Thus, in the formulation of chemical engineering problem equations, we obtain a set of partial differential equations in the x, y, z, and t domains. If the x, y, and Z gradients are ignored, the system is lumped and all the physical properties are considered to be equal in space. In such, an analysis in which only the time gradients are considered, the chemical engineering system equations are represented by a set of ordinary differential equations (ODE's). This method saves calculation time and provides a solution that is reasonably close to the distributed model solution. Aspen HYSYS uses lumped models for all unit operations. For instance, in the development ofthe equations describing the separator, it is assumed that there are no thermal, pressure or concentration gradients present. In other words, the temperature, pressure, and component gradients are the same throughout the entire separator. Aspen HYSYS does take into account the static pressures in the fluid and vapour phases. This can result in a dP/dz effect in a vessel. However, Aspen HYSYS does not solve any partial differential equations.

7

8

Pressure Flow Theory

Pressure-Flow Relationship for Valves In any Flowsheet, the valve unit operation describes the resistance to flow between two material streams by the Turbulent Flow equation: (4)

where: PI

=

upstream pressure (pressure of stream I)

P 2 = downstream pressure (pressure of stream 2)

Cy = the valve coefficient, Aspen HYSYS will calculate value on request

Pressure-Flow Relationship for Other Operations More generally, flow rates in Aspen HYSYS Dynamics are related to delta P. All process equipment relates the flow between its feed and product streams with flow equations that are similar to the turbulent flow equation. The form ofthese equations IS:

F = kJpM where:

(5)

k = Conductance, which is a constant representing the reciprocal of resistance to flow

p = Stream bulk density M

=

Pressure gradient across the operation

Specifying Cy or k values (rather than a fixed delta P) across valves and process equipment provides for a more realistic simulation. By specifying these variables, the pressure drop through valves and process equipment can change with changes in flow, as it would happen in an actual plant. This allows the Dynamic simulator to model the actual operating conditions of the plant more accurately.

8

Pressure Flow Theory

9

Pressure/Flow Networks In Aspen HYSYS Dynamics, the pressure/flow network is described in terms of nodes, resistance, and conductance. Flow takes place in streams from one node to another. Thus there are two basic sets of equations that define the pressure/flow network: The resistance to flow through valves and the conductance through process equipment determines stream flow rates between nodes.

1.

Equations that define the material balance at the nodes

2.

Equations that define the flow-conductance and resistance to flow

The simplest case is that of incompressible flow with no accumulation at the nodes. In this situation, the flow equations are a function of the pressure gradient and equipment parameters such as the pipe diameter and roughness. The material balance at the nodes is simply that the accumulation is zero. In a more comprehensive dynamic simulation, the pressure flow equations are more complex. They account for: •

Multi-phase flow with the potential for slippage between phases

•

The rate of change of pressure at the nodes as a function of the equipment geometry, hold-up, and enthalpy of the phases

•

Flow rates that are determined not only by pressure gradient, but also by weir heights (columns) and density differences

9

10

Pressure Flow Theory

Simultaneous Solution Approach to Pressure Flow Balances Since pressures at nodes are a function of the flow rates into and out of the nodes, and the flow rates through equipment are functions of the upstream and downstream pressures, the relationships between pressure and Flowrate equations in Aspen HYSYS Dynamics are significantly coupled. To find a solution to the pressure-flow relationships in Aspen HYSYS Dynamics, a simultaneous solution of the Flowsheet is performed. Solving for the flows and pressures requires the simultaneous solution of a set of linear and non-linear equations. Figure 2

•

PI, P2, P3, etc. represent Pressure Nodes (Vessels with hold up)

•

FI, F2, F3, etc. represent streams with flow rates

Moreover, in order to epitomize computational effort, Aspen HYSYS Dynamics partitions the equations describing any unit operation into three classes: •

Pressurelflow relationships

•

Energy relationships

•

Compositional relationships

These groups of equations can then be integrated/solved with different frequencies. Typically, the pressure flow relationships will have the smallest step size and the composition relationships the largest. The grouping of the equations also permits a different solution strategy to be applied to each group. In particular, it is possible to solve the pressurelflow relationships simultaneously across the entire Flowsheet while the other equations (composition, enthalpy) are solved on a module-by-module basis.

10

Pressure Flow Theory

11

~-

If you suspect that the P/F solver is failing because ofthe interaction with the VLE correlation, then you can do one ofthe following: •

Reduce the integration step size - this can be accessed from the menu bar: Simulation I Integrator I General.

•

Change the frequency of integration steps per step size (composition and enthalpy). This can be accessed from the menu bar: Simulation I Integrator I Execution.

11

12

Pressure Flow Theory

Degrees of Freedom Analysis In Module 2, we introduced the concept of dynamic specifications. The simultaneous

solution of the pressure-flow relationships within the Flowsheet requires the user to make a number of dynamic operating specifications. •

P = Pressure

•

F=Flow

Figure 4

V-100 Separator

In this Flowsheet, there are 7 variables in total that will define the system. These are

•

•

• •

Feedl

a

One variable for pressure

a

One variable for flowrate

Product!

a

One variable for pressure

a

One variable for flowrate

Product2

a

One variable for pressure

a

One variable for flowrate

V-IOO

a

One variable for pressure

In addition, four equations define the pressure-flow relationships in the Flowsheet: •

VLV-IOO: Resistance to Flow equation FVLV-IOO = fn(C v, PI, P 2)

•

VLV-lOl: Resistance to Flow equation FVLV-IOI = fn(Cv,PJ, P2)

•

VLV-102: Resistance to Flow equation FVLV-I02 = fn(C v, PI, P 2)

•

V-IOO: Pressure Node Relationship dP/dt = fn(V,F,T)

With 7 variables and 4 equations, the DOF = 7 - 4 = 3. Therefore, 3 PIF specifications need to be made to define this system.

12

c Pressure Flow Theory

13

Understanding the Placement of P/F Specifications Why do some PIF specifications work while others don't? Aspen HYSYS Dynamics is equipped with a Dynamics Assistant that analyzes the Flowsheet to identify problems. (We will discuss this simulation aid later in this module). However, with a greater understanding of the role ofthe PIF solver and the PIF calculations you will be better able to: •

Specify the process Flowsheet correctly

•

Troubleshoot the process Flowsheet to identify PIF problems

•

Use the power of Aspen HYSYS Dynamics to its full capabilities

Making Consistent Pressure or Flow Specifications As mentioned earlier, Aspen HYSYS Dynamics users can select from a variety of pressure-flow specification combinations to solve the process Flowsheet. These include: •

Pressure specifications on material streams

•

Flow specifications on material streams

•

Fixed pressure drop specifications across equipment

•

PressurelFlow calculations for valves - resistance to flow (Cv)

•

Conductance calculations (k) for process equipment

In the previous example, we had three Degrees of Freedom, requiring that three specifications be made to define the system.

13

14

Pressure Flow Theory

One Possible Solution Figure 5

t,,,,,,,,,~~~~,,,,,··,,,,,,,,~,,,,,,,,,,~~~::,,""-+~~"""""""""~··"""~.::.:~4""""''''''''''''''''''''"""..,,,.t:::::......-.:::pj

Vapour

Product 1

VLV-101

P, F?

V-100

Separator

P? L ~

~ u;:'::.::::.:::~~~

.-...~~

Liquid

. VLV-102

,

·.·~

~:::::::::-::t..

Product 2 P, F?

Specify: •

Feed I Pressure

•

Stream I Pressure

•

VLV-IOO Delta P

Although making these three specifications will satisfy the DOF analysis, the choice of specifications would not make sense. PFeedl, PI and PVLV-IOO are all related by the following equation: PFeedi -

PI -

MVLV-IOO =

0

(6)

Specifying the Flowsheet in this manner would lead to an inconsistent solution. In fact, the Flowsheet would be under-specified because one ofthe specifications is redundant.

14

o~

__

~

_

~~-~ ----~-

(~-

'-----------------------------

-

Pressure Flow Theory

15

Another Possible Solution Figure 6

\/-100 Separator

SpecifY •

Feed1 Pressure

•

Product! Pressure

•

Product2 Pressure

Consider the same Flowsheet with pressure specifications made on all the boundary streams. This solution is consistent because the pressure in the vessel is calculated by the hold-up equation. (The stream flow rates were calculated using the turbulent equation or the resistance to flow equation).

15

16

Pressure Flow Theory

Summary of P/F Theory and Specifications: •

The flow through the plant, or operation, is driven by the pressure gradient.

•

P/F theory defines the relationship between flow and pressure.

•

The Aspen HYSYS P/F solver solves a set oflinear and non-linear equations simultaneously to determine the P/F relationship.

•

In order for the P/F solver to solve the Flowsheet, there must be a pressure

gradient established over the entire Flowsheet. •

The pressure gradient exists due to a specified pressure flow relationship (or a specified pressure drop) over all operations in the Flowsheet.

•

The P/F solver works by finding P from F, according to the P/F theory, or by solving the pressure node equation.

•

Following any flow path through the Flowsheet, the user should be able to see the pressure gradient or expect to see a pressure gradient established along the path. If the pressure gradient cannot be seen, an additional pressure specification may be needed.

•

Ifthe user follows a flow path to the boundary of the Flowsheet, they should see that at such a location, a pressure gradient does not exist, nor can it be established. This means that a pressure (or flow) specification is always needed on boundary streams.

Other Possible Solutions If we modeled the same unit operation without using valves on all product streams, then we could not make P specifications on all boundary streams. Remember the lumped parameter model - the model assumes there are no pressure gradients inside the unit operation. Thus, if a pressure specification is made on the vapour product stream it is best not to make pressure specifications on the other unit operation streams. This can lead to an inconsistent solution because once one stream pressure is known they all become known, resulting in no pressure gradients in the unit operation.

16

(

c-------------------------r

Pressure Flow Theory

17

Figure 7

Vapour Product

Feed 1

Separator

Liquid Product

It is possible to have flow specifications on all unit operation streams as long as the vessel pressure is controlled. Figure 8

,--

+ Vapour A

Product

Feed

1 Separator

Sep PC

Liquid Product

17

18

18

Pressure Flow Theory

r----

Dynamic Details

Dynamic Details

© 2007 AspenTech - All Rights reserved. EB1105.06.04 04_DynamicDetails.doc

1

2

2

Dynamic Details

"

A;;;h~ PIPE-100

Q-AIPha

rav

.... -;::~".~.~}

-_

VLV-101

VLV-100

~ .- "" ~_ Ai~-A'Ph~ .

Q-Bravo

.nCh~ie1 ~~2

in~~~~,'~'~;liE!

/,

-~ 1~~-8;~~

Bravo

PIPE-1 02

n, n n n

b;~~1i~h

.~~~u~....:··

Q-Charlie

~h'i'~'li'~nn"'nnnn"'H',': ,ii("""'""""",iiii:J

TRF_1

VLV-104

VLV-104-2

Alpha2

B

Charlie2

«TY'~~~."",. "";!."'i>~~"'''''-'<''5}

.......

PC-HPSep

ToSep

In;}---~

MIX-100

(::::

r;;;'~'~!~~1

VLV-105

!··"·,>wt""'Jj{V~~,,-,,u,¥/t·

h,

IHPva p HPSep

..} .. ,:",

qcool

Hot\i'ap

I''''''''''''''''''.,;mr:::'".f-

E-100

~

...

)

L-..., ~"'''''''''HP~rq1 HPLiq VLV-103

\

,/

TC-HoM~p

'.

(")

,--~.,

4

Dynamic Details

Workshop This module examines some of the detailed parameters available in Aspen HYSYS. These details include actuator characteristics, nozzle locations, and heat loss parameters. Starting with the dynamic module that you prepared in Module Transitioning, you will add the necessary information to provide for a more exacting model.

Learning Objectives After you have completed this module, you will be able to: •

Add valve characteristics

•

Add heat loss models

•

Understand nozzle location

•

Make changes/additions to the dynamic model

Prerequisites Before beginning this section you need to know how to:

4

•

Set up Strip Charts

•

Understand the Flow-Pressure network

Dynamic Details

5

Aspen HYSYS Fidelity Fidelity is an extension of Aspen HYSYS that provides advanced dynamic features to your simulation. Fidelity allows you to put together very detailed models for operator training work or detailed dynamic studies. The capabilities exclusive to Fidelity are as follows: •

Static head included in the pressure relationships. You also have the ability to modify equipment elevations.

•

Nozzle locations can be modified. For example, an overhead vapour nozzle is somewhere near the top of the vessel.

•

Detailed valve actuator dynamics. The dynamics of the valve opening and closing are included in the model.

•

A detailed heat loss model to take into account heat loss from vessels with holdup to the environment. For example, you can supply details about the equipment and insulation to take into account heat transfer from the vessel to the environment.

•

Details on rotating equipment. Inertia terms account for the start-up and shut-down of rotating equipment.

Go to the Simulation Integrator window (Ctrl+I) and on the Options tab enable the use of Aspen HYSYS Fidelity.

Characteristics of Valve and Its Actuator The following behaviors on the valves have been observed in the field: •

Valves on feeds go from fully open to closed in 1 minute.

•

Valve on Charlie is not responding instantly; take 5-second time constant.

•

Valve on Bravo does not shut off completely; there remains a 2% leak.

•

The HPSep liquid valve is an Equal Percentage Valve.

Aspen HYSYS has the ability to model each of the above observations.

5

6

Dynamic Details

Prior to entering the appropriate infonnation, we will make a simple model consisting of two identical valves with two identical controllers.

Stream dynamic specification can be introduced on the Worksheet PF Specs window of each unit operation.

PeMPr~ssDte

1.

Open up the ftle EM2-Dyn1.hsc and save it as EM4-Valvel.hsc. This model should be in dynamic mode with all PVs lined out at the desired SPs.

2.

Add two valves to the simulation. Name Inlet streams as 1 and 3 and Output streams 2 and 4.

3.

Enter a dynamic pressure specification for feed streams 445 kPa (64.5 psia) and for product streams 101.3 kPa (14.7 psia).

4.

Size the valves to handle a 4535 kg/h (10,000 lblhr) flow with a 345 kPa (50 psi) pressure drop and a valve opening of 50%.

5.

Add two flow controllers; the PVs are the mass flow to the valve and the Ops are the valves themselves. Provide identical controller tunings and ranges according with the following table. Place the controllers in Automatic.

A4s:0l
Pt()ductW~gSUre1Ql,3

MassFlaWA~;3~ Resisti'fM~iGv,QtK)

Action

Reverse

Range PV Minimum

o kg/h (0 Ib/hr)

Range PV Maximum

9000 kg/h (19800 Ib/hr)

Mode

Auto

OP

50%

Kc

0.25

Ti

0.10

Td

---

o

6

~·I;.,~<>10

Dynamic Details

7

Actuator Lineal Rate The Actuator page, located on the Dynamics tab ofthe Valve property view, allows you to model valve dynamics. This page also contains information regarding the dynamic parameters of the valve and the per cent open positions of the actuator and the valve. In reality, changes that occur in the actuator are not observed instantaneously in the valve. Moreover, changes in the output signal of a controller (OP) do not instantaneously translate to changes in the actuator. Because the actuator and valve are physical items, they take time to move to their respective desired positions. This causes dynamic behaviour in actual control valves. An actuator is a device which applies the force required to cause movement in the valve. The valve opening has a direct impact on the flow through the valve. This relationship is a function of the valve type and the pressure of the surrounding pieces of equipment.

The valve mode defines the relationship between the desired actuator position and current actuator position. The desired actuator position can be set by a PID Controller or Spreadsheet operation. A controller's output, (OP), for instance, is exported to the desired actuator position. Depending on the valve mode, the current actuator position can behave in one of the following three ways: •

Instantaneous Mode

•

First Order Mode

•

Linear Mode

In instantaneous mode, the actuator moves instantaneously to the desired actuator position defined by the controller. If the First order mode has been selected, the actuator responds with a first order lag according to the Actuator Time Constant cell value. Activating the linear mode, the actuator moves according with the Actuator Linear Rate value (%/second).

7 \_.

8

Dynamic Details

1.

Open up the view for VLV-I05. Select the Dynamics-Actuator tab. Change the mode to Linear and keep the default rate of 0.01 %/sec for the Actuator Linear Rate.

2.

Set up a strip chart to monitor the Actuator Desired Position, the Percentage Open for each of the two valves, and the Mass Flow of each of the two streams. (Set-up logger 3600 samples and 5 s for the sample interval). Run the integrator until you get a steady state.

3.

Change the setpoints for FC-l04 and FC-105 to 2250 kg/h (5000 lb/hr).

4.

Start the integrator and observe the response. Figure 2

Figure 3

8

Dynamic Details

9

Valves on feeds go from fully open to closed in 1 minute: What should be the desired Actuator Linear Rate value for Alpha, Bravo, and Charlie control valves?

5.

Change the Actuator Linear Rate ofVLV-105 to the desired 100% in 1 minute (= 1.666 %/sec) and change the setpoints for FC-l04 and FC 105 to 4535 kg/h (10,000 1b/hr).

6.

Start the integrator and observe the response.

7.

Enter the desired values for the Alpha, Bravo, and Charlie Actuator Lineal Rates (1.667 %/sec) and switch from Instantaneous mode to Linear.

Stickiness Time Constant In reality, the valve does not respond instantaneously to changes in the actuator. A first order lag can be modeled in the response of the actual valve position to changes in the actuator position. The Valve Stickiness Time constant allows you to specify the time constant used to model the time offset caused by a sticky actuator. The offset can be specified in the Actuator page of the Dynamics tab in the Valve property view. 1.

Set up VLV-I 04 and VLV-I 05 so that they are both linear with a rate of 1.667 %sec and enter a time constant of20 sec on the Valve Stickiness cell ofVLV105. Figure 4

',,-:-

-

9

10

Dynamic Details

2.

Change the setpoints for FC-104 and FC 105 to 2250 kg/h (5,000 lb/hr) and 6800 kg/h (15,000 lb/hr) respectively.

3.

Set up the strip chart to include the Actuator Current Position. Compare this to the Percentage Open. Start the integrator and observe the response. Figure 5

4.

10

The valve on Charlie is not responding instantly, taking a 5-second time constant. What should be the Valve Stickiness Time Constant for the Charlie control valve? Make the desired change to VLV-102 for Charlie (5-second time constant).

Dynamic Details

11

Leaky Valves Leaky valves can be modeled by specifying a non-zero value for the minimum valve position. 1.

Make valves 104 and 105 identical, change the setpoints of both flow controllers to 4535 kg/h (10,000 lblhr), and run the integrator until you get the new mass flow.

2.

Enter 2% for the minimum valve position for valve V-105.

3.

Change the setpoints for the controllers to a small number such as 135 kg/h (300 lblhr) and observe the response. Figure 6

4.

Make the appropriate change to VLV-101 for Bravo, say 2% ofleak

5.

Save case and then save case as EM4-Valve2.hsc

11

12

Dynamic Details

Inherent Flow Characteristics - Valve Operating Characteristics We define inherent flow characteristic ofa valve f(x) as the relation between the tap position, x, and the product flow rate that flows through it as a fraction to the maximum flow rate when the delta P, AI> in the valve is constant.

Q = Cv· J(x) -l~P PL

F A f(x)=-=Fmax A max PL

=

specific density of the low viscosity liquid

p(1b/ft 3 ) PL = 62.4 Cv = flow of water at 60° F in US Gallons per minute that goes through a control valves when the delta P is 1 psi and f(x) is equal to 1 (100% open).

Q(USGPM) = Cv .1. ~1(~Si) Kvs = maximum flow in m 3/h that flows through a valve when the Delta Pis 1 bar. Kv = flow in m 3/h that flows through a valve when the Delta P is I bar.

K vs = O.86C v K v = f(X)K vs

What would be the Cv ofa valve to have a maximum flow of 5 USGPM ofwater when it _ looses 2 psi fully open? What is the Cv if the valve outlet is stream HpLIQ in the model and yet the maximum flow rate is still 5 USGPM?

12

Dynamic Details

13

The most frequent encountered valves in terms of operating characteristics are Lineal, Equal Percentage, and Quick opening. The first two are the most popular in control systems. Aspen HYSYS implements all three. 1.

Defme stream 1 as a pure water stream at 15.6 C (60 OF) and 108.2 (15.7 psia).

2.

Size the Cv for VLV-104 equal to 1 USGPM and tum off the FIC-104 controller.

3.

Add a Transfer Function and select the Percentage Open ofVLV-104 as selected OP.

Transfer Function icon

Figure 7

(

L

4.

On the Parameters tab, select the Configuration page, define 0% and 100% as minimum and maximum values respectively for both the PV and OP ranges, and enter 0 as PV.

(

( (

13

(

14

Dynamic Details

5.

Select Ramp from the left column menu and select Ramp as Active Transfer Function.

6.

Type 100% for Ramp Magnitude and 10 minutes for the Ramp Duration. Figure 8

7.

Add a chart to follow the Std Ideal Liq Vol Flow ofYLV-I 04.

8.

Run the integrator and press the Start Ramp button.

9.

Press the Reset Ramp button and repeat steps 6 and 9 for an equal percentage valve.

10. Repeat Step 9 for a quick opening valve.

14

-._-.-._--=-=

-c=----'-

(~

" _---------------------------("

Dynamic Details

15

'- . .

(

You will get a plot similar to the following one for each type of valve:

( Figure 9

( (

( (

(

Installed Flow Characteristics The installed flow characteristic F(x) of a valve depends on the pressure drop of the flow circuit where the valve is installed and how it will react to flow rate changes. The calculation of the static gain in this case will depend on the overall flow system where the valve is installed. It is convenient for calculation purposes to define a parameter r as the ratio of the minimum delta P of the valve (IOO%open) divided to the total delta P of the flow circuit.

r =

M valve ~ystem

1.

Add a valve to stream 2 by copying/pasting VLV-I 04. Define it as linear with OP equal to 25%.

2.

Move the atmospheric pressure spec from stream 2 to the new stream and set the pressure spec of stream 1 equal to 115.1 kPa (16.7 psia).

{

'.

15

18

Dynamic Details

Equipment Location Static Head By default, the static head is not included in any calculation. Look at the pressures around the high-pressure separator; they are all equal. For any unit operations with hold-up, Aspen HYSYS calculates the static head by considering the equipment hold-up, geometry, and elevation of any attached nozzles. In order for Aspen HYSYS to calculate the static head for any unit operation, you need to enable the calculations. This is done on the Options page of the Integrator property view.

1.

Open the Integrator view (Simulation I Integrator).

2.

Enable the Static Head Contribution checkbox on the Options tab. Figure 12

18

Dynamic Details

3.

19

Run the simulation and observe the slight changes in the controller outputs. Look at the pressures surrounding the separator.

What's the Pressure ofStream HP Liq?

Nozzles By default, all unit operations are placed on the ground, and for most simulations, this is fine. The Nozzles page, located on the Rating tab, contains information regarding the elevation and diameter of the nozzles.

( The elevations of each nozzle attached to the equipment are displayed relative to several reference points:

C ( (

1.

(

•

The Ground is a common reference point from which all equipment elevations can be measured.

•

The Base is defmed as the bottom of the piece of equipment.

Plot on a chart: •

The aqueous and liquid mass flows of Stream HPLiq

•

The Liquid Percent Level ofHPsep

•

The HvyLiquid Percent Level ofHPsep

' ..

~ 19

( '-

20

Dynamic Details

2.

Run simulator.

3.

Change the nozzle location for the stream Hp Liq from 0% to 25% and observe.

4.

Change the set point for the level controller LC-HP Sep and observe.

Can you explain the Results?

_

Could the controller achieve the SP all the time?

The nozzle diameter, nozzle elevation, and the levels of the phases inside the vessel determine what flows out through the nozzle. The nozzle elevation refers to the center ofthe nozzle opening, not to the bottom or top of it. If the product nozzle is located above the liquid level, the exit stream draws material from the vapor holdup. If the liquid level sits across the nozzle, the mole fraction ofliquid in the product stream varies linearly with how far up the nozzle the liquid is.

Feed-..j

Move the nozzle location back to the original result (0%), move the SP for LC-HP Sep to 50%.

Save the case once and then save as: EM4-Heatloss.hsc

20

Dynamic Details

21

Detailed Heat Model The vessel unit operation also has the ability to take into account heat loss from vessels with hold-up to the environment. For example, you can supply details about the equipment and insulation to take into account heat transfer from the vessel to the environment. There are several underlying assumptions that are considered during a heat loss calculation: •

•

•

There is a heat capacity and a thermal conductivity associated with the wall and insulation housing the fluid.

The detailed heat model is located on the Heat Loss page of the Rating tab. You can choose to neglect the heat loss calculation in the energy balance by selecting the None radio button. There are two heat loss models available: Simple and Detailed.

Simple Model The Simple model allows you to either specify the heat loss directly or have the heat loss calculated from the following variables:

The temperature across the wall and insulation is assumed to be constant (lumped parameter analysis).

•

Overall U value is specified by the user.

•

Ambient Temperature is specified by the user.

•

Heat transfer area, A, is calculated by Aspen HYSYS.

The calculation uses convective heat transfer on the inside and outside of the vessel.

•

Fluid temperature, Tf, is calculated by Aspen HYSYS.

T wali IN

The heat loss is calculated using:

T Fluid

Detailed Model In the detailed model, the user supplies both conductive and convective information.

The program calculates the heat loss and supplies a temperature profile from the fluid to the ambient. 1.

Enter the following conductive and convective data for the HPSep vessel and insulation.

rC·

Material

Metal

Insulation

Thickness

0.051 m (0.167 ft)

0.030 m (0.098 ft)

Cp

0.473 kJ/kg-C (0.113 Btu/lb-F)

0.82 kJ/kg-C (0.196 Btu/lb-F)

Density

7801 kg/m3 (487 Ib1ft3)

520 kg/m3 (32.46 Ib/ft3)

Conductivity

45 W/m-K (26 Btu/hr-ft-F)

0.15 W/m-K (8.7e-2 Btu/hr-ft-F)

21

(

22

Dynamic Details

Inside Liq Phase U

180 kJ/h-m2-C (8.806 Btu/hr-ft2-F)

Outside U

36 kJ/h-m2-C (1.761 Btu/hr-ft2-F)

Vapor to liquid U

18 kJ/h-m2-C (0.8806 Btu/hr-ft2-F)

The ambient temperature is set with the integrator settings. From the tool bar select Simulation I Integrator and open the Heat Loss tab. Figure 14

Maintain the default ambient temperature of25 C (77 F).

~ ~

2.

Click "Temperature Profile" and then click the "Initialize Temperature" button.

3.

Start the integrator.

Observe the temperature profile ofthe vessel. Did it immediately go to a steady state value?

Do you see any changes?

Save your case.

22

c (

(

Expanding the Model

Expanding the Model ,(

© 2007 AspenTech - All Rights reserved. EB1105.06.04 05_ExpandingTheModel.doc

1

2

2

Expanding the Model

c J.-"-'.

c"""-

_

Expanding the Model

3

(

3

4

Expanding the Model

Workshop In this module, the High Pressure Separator will be expanded with a Knockout Drum and a three-phase Low Pressure Separator. The control system will also be modified with Cascade, turn-off, and split range controllers. Furthermore, a pressure relief valve will be added to the simulation. There are several ways to add equipment. Some prefer switching the model back to steady state, making the changes, then switch back to dynamics. Others prefer adding the equipment directly to the dynamic model. Aspen HYSYS supports either method. In this workshop, the unit operations are going to be added in the Dynamic Mode.

Learning Objectives After you have completed this module, you will be able to: •

Add unit operations and controllers in the Dynamic mode

•

Make necessary P/F Specs for the system

•

Implement appropriate control strategies

•

Install a relief valve

Prerequisites Before beginning this section, you need to know how to:

4

•

Use the Dynamic Assistant

•

Set up Strip Charts

•

Understand the Flow-Pressure network

Expanding the Model

5

Build the Flowsheet In this module, you will install several unit operations in the Dynamic mode. You may also construct the simulations in the Steady State mode and then move them into the Dynamic mode. However, the goal of this module is to teach you how to built simulations in the Dynamic mode.

Pressure Control for HP Separator Open the saved case from Module 4, EM4-HeatLoss. This case should still be in the Dynamic mode. If it is not, enter the Dynamic mode, run the Integrator until Steady State is reached, and turn off the Integrator. Save it as EM5-l.hsc Before adding a Knockout Drum, we are going to change the pressure control scheme for the lIP Separator. A control valve will be placed on the HotVap line. 1.

Add a valve to the simulation. The inlet stream is HotVap and the product stream is HotVap Out. Name the valve Knockout Valve.

2.

On the Rating-Sizing tab for the valve, size the valve based on a 10 psi pressure drop.

What Cv value does Aspen HYSYS calculate/or the new valve?

3.

Run the Integrator for a few minutes to propagate the values to the boundary streams and save the case.

4.

Move the P/F specs to the boundary stream - From the Worksheet-PF Specs tab move the pressure specification from the stream HotVap to HotVap Out.

5

6

Expanding the Model

• ~ ~

~.":"."' .. ..

6

5.

Change the output connection for PC-HPSep from FC-Alpha to the Knockout valve.

6.

Place FC-Alpha in Auto and copy PV value into SP for smooth transition.

What else needs to be done to the pressure controller?

7.

Change the setpoint to the controller and test if the model is still okay.

8.

Change the setpoint back to its original value.

.,.--

c

Expanding the Model

7

Adding the Knockout Drum A Separator unit operation will be used as the Knockout Drum. The cooled overhead stream HotVapOut will be the feed for the separator unit.

Remember the rules for P/F Specifications. You cannot place Pressure Specifications on both Accumulator product streams.

1.

Turn off the Integrator.

2.

Add a Separator and provide the following information.

3.

Name

KnockOut Drum

Feed

HotVap Out

Vapor Outlet

KOVap

Liquid Outlet

KO Liq

By adding the Separator you have added an additional boundary stream. Supply a Molar Flow specification for stream KO Liq.

What value shouldyou use for the KO Liq Molar flow specification?

4.

Go to the Dynamics tab of the Knockout Drum. By default, the simulation will initialize the vessel in a Dry Startup condition (empty). Change the level to 0%.

7

8

Expanding the Model

5.

Run the Integrator for a few minutes to propagate the values to the boundary streams.

6.

Move the pressure specification from stream HotVap Out to the boundary stream KO Yap.

Reverse Flow In the previous rules, there should be a resistance device on both of the KG streams,

and pressure specs should be used (they are boundary streams). When a pressure spec is given to the exit stream of a vessel, then that stream will do whatever is required to maintain the vessel pressure. The exit stream will either let flow out of the vessel or let flow into the vessel (negative flow). For this reason, all boundary product streams have a product block. The product block sets the conditions ofthe stream if there is negative flow. By default, it has the same composition as the vessel, but you can make the appropriate changes (for example, N2 or CH3 for atmosphere or fuel gas blankets).

8

( (

(-------------------------

Expanding the Model

9

C (

c

re-·

In the case of the flow spec, the flow is always equal to the specification supplied, no matter what the pressure or level of the vessel.

These types of specs are often used when it is desired to simplify problems, but they can lead to some strange results. While they help us simplify a problem, we often want to eliminate these specifications whenever possible.

Adding a Valve on the Liquid Outlet A Compressor unit operation will be added to the flowsheet to increase the pressure of the KO Yap stream in the next module. For the KO Liq stream, a valve will be added to the flowsheet. This valve allows manipulating the liquid flow from the vessel, and it will be the control valve for the vessel level controller. 7.

8.

Add a Valve and provide the following information.

Name

LP Separator Valve

Feed

KO Liq

Product

To LP Separator

Move the PIP specifications to their proper locations. The stream To LP Separator will have the same pressure value as stream HP Liql.

What Cv value does Aspen HYSYS calculate/or the new valve?

9

10

Expanding the Model

9.

Start the Integrator for a few seconds to propagate values to the boundary stream.

10. Add a level controller in order to maintain the level ofthe KnockOut Drum. Open up the view for KnockOut Drum. Go to the Dynamics tab and press the Add/Configure Level Controller button. Valves normally operate between 25 and 75% of the Valve Opening %.

11. Set the SP level equal to 50%. Once at steady-state conditions, what's the Valve Opening %?

Adding the LP Separator In this section, we will add the three-phase separator. The inlets are the two liquid streams, HPLiq-l and To LP Separator, and the vessel size is the same as the High Pressure Separator. 1.

2.

Add a Three-Phase Separator with boot and provide the following information:

Name

LP Separator

Feed 1

HP Liq 1

Feed 2

To LP Separator

Vapor Outlet

LPVap

Light liquid Outlet

LP Liq

Heavy Liquid Outlet

Waste Water

Move the P-F Specifications to the new boundary streams. Supply a Pressure Specification for stream LP Yap and Molar Flow specifications for streams LPLiq and Waste water.

What value should you use for the Liq Molar flow specifications?

10

3.

Go to the Dynamics tab ofthe LP Separator. By default, the simulation will initialize the vessel in a Dry Startup condition (empty). Change the level to 0%.

4.

Start the Integrator and let it run for a short period of time. Do not let the level go to 100%. Check the Rating-Nozzle and the Dynamics Holdup windows.

(

(

c-------------------

c

Expanding the Model

11

Are the liquid level and the Nozzle elevation OK/or both liquid phases?

Adding the Valves and Controllers ( \

The LP separator requires controllers to stabilize the process. Three controllers are going to be added. One PI controller for the vessel pressure, one PI controller for the light liquid level on the LP Separator, and one on-off controller for the waste water stream valve.

Pressure Controller 5.

Add a valve for the LP Separator Vapor Stream. Name it LP Pressure Valve and name the outlet stream LP Vap-l.

6.

The boundary pressure is 1380 kPa (200 psia).

7.

Add the following Pressure Controller to the flowsheet.

8.

Action

Direct

Mode

Auto

PVMinimum

1380 kPa (200 psia)

PV Maximum

3100 kPa (450 psia)

Kc

3

Ti

2

SP

2690 kPa (390 psia)

Run the Integrator for a few minutes.

Level Controller 1.

Add a valve for the LP Liq stream. Name it LP Level Valve and name the outlet stream LP Liq-l.

2.

The boundary pressure is 1380 kPa (200 psia).

11

12

Expanding the Model

3.

4.

Add the following Level Controller to the flowsheet:

Action

Direct

Mode

Auto

PVMinimum

0%

PV Maximum

100 %

Kc

2

Ti

10

SP

65%

Run the Integrator for a few minutes.

On-Off Controller For safety reasons, water phase level cannot be over 1.98 meters (65ft) and under 0.45 m (1.5 ft). An On-Off Controller will fully open a control valve when the level is higher than 1.98 m and it will fully close the valve when the level is lower than 0.45 m. ..................-...-•.•.•.•.•.•.........• •".

•••·.••••••••••••••.;-c---.·;

on-Off~

~" '"

~

Contrfller

•••>.-.>.•.•" ••••• " •••••••••

Waste Water-1

LP Waste Valve

12

•.••••••••••...•.•

1.

Add a valve for the Waste Water stream. Name it LP Waste Valve and name the outlet stream as Waste Water-1.

2.

The boundary pressure is 1380 kPa (200 psia). Run the Integrator for a few minutes.

3.

Make an On-Off Controller for the LP Separator. This can be done adding a Digital Point operation. Name it On-Off Controller.

Waste Water

(, ,,---,

t

(

Expanding the Model

13

(

c (

4.

If the level is above the maximum level, the desired actuator position is 100%. Open the first digital point (Full Open) and add the following information:

( (

(

c

Name

On-Off Controller

Process Variable Source

LP Separator, Phase Level, Liquid 2

Output Target

LP Waste Valve, Actuator Desired Position

Mode

Auto

Output Cold Init OP

Default

Auto Operational Parameters

Latch

Threshold

1.98 m (6.5 ft)

Higher Dead Band

0.00

Lower Dead Band

1.53 m (5.0 ft)

OP is On when

PV >= Threshold

PVMinimum

o

PV Maximum

7.108 m (23.32 ft)

"r

--

When PV is higher than Threshold, OP is ON and Actuator Desired Position is Full Open. When the PV is lower than the Threshold minus the Lower Dead Band, (1.981.53 = 0.45), the OP State is OFF and the Actuator Desired Position is OFF.

13

14

Expanding the Model

Figure 4

Adding Strip Charts

14

1.

Create strip charts to monitor primary variables in your flowsheet.

2.

Start the Integrator and allow the model to stabilize.

3.

Once the model has reached steady state, create a disturbance to test your model.

4.

Save your case as EM5-Mods.hsc.

C' E C (

G-'- - - - - - - - - - - - - - - C

Expanding the Model

15

(

Exercises

( ( (

Adding Cascade Controllers

C

C

For slow acting control loops (such as level controllers), it is desirable to place a fast acting controller (such as a flow controller) into the loop.

c

Place flow controllers on the HPLiq, LPLiq, and KOLiq streams. Set up these controllers as slaves to the level controllers.

(

Fail Open Valve

c

Actuators usually have a fail-safe function. If there is a disruption to the power source driving the valve, the actuator places the valve in a safe position, either Fail Open or Fail Close. In Aspen HYSYS, if the Fail Shut Mode is selected, the valve becomes fully closed (ActDesired% = 0) in the event that the signal from the controller is cut off from the valve. If the Fail Open radio button is selected, the signal received by the valve is modified as follows:

( r

ActDesired%(Valve) = 100% - ActDesired%(Controller) Ifthe signal from the controller is cut off from the valve, the valve becomes fully open. Select the Fail Open option for the Fail Position ofYLV-103 in the Dynamics Actuator window.

···GP

Does the Fail Open mode selection have an effect on the direction ofthe controller?

~)('~,-,_.j...-,,-'Y'-

"1

f-

r

o/'S ,'>b.'..v/f\

r- c"-\'z ~~ cl~~

,Jy~r-~e 0," ...._~

}vl..,L.

c.. . j---t-~,'..e,·.'\

tJiS.~",

fa

-7"

'1: Y\.

J....,--t£/

~ 1

'::>-v-o

~J--"'" h.L, d....._~.(

')4~ •

o:-vt.-;'v{ ..

~

',\,"O.> d:.. W

15

~v

\P

i~

l.-

16

Expanding the Model

Adding a Split Range Controller In this exercise you will add a split range controller. 1.

Disconnect the LP Separator's vapor outlet stream LP Vap from the control valve LP Pressure Valve.

2.

Add a Tee and connect LP VAP as inlet and the valve inlet stream as Outlet of the Tee.

3.

Add a second stream to the tee outlets and a new valve. You should get a PPD as follows: Figure 5

TEE-100

16

* Cg

4.

Make the size ofthe vapor valve about 50% of what it was (Cg new = 0.5 old). Make new valve the same size as the previous one.

5.

Make the appropriate pressure specifications for the new boundary streams. Set their pressures at 1380 kPa (200 psia).

(

Expanding the Model

17

(

(

c (

6.

Replace the Pressure Controller for the LP Sep by a Split Range Controller with the following infonnation:

(

(

c

Split Range Controller icon

(

Action

Direct

ControllerMode Kc

Ti

7.

Run simulation, change the ranges and observe the behaviors.

17

18

Expanding the Model

Adding a Pressure Relief Valve The Relief Valve operation is used in many situations in which there has been excess pressure build up. Although it is available in Steady State Mode, its purpose is to avert situations that occur in a dynamic environment.

18

1.

Add a Relief Valve and enter To Flare as the Outlet stream and enter From LP as the Inlet stream.

2.

Select From LP as a third outlet of the Tee.

3.

On the Dynamics tab, Specs page of stream To Flare, activate the Pressure Specification. The pressure of this stream should be atmospheric.

4.

The Relief Valve requires a value for the Orifice Area to initialize. Go to the Sizing page of the Relief Valve under the Ratings tab and enter 25.81 mm2 (0.04 iIi).

5.

On the Parameters page of the Design tab, enter 2830 kPa (410 psia) as the set pressure ofthe relief valve and 2890 kPa (420 psia) as the Full Open Pressure.

6.

Save the case as EM5-RV.hsc.

7.

Run the model, build up pressure in LP separator (How?) and observe how the valve works. Is the valve big enough?

Compressor

Compressor

© 2007 AspenTech - All Rights reserved. EB 11 05.06.04 06_Compressor.doc

1

2

Compressor

( (

\

(

2

Compressor

3

(-

Back line .... }:.:.::.

Kic back .... • Val e stage 1: Surge Controller

HotVapOut

KnockOut Drum ....

.....-...-•.•.•.•..-..•-.-...-...-.-.-...-.,.':~:;:.\

.,'

•....................

\(

",

KO

.-......

. ?~~

~.N

.

DrumK~'

-LC

'"_.~,,

3

4

Compressor

Workshop In this module, compressor curves will be used to model the behavior of compressors for the high-pressure gas. Using curves to model these unit operations allows Aspen HYSYS to accurately simulate actual plant equipment.

Learning Objectives After you have completed this module, you will be able to: •

Specify head and efficiency curves to compressors

•

Use multiple curves to model compressors

•

Make changes and additions to the dynamic mode

•

Move Pressure-Flow Specs around

•

Accurately model existing plant equipment with Aspen HYSYS

Prerequisites Before beginning this section, you need to know how to:

4

•

Set up Strip Charts

•

Understand the Flow-Pressure network

I',

C1I

:

\11..\t 102

Bravo

Charlie 1 \11..\t 103

i

Charlie

F~:~~~~il ...,.i., .~~;:W-¥~W~~

Bravo 1

W'k$mm... "

' .

~

"'~~...... •. ,

,FC-Bravo

....:.;.:

~+

. . . "*'I>
r~:~;~:

... ;:::::::.

......

\I1..\tl00

,.l>
Ll;::'! q

LC·HP Sep

....

HPUql

PC-HPSep

q¢col TC- Hot! , Vap .

·,1

( 'I

L1C-l00 !

:7~ee

Water

V~\fJIaste

LP

N~

I

,

LP Uq

IIIIIl.+

1

LP i

Valvi~

~vel

Valve

)II~

Uq-FC',

L;·\···.•

B

Pres:::ure

LP Uq-l

4'

Vap-2

'~''''''LP'+

;n~:"""i;

VaIVf'

LPSep~PC Splh "."

2·

········l·········~;;-~~··

I

TEE-100

{

;+

l

Flare

Th

To TEG Tower

..... «......4-

L:.-_:~j..

From

Coinp Val\ie

TE~'"

t!: p •',."

KO

PC

J'" · . · ·.

S

\JIJaste

'~7//ff

BaSiine

1:~~~: comroller~ ~

g~~~ler ~

Stage

I>,<{-

I(jck~ack

Water-j

I

\.'

['

~,

6

Compressor

Compressor Curves Using compressor curves in your Aspen HYSYS simulation allows you to accurately model existing plant equipment. With an accurate simulation, you can determine if an existing compressor is able to meet the specifications of your process. The compressor curves also allow Aspen HYSYS to calculate heads and efficiencies that are dependant on the flow rate. •

If the flow rate through the compressor is known to be constant, a single flow rate and efficiency can be supplied.

•

If the flow rate is expected to change, using a compressor curve will allow Aspen HYSYS to calculate new heads and efficiencies based on the current flow rate.

This results in greater accuracy in the simulation and allows Aspen HYSYS to more closely model actual plant equipment.

Adding the Compressor Included in your course files folder there is a case file EM6Comp_with_curves.hsc, which contains a compressor with a series of compressor curves for different speeds.

6

1.

Open the Aspen HYSYS file EM6-Comp_with_curves.hsc.

2.

Click the compressor and copy it (via Edit menu), select NO when asked whether to copy streams info.

3.

Close the case.

4.

Open the Aspen HYSYS file from the last exercise (EM5-RV.hsc).

5.

Save the case as EM6 COMP.hsc.

6.

From the Edit menu bar choose Paste.

7.

Connect the compressor to the KO Yap stream, rename it as Stage 1. The outlet stream is Comp Out and the energy stream is q Compo

8.

On the Rating tab, Curves page, notice the curves are checked and the Enable Curves checkbox is also checked) and on the Dynamics tab, Specs page, notice Use Characteristic Curves is checked and speed spec is set as 5000.

- ----- ---------------=---=--- - -- -

--~--_.------.

(: ( (,

( ;;------------------(=

Compressor

7

(

9.

Run the Integrator for about 2 minutes to propagate information to stream Comp Out.

What is the pressure at the compressor discharge?

10. Add a discharge valve to the compressor outlet. Name the valve Comp Valve and name the outlet stream To TEG Tower.

7

8

Compressor

11. Remove the pressure specification of the compressor suction and move it to the outlet of the discharge valve. Set it to 6900 kPa (1000 psia). 12. Size the valve. What's the Cg value you have got for such Valve?

13. Add a pressure controller to control the compressor discharge pressure (recorder in step 11).

Action

Direct

Range PV Minimum

6000 kPa

Range PV Maximum

10500 kPa

You may want to start the pressure controller in manual. Start the Integrator and tune the controller. Use the pressure recorded in the previous step as the setpoint.

Compressor Surge Control If a centrifugal compressor goes below a certain throughput at a given head, the performance of the compressor becomes erratic and the vibrations associated with this erratic behavior can be extremely damaging for the compressor. Hence most centrifugal compressors have a protection mechanism that ensures a minimum throughput by opening a recycle valve. As the drop in throughput can sometimes be very fast, the controller must also react quickly when surge is approached and smooth control when the compressor is operating normally but not too far from surge. The surge controller is an extremely rapidly acting controller. The control algorithms used to prevent a compressor going into surge are extensions ofthe usual Pill algorithms.

8

-----~=-_.

-----

(

c----------------c··.• ( (

(

Compressor

9

Add Anti-surge Loop 1.

Add a valve to the simulation, and make sure the valve opening is 0%. Name it Kickback Valve.

2.

From the compressors Dynamics tab, Specs page, add a Surge Controller.

3.

The output for the Surge controller is Kickback Valve.

4.

Disconnect the compressor discharge stream from the inlet of the compressor discharge valve. Add a tee. The inlet stream to the tee is the compressor discharge stream. Connect the outlets of the tee to the compressor's discharge (TEG Line stream) and kickback valves (Back Line stream).

5.

Connect the outlet of the Kickback valve to the knockout separator.

( i

'--

The surge controller is used exclusively for compressors. When in Auto, the setpoint is calculated by the controller itself; it is not set by the user. The surge curve ofthe compressor model could be explicitly entered as a Speed (rpm) vs. Flow rate curve in a tabular form on RatingIFlow Limits page. Otherwise surge data are implied by the lower ends of the regular curves. For the purpose of the surge controller, the surge data is re-interpreted from the following formula, that is, for any given head at left hand side under current operating condition, the minimal flow (where surging would ta~~pla.£~) is found from right hand side. Note that users need to regress the constants.-~-\

,~

~-

He'id(m) 6.

L--r

~X

{Yr S-[O m3/s )]+ CIO(m3/s)]' + OIQ(m3/s)r

Inspect the compressor curves and fill the cells of the following table, pay attention to unit of measure: Speed

Min Vol Flow (m3/s)

Head (m)

Linear regression of Head vs. the quadratic value of Vol Flow provides parameters A(m) and C(s2/m5) with Band D equal to zero. In addition to the surge line, the user also inputs the control line and the backup line.

9

10

Compressor

Control Line The control line determines the "set point" for the surge controller. This line is set up at some % above the surge flow (typically 10%) and the controller tries to maintain the compressor flow above this control line. If the flow is above the backup line then the control action is the "normal" Pill action. The control line setpoint can be determined by multiplying the flow from the surge line equation by (l OO+x)/l 00, where x is the flow % above the surge line.

Q(m 3 /s)= 100+x ~Head(m)-A 100 C

Backup Line The backup line is located between surge line and the control line. If the process variable (suction flow rate) is above the backup line, then a Pill algorithm is used. However, when the flow goes below the backup line more aggressive control action is taken to prevent surge. The backup line setpoint can be determined by multiplying the flow from the surge line equation by (l OO+x)/1 00, where x is the flow % above the surge line. The relative positions of the 3 lines are illustrated in the following figure. For a certain head, the flow rate drops until it hits control line, then the surge controller starts to open up the surge valve, using the regular Pill mechanism. If flow rate keeps dropping and hitting back up line, the back up control mechanism replaces the regular Pill mechanism and more aggressive action, that is, if the process variable drops below the backup line, then a Quick Opening Algorithm replaces the Pill algorithm, and the recycle valve will be opened by a step of open percentage per second (%/s), (which presumably, is more aggressive).

10

(

c------------------

Compressor

11

Figure 3

Y=774.88X~

Surge Flow (

5000

(

4500

(

4000

..~

3500

+

E3000

, (

"C

~ 2500 2000

Surge Line

-;ss-.

Back Up Line

~. .~

Control Line

-

Linear (Surge Line)

1500

.'

"'-

1000 500

"-

0

,.

0

2

4

6

8

Flow (m6/s2)

7.

Add the control Surge Parameters. A and C can be calculated from the linear regression ofthe above table data.

8.

The control line and back up line are set up at 10% and 5% respectively above the surge line. The recycle valve will be opened by a step of3% per second.

9.

Select kickback valve as Output of the Surge Controller.

10. Add the Pill parameters for the controller and the minimum and maximum, say Kp=O.25, Ti=O.l, PV Max = 20000m3/h.

11

12

Compressor

Figure 4

11. Size the recycle valve such that Cg is 30% of the discharge Cg value. 12'JConstruct a strip chart that shows the actual volume flow for the suction flow . (KO Vap), the net flow through the system (HotVap Out), the control line (= surge controller SP), the back up line and the surge line. The last two will have to be calculated in a spreadsheet. Also display the kickback flow and the surge valve % opening. 13. Decrease the flows of Alpha, Charlie, and Bravo. 14. Observe the valve open as the control line is hit. If the backup line is hit, one should be able to observe the valve opening, go from a Pill control to the Quick Opening. 15. Bring the Set points back to their original values. 16. Save the case as EM6 COMP.hsc.

12

- ----------

- - - ----------------------- - - - -

-----------_. -----_._----------

TEG Dehydration Tower

TEG Dehydration Tower

--/

© 2007 AspenTech - All Rights reserved. EB 11 05.06.04 07_TEGDehydrationTower.doc

1

2

2

TEG Dehydration Tower

I

I

I \,,1

i

(

/

"

II

-"""'ill' TEG Uoe

........

• FC ...

: TEG : • Rec •

TEG reCRecycle

TEc0::::::::::::::::······+f>i~CL from Regen

Towe,

~i~ ....., """;~OTEG Com, Valve

I TEG

~I Tower

.~...

' I,

~·

, '.

~~

pIC-10(:

~

Gas Valve

••.••••••••••••••••v •••••••••v . w•••.

TEG Valve

=&

G~i

TEG

........~~-------.,,""',¥

Sump LC @COL1 :

I \

('f)

4

TEG Dehydration Tower

Workshop The Column Dynamics module introduces you to the process of setting up a distillation column with Aspen HYSYS. Some of the concepts you have learned in the previous modules are applied here. However, the distillation column is one of the most complex unit operations in Aspen HYSYS and deserves special attention. Starting with the previous module, you will expand the Flowsheet by adding a TEG Dehydration Tower to the high-pressure gas stream.

Learning Objectives After you have completed this module, you will be able to: •

Configure a distillation column in steady state

•

Prepare a distillation column for dynamic simulation

Prerequisites Before beginning this section you need to know how to: •

Set up Strip Charts

•

Understand the Flow-Pressure network

4

- - ---~--~

----~~~--~~-

CJ1

I . VD.:'-IID Brooo:

C.arl~·1

\:'LV-toJ

e.;3r1J;:

!Fc-,aarl!~ , ,!"¥ ,,.I>
E!I~(

~m+"""" •.l>l.r..

;

::

VLV-ID'I ~·la"

.,[><1-:

1FC-5r~:~

Al)~al

' 0'.;/+

FC-···l)~a

"."

~U;;-100

T~:l'p

r=::r--r-

-

HPLtl

IUU

L""""

:):.

VLu-l00

-1><1

um;~~~cl

t1e:p

:.. \ ..../ .. LC-HP

kl<>.}.D1!

'cl..:l~

~ HPLkil

PC-HPr<1'

.:...~.... :.

.....

~~\l..:lt!

r-l

~\tat~

%*"11'&:::...

A

H~A)·~~,·· Mf~' 5-11])

' :::

'i~'~~;HO'\

~:::;.;

",'.1 i

\

J..

8

~~

...... l>~1b.m'P ~lr"'l! ~',,*r

I!IDWi cO.1mn~1 ~ o.")!t

I

r ..... LP L~l

I~ ,!,.

~~

t:'~1

1

t~-I

t~l-FCJ .)1} .1>
tfr
",

S.~:rLC

j:J~ . . . LP

TEG Fl:-'

PC

TI;?«'I!'.t

TEG

[

' , ." '; '

"\

')

.'~ ...... \

!

,

6

TEG Dehydration Tower

Adding the Column The remaining moisture in the high-pressure gas stream is removed by contacting the stream with TEG. For this simulation, we are only going to model the TEG dehydration column. The same techniques could be used to add the regeneration column as well. In previous cases, we added to our model while staying in dynamic mode. In this case, we will switch back to steady state mode, make the additions, and then switch back to dynamics. It is not mandatory to go back to steady state to add a column, but it is usually easier to do so because the column internal flows and thermo conditions are readily populated in steady state after it is converged which makes a good starting point for dynamic simulation. The other easier approach is to make a copy of the column inlet streams from the master dynamic case into a temporary steady state case and after finishing the column work there, copy converged column over into the master dynamic case. It is useful when the dynamic case becomes large and one does not want to switch in between steady state and dynamic state any more.

Prepare Model for Steady State Operation 1.

Open the case stored at the end of the last simulation (EM6 Comp.hsc). Make sure it has lined out.

2.

Change the mode from Dynamics to Steady State (use the icons in the tool bar on the Aspen HYSYS desktop). When the mode has changed to Steady State, the Steady State solver will be in the Holding mode.

Can you think ofany changes that need to be made before the Steady State solver is turned on?

Here are some things to check out: In dynamics, we do not need recycle blocks; in steady state, they are required. In our case, the compressor kickback is considered a recycle stream. However, because there is no flow, it does not affect our steady state results.

6

-----

--

~

------~-~-~-- -~~----~-------=- ~-'----"-~~

TEG Dehydration Tower

3.

7

Disconnect the recycle stream where it reenters the KG drum. Figure 1

Back line

The Tees will need some sort of specification around them in steady state, either a specified flow or a specified ratio. 4.

Enter a ratio for the LP Separator tee. For the compressor discharge tee, set the ratio of kickback to O.

There may be some problems around the product streams. Both the pressure drops for the valves and the product pressures are specified. This will lead to consistency errors. f

5.

Remove the pressure drop specifications on the valves.

6.

Check the specifications against those that we originally entered. Make sure the vessels have the correct pressures.

7.

The compressor curves will supply the compressor efficiency. Remove the specified efficiency and duty from the compressor's Design-Parameters tab.

8.

If any pipe is not solved, delete its heat loss value from Rating/heat transfer page.

9.

Turn the steady state solver on. The simulation should converge.

7

8

TEG Dehydration Tower

Adding the TEG Tower After you are satisfied with the steady state solution, add the TEG tower.

Absorber icon

Aspen HYSYS is equipped with a distillation column Input Expert option that allows you to have properly configured the column. The Use Input Expert option can be accessed from the Tools-PreferencesSimulation menu

1.

Enter the Simulation Basis Manager and add a new component, TEG. Return to the Simulation Environment (make sure the solver is on).

2.

Add a new stream to the simulation. Name it TEG from Regen. Supply the following information.

H20

0.07

TEG

0.93

3.

Add a valve to the simulation. Name it TEG rec. The outlet pressure is 6900 kPa (1000 psia). Enter TEG Recycle for the name of the output valve stream.

4.

Add an Absorber to the Flowsheet. Configure the Absorber with the following information.

Name

TEG Tower

Number of Stages

4

Top Stage Inlet Stream

TEG Recycle

Bottom Stage Inlet Stream

ToTEG Tower

Ovhd Vapor Outlet

Dry Gas

Bottoms Liquid Outlet

Btm Liq

Top Pressure

6900 kPa (1000 psia)

Bottom Pressure

6900 kPa (1000 psia)

5.

Run the Column and converge the simulation.

Save your case as EM7_TEGtower_ss.hsc.

\••. · · · .· . ·. . .\/·•1• 11111••1••1• ·• ·• . . 8

TEG Dehydration Tower

9

Sizing the Column Tray Section In steady state mode you are free to specify the column pressure profile. In fact, we just specified a column pressure drop of 0 psi. In dynamics mode, the column pressure profile is calculated by the hydraulic calculations on each stage. Thus the calculated pressure drop on each tray section is a function ofthe tray geometry (diameter, weir height, weir length, and tray spacing). The dynamic column pressure profile of each column in your flow sheet can be estimated with the Tray Sizing utility. The Tray Sizing utility performs the hydraulic calculations on each stage of the column. If you do not know the actual dimensions of your column tray section, the tray section geometry can be used to estimate the size of the tray section.

Running the Tray Sizing Utility To run the Tray Sizing utility, do the following: 1.

From the menu bar, select Tools I Utilities, and from the Available Utilities view, select Tray Sizing. Click the Add Utility button.

2.

Click the Select TS button and choose the tray section to size. In this case, select the TEG Tower Flowsheet and TS-l as the object. Click the OK button.

3.

Click the Add Section button. This adds the selected tray section to the utility, allowing the sizing calculations to be performed.

4.

Select the Performance tab, then the Results page. The hydraulic calculations based on the recommended tray geometry for the column are displayed. This page also displays the estimated Tray Section Pressure Drop. Notice the Tray Sizing utility estimates approximately a 2.2 kPa (0.25 psi) pressure drop through the 4 stages of the column.

When we initially designed our column in Steady State mode, we entered a pressure drop of 0 psi through the column. Additionally, the high-pressure gas was specified to be operating at 1000 psia (6900 kPa). In dynamic mode, the pressure drop at each stage will be calculated based on the tray geometry. Moreover, the feed stream and any connected equipment pressures will be based on this dynamic pressure profile. If the steady state pressure profile does not match the dynamic pressure profile, the column tray pressures and flow rates will oscillate until an equilibrium pressure profile is established. As the column tries to adjust to an equilibrium pressure profile, the column can possibly go unstable. Thus a proper column pressure profile based on the hydraulic calculations should be entered for the column before moving to dynamic simulation analysis.

9

10

TEG Dehydration Tower

Size the Column Tray Section 1.

From the Results page on the Performance tab of the Tray Sizing utility, fill in the appropriate tray dimensions in the table below.

Section Diameter Weir Height Tray Space Weir Length Pressure Drop

10

2.

Return to the Design-Setup tab of the Tray Sizing utility and make the section active by selecting the check box.

3.

Return to the Performance tab-Results page and click the Export Pressures button. This will transfer the pressure profile and the geometric information to the column.

4.

Go into column environment and double-click the tray section icon and go to the Rating tab-Sizing page and Performance tab-Pressure page ofthe column property to verify that the information has been passed.

5.

Click the Non Uniform Tray Data button and change stage 4's type to be Sump and rerun the column. You should have something similar to this:

I~

TEG Dehydration Tower

11

The liquid flow through the column is based on the Frances weir equation. All the parameters for this equation are calculated based on geometry. There are no user inputs. Vapor flow through the column is based on a resistance. The resistances can be based on geometry or back calculated from steady state values. 6.

Go to the Dynamics tab, Specs page. Deselect the Use Tower Diameter check box and then click on All Stages button to calculate k-values (resistances).

If you do not know the dimensions of your process equipment, calculate the vessel size based on an appropriate residence time: •

10 minutes is typically a suitable residence for liquid phase hold ups.

•

2 minutes is typically a suitable residence time for vapour phase hold

Add Valves to the Product Streams 1.

Add a valve to the boundary stream Dry Gas. Name it Gas Valve. Enter 6200 kPa (900 psia) for the outlet valve pressure stream, Gas.

2.

Add a valve to the boundary stream Btm Liq. Name it TEG Valve. Enter 5500 kPa (800 psia) for the outlet valve pressure stream, TEG.

3.

Size the three valves added to the simulation.

4.

Save your case again and move the model to dynamics.

Review the pressure-flow specification; are they correct?

_

Add Control System 5.

Reconnect the compressor kickback stream to the KO Drum Separator.

11

12

TEG Dehydration Tower

6.

Add a flow controller on the TEG from Regen stream, a pressure controller on the DryGas stream and a level controller on the sump, using "Total Liquid Volume Percent" for last stage. Provide appropriate ranges.

7.

Start the Integrator.

8.

Save your case as EM7 TEG.hsc.

Congratulations, you have completed the simulation ofour offshore platform.

12

'-------------------

The Event Scheduler and Spreadsheet

The Event Scheduler, Cause and Effect Matrix, and Spreadsheet

© 2007 AspenTech - All Rights reserved. EB1105.06.04 08_TheEventScheduler.doc

1

2

2

The Event Scheduler and Spreadsheet

,/.--

The Event Scheduler and Spreadsheet

3

Workshop With the Event Scheduler, it is possible to have Aspen HYSYS perform given tasks while a simulation is running in dynamics. The tasks can be triggered by a predetermined simulation time, a logical expression becoming true, or a variable stabilizing to within a given tolerance for a set amount of time. Examples of tasks that can be performed with the Event Scheduler are: emergency shutdown, start-up and some batch processes, among others. The Spreadsheet Operation in Aspen HYSYS allows users to perform calculations that may otherwise not be done in Aspen HYSYS. Virtually any process variable can be transferred into the spreadsheet. Any user specifiable process variable can be calculated and exported. This functionality is particularly useful in Dynamics because it provides an easy way to customize the flowsheet. The cause and effect matrix operation was introduced later than the event scheduler, and it replicates a cause and effect matrix commonly used in designing and operating the safety system of many processing plants. It looks at process values throughout the process and, based upon safety thresholds, determines if certain equipment and/or valves should be shutdown. The unit operation is similar to a spreadsheet. It takes inputs called causes, and sends outputs called effects.

In this workshop, we will work with the Spreadsheet, the cause and effect matrix, and the Event Scheduler Manager.

Learning Objectives Once you have completed this section, you will be able to:

• • • •

Use the Spreadsheet

•

Create Sequences

•

Create Event Conditions

Use the cause and effect matrix Use the Event Scheduler Manager Create Event Schedules

3

.j>.

Process Overview

...... Feed

VLV-100

>1,::>'<1]

FIC-100

Tank

To

.....

Q Fire

»0

L1C-100

Drain

To

V-100

Vent

To

VLV-102

"t

i:~;:~:'=''''''

PIC-100

Drain

Fire Calculations

Vent

Flare

The Event Scheduler and Spreadsheet

5

,~,,----------------------------r'

The Event Scheduler Manager The Event Scheduler Manager contains all the Event Schedules in the current Aspen HYSYS case. Each Schedule is comprised of Sequences, which in turn are made up of Events. An Event must have a Condition, which determines when the Event will be triggered. At least one action must be fully defined for each Event. Once an Event has been triggered, a series of pre-defined Actions will be executed. A variety of Actions can be performed including closing valves, switching controller modes, changing variable values and switching to other sequences. The following chart illustrates the relationship between these various items: Figure 1

Event 1 lr: \M1en Condition Met

.

t

Action A Close Valve

Action B Change Con/roller Mode

Event 2

L

Action C

....

Jump /0 Sequence B ....

The Event Scheduler is accessed from the Simulation Menu Bar, or by using the hot key CTRL E. When you open the Event Scheduler, Aspen HYSYS displays the Event Scheduler Manager, which contains a list of the Schedules in the case.

5

6

The Event Scheduler and Spreadsheet

The buttons in the Schedule Options group are used to manage the schedules for the current case:

6

•

Add - Allows you to add new schedules to the case.

•

Copy - Allows you to make a copy of the selected schedule. This is only active when a schedule exists in the case.

•

Delete - Allows you to delete the selected schedule. This is only active when a schedule exists in the case.

•

Import - Allows you to import a saved schedule from disk. Schedules have the extension .sch.

•

Export - Allows you to export the selected schedule to disk. Once exported, a schedule can be retrieved using the Import button. This is only active when a schedule exists in the case.

•

Sort - Allows you to arrange the schedules. This is only active when at least two schedules exist in the case.

",------------------------------

The Event Scheduler and Spreadsheet

7

The Cause and Effect Matrix The operation's input may be any simulation variable from the users' case or a simple switch which is not required to be connected to an object's variable. Each input generates either a Healthy (1) or Tripped (0) state. The output is a Boolean (lor 0) result from processing one of the Cause and Effect Matrix columns. The output may write or export its result to any simulation variable within the users' case. The user must specify a variable of discrete type (lor 0). The output is not required to be connected to an object or variable. The same 1 or 0 result is produced from the matrix column, and then any other object in the simulation may read or use this value. The matrix is processed one column at a time to determine the resultant state of the output associated with that column. The associated input state is reviewed for each element (or row entry) of a particular column having a non-blank user specified matrix element. All of the matrix elements of that column (and their associated input state) are compared based upon their respective and collective meaning to determine the cause result. Each matrix element type is described in the following table. X

R

T

c

p

S

TRIP One or more zero input(s) causes a zero output. RESET One or more 1 inputs causes a 1 output (as long as there are no X, Tor C active and ALL P must be 1) There is no requirement to have a reset on a particular output. If you want a reset, this can either be done with one or more R matrix element entries or a local reset switch. In the case of both R and a local reset, then both reset features must be reset for the output to return to normal, and the local reset must be done last. TIMED TRIP Same as the TRIP but the input must have remained zero for at least the time period. The T matrix element should be followed by an integer representing the number of seconds of time delayed trip. COINCIDENT TRIP In contrast to all other trips, a zero input for ALL the coincident signals of the same grouping causes a zero output. The C matrix element should be followed by an integer representing the Coincident group number. There should be more than one in each group. PERMISSIVE All P inputs must be 1 to permit an R to have the desired effect. It is also required for a STANDBY 1 effect, a local reset and a local switch ON. INHIBIT A 1 will inhibit any trip of the output which would normally be caused by an X,T or C. STANDBY A 1 will cause a 1 (as long as there are no X, T or C active and ALL P must be i), and a zero will cause a zero output (no INHIBIT applicable). Normally one would not want more than one Standby input designation per output. If you have more than one Standby, ALL Standby inputs must have a 1 for the output to be started (1 result). OtherWise, a zero output result is produced. All Permissive inputs must be 1 for the Standby 1 action to occur.

7

8

The Event Scheduler and Spreadsheet

The Aspen HYSYS Spreadsheet The Spreadsheet's ability to import and export variables means that seamless transfer of data between the Simulation Enviromnent and the Spreadsheet is a simple matter. Any changes in the Simulation Enviromnent are immediately reflected in the Spreadsheet, and vice-versa. The Spreadsheet has several common applications. For example, the Spreadsheet can be used to: •

Transfer variables between flowsheet objects.

•

Perform mathematical operations using variables from the simulation.

Importing and Exporting Variables The contents of any cell in the Simulation Enviromnent can be added to the Spreadsheet. The contents of any Spreadsheet cell can be exported to any specifiable (blue) cell in the Simulation Enviromnent. Note that the contents of any Spreadsheet cell cannot be simultaneously imported and exported. There are three ways of importing values into the Spreadsheet:

8

•

Drag-and-Drop - Position the cursor over the desired item; then press and hold the right mouse button. Move the cursor over to the Spreadsheet. Once over the Spreadsheet, the cursor's appearance will change to a "bull's eye" type. Release the right mouse button when the "bull's eye" cursor is over the desired cell. The specific information about the imported variable will appear in the Current Cell group.

•

Variable Browsing - A variable may also be imported into the Spreadsheet by placing the cursor on an empty cell in the Spreadsheet and pressing (and releasing) the right mouse button. Choose Import Variable from the list that appears, and select the variable using the Variable Navigator.

•

Connections Page - On the Connections page, press the Add Import button and select the desired variable using the Variable Navigator. After selecting the variable, choose the desired cell from the drop-down list.

c (

,,-,.-------------------------(~

The Event Scheduler and Spreadsheet

9

(-

Exporting variables from the Spreadsheet into the Simulation environment is also a simple procedure. The methods for doing this are very similar: •

Drag-and-Drop - Position the cursor over the Spreadsheet cell that is to be exported. Press and hold the right mouse button; the cursor should now change to the ''hull's eye" type. Move the "bull's eye" cursor over to the desired cell. Release the right mouse button and the transfer should be completed.

•

Variable Browsing - A variable may be exported from the Spreadsheet into the Simulation environment by placing the cursor on the exportable cell in the Spreadsheet and pressing (and releasing) the right mouse button. Choose Export Formula Result from the list that appears, and select the desired location

•

Connections Page - On the Connections page, click the Add Export button and select the desired variable using the Variable Navigator. After selecting the variable, choose the desired cell from the Drop Down list.

,/--

The value of any spreadsheet cell can be exported, except if it is an imported value.

Adding Spreadsheet Functions

To view the available Aspen HYSYS functions at any time, click the Function Help button.

The Aspen HYSYS Spreadsheet has extensive mathematical and logical function capabilities. Users familiar with common Spreadsheet programs will inunediately recognize the form of the Aspen HYSYS functions as similar to the form used by these other programs. All functions in the Aspen HYSYS Spreadsheet must be proceeded by either a "= or an "@" depending on the type of function. Equal signs (=) are used for straight mathematical functions: addition, subtraction, multiplication, and division. The ampersand (@) is used before special functions such as logarithmic, trigonometric, and logical functions. Some examples of the Aspen HYSYS functions and their form:

A cell's numerical value can be copied to another cell by using a simple formula (e.g., =A1). Parenthesis is mandatory in many of the advanced spreadsheet functions and can also be used to designate calculation order.

•

Addition - Uses the "=" sign, e.g., =A1+A2

•

Subtraction - Uses the "-" sign, e.g., =AI-A2

•

Multiplication - Uses the "*,, sign, e.g., =Al *A2

•

Division - Uses the "f" sign (not the "\"), e.g., =Al/A2

•

Power - Uses the """ sign, e.g., =A2"4

•

Factorial- Uses the "!" sign, e.g., =A2!

•

Square Root - Uses the "@SQRT" function, e.g., @SQRT(A2)

9

10

The Event Scheduler and Spreadsheet

•

Sine, Cosine, and Tangent - Use the @sin, @cos, and @tan functions, e.g. @sin(A2). Inverse trigonometric functions are also available, @asin, @acos, and @atan. Hyperbolic functions can also be represented in Aspen HYSYS, they use the form @sinh, @cosh, and @tanh.

•

Logarithmic Functions - Represented in Aspen HYSYS with the following forms: @In, @log, and @exp.

•

Pi - Simply enter "=pi" to represent the number 3.1416....

Logical Operators The Aspen HYSYS Spreadsheet supports Boolean logic, essentially a true/false logic. A true statement has a value of 1, and a false statement has a value ofO. For example, suppose that the cell Al has a value of 10, and the cell A2 has a value of 5. If the logical statement +A1
Equal to - Uses "=", e.g., =A1=A2

•

Not Equal to - Uses "!=", e.g., =Al!=A2

•

Greater than - Uses " rel="nofollow">", e.g., =AI>A2

•

Less than - Uses "<", e.g., =A1
•

Greater than or Equal to - Uses " rel="nofollow">=", e.g., =A1>=A2

•

Less than or Equal to - Uses "<=", e.g., =AI<=A2

IFfTHEN/ELSE Statements The Aspen HYSYS Spreadsheet also supports the basic IF/THENIELSE Statement. The form of this statement is: @if(condition) then (iftrue) else (iffalse) The else portion of the statement is not optional. An if then statement is not valid.

10

The condition is a logical expression, such as "B 1<= 10". The iftrue represents what the cell will show if the condition is true; it can be a number or a formula. The if false represents what the cell will show if the condition is false; it can also be a number or a formula.

-

-

-

The Event Scheduler and Spreadsheet

11

An example of the completed statement follows:

@if(Bl<=lO) then (Bl*2) else (Bl/lO) Suppose that the value in cell BI is 8. What will the IF/ THEN/ELSE Statement shown above calculates?

Building the Simulation Open the Aspen HYSYS case, Evt-C&E-Sprdsht-starter.HSC on the disk that was supplied with this module. Under normal operation, the flow into the vessel is controlled by FIC-I00, the level is controlled by LIC-l 00 and the pressure is controlled by PIC-l 00. In this module, a series of events and a cause and effect matrix will be added to simulate a ftre in the vicinity of the unit. When the pressure reaches the set pressure, the relief valve will open to relieve the pressure built up inside the system. Since the fIre is too strong, pressure keeps building up and it causes the closing ofthe control valves. When ftre is put out, the plant is resumed to operate. A spreadsheet will be used to calculate the heat added to the vessel under ftre conditions. The spreadsheet has been included in the starter case but no variables or calculations have been added.

Spreadsheet Variables and Equations 1.

Position the cursor in cell B2 of the Fire Calculations Spreadsheet.

2.

Click the right mouse button.

3.

Select Import Variable from the drop-down menu.

4.

Select V-tOO from the Object list.

5.

Select Vessel Diameter as the Variable.

11

12

The Event Scheduler and Spreadsheet

Spreadsheet formulas can be prefixed with either "+" or "=". Logicals are prefixed with"@".

12

6.

ClickOK.

7.

Import the following variables into the Fire Calculations spreadsheet:

8.

Cell

Object

Variable

B3

V-100

Vessel length or height

B4

RV-100

Set Pressure

B6

To Tank

Temperature

Add the following formulas to the Fire Calculations spreadsheet:

Cell

Formula

Variable Type

Purpose of Formula

B5

=b4-14.696

Pressure

Converts the Set Pressure on the relief valve from absolute to gauge.

B?

=(3. 14*b3*b2)+2*3.14*(b2 A2)/4*1.66

Area

Calculates the wetted area of the vessel.

B8

=21000*(b7 AO.82)

Energy

Determines the amount of energy to be applied to the vessel under fire conditions

02

@if(b9=1)then(O)else(b8)

Energy

This logical operation determines whether the heat should be applied to the vessel.

C

E--

(

C

c-------------------(

c"

The Event Scheduler and Spreadsheet

13

Figure 4

The heat input to the vessel is determined by equation A-2 from API Recommended Practice

521.

9.

The contents of cell D2 need to be exported to Q-Fire.

10. Place the cursor in D2 and press the right mouse button. 11. Select Export Formula Result from the drop-down menu. 12. Select Q-Fire as the object and Heat Flow as the variable.

13. Click OK. 14. Note that the Event Scheduler will change the value shown in cell B9. Ifit is equal to anything other than 1, the energy calculated in cell B8 will be exported to Q Fire via the logical expression in cell D2. Otherwise cell D2 exports a value of zero to Q Fire.

13

14

The Event Scheduler and Spreadsheet

Adding the Cause and Effect Matrix 15. Create a cause and effect matrix, CEM-1. 16. On connections page, click Add Input and select V-I 00, Vessel Temperature, with description as Tank Temperature and tag as HHT-001. 17. Click Add Output, select VLV-I 00, Actuator Failed, with description as "Inlet Valve Shutdown" and tag as "LX-OOI" 18. Go to parameters tab, check Inv column for output. Assign alarm to be 50 and trip to be 60 for input. 19. Go to matrix tab, click Add Switch with description "CEM Reset". 20. On the matrix elements section, put X on row Tank Temperature and R for CEM Reset.

Adding the Event Scheduler 21. Open the Event Scheduler. On the Schedule Options group, click the Add button to create a new Schedule.

It is important to give all objects in the Event Scheduler logical names! This will make the Schedule easier for you and anyone else using your simulation to understand.

14

c (~

(

c-----------------

The Event Scheduler and Spreadsheet

15

22. In the bottom right hand comer of the Schedule view, rename the new Schedule to Fire Relief Scenario. In the Schedule Sequences group, seven buttons allow you to organize all the Sequences for the current Schedule:

•

View - Allows you to view the selected sequence. This is only active when a sequence exists in the schedule.

•

Add - Allows you to add new sequence to the schedule.

•

Delete - Allows you to delete the selected sequence. This is only active when a sequence exists in the schedule.

•

Copy - Allows you to make a copy ofthe selected sequence. This is only active when a sequence exists in the schedule.

•

Import - Allows you to import a saved sequence from disk. Are saved in *.seq format.

•

Export - Allows you to export the selected sequence to disk. Once exported, a sequence can be retrieved using the Import button. This is only active when a sequence exists in the schedule.

•

Sort - Allows you to arrange the sequences. This is only active when at least two sequences exist in the schedule.

The Sequence field contains the name of the Sequence. The Run Mode field displays the sequence mode, which can be either One Shot or Continuous. A One Shot Sequence executes all its Events in order and then its Status will change to Complete. A Continuous Sequence returns to the first Event after the last Event has been executed, in a continuous loop. The Event, Waiting For, and Pending Actions fields display the current event number with its corresponding Condition name and Action List name. 23. On the Schedule view, click the Add button to create a new sequence. 24. Rename the Sequence to Fire Relief.

15

16

The Event Scheduler and Spreadsheet

The Schedule of Events tab shows the list of events for the schedule. The fields on this view are defined as follows: •

Specified - TruelFalse indicates whether the event has been fully defined

•

Condition - Displays the condition name of the event

•

Action List - Shows the name ofthe events action list

•

Jump When - Displays whether the Event will Jump over any Events and, if so, under what condition

•

Jump To - Displays the event to jump to.

At the bottom ofthe view the sequence name, current event number and sequence status are showu. If the event is not fully specified, the Analyze button is enabled and provides additional feedback. On the Settings tab, the sequence Run Mode is selected, a unit set global to the sequence can be chosen and different Status Window Reporting options are available. Selecting the Synchronize All Time Sensitive Conditions check box will assure execution of a particular event at an exact simulation time. This applies only when one ofthe two time conditions has been selected for the Event. Default Time Out Behaviour is specified here and applies to all the events unless a specific event overrides the behaviour.

16

f (( (

(=--------------------

The Event Scheduler and Spreadsheet

17

(

C (

(

The Fire Relief Sequence created in this module will hold 2 Events to be executed at 2 predetermined times. The first one is to start the fire and the second puts out the fire.

Event 1

\

c

Purpose

To start a simulated fire under V-1 00. This Event will change the value of cell 89 such that the logic in the spreadsheet will evaluate to false and the heat flow to V100 will be determined using the equations in the spreadsheet.

Condition Name

Start Fire

Condition

Elapsed Time

Action List Name

Start Fire

Action

Name

Add Heat to Vessel

Type

Specify Variable

Target

Fire Calculations, 89

Value

2

=2 min

Event 2 Purpose

Once the vessel pressure has stabilized, the fire will be extinguished.

Condition Name

Post Fire

Condition

Elapsed Time

Action List Name

Post Fire Actions

Action 1

Name

Fire Off

Type

Specified Variable

Target

Fire Calculations, 89

Value

1

=10 min

25. Once all of the Events have been added, return to the Fire Relief Scenario Schedule view and press the Start button at the bottom of the view. This will activate the selected sequence. You should notice that the status of Fire Relief would change from Inactive to Waiting.

26. Start the Integrator. Be prepared to stop the Integrator as the system runs very quickly. 27. When events finish, the inlet flow rate remains as zero. To resume feed flow, wait for vessel temperature to drop under 60C, and on CEM-I matrix tab, click

17

18

The Event Scheduler and Spreadsheet

on matrix element at CEM Reset row, and then click state on via the radio box within the input section to open up the feed flow. 28. Stop the Integrator once a steady state condition is reached. 29. Observe the strip charts, the controller faceplates and the Trace window to see how the model behaves.

18

le- -

Upstream Options - Hydraulics

Upstream Options · Hydraulics

© 2007 AspenTech - All Rights reserved. EB1105.06.04 09_UpstreamOptions_Hydraulics.doc

1

(

2

Upstream Options - Hydraulics

(

(

2

--------------------------~--~~-~~-~-~--~---==--==~--~- ~,~-

---

-----------

~ ( ( Upstream Options· Hydraulics

3

L \

Workshop This module is optional since Aspen Hysys upstream options require extra upstream dynamics license. One of the major features in upstream option is the Aspen Hydraulics. It is intended for use within the Aspen HYSYS Oil & Gas Option and in particular with the Aspen Dynamic Pipeline Solver embedded within Aspen Hydraulics. The Aspen Dynamic Pipeline Solver is a code for modeling transient multiphase hydrocarbon flows in wells, pipelines, and components. The Aspen Dynamic Pipeline Solver solves mass, momentum and energy equations for each phase using a one-dimensional finite difference scheme. Appropriate flow pattern maps and constitutive relationships are provided for wall and interfacial friction and heat transfer, and a model for multi-component phase-change is included. The module is not intended to go through the construction of Hydraulics sub flow sheet or any other upstream options. Users should go through Aspen HYSYS Upstream course work for that purpose. It only illustrates some of Hydraulics' basic features and how to integrate with rest of Aspen HYSYS model.

Learning Objectives After you have completed this module, you will be able to: •

Configure and use the pigging operations within Hydraulics.

•

Read and analyze the dynamic behavior of Aspen Hydraulics feature in Aspen HYSYS.

•

Use a Hydraulics sub flow sheet within a case.

Prerequisites Before beginning this section you need to know how to: •

Set up Strip Charts

•

Understand the Flow-Pressure network

3

-- ------- -- -- --------------

4

4

Upstream Options - Hydraulics

G (

( Upstream Options· Hydraulics

5

Practice with the Base Hydraulics Case In this section, we are only going to inspect and run the hydraulics model that is shipped with the class. We will configure pigging operations and observe how it affects the liquid holdup and other variables in the pipes. The hydraulics pipes are supposed to replace the 3 pipe segments used in our earlier modules. 1.

Open the case shipped with our class (EM9_Hydro_base.hsc).

2.

Open up PPD and double-click the hydraulics operation AH-100.

3.

On the hydraulics view, show pressure profile across flow sheet. Notice that 3 incoming flows merge into 1 single outgoing flow.

Look at Swage-100, pressure is increases going across it, why?

_

Look at Pipe-102, again, pressure is increasing, why?

4.

Inspect the elevation landscape of the pipe network by looking into each pipe's Design/Data page, pay closer attention to Pipe-104.

5.

Compare pipe-l 04's liquid hold-up in its cells to that of pipe-l 02 and pipe-lOS.

6.

Now go to dynamics tab of the hydraulics window. Click "View Pig Options".

7.

Delete the pig already shown there and add a new one.

8.

On the pigging view, choose option "Pig moves with constant speed".

9.

Select entry Pipe 104.

10. Entry location O. 11. Exit pipe is Pipe 104. 12. Exit location 100. 13. Leave leakage as is. 14. Assign velocity = 0.4. 15. Now the pig should be ready to launch once integrator starts. The window should looks like the following:

5

-----------------

6

Upstream Options - Hydraulics

16. Kick off integrator and view chart "Waves". 17. Click "Launch Pig" on the pigging window. 18. Explain the liquid holdup curves on pipe-104 as you observe along the way. 19. Once the pigging operation is finished, reset the pig and experiment with different speed and different span of the pigging and observe and explain the results. You may need to track other variables in charts along the way.

Replace the 3 Pipelines with the Hydraulics Sub flow-sheet in the Model Now we are going to merge the hydraulics sub flow sheet with the model saved in module 7:

6

1.

Bring up case EM9_Hydro_base.hsc if it is not up yet.

2.

Switch back to steady state and solve it.

3.

Open up case EM7_TEG.hsc

4.

Switch back to steady state as well and resolve the issues just as discussed in module 7 and make sure the case converges.

5.

Bring up PPD view, make sure all the objects are in view and then select all of them, and from Menu select Copy.

6.

Go back to case EM9_Hydro_base.hsc and with PPD showing, select paste from menu. Rearrange the PPD view as you wish.

---------

(

------------------------

Upstream Options - Hydraulics

7.

Make sure it converges and save case as EM9_Hydro_merged.hsc

8.

Make sure ABCD's pressure spec is in blue color. Ifnot, go into hydraulics sub flow sheet and delete the blue number there and retype it into ABCD pressure field.

7

Congratulations, you have just completed the integration between Hydraulics and regular Aspen HYSYS modeL

7

8

8

Upstream Options· Hydraulics

( Basic Control Theory

Basic Control Theory

© 2007 AspenTech - All Rights reserved. EB1105.06.04 10_AppendixA_BasicControITheory.doc

1

2

2

Basic Control Theory

~~~-

----,--

-

--,'-~

-

Basic Control Theory

3

Workshop ( "

Process control, on a working level, involves the control of variables such as flowrate, temperature, and pressure in a continuously operating plant. Process control, in a general sense, attempts to maximize profitability, ensure product quality, and improve the safety and operability of the plant. While a steady state simulation in Aspen HYSYS allows the design engineer to optimize operating conditions in the plant, dynamic simulation allows you to: •

Design and test a variety of control strategies before choosing one that may be suitable for implementation

•

Stress the system with disturbances as desired to test for plant performance

Even after a plant has started operation, process engineers may look for ways to improve the quality of the product, maximize yield, or reduce utility costs. Dynamic simulation using Aspen HYSYS allows the process engineer to compare alternative control strategies and operating schemes in order to improve the overall performance of the plant. In short, the engineer can accomplish a lot of analysis off-line on a dynamic simulator, instead of disturbing the actual process. Three topics will be covered in this module. First, the characteristic parameters of a process will be discussed in the Process Dynamics section. Next, the control strategies available in Aspen HYSYS will be discussed in the Controller Setup section. Finally, the General Guidelines section will outline some steps you can follow to implement a control strategy in Aspen HYSYS. Included in this section are several techniques that may be used to determine possible initial tuning values for the controller operations.

Learning Objectives In this module, you will: •

Learn the basics of Process Control Theory

•

Explore the development of Control Strategies

•

Examine General Guidelines for implementing appropriate control strategies.

3

4

Basic Control Theory

Process Dynamics As a precursor to understanding the concepts of process control, the dynamic characteristics of the process will be discussed. The task of designing a control scheme is best carried out ifthere is a good understanding of the process system being studied. A process response to a change can vary considerably depending on the manner in which the input is applied to the system, and the nature of the system itself. Therefore, it is important to understand the dynamic characteristics of the process system before proceeding with the process control design. Many chemical engineering systems are non-linear in nature. However, it is convenient to define some essential characteristic parameters of a process system by approximating the system as linear.

Characteristic Parameters of the Process System It is easiest to define a chemical process system using the general conservation principle that states:

Rate of Accumulation = Input - Output + Internal Generation

(1)

In order to describe some characteristic parameters of a chemical process system, the general conservation principle is applied to a flow relation first-order liquid level system: Figure 1

H

4

Basic Control Theory

5

The conservation of material in the tank is expressed as follows:

dV

dt where:

A dH dt

F.-F I 0

(2)

V = the liquid volume in the tank H = the liquid height in the tank A = the cross-sectional area of the tank, assumed constant F i = the inlet flow rate

F o = the exit flow rate

There is a non-linear relationship describing the flow out of the bottom of the tank, F0' and the liquid height in the tank, H. In order to express Equation 1 as a first order linear differential equation, it must be assumed that the exit flow varies linearly with height. Linearity can be assumed in situations in which the flow does not vary considerably over time. The exit flow, F o' can be expressed in terms ofthe linearity constant, R (the valve resistance):

H

(3)

R Equation 2 can therefore be expressed as:

H F.-I R

RA dH +H dt

RF.I

(4)

(5)

Equation 4 is a general first-order differential equation, which can be expressed in terms of two characteristic parameters: the steady-state gain, K, and the time constant, 't:

r;dY + yet) = Ku(t) dt where: y(t)

=

(6)

the output of the system

u(t)= input to the system K

=

1: =

the steady-state gain the time constant of the system

5

6

Basic Control Theory

The change in liquid level, H, is the change in the output of the system, y(t). The change in the input to the system, u(t), is the change in flow into the tank, Fi . Similarly, the time constant, '1:, and the steady state gain, K, can be expressed as:

AR and K

't =

=

R

(7)

When a step function of magnitude U is applied to the general fIrst-order system, the output response, y(t), is as follows: Figure 2 A

If, II'<

e

: 6J %

Process V allahl e

APV:

of APV

:

-~'-------'-l~-.---- ---- ---¥--- ------------ -------- ----- ---- ---- ---- i __ ---- ---- -----._. - .. .... ~-

~-_

~

J

l' A. U (Coniroollir

o...ip...t)

llalw

i

-----------------------------------------------------------------*--------------TIME

As shown, the output, y(t), attains 63.2% of its fInal steady state value in one time constant, '1:. The Process Variable (PV) can be assumed to equal its fInal value after approximately four time constants (4'1:) have passed.

6

-

-

-

-------".~"

\"-

,-

'

------------------------------

Basic Control Theory

7

The dead time of the process is represented by the Greek letter, 8. The dead time is defined as the amount of time that passes between the time of the change in the Controller Output (U), and the time that the fIrst change is seen in the Process Variable (PV). In the flow example given above, the dead time will be virtually nonexistent; however, it can become significant for other systems. The following is a list of characteristic parameters that may be defined in terms of the first-order response illustrated in the previous example.

Process Gain The process gain is defIned as the ratio of the change in the process output to the change in the process input. The change in the process input is defined in Equation 6 as u(t). The change in the process output is defined as yet). The fIrst term in Equation 5 is transient and becomes zero at steady state. Therefore, the gain can be calculated as shown in Equation 8.

Steady-state gain

y(t) u(t)

K

(8)

Time Constant The time constant, 't, defines the speed ofthe response. The response ofthe system will always follow the proftle shown on the previous page. After 't time units, the response yet) equals 0.632 ~PV or 63.2% ofthe fInal PV value. This will always be true for first-order systems.

Dead Time While capacitance is a measure of how fast a system responds to disturbances, dead time is a measure of the amount of time that elapses between a disturbance to the system and the observed response in the system. Time delays in a system can become significant depending on the nature of the process and the location of measuring devices around the process. It is usually the time associated with the transport of material or energy from one part of the plant to another that contributes to time delays observed in a system. The dead time of a process may be easily modelled using the Transfer Function block operation.

7

8

Basic Control Theory

Capacity Definition 1 Capacity can be defined simply as the volume or storage space of a system. The capacitance of a system dampens the output causing the response to take time to reach a new steady state. For electrical systems, the capacity is defined in terms of the resistance ofthe system and the time constant of the response:

c=-R't

(9)

Since the capacity of a system is proportional to the time constant, 'r, it can be concluded that the larger the capacity, the slower the response of the system for a given forcing function. In first order systems, the capacity of a system has no effect on the process gain. However, the capacity varies in direct proportion with the time constant of a system.

Definition 2 A system's capacity is also defined as its ability to attenuate an incoming disturbance. Attenuation is defined as: Attenuation

=

Attenuation =

8

I _ Response Amplitude out of the system Disturbance Amplitude into the system I-Amplitude Ratio

(10)

( f \

,-

'-------------------------------

Basic Control Theory

9

Controller Setup The PID Controller operation is the primary tool that you can use to manipulate and control process variables in the dynamic simulation. You can implement a variety of feedback control schemes by modifying the tuning parameters in the Pill Controller operation. Tuning parameters can be modified to incorporate proportional, integral, and derivative action into the controller. A Digital On/Off control operation is also available. Cascade control may be modelled using interacting Pill Controller operations. Feedforward control can be implemented into the simulation model using the Spreadsheet operation. Instrumentation dynamics can also be modelled in Aspen HYSYS, increasing the accuracy of the simulation with real valve dynamics. Final control elements can be modelled with hysteresis (lag). The valve response to controller input can be modelled as instantaneous, linear, or first order. Dead time, lags, leads, whether they originate from disturbances or within the process control loop may be modelled effectively using the Transfer function operation.

Terminology Before reviewing the major control operations that are available in Aspen HYSYS, it is useful to describe some terms.

Disturbances A disturbance upsets the process system and causes the output variables to move from their desired setpoints. Disturbance variables cannot be controlled or manipulated by the process engineer. The control structure should account for all disturbances that can significantly affect a process. The disturbances to a process can . either be measured or unmeasured.

Open Loop Control Varying the input to a system and measuring the output's response determine an open loop response from a process. An open-loop system is shown below. In openloop control, the controller sets the input to the process without any knowledge of the output variable that closes the loop in feedback control schemes.

9

10

Basic Control Theory

Figure 3

CONTROLLER

I I I I I

., INPUT

PROCESS

OUTPUT

Control Valve

OPEN LOOP CONTROL

A common example of open-loop control is the control of traffic in a city. The traffic lights change according to a set of predetermined rules.

10

c c c-----------------

11

Basic Control Theory

c

Feedback Control (Closed Loop) Feedback control is achieved by "feeding back" process output information to the controller. The controller makes use of the current information about the process variable in order to determine what action to take to regulate the process variable. This is the simplest and most widely used control structure in chemical process systems. Figure 4

CONTROLLER < C - - - - - - - - - - - l

I I

Disturbance

I

I

,..I INPUT

PROCESS Control Valve

I I I I I I I

.J.. OUTPUT

-

Sensor

CLOSED LOOP CONTROL Feedback control attempts to maintain the output variable, PV, at a user-defmed setpoint, SP. There are some basic steps that are carried out by the controller in order to achieve this task: 1.

Measure the output variable, PV.

2.

Compare the measured value, PV, with the desired setpoint value, SP. Calculate the error, E(t), between the two values. The definition of error depends on the whether the controller is direct or reverse-acting.

3.

Supply the error, E(t), to the general control equation. The value of the desired percent opening ofthe control valve, OP%, is calculated.

4.

The value ofOP% is passed to the final control element which determines the input to the process, D(t).

5.

The entire procedure is repeated.

11

12

Basic Control Theory

The general control equation for a PID controller is given by:

OP(t) = KcE(t) + ~~ JE(t)dt+KcTd d~~t)

(11)

1

where:

OP(t) = Controller output at time t E(t) = Error at time t

Kc = Proportional gain of the controller Ti = Integral (reset) time ofthe controller Td = Derivative (rate) time of the controller Figure 5

Disturbance

SP +

PV

FEEDBACK CONTROL

12

r-c= (

c c----------------c

Basic Control Theory

13

Direct and Reverse Acting The input to the feedback controller is called the error. It is the difference between the output process variable and the setpoint. The error is defmed differently depending on whether the process has a positive or negative steady state gain. For a process with a positive steady state gain, the error should be defmed as reverse acting.

E(t) where:

SP(t) - PV(t)

(12)

SP(t) = setpoint PV(t) = measured output process variable

That is, if the PV rises above the SP, the OP, or input to the process, decreases. Ifthe PV falls below the SP, the OP increases. For a process with a negative steady state gain, the error should be set as direct acting:

E(t)

=

P V(t) - SP(t)

(13)

That is, if the PV rises above the SP, the OP, or input to the process, increases. If the PV falls below the SP, the OP decreases. A typical example of a reverse-acting controller is in the temperature control of a Reboiler. In this case, as the temperature in the vessel rises past the SP, the OP decreases, in effect closing the steam valve and reducing the flow of heat.

":.

. ~

......:..;:.

Think about the tank example presented at the beginning ofthis module. Assume that the flow out ofthe tank is controlled at a constant value and a PID controller is used to keep the level in the tank at a certain SP by opening or closing a valve on the inlet stream. Should the controller be Reverse or Direct acting?

13

14

Basic Control Theory

Stability The stability of a system is a very important aspect to consider when designing control schemes. Many systems have oscillatory responses, depending on its controller tuning parameters. When a process is upset by a bounded disturbance or bounded change in the input forcing function, the output typically will respond in one of three ways: 1.

The response will oscillate with decreasing amplitude and eventually reach steady state and stabilize.

2.

The response will oscillate continuously with constant amplitude.

3.

The response will grow continuously and never reach steady state conditions. Figure 6

yet)

t. time

Stability Response Cases The system is generally considered stable if the response proceeds to a steady state value and stabilizes. It is considered unstable if the response continues to fluctuate. A stable open-loop response is said to be self-regulating. If the open loop response of a system is not stable, it is said to be non-self-regulating. For instance, a pure integrating process, such as a tank with a pumped (constant) exit flow, is non-selfregulating since a bounded increase in the flow input to the system from steady state will result in the response (liquid height) to increase continuously.

14

-

-

-

--:-:;:::-----~~.

--

(-

G (-

(

c- - - - - - - - - - - - - - - c

c ( (

Basic Control Theory

15

A prerequisite for closed-loop control is that the closed-loop response is stable. The closed-loop response can vary considerably depending on the tuning parameters used in the feedback control equation. In general, a higher controller gain gives tighter control. However, the value ofKc cannot increase indeftnitely. The response will remain stable up to a certain value ofKc. Increasing Kc beyond the stability limit will cause the closed-loop response to become unstable.

"

,(

A number of factors can affect the stability of a closed-loop system: •

Tuning parameters

•

Non-linearities in the process

•

Range and non-linearities in the instruments

•

Interactions between control loops

•

Frequency of disturbance

•

Capacity of process

•

Noise in measurement of process variables

Available Control Operations Modelling Hardware Elements Modelling the hardware elements of the control loop may simulate the plant more accurately. Non-linearities may be modelled in the VALVB operation in the Actuator page ofthe Dynamics tab.

Sensors Sensors are used to measure process variables. In Aspen HYSYS, the sensing instrument is incorporated directly in the PID Controller operation. You can choose the range of the sensing instrument in the Minimum and Maximum PV parameters in the controller operation. It is assumed in Aspen HYSYS that the PID controller is perfectly accurate in its measurement of the process variable.

15

16

Basic Control Theory

Final Control Element - Valve Dynamics You have the option of specifying a number of different dynamic modes for the valve. If valve dynamics are very quick compared to the process, the instantaneous mode may be used. The following is a list of the available dynamic modes for the VALVB operation: Valve Mode

Description

Instantaneous

In this mode, the actuator moves instantaneously to the desired OP% position from the controller.

First Order

A first order lag can be modelled in the response of the actuator position to changes in the desired OP%. The actuator time constant can be specified in the Parameters group. Similarly, a first order lag can be modelled in the response of the actual valve position to changes in the actuator position. The valve stickiness time constant is specified in the Parameters group. In effect, a second order lag can be modelled between the valve position and the desire OP%.

Linear

The actuator can be modelled to move to the desire OP% at a constant rate. This rate is specified in the Parameters group.

Final Control Element - Valve Type The flowrate through a control valve varies as a function of the valve percent opening and the Valve Type. Valve type may be defined more easily by expressing flow as a percentage, Cy (0% representing no flow conditions and 100% representing maximum flow conditions). The valve type can then be defined as the dependence on the quantity of %Cy as a function of the actual valve percent opening. There are three different valve characteristics available in Aspen HYSYS. The valve types are specified in the Ratings tab in the Valve Type and Sizing Methods group.

16

Valve Type

Description

Linear

A control valve with linear valve characteristics has a flow, which is directly proportional to the valve % opening.

Quick Opening

A control valve with quick opening valve characteristics obtains larger flows initially at lower valve openings. As the valve opens further, the flow increases at a smaller rate.

Equal Percentage

A control valve with equal percentage valve characteristics initially obtains very small flows at lower valve openings. However, the flow increases rapidly as the valve opens to its full position.

( ( Basic Control Theory

17

The valve characteristics are shown graphically as follows: Figure 7 CONTROL VALVE FLOW CHARACTERISTICS

100 r-------------:::::::====7f~

80

Linear

60 %Cv

40

Equal Percentage

20

o o

20

40

60

80

100

% Valve Position

Feedback Control Digital On/Off Digital On/Off control is one ofthe most basic forms of regulatory control. In Aspen HYSYS, it is implemented using the DIGITAL POINT operation. An example of On/Off control is a home heating system. When the thermostat detects that the temperature is below the setpoint, the heating element turns on. When the temperature rises above the setpoint, the heating element turns off. Control is maintained using a switch as a final control element (FCE). On/Off control parameters are specified in the Parameters page of the DIGITAL POINT operation in Aspen HYSYS. If the OP is ON option is set to "PV < Threshold", the controller output turns on when the PV falls below the setpoint. This is similar to the thermostat example given above

OP = 0% for PV> SP or OP = 100% for PV < SP

(14)

17

18

Basic Control Theory

The opposite is true when the OP is ON option is set to "PV > Threshold". This setting can be used for pressure relief valves; the valve is open (on) when the PV is greater than the threshold pressure.

OP

=

O%forPV<SP or OP

= lOO%forPV>SP

(15)

One main characteristic of the On/Off controller is that the PV will always cycle about the setpoint. Figure 8

IOO%J------,

OP 0% I - - - - - ' ' - - - - - - ' - - - - - - L - - - - ' - - - - J . -

_

PV

t, time

On-Off Controller Response

The cycling frequency will depend on the dynamics of the process. Those systems with a large capacity (large time constant) will cycle less frequently. The On/Off controller is an appropriate controller ifthe deviation from the setpoint is within an acceptable range and the cycling does not destabilize the rest of the process.

18

L G

c

C

C,:------------------

Basic Control Theory

19

(Proportional Control (P·only)

( (

(

(

C

Leaving the values for Td and Ti as <empty> will also result in P-only control.

Unlike On/Off control, proportional control can damp out oscillations from disturbances and stop the cycling ofthe process variable. P-only control is implemented in Aspen HYSYS by setting the value of Td to zero and the value for T; to a large value (1000*Kc) in the PID Controller operation. With P-only control, oscillations that occur in the process variable due to disturbances or changes in the setpoint dampen out the quickest (have the smallest natural period) among all other simple feedback control schemes. The output of the proportional control is defmed as:

( \

(16) ( \

The value of the bias, OP ss, is calculated when the controller is switched to Automatic mode. The setpoint is defaulted to equal the current PV. In effect, the error becomes zero and OP ss is then set to the value of OP(t) at that time. A sustained offset between the process variable and the setpoint will always be present in this sort of control scheme. The error becomes zero only if:

1"---

•

The bias, OPs" equals the operating variable, OP

•

Kc becomes infinitely large

However, Kc cannot practically become infinitely large. The magnitude ofKc is restricted by the stability of the closed-loop system. In general, a higher controller gain gives tighter control. However, the value ofKc cannot increase indefinitely. The response will remain stable up to a certain value of Kc. Increasing Kc beyond the stability limit will cause the closed-loop response to become unstable.

19

20

Basic Control Theory

The following shows the effect of the magnitude ofKc on the closed loop response of a first order system to a unit step change in the setpoint. Figure 9

1.0

Offset

Increasing Kp

Yet)

t, time

Closed loop response of a system under P-Only Control

Proportional only (P-only) control is suitable when a quick response to a disturbance is required. P-only control is also suitable when steady-state offsets are unimportant, or when the process possesses a large integrating process (has a large capacity). Many liquid level control loops are under P-only control. If a sustained error is undesirable, integral action is required to eliminate the offset.

20

(~

G;.. ( (

c----------------

Basic Control Theory

21

c (

Proportional + Integral Control (PI)

c

Unlike P-only control, proportional + integral control can dampen out oscillations and return the process variable to the setpoint. Despite the fact that PI control results in zero error, the integral action of the controller increases the natural period of the oscillations. That is, PI control will take longer to line out (dampen) the process variable than P-only control. The output of the proportional controller + integral controller is defined as:

(

(

c ( (

OP(t)

(

= KcE(t) + ~~

fE(tJdt

(17)

1

The integral term serves to bring the error to zero in the control scheme. The more integral action there is, the slower the response of the controller will be. The integral term continuously moves to eliminate the error. The closed loop response of a process with PI control and P-only control is shown as follows: Notice that the time it takes to reach steady state is no longer for the PI controller. The integral action slows the controller's response.

Figure 10-

y(t) P-Only Control

t, time

Proportional and PI Control

21

22

Basic Control Theory

The integral time, Ti> is defmed as the amount oftime required for the controller output to move an amount equivalent to the error. Because the relationship between T j and the control action is reciprocal, increasing T j will result in less integral action, while decreasing T j will result in greater integral action. The integral time should be decreased (increased integral action) just enough to return the process variable to the setpoint. Any more action will only serve to lengthen the response time. Due to the reciprocal effect, setting T; to zero means that there will be infinite integral effect. To minimize the integral effect set T; to a large value (1000*Kc).

PI control is suitable when offsets cannot be tolerated. The majority of controllers in chemical process plants are under PI control. They combine accuracy (no offset) with a relatively quick response time. However, the added integral action acts as a destabilizing force, which may cause oscillations in the system and cause the control system to become unstable. The larger the integral action the more likely it will become unstable.

Proportional Integral Derivative Control (PID) If the response of a PI controller to a disturbance is not fast enough, the derivative action in a PID controller can reduce the natural period of oscillations even further. By measuring the rate of change in error, the controller can anticipate the direction of the error and thus respond more quickly than a controller without derivative action. The output of the proportional + integral + derivative controller is defined as:

OP(t) = KcE(t) + K c fE(t)dt + KcTd dEd(t)

T.1

t

(18)

T d is defined as the time required for the proportional action to reach the same level as the derivative action. It is, in effect, a lead term in the control equation. For a ramped input, the proportional only response will be ramped, as well. For the same ramped input the derivative only response will be constant. As the slope of the measured error increases to infinity, so does the derivative action. While a perfect step change with a slope of infmity in either the setpoint or the measured process variable is not physically possible, signals which have short rise and fall times can occur. This adversely affects the output of the derivative term in the control equation, driving the controller response to saturation. Derivative action cannot be used in systems where the PV signal will contain noise.

22

(

c

c---------------c ( ( ( (

Basic Control Theory

23

Derivative action control is best for processes that have little or no dead times and large capacities. Processes such as these, having large lags, benefit from the additional response speed that derivative action provides. While the integral term in Pill control schemes reduces the error to zero, it also adds a considerable lag to the response compared to P-only control. It is the derivative action in Pill control that shortens the controller's response to be comparable to the response of a P-only controller. However, if a controller has a very noisy input that cannot be filtered or minimized in the process, Pill control is not a suitable control scheme.

c (

Figure 11

(

"

(

c (

"-. (

'-

Notice here that the time to steady state is shorter for the PIO controller as compared to the PI control. This is due to the derivative action.

yet)

t. time

PI and PID Control

Feedforward Control Feedforward control may be used in cases for which feedback control cannot effectively control a process variable. The main disadvantage of feedback control is that the controller must wait until disturbances upset the process before responding. With feedforward control, the controller can compensate for disturbances before the process is affected. Cascade control is useful when measured disturbances significantly affect the input to a process. On the other hand, feedforward control is useful if there are measured disturbances that affect the output ofthe process. With feedback control, the controller requires information about the controlled process variable(pV) and the setpoint (SP), in order to determine the value ofthe desired valve percent opening (OP%) ofthe input to the process. In order to determine the value of OP%, the feedforward controller requires information from two variables: the setpoint ofthe process variable (SP) and the disturbance affecting the process. A steady-state process model is used in the feedforward controller to determine the value of OP%.

23

24

Basic Control Theory

Figure 12

Disturbance

I ~

Feedforward Controller

~

Final Control Element

I

ProcessD

+

IU(t)

.. Process

~PV '<..Y

FEEDFORWARD CONTROL

Consider an example of a liquid stream being heated in a steam heat exchanger. Figure 13

Exit

T2 F

..

T1

Condensate

Steam

Feedforward Control of a Heat Exchanger

24

(~

~:

C

C

c---------------c( (

( (

Basic Control Theory

It is desired to control the Exit stream temperature, T2, at a certain setpoint, SP, using

the Steam flow as the manipulated variable. However, the process suffers from frequent changes in the Feed temperature, T 1. In order to detennine the value of OP%, the values of SP and T 1 are required by the controller. At steady state, the overall energy balance relates the steam flow to the disturbance of the process, T 1, and the temperature of stream Exit, T2:

c

o

( ( (

"

c

25

(19)

where: F s = the steam flow

A = the heat of condensation for steam F = the flow of stream Exit Cp = the specific heat of stream Exit

From this process model, the desired value of steam flow into the heat exchanger can be calculated. The flow of steam must be calculated such that the temperature of stream Exit, T2, equals the desired temperature, SP. Therefore, Equation 19 becomes: (20)

In order to calculate the feedforward controller output, a linear relation is assumed to exist between the steam flow and the valve opening ofthe steam valve. Therefore, the final form of the feedforward controller equation is:

oP(t)

C -EF(SP _ T )steam valve span A

1

(21)

100%

25

26

Basic Control Theory

There are some points to consider in order to successfully implement a feedforward control system: 1.

It cannot be implemented if the disturbance is not measurable. Ifunexpected disturbances enter the process when pure feedforward control is used, no corrective action is taken and the errors will build up in the system.

2.

A fairly accurate model of the system is required.

3.

The feedforward controller contains the reciprocal of the process model. Even if the process model is accurate, a time delay in the process model implies that a predictor is required in the feedforward controller. Unfortunately, it is impossible to predict the nature of disturbances before they occur.

It is important to note that the process variable to be controlled is not measured using feedforward control. There is no way of confmning that the process variable is attenuating disturbances or maintaining a desired setpoint. Considering that an accurate model of the process is usually not available, that the process or valve dynamics are not accounted for in this control scheme, and that the valve opening percent is not related linearly to the flow in most dynamic simulation applications, there will probably be an offset between the actual controlled variable and its desired setpoint. Therefore, feedback control is often used in conjunction with feedforward control to eliminate the offset associated with feedforward-only control.

Feedforward control in Aspen HYSYS can be implemented using the Spreadsheet operation. Variables can be imported from the simulation flowsheet. A feedforward controller can be calculated in the spreadsheet and the controller output exported to the main flowsheet. Ifthe operating variable, OP, is a valve in the plant, the desired controller output calculated by the Spreadsheet should be exported to the Actuator Desired Position of the valve.

26

(

c

c-----------------

c (

Basic Control Theory

27

General Guidelines

(

( (

c ( (

c

Effect of Characteristic Process Parameters on Control The characteristic parameters of a process have a significant effect on how well a controller is able to attenuate disturbances to the process. In many cases, the process itself is able to attenuate disturbances and can be used in conjunction with the controller to achieve better control. The following is a brief discussion outlining the effect of capacity and dead time on the control strategy of a plant.

(

Capacity The ability of a system to attenuate incoming disturbances is a function of the capacitance ofa system and the period of the disturbances to the system. Attenuation is defined as:

Attenuation = 1 _

K J(ffi'T;)2 + 1

(22)

The time constant, 't, is directly proportional to the capacity of a linear process system. The higher the capacity (time constant) is in a system, the more easily the system can attenuate incoming disturbances since the amplitude ratio decreases. The frequency of incoming disturbances affects the system's ability to attenuate these disturbances. High-frequency disturbances are more easily attenuated than lowfrequency disturbances.

27

28

Basic Control Theory

Dead Time The dead time has no effect on attenuating disturbances to open loop systems. However, it does have a significant negative effect on controllability. Dead time in a process system reduces the amount of gain the controller can implement before encountering instability. Because the controller is forced to reduce the gain, the process is less able to attenuate disturbances than the same process without dead time. Tight control is possible only if the equivalent dead time in the loop is small compared to the shortest time constant of a disturbance with a significant amplitude. It is generally more effective to reduce the dead time of a process than increase its

capacity. To reduce dead time: •

Relocate sensor and valves in more strategic locations

•

Minimize sensor and valve lags (lags in the control loop act like dead time)

To reduce the lag in a system and therefore reduce the effects of dead time, you can also modify the controller to reduce the lead terms to the closed-loop response. This can be achieved by adding derivative action to a controller. Other model-based controller methods anticipate disturbances to the system and reduce the effective lag of the control loop.

Choosing the Correct Controller You should consider what type ofperformance criteria is required for the setpoint variables, and what acceptable limits they must operate within. Generally, an effective closed loop system is expected to be stable and cause the process variable to ultimately attain a value equal to the setpoint. The performance of the controller should be a reasonable compromise between performance and robustness. A very tightly tuned or aggressive controller gives good performance but is not robust to process changes. It could go unstable if the process changes too much. A very sluggishly-tuned controller delivers poor performance but will be very robust. It is not likely to become unstable. •

If an offset can be tolerated, a proportional controller should be used.

•

If there is significant noise, or if there is significant dead time and/or a small capacity in the process, the PI controller should be used.

•

If there is no significant noise in the process, and the capacity of the system is large and there is no dead time, a PID controller may be appropriate.

It is apparent why the PI controller is often the most common controller found in a plant. There are three possible conditions that a PI controller can handle, whereas the Pill controller requires a specific set of conditions in order to be used effectively.

28

c{;.-

c (

c------------------

Basic Control Theory

29

F""-

l-

(

Choosing Controller Tuning Parameters

(

The following is a list of general tuning parameters appropriate for various processes l . Keep in mind that there is no single correct way of tuning a controller. The objective of control is to provide a reasonable compromise between performance and robustness in the closed loop response.

( (

( (

The following rules are approximate. They will provide you very close to tight control. You can adjust the tuning parameters further if the closed loop response is not satisfactory. Tighter control and better performance can be achieved by increasing the gain. Decreasing the controller gain results in a slower but more stable response.

c c

Generally, proportional control can be considered the principal controller. Integral and derivative action should be used to trim the proportional response. Therefore, the controller gain should be tuned fIrst with the integral and derivative actions set to a minimum. If instability occurs, the controller gain should be adjusted first. Adjustments to the controller gain should be made gradually. Typical Controller Tuning Parameters:

These values are estimates only. They will generally provide adequate. control under most circumstances.

System

Kc

Ti (minutes)

Td (minutes)

Flow

0.1

0.2

0

Level

2

10

0

Pressure

2

2

0

Temperature

1

20

0

Flow Control Most controllers must be Otuned by taking the control loop dynamics into account.

Since the flow control is fast responding, it can be used effectively as the secondary controller in a cascade control structure. The non-linearity in the control loop may cause the control loop to become unstable at different operating conditions. Since flow measurement is naturally noisy, derivative action is not recommended.

Liquid Pressure Control The liquid pressure loop is typically very fast. The process is essentially identical to the liquid flow process except that liquid pressure instead of flow is controlled using the fInal control element. The liquid pressure loop can be tuned for PI control, depending on your performance requirements.

29

30

Basic Control Theory

Liquid Level Control Liquid level control is essentially a single dominant capacity without dead time. In some cases, level control is used on processes which are used to attenuate disturbances in the process. In this case, liquid level control is not as important. Such processes can be controlled with a loosely tuned P-only controller. If a liquid level offset cannot be tolerated, PI level controllers should be used. There is some noise associated with the measurement oflevel in liquid control. If this noise can be practically minimized, then derivative action can be applied to the controller. It is recommended that Kc be specified as 2 and the bias term (OP ss ) be specified as 50% for P-only control. This ensures that the control valve is wide open for a level of75% and completely shut when the level is 25% for a setpoint level of 50%. If PI control is desired, the liquid level controller is typically set to have a gain, Kc, between 2 and 10. The integral time, Tj, should be set between 1 and 5 minutes.

Gas Pressure Control Gas pressure control is similar to the liquid level process in that it is capacity dominated without dead time. Varying the flow into or out of a vessel controls the vessel pressure. Because of the capacitive nature of most vessels, the gas pressure process usually has a small process gain and a slow response. Consequently, a high controller gain can be implemented with little chance of instability.

Temperature Control PI controllers are widely used in industry; however Pill control can be used to improve the response time if the loop is slow.

30

----- --(~

(;;:-

C

C (,-----------------

Basic Control Theory

31

C

C C ( ( ( ( (

C (

(

C (

Tuning Methods An effective means of determining controller tuning parameters is to bring the closed-loop system to the verge of instability. This is achieved by attaching a P- only controller and increasing the gain such that the closed-loop response cycles with a constant amplitude. At a system's stability margins, two important system parameters, the ultimate period (Pu) and the ultimate gain (Ku), allow the calculation of appropriate proportional gain, integral time, and derivative time values.

ATV Tuning Technique The ATV (Auto Tuning Variation) technique is used for processes which have significant dead time. A small limit cycle disturbance is set up between the manipulated variable (OP%) and the controlled variable (PV). The ATV tuning method is as follows: I.

Determine a reasonable value for the OP% valve change (h = fractional change in valve position).

2.

Move valve +h%.

3.

Wait until process variable starts moving, and then move valve -2h%.

4.

When the PV crosses the setpoint, move the valve position +2h%.

5.

Continue until a limit cycle is established.

6.

Record the amplitude of the response, A, expressed as a fraction of the PV span. Figure 14 ............ -..... --. .................. - .. .... . _-- ........ .- ... ---- ---- -_ ........

I

I

Controller Output, OP

~

I I

/

~I

!I

/

I I

Ultimate Period:

14-it'

,

/ ./

/

\

\

V

\

/

"-

/

I_ f.!

/ V\

I'u

V'\

r.J

\

\

1/

1\

Ultimllte Gain: Kt>.°o 4h!{J 1411)

"1\

\

I

\

\

[\ / :; Process Variable, PV

t, time

ATV Stability Limit Parameters

31

---------_.--_

32

.._~

Basic Control Theory

7.

The tuning parameters are calculated as follows: Tuning Parameter

Equation

Ultimate Gain

K u

Ultimate Period

Pu

=

=

4h

na

Period taken from limit cycle

Ku K c = 3.2

Controller Gain

Controller Integral Time

Ti = 2.2 x P u

Ziegler-Nichols Tuning Technique The Ziegler-Nichols2 (Z-N) tuning method is another method which calculates tuning parameters. The Z-N technique was originally developed for electromechanical system controllers and is based on a more aggressive "quarter amplitude decay" criterion. The Z-N technique can be used on processes without dead time. The procedure is as follows: 1.

Attach a proportional-only controller (no integral or derivative action).

2.

Increase the proportional gain until a limit cycle is established in the PV.

3.

The tuning parameters are calculated as follows: Tuning Parameter

Equation

Ultimate Gain

K u = Controller gain that produces limit cycle Ultimate Period

p u = Period taken from limit cycle Controller Gain

Controller Integral Time

32

I

~~ __ ._---

Ku K - 2.2 c

Ti = P u /1.2

----~

-

---_._--- -_._.....

C

G ( "

c

c-----------------c C

Basic Control Theory

33

Autotuner

(

'.

The new autotuner function provides tuning parameters for the PID controller based on gain and phase margin design. The autotuner itself can be viewed as another controller object that has been embedded into the PID controller. The autotuner is based on a relay feedback technique and by default incorporates a relay with hysteresis.

( (

(

c

The PID controller parameters that are obtained from the autotuner are based on a design methodology that makes use of a gain margin at a specified phase angle. This design is quite similar to the regular gain and phase margin methodology except that it is more accurate since the relay has the ability to determine points in the frequency domain accurately and quickly. Also, the relay experiment is controlled and does not take a long time during the tuning cycle.

(

(

c c

In the present version of the software there are default values specified for the PIO tuning. Before starting the autotuner, you must ensure that the controller is in the manual or automatic mode and the process is relatively steady. If you move the cursor over the tuning parameters field, the Status Bar will display the parameters range.

In the present autotuner implementation, there are four parameters that you must supply: Parameter

Range

Ratio (TIlTd) (a)

3.0 :s; a. :s; 6.0

Gain ratio (b)

0.10 :s;

Phase angle (f)

30 0 :s; 4J :s; 65 0

f3

:s; 1.0

Relay hysteresis (h)

0.01 % :s; h :s; 5.0 %

Relay amplitude (d)

0.5 % :s; d :s; 10.0 %

33

34

Basic Control Theory

Setting up a Control Strategy in Aspen HYSYS This section outlines a possible way to create a control strategy in Aspen HYSYS. You should first follow the guidelines outlined in the Dynamic Modelling Manual in Section 1.4.2 - Movingfrom Steady State to Dynamics in order to setup a stable dynamic case. In many cases, an effective control strategy will stabilize the model. PID and Digital On/Off controllers are not active while Aspen HYSYS is in Steady State mode.

You can install controllers in the simulation case either in Steady State or Dynamic mode. There are many different ways to setup a control strategy. The following is a brief outline of some of the more essential items that should be considered when setting up controllers in Aspen HYSYS.

1. Select the Controlled Variables in the Plant Plan a control strategy that is able to achieve an overall plant objective and maintain stability within the plant. Either design the controllers in the plant according to your own standards and conventions or model a control strategy from an existing plant. In Aspen HYSYS, there are a number of variables which can either be set or controlled manually in a dynamic simulation case. You should distinguish between variables that do not change in a plant and those variables which are controlled. Set variables do not change in the dynamic simulation case. Variables such as temperature and composition should be set at each flowsheet boundary feed stream. One pressurelflow specification is usually required for each flowsheet boundary stream in the simulation case. These are the minimum number of variables required by the simulation case for a solution. Reserve these specifications for variables that physically remain constant in a plant. For example, you can specify the exit pressure of a pressure relief valve since the exit pressure typically remains constant in a plant.

In some instances, you can vary a set variable such as a stream's temperature, composition, pressure or flow. To force a specification to behave sinusoidally or ramped, attach the variable to the Transfer Function operation. A variety of different forcing functions and disturbances can be modelled in this manner. The behaviour of controlled variables is determined by the type of controller and the tuning parameters associated with the controller. Typically, the number of control valves in a plant dictates the possible number of controlled variables. There will be more variables to control in Dynamic mode than in Steady State mode. For instance, a two-product column in Steady State mode requires two steady state specifications. The simulator will manipulate the other variables in the column in order to satisfy the provided specifications and the column material and energy balances. The same column in Dynamic mode requires five specifications. The three new specifications correspond to the inventory or integrating specifications not fixed in steady state. The inventory variables include condenser level, reboiler level, and column pressure.

34

====---- ---

--=--.-:;._-~ -

---------'~=-'-'--'-

c

(, ( '.

(

C-----------------C

Basic Control Theory

35

2. Select Controller Structures for Each Controlled Variable

C ( (

You should choose appropriate controller structures for each controlled variable in the simulation case.

(

The controller operations can be added in either Steady State or Dynamic mode. However, controllers have no effect on the simulation in Steady State mode. You must specifY the following in order to fully define the PID Controller operation.

C (

Connections Tab

(

( (

C

C C

The Process Variable (PV) can be specified in the Connections tab by clicking the Select PV button. The controller measures the process variable in an attempt to maintain it at a specified setpoint, SP. The output of a controller is always a control valve, unless the controller is the primary controller in a Cascade control setup.

The Operating Variable (OP) can be specified in the Connections tab by clicking the Select OP button. The output of the controller is a control valve. The output signal, OP, is the percent opening of the control valve. The operating variable may be specified as a physical valve in the plant, a material stream, or an energy stream.

It is possible to have a flow reversal occur in a valve if the pressure drop. across the valve becomes negative. The flow reversal can be avoided by checking the Check Valve.

Operating Variable

Description

Physical Valve

It is recommended that a physical valve be used as the operating variable for a controller. The controller's output signal, OP, is the desired actuator position of the physical valve. With this setup, a more realistic analysis of the effect of the controller on the process is possible. Material flow through the valve is calculated from the frictional resistance equation of the valve and the surrounding unit operations. Flow reversal conditions are possible and valve dynamics may be modelled if a physical valve is chosen.

Material Stream

If a material stream is chosen as an operating variable, the material stream's flow becomes a P-F specification in the dynamic simulation case. You must specify the maximum and minimum flow of the material stream by clicking the Control Valve button. The actual flow of the material stream is calculated from the formula:

Flow

=

OP(%)(FIoW 100 max -

Fl oW

min

). + Fl oW

min

Aspen HYSYS varies the flow specification of the material stream according to the calculated controller output, OP. (Therefore, a nonrealistic situation may arise in the dynamic case since material flow is not dependent on the surrounding conditions.)

35

36

Basic Control Theory

Operating Variable

Description

Energy Stream

If an energy stream is chosen as an operating variable, you may choose a Direct Q or a Utility Fluid Duty Source by clicking the Control Valve button. If the Direct Q option is chosen, you must specify the maximum and minimum energy flow of the energy stream. The actual energy flow of the energy stream is calculated similarly to the material flow:

Energy Flow = OP(%) lOO(Flow

max

_ Flow

min )

+ Flow

min

If the Utility Fluid option is chosen, you need to specify the maximum and minimum flow of the utility fluid. The heat flow is then calculated using the local overall heat transfer coefficient, the inlet fluid conditions, and the process conditions.

Parameters Tab The action of the controller, the controller's PV range, and the tuning parameters can be specified in the Parameters tab. A controller's action (direct or reverse) is specified using the Action radio buttons. A controller's PV span is also specified in the PV Range field. A controller's PV span must cover the entire range of the process variable the sensor is to measure. Tuning parameters are specified in the tuning field.

36

C' ( (

(

c.·-----------------

Basic Control Theory

37

c

c

(

( (

( ( (

( ( (

( (

3. Final Control Elements Set the range on the control valve at roughly twice the steady state flow you are controlling. This is achieved by sizing the valve as Linear with an opening of 50% at the Steady State pressure drop and flowrate. If the controller uses a material or energy stream as an operating variable (OP), the range of the stream's flow can be specified explicitly in the FCV view of the material or energy stream. This view is displayed by clicking the Control Valve button in the PID Controller view. The fmal control element can be characterized as a linear, equal percentage, or quick opening valve. Control valves also have time constants which can be accounted for in Aspen HYSYS. It is suggested that a linear valve mode be used to characterize the valve dynamics offmal control elements. This causes the actual valve position to move at a constant rate to the desired valve positions much like an actual valve in a plant. Since the actual valve position does not move immediately to the OP% set by the controller, the process is less affected by aggressive controller tuning and may be more stable.

(

4. Set up the Databook and Strip Charts Setting up strip charts for your model allows you to easily view several variables while the simulation is running. The procedure for setting up these charts is straightforward. 1.

Enter the Databook property view and select the variables that are to be included in the strip chart in the Variables tab.

37

(

38

Basic Control Theory

2.

From the Strip Charts view, add a new strip chart by clicking the Add button and activate the variables to be displayed on the strip chart. Figure 16

No more than six variables should be active on any given strip chart; having more than six active variables will make the strip chart difficult to read.

Click the Strip Chart button in the View group box to see the strip chart. Size it as desired, and then double-click the strip chart. There are three tabs where you can set: •

The numerical ranges of the strip chart for each variable

•

The nature of the lines for each variable

•

How the strip chart updates and plots the data

Add additional strip charts as desired by going back into the Databook property view and going to the Strip Charts tab.

38

1

_

-~~-_ --'-''::''__ ~ _ ' ' _, __

-

~_.~.-

_.-----....-C .•.

-~­

.._ . __

~.~'-'-',...-

:~

__

(~

'-

~...

(

c

c----------------C~ (

C ( (

c c

Basic Control Theory

39

5. Set up the Controller Faceplates Click the Faceplate button in the PID Controller view to display the controller's faceplate. The faceplate displays the PV, SP, OP, and mode of the controller. Controller faceplates can be arranged in the Aspen HYSYS work environment to allow for monitoring of key process variables and easy access to tuning parameters.

(

( (

c

(

( (

You may edit the setpoint or mode directly from the Face Plate.

Clicking the Face Plate button leads to the Face Plate view.

39

( \

40

Basic Control Theory

6. Set up the Integrator

During start up of a dynamic simulation, it is recommended that a small step size be used. However, once the system has a stabilizer, a larger value can be used.

40

The integration step size can be modified in the Integrator view in the Simulation menu. If desired, change the integration step size to a smaller interval. The default integration time step is 0.5 seconds. Reducing the step size will cause the model to run slower, but during the initial switch from Steady State to Dynamic mode, the smaller step size allows the system to initialize better and enables close monitoring of the controllers to ensure that everything was set up properly. A smaller step size also increases the stability of the model since the solver can more closely follow changes occurring in the plant. Increase the integration step size to a reasonable value when the simulation case has achieved some level of stability. Larger step sizes increase the speed of integration and may be specified ifthe process can maintain stability.

-e c=.-

c (

c---------------r\~

(

C ( ( ( ( (

Basic Control Theory

41

7. Fine Tuning of Controllers Before the Integrator is run, each controller should be turned off and then put back in manual mode to initialize the controllers. Placing the controllers in manual will default the setpoint to the current process variable and allow you to "manually" adjust the valve % opening of the operating variable.

(

If reasonable pressure/flow specifications are set in the dynamic simulation and all the equipment is properly sized, most process variables should line out after the Integrator is run. The transition of most unit operations from Steady State to Dynamic mode is very smooth. However, controller tuning is critical if the plant simulation is to remain stable. Dynamic columns, for instance, are not open loop stable like many of the unit operations in Aspen HYSYS. Any large disturbances in the column may result in simulation instability.

C

Once the Integrator is running:

C ( (

1.

Slowly bring the controllers online starting with the ones attached to upstream unit operations. The control of flow and pressure of upstream unit operations should be handled initially since these variables have a significant effect on the stability of downstream unit operations.

2.

Concentrate on controlling variables critical to the stability of the unit operation. Always keep in mind that upstream variables to a unit operation should be stabilized first. For example, the feed flow to a column should be controlled initially. Next, try to control the temperature and pressure profile of the column. Finally, pay attention to the accumulations of the condenser and reboiler and control those variables.

3.

Start conservatively using low gains and no integral action. Most unit operations can initially be set to use P-only control. If an offset cannot be tolerated initially, then integral action should be added.

4.

Trim the controllers using integral or derivative action until satisfactory closedloop performance is obtained.

5.

At this point, you can concentrate on changing the plant to perform as desired. For example, the control strategy can be modified to maintain a desired product composition. If energy considerations are critical to a plant, different control strategies may be tested to reduce the energy requirements of unit operations.

(

41

(

( 42

Basic Control Theory

Stability It has been shown that the stability of a closed loop process depends on the controller gain. If the controller gain is increased, the closed loop response is more likely to become unstable. The controller gain (Kc) input in the PID Controller operation in Aspen HYSYS is a unitless value defined in Equation 23.

K

c

=

OP % f!.PV / PVRange

(23)

In order to control the process, the controller must interact with the actual process. This is achieved by using the effective gain, K.em which is essentially the controller gain with units. The effective gain is defined as:

Kejf = Kc(F10Wmax- Flow min) PVRang

(24)

The process gain has units that are reciprocal to the effective gain.

It is possible to achieve tight control in a plant and to have the simulation case become unstable due to modifications in the PV range or Cy values of a final control element.

You should also consider the effect of interactions between the control loops existing in a plant. Interactions between the control loops change the effective gain of each loop. It is possible for a control loop that was tuned independently ofthe other control loops in the plant to become unstable as soon as it is put into operation with the other loops. It is therefore useful to design feedback control loops, which minimize the interactions between the controllers.

42

C f (

C

C------------------( (

Basic Control Theory

43

References

(

(

1.

( (

Svrcek, W.Y., Mahoney, D.P., andB.R. Young. A Real Time Approach to Process Controls John Wiley & Sons Ltd, Chichester (2000) p. 125

2.

Ogunnaike, B.A. and W.H. Ray. Process Dynamics, Modelling, and Control Oxford University Press, New York (1994) p. 531

(

C ( ( (

"(

r :p./""

:t~'I"~ . ~/,,;oJ; .-i

43

44

44

Basic Control Theory

{;9;'~>~""'.e-;;:;,,,,,,,,\,(_

II!)aspent

Ii

AspenTech Customer Education North and South America Training

us Training Coordinator [Phone toll-free] + 1-888-996-7100 [Phone- outside North America:] + 1-281-584-4357

[Fax:] +1-281-584-1064 [e-mail:][email protected]

Europe, Middle East and Africa Training Europe Training Coordinator [Phone toll-free:] + 32-2-701-9555 For a list of toll-free numbers in Northern Europe, please visit web site hUp://support.aspentech.com/training

[Fax:] +32-2-403-1771 [e-mail:][email protected]

Japan Training Japan Training Coordinator [Phone:] + 81-3-3262-1743

[Fax:] + 81-3-3262-1744 [e-mail:][email protected]

Asia Training Asia Training Coordinator [Phone:] + 65-6395-3900

[Fax:] + 65-6395-3950 [e-mail:][email protected]

his training manual contains AspenTech proprietary and confidential Iformation and may not be disclosed, used, or copied without the prior onsent of AspenTech or as set forth in the applicable license agreement. sers are solely responsible for the proper use of the software and the pplication of the results obtained.

)ugh AspenTech has tested the software and reviewed the training materials, the sole warranty for the software may be found in the applicable license agreement between anTech and the user. ASPENTECH MAKES NO WARRANTY OR REPRESENTATION, EITHER EXPRESSED OR IMPLIED, WITH RESPECT TO THIS TRAINING MATERIAL, QUALITY, PERFORMANCE, MERCHANTABILITY, OR FITNESS FOR A PARTICULAR PURPOSE. Copyright© 2006. AspenTech· aspenONETM, Aspen ADSIM®, Aspen JS Project Manager™, Aspen Advisor™, Aspen IQTM, Aspen POLYMERS PLUS®, Aspen B-JAC®, Aspen MIMITM, Aspen Transition Manager™, Aspen Calc™, Aspen PIMSTM, an Watch™, Aspen SQLplus™, Aspen ChromatographyTM, Aspen Pinch®, Aspen Water™, Aspen Web.21™, Aspen Custom Modeler®, Aspen PluS®, Aspen Zyqad™, Aspen amicS®, Aspen Process Explorer™, Aspen Batch.21™, Aspen eBRSTM, Aspen Process Order™, Aspen Process Recipe™, Aspen Framework, Aspen Icarus 2000®, Aspen ;PluS®, Aspen Icarus Process Evaluator™, Aspen InfoPlus.21®, Aspen QTM are trademarks or registered trademarks of Aspen Technology Inc., Cambridge, Massachusetts

.,

ghts reserved, All other brand and product names are trademarks or registered trademarks of their respective companies.