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Copyright (c) 1981-2008 by Aspen Technology, Inc. All rights reserved. 
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ThTechnical Supportv
  Online Technical Support Center ........................................................ vi
  Phone and E-mail .............................................................................. vii
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ThOnline Technical Support 
Center
AspenTech customers with a valid license and software 
maintenance agreement can register to access the Online 
Technical Support Center at:
http://support.aspentech.com
You use the Online Technical Support Center to: 
• Access current product documentation
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Registered users can also subscribe to our Technical Support e-
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ThPhone and E-mail
Customer support is also available by phone, fax, and e-mail for 
customers who have a current support contract for their 
product(s). Toll-free charges are listed where available; 
otherwise local and international rates apply. 
For the most up-to-date phone listings, please see the Online 
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Thviii
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Table of ContentsTechnical Support..................................................... v
Online Technical Support Center ............................vi
Phone and E-mail................................................ vii
1  Dynamic Theory....................................................1-1
1.1 Introduction .................................................... 1-3
1.2 General Concepts ............................................. 1-6
1.3 Holdup Model .................................................1-13
1.4 Pressure Flow Solver .......................................1-29
1.5 Dynamic Operations: General Guidelines ............1-42
1.6 Aspen HYSYS Dynamics ...................................1-49
1.7 References .....................................................1-71
2  Dynamic Tools ......................................................2-1
2.1 Introduction .................................................... 2-3
2.2 Dynamics Assistant .......................................... 2-4
2.3 Equation Summary Property View......................2-28
2.4 Integrator ......................................................2-38
2.5 Event Scheduler..............................................2-45
2.6 Control Manager .............................................2-72
2.7 Dynamic Initialization ......................................2-73
3  Control Theory......................................................3-1
3.1 Introduction .................................................... 3-2
3.2 Process Dynamics ............................................ 3-3
3.3 Basic Control ..................................................3-10
3.4 Advanced Control ............................................3-32
3.5 General Guidelines ..........................................3-39
3.6 References .....................................................3-58
 Index.................................................................... I-1ix
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Dynamic Theory 1-1
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Th1  Dynamic Theoryw.cadfamily.com    EMa
e document is for study 1.1  Introduction................................................................................... 3
1.2  General Concepts ........................................................................... 6
1.2.1  Mathematical Model Classification............................................... 6
1.3  Holdup Model ............................................................................... 13
1.3.1  Assumptions of Holdup Model .................................................. 14
1.3.2  Accumulation......................................................................... 14
1.3.3  Non-Equilibrium Flash ............................................................. 15
1.3.4  Heat Loss Model..................................................................... 20
1.3.5  Chemical Reactions................................................................. 23
1.3.6  Related Calculations ............................................................... 24
1.3.7  Advanced Holdup Properties .................................................... 24
1.4  Pressure Flow Solver ................................................................... 29
1.4.1  Simultaneous Solution in Pressure Flow Balances........................ 30
1.4.2  Basic Pressure Flow Equations.................................................. 31
1.4.3  Pressure Flow Specifications .................................................... 34
1.5  Dynamic Operations: General Guidelines ..................................... 42
1.5.1  Differences between Dynamic & Steady State Mode .................... 43
1.5.2  Moving from Steady State to Dynamics ..................................... 44
1.6  HYSYS Dynamics.......................................................................... 49
1.6.1  Detailed Heat Model ............................................................... 51
1.6.2  Nozzles................................................................................. 53
1.6.3  Control Valve Actuator ............................................................ 57
1.6.4  Inertia .................................................................................. 62
1.6.5  Static Head ........................................................................... 68
1.6.6  Design.................................................................................. 711-1
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1-2 Dynamic Theory 
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Dynamic Theory 1-3
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Th1.1 Introduction
Dynamic simulation can help you to better design, optimize, and 
operate your chemical process or refining plant. Chemical plants 
are never truly at steady state. Feed and environmental 
disturbances, heat exchanger fouling, and catalytic degradation 
continuously upset the conditions of a smooth running process. 
The transient behavior of the process system is best studied 
using a dynamic simulation tool like Aspen HYSYS.
The design and optimization of a chemical process involves the 
study of both steady state and dynamic behavior. Steady state 
models can perform steady state energy and material balances 
and evaluate different plant scenarios. The design engineer can 
use steady state simulation to optimize the process by reducing 
capital and equipment costs while maximizing production.
With dynamic simulation, you can confirm that the plant can 
produce the desired product in a manner that is safe and easy to 
operate. By defining detailed equipment specifications in the 
dynamic simulation, you can verify that the equipment functions 
as expected in an actual plant situation. Offline dynamic 
simulation can optimize controller design without adversely 
affecting the profitability or safety of the plant. 
You can design and test a variety of control strategies before 
choosing one that is suitable for implementation. You can 
examine the dynamic response to system disturbances and 
optimize the tuning of controllers. Dynamic analysis provides 
feedback and improves the steady state model by identifying 
specific areas in a plant that have difficulty achieving the steady 
state objectives.
In Aspen HYSYS, the dynamic analysis of a process system can 
provide insight into the process system when it is not possible 
with steady state modeling. 
Note: Aspen HYSYS Thermodynamics COM Interface is not 
optimized for Dynamics mode and can result in performance 
issues if used in Dynamics mode.
Contact your Aspentech 
agent for more 
information, or e-mail us 
at 
info@aspentech.com.1-3
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1-4 Introduction
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ThWith dynamic simulation, you can investigate:
• Process optimization
• Controller optimization
• Safety evaluation
• Transitions between operating conditions
• Startup/Shutdown conditions
The Aspen HYSYS dynamic model shares the same physical 
property packages as the steady state model. The dynamic 
model simulates the thermal, equilibrium, and reactive behavior 
of the chemical system in a similar manner as the steady state 
model.
On the other hand, the dynamic model uses a different set of 
conservation equations which account for changes occurring 
over time. The equations for material, energy, and composition 
balances include an additional “accumulation” term which is 
differentiated with respect to time. Non-linear differential 
equations can be formulated to approximate the conservation 
principles; however, an analytical solution method does not 
exist.
Therefore, numerical integration is used to determine the 
process behavior at distinct time steps. The smaller the time 
step, the more closely the calculated solution matches the 
analytic solution. However, this gain in rigour is offset by the 
additional calculation time required to simulate the same 
amount of elapsed real time. A reasonable compromise is 
achieved by using the largest possible step size, while 
maintaining an acceptable degree of accuracy without becoming 
unstable.
The Aspen HYSYS dynamic simulation package has the capacity 
to reach a wide audience by offering the following features 
demanded by industry:
• Accuracy. The Aspen HYSYS dynamic model provides 
accurate results based on rigorous equilibrium, reaction, 
unit operations, and controller models. You must be able 
to trust the results if the dynamic tool is to prove useful.
• Ease of Use. The Aspen HYSYS dynamic package uses 
the same intuitive and interactive graphical environment 
as the Aspen HYSYS steady state model. Streams and 
unit operations in the flowsheet can be added to the 1-4
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Dynamic Theory 1-5
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Thdynamic simulation environment as easily as in steady 
state. All flowsheet information from a steady state 
simulation case transfers easily to the dynamic 
simulation environment.
• Speed. The dynamic modeling options in Aspen HYSYS 
were developed to provide a balance between accuracy 
and speed. Aspen HYSYS uses the Implicit fixed step size 
Euler method. Volume, energy, and composition balances 
are solved at different frequencies. Volume (pressure-
flow) balances are defaulted to solve at every time step, 
whereas energy and composition balances are defaulted 
to solve at every second and tenth time step. This 
solution method allows Aspen HYSYS to perform quick, 
accurate and stable calculations in your simulation case.
• Detailed Design. You can provide specific rating details 
for each piece of equipment in the plant and confirm that 
the specified equipment can achieve desired product 
specs and quality. Rating information includes the 
equipment size, geometry, nozzle placement, and 
position relative to the ground. A comprehensive holdup 
model calculates levels, heat loss, static head 
contributions, and product compositions based on the 
rating information of each piece of equipment.
• Realism. A new level of realism with regards to material 
flow within the simulation is achieved with the Pressure 
Flow solver. With the Pressure Flow option, the flow rate 
through any unit operation depends on the pressures of 
the surrounding pieces of equipment. Material flow 
through an actual plant can be more accurately modeled 
using the Pressure Flow solver.
• Customizable. The Aspen HYSYS dynamic model is 
customizable. Many organizations have proprietary 
information that they want to integrate into their 
commercial simulator platform. Aspen HYSYS allows you 
to add your own OLE modules to the Aspen HYSYS 
dynamic simulation environment.1-5
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1-6 General Concepts
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Th1.2 General Concepts
1.2.1 Mathematical Model 
Classification 
Distributed & Lumped Models
Most chemical engineering systems have thermal or component 
concentration gradients in three dimensions (x,y,z) as well as in 
time. This is known as a distributed system. If you were to 
characterize such a system mathematically, you would obtain a 
set of partial differential equations (PDEs).
If the x, y, and z gradients are ignored, the system is “lumped”, 
and all physical properties are considered to be equal in space. 
Only the time gradients are considered in such an analysis. This 
consideration allows for the process to be described using 
ordinary differential equations (ODEs) which are much less 
rigorous than PDEs, thereby saving calculation time. For most 
instances, the lumped method gives a solution which is a 
reasonable approximation of the distributed model solution.
Aspen HYSYS uses lumped models for all of the unit operations. 
For instance, in the development of the equations describing the 
separator operation, it is assumed that there are no thermal or 
concentration gradients present in a single phase. In other 
words, the temperature and composition of each phase are the 
same throughout the entire separator.
In the solution algorithm, the PFR reactor is subdivided into 
several sub-volumes which are considered to be lumped; that is, 
the reaction rate, temperature and compositions are constant 
through each sub-volume, varying only with time. In essence, 
Notice that by definition, the PFR has thermal and 
concentration gradients with respect to the length of the 
vessel.1-6
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Dynamic Theory 1-7
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Ththerefore, the PFR model, though inherently distributed (with 
respect to the length of the vessel), still uses a lumped analysis 
to obtain the solution.
Linear & Non-Linear Systems
A linear first-order ODE can be described as follows:
In a non-linear equation, the process variable Y is displayed as a 
power, exponential, or is not independent of other process 
variables. Refer to the following two examples: 
The great majority of chemical engineering processes occurring 
in nature are nonlinear. Nonlinearity arises from equations 
describing equilibrium behavior, fluid flow behavior, or reaction 
rates of chemical systems. While a linear system of equations 
are solved analytically using matrix algebra, the solution to a 
non-linear set of equations usually requires the aid of a 
computer.
Conservation Relationships
Material Balance
The conservation relationships are the basis of mathematical 
modeling in Aspen HYSYS. The dynamic mass, component, and 
energy balances that are derived in the following section are 
similar to the steady state balances with the exception of the 
accumulation term in the dynamic balance. It is the 
(1.1)
(1.2)
(1.3)
τ Yd td
----- Y+ Kf u( )=
τ Yd td
----- Y3+ Kf u( )=
τ Yd
td
----- YY2+ Kf u( )=1-7
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1-8 General Concepts
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Thaccumulation term which allows the output variables from the 
system to vary with time.
The conservation of mass is maintained in the following general 
relation:
For the simple case of a perfectly mixed tank with a single 
component feed, the mass balance is as follows: 
where:  
Fi = flowrate of the feed entering the tank
 = density of the feed entering the tank
Fo = flowrate of the product exiting the tank
 = density of the product exiting the tank
V = volume of the fluid in the tank
Rate of accumulation of mass = mass flow into system - 
mass flow out of system
(1.4)
 Figure 1.1
(1.5)ρoV( )d
td
----------------- Fiρi Foρo–=
ρi
ρo1-8
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Dynamic Theory 1-9
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ThComponent Balance
Component balances can be written as follows: 
Flow into or out of the system can be convective (bulk flow) 
and/or molecular (diffusion). While convective flow contributes 
to the majority of the flow into and out of a system, diffusive 
flow becomes significant if there is a high interfacial area to 
volume ratio for a particular phase.
For a multi-component feed for a perfectly mixed tank, the 
balance for component j would be as follows:
where:  
Cji = concentration of j in the inlet stream
Cjo = concentration of j in the outlet stream
Rj = reaction of rate of the generation of component j
For a system with NC components, there are NC component 
balances. The total mass balance and component balances are 
not independent; in general, you would write the mass balance 
and NC-1 component balances.
Equation (1.5) is a simplification of the more rigorous 
equation used inside Aspen HYSYS which also considers 
other phenomena such as vapourization, reactions, density 
changes, etc.
Rate of accumulation of component j = Flow of component 
j into system - Flow of component j out of system + 
Rate of formation of component j by reaction 
(1.6)
(1.7)
CjoV( )d
td
------------------- FiCji FoCjo– RjV+=1-9
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1-10 General Concepts
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ThEnergy Balance
The energy balance is as follows:
The flow of energy into or out of the system is by convection or 
conduction. Heat added to the system across its boundary is by 
conduction or radiation.
For a CSTR with heat removal, the following general equation 
applies:
where:  
u = internal energy (energy per unit mass)
k = kinetic energy (energy per unit mass)
 = potential energy (energy per unit mass)
V = volume of the fluid
w = shaft work done by system (energy per time)
Po = vessel pressure
Pi = pressure of feed stream
Q = heat added across boundary
Qr = , heat generated by reaction 
Several simplifying assumptions can usually be made:
• The potential energy can almost always be ignored; the 
inlet and outlet elevations are roughly equal.
• The inlet and outlet velocities are not high, therefore 
kinetic energy terms are negligible.
• If there is no shaft work (no pump), w=0.
Rate of accumulation of total energy = Flow of total energy 
into system - Flow of total energy out of system + 
Heat added to system across its boundary + Heat 
generated by reaction - Work done by system on 
surroundings
(1.8)
(1.9)
td
d u k φ+ +( )V[ ] Fiρi ui ki φi+ +( ) Foρo uo ko φo+ +( ) Q Qr w FoPo FiPi–+( )–+ +–=
φ
DHrxnrA1-10
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Dynamic Theory 1-11
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ThThe general energy balance for a 2-phase system is as follows:
Solution Method
Implicit Euler Method
Yn+1 is analytically calculated to equal:
where:  
Ordinary differential equations are solved using the Implicit 
Euler method. The Implicit Euler method is simply an 
approximation of Yn+1 using rectangular integration. Graphically, 
a line of slope zero and length h (the step size) is extended from 
tn to tn+1 on an f(Y) vs. time plot. The area under the curve is 
approximated by a rectangle of length h and height fn+1(Yn+1):
(1.10)
 
(1.11)
(1.12)
td
d ρvVvH ρlVlh+[ ] Fiρihi F– lρlh FvρvH– Q Qr+ +=
Yn 1+ Yn f Y( ) td
tn
tn 1+
∫+=
dY
dt
------ f Y( )=
Yn 1+ Yn hfn 1+ Yn 1+( )+=1-11
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1-12 General Concepts
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ThThe figure below shows the integration of f(Y) over time step, h, 
using exact integration and the Implicit Euler approximation: 
The Implicit Euler method handles stiff systems well. This is an 
implicit method because information is required at time tn+1. 
Integration parameters such as the integration time step can be 
specified in the Integrator property view from the Simulation 
menu in Aspen HYSYS. The integration time step can be 
adjusted to increase the speed or stability of the system.
Integration Strategy
In Aspen HYSYS, dynamic calculations are performed at three 
different frequencies:
• Volume (pressure-flow)
• Energy
• Composition
These relations are not solved simultaneously at every time 
step. This would be computationally expensive. The compromise 
is to solve the balances at different time step frequencies. The 
default solution frequencies, which are multiples of the 
integration time step, are one, two, and ten for the pressure 
flow equations, energy, and composition balances. 
Meaning pressure flow equations are solved at every time step 
while composition balances are solved at every 10th time step. 
Since composition tends to change much more gradually than 
the pressure, flow, or energy in a system, the equations 
associated with composition can be solved less frequently. An 
 Figure 1.2
tn tn+1
f(Y)
tn tn+1
f(Y)
Area f Y( )dt
fn
fn 1+
∫= Area fn 1+( )h=
Exact Integration Rectangular Integration (Implicit Euler)1-12
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Dynamic Theory 1-13
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Thapproximate flash is used for each pressure flow integration 
time step. A rigorous flash is performed at every composition 
integration time step.
1.3 Holdup Model
Dynamic behavior arises from the fact that many pieces of plant 
equipment have some sort of material inventory or holdup. A 
holdup model is necessary because changes in the composition, 
temperature, pressure or flow in an inlet stream to a vessel with 
volume (holdup) are not immediately seen in the exit stream. 
The model predicts how the holdup and exit streams of a piece 
of equipment respond to input changes to the holdup over time.
In some cases, the holdup model corresponds directly with a 
single piece of equipment in Aspen HYSYS. For example, a 
separator is considered a single holdup. In other cases, there 
are numerous holdups within a single piece of equipment. In the 
case of a distillation column, each tray can be considered a 
single holdup. Heat exchangers can also be split up into zones 
with each zone being a set of holdups.
Calculations included in the holdup model are:
• Material and energy accumulation
• Thermodynamic equilibrium
• Heat transfer
• Chemical reaction
The new holdup model offers certain advantages over the 
previous Aspen HYSYS dynamic model:
• An adiabatic PH flash calculation replaces the bubble 
point algorithm used in the previous holdup model. 
Adiabatic flashes also allow for more accurate 
calculations of vapour composition and pressure effects 
in the vapour holdup.
• Flash efficiencies can be specified allowing for the 
modeling of non-equilibrium behavior between the feed 
phases of the holdup.
• The placement of feed and product nozzles on the 
equipment has physical meaning in relation to the 
holdup. For example, if the vapour product nozzle is 
placed below the liquid level in a separator, only liquid 
exits from the nozzle.1-13
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1-14 Holdup Model
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Th1.3.1 Assumptions of Holdup 
Model
There are several underlying assumptions that are considered in 
the calculations of the holdup model:
• Each phase is assumed to be well mixed.
• Mass and heat transfer occur between feeds to the 
holdup and material already in the holdup.
• Mass and heat transfer occur between phases in the 
holdup.
1.3.2 Accumulation
The lagged response that is observed in any unit operation is the 
result of the accumulation of material, energy, or composition in 
the holdup. To predict how the holdup conditions change over 
time, a recycle stream is added alongside the feed streams. For 
example, the material accumulation in a holdup can be 
calculated from: 
The recycle stream is not a physical stream in the unit 
operation. Rather, it is used to introduce a lagged response in 
the output. Essentially, the recycle stream represents the 
material already existing in the piece of equipment. It becomes 
apparent that a greater amount of material in the holdup means 
a larger recycle stream and thus, a greater lagged response in 
the output. 
Material accumulationnew = material flow into system + 
material accumulationold (recycle stream) - material 
flow out of system
(1.13)1-14
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Dynamic Theory 1-15
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ThThe holdup model is used to calculate material, energy, and 
composition accumulation. Material accumulation is defaulted to 
calculate at every integration time step. The energy of the 
holdup is defaulted to calculate at every second time step. The 
composition of the holdup is defaulted to calculate at every 
tenth time step.
1.3.3 Non-Equilibrium Flash
As material enters a holdup, the liquid and vapour feeds can 
associate in different proportions with the existing material 
already in the holdup. For instance, a separator’s vapour and 
liquid feeds can enter the column differently. It is very likely that 
the liquid feed mixes well with the liquid already in the holdup.
The vapour feed is not mixed as well with the existing material 
in the vessel since the residence time of the vapour holdup is 
much smaller than that of the liquid. If the feed nozzle is 
situated close to the vapour product nozzle, it is possible that 
even less mixing occurs. In the physical world, the extent of 
mixing the feeds with a holdup depends on the placement of the 
feed nozzles, the amount of holdup, and the geometry of the 
piece of equipment.
 Figure 1.31-15
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1-16 Holdup Model
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ThEfficiencies
In Aspen HYSYS, you can indirectly specify the amount of 
mixing that occurs between the feed phases and the existing 
holdup using feed, recycle, and product efficiencies. These feed 
efficiency parameters can be specified on the Efficiencies tab of 
the unit operation’s Advance property view. Click the Advance 
button on the Holdup page under the Dynamics tab to open the 
Advance property view.
Essentially, the efficiencies determine how rapidly the system 
reached equilibrium. If all efficiencies are 1, then all feeds reach 
equilibrium instantaneously. If the values are lower, it takes 
longer and the phases cannot be in equilibrium and can have 
different temperatures.
 Figure 1.41-16
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Dynamic Theory 1-17
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ThA flash efficiency can be specified for each phase of any stream 
entering the holdup. A conceptual diagram of the non-
equilibrium flash is shown for a two phase system in the figure 
below:
As shown, the flash efficiency, , is the fraction of feed stream 
that participates in the rigorous flash. If the efficiency is 
specified as 1, the entire stream participates in the flash; if the 
efficiency is 0, the entire stream bypasses the flash and is mixed 
with the product stream. 
The recycle stream (and any streams entering the holdup) 
participates in the flash. You can specify the flash efficiency for 
each phase of the recycle stream and any feed entering the 
holdup. The flash efficiency can also be specified for each phase 
of any product streams leaving the holdup.
The default efficiencies for the feed, product, and recycle 
streams is 1, and this value is sufficient in the vast majority of 
cases. The flash efficiencies can be changed to model non-
 Figure 1.5
Product flash efficiencies are only used by the holdup model 
when reverse flow occurs in the product flow nozzles. In 
such cases, the product nozzle effectively becomes a feed 
nozzle and uses the product flash efficiencies that you 
provided.
η
1-17
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1-18 Holdup Model
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Thequilibrium conditions. For example, the efficiency of vapour 
flowing through a vessel containing liquid can be reduced if the 
residence time of the vapour is very small and there is little time 
for it to reach thermodynamic equilibrium with the liquid. Also, 
in some narrow boiling systems, lower efficiencies can be used 
to reduce the rate at which material can condense or evaporate. 
This can help to stabilize the pressure in certain difficult cases 
such as narrow boiling systems like steam.
For example, a water system is heated by pure steam (no 
inerts) can encounter problems if the stream efficiency is 
specified as 1. If the holdup material is significantly larger than 
the stream flow, all the steam condenses and the holdup 
temperature increases accordingly. No vapour is present which 
can complicate pressure control of the system. In the physical 
world, typically not all of the steam condenses in the water and 
there are also some inerts (e.g., nitrogen or air) present in the 
system. Using lower efficiencies can help to model this system 
better.
Nozzles
In Aspen HYSYS, you can specify the feed and product nozzle 
locations and diameters. These nozzle placement parameters 
can be specified in the unit operation’s Nozzles page under the 
Rating tab. 
 Figure 1.61-18
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ThThe placement of feed and product nozzles on the equipment 
has physical meaning in relation to the holdup. The exit stream’s 
composition depends partially on the exit stream nozzle’s 
location in relation to the physical holdup level in the vessel. If 
the product nozzle is located below the liquid level in the vessel, 
the exit stream draws material from the liquid holdup. If the 
product nozzle is located above the liquid level, the exit stream 
draws material from the vapour holdup. If the liquid level sits 
across a nozzle, the mole fraction of liquid in the product stream 
varies linearly with how far up the nozzle the liquid is.
Static Head Contributions
When the Static Head Contributions checkbox is selected on 
the Options tab of the Integrator property view, Aspen HYSYS 
calculates static head using the following contributions:
• Levels inside separators, tray sections, and so forth
• Elevation differences between connected equipment
For unit operations with negligible holdup, such as the valve 
operation, Aspen HYSYS incorporates only the concept of 
nozzles. There is no static head contributions for levels, unless 
the feed and product nozzles are specified at different 
elevations. 
You can specify the elevation of both the feed and product 
nozzles. If there is a difference in elevation between the feed 
and product nozzles, Aspen HYSYS uses this value to calculate 
the static head contributions. It is recommended that static 
head contributions not be modeled in these unit operations in 
this way since this is not a realistic situation. Static head can be 
better modeled in these unit operations by relocating the entire 
piece of equipment.
Static head is important in vessels with levels. For example, 
consider a vertical separator unit operation that has a current 
liquid level of 50%. The static head contribution of the liquid 
Including static head contributions in the modeling of 
pressure-flow dynamics is an option in Aspen HYSYS.1-19
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1-20 Holdup Model
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Thholdup makes the pressure at the liquid outlet nozzle higher 
than that at the vapour outlet nozzle. Nozzle location also 
becomes a factor. The pressure-flow relationship for the 
separator is different for a feed nozzle which is below the 
current liquid holdup level as opposed to a feed which is 
entering in the vapour region of the unit.
It is important to notice that exit stream pressures from a unit 
operation are calculated at the exit nozzle locations on the piece 
of equipment and not the inlet nozzle locations of the next piece 
of equipment.
1.3.4 Heat Loss Model
The heat loss experienced by any pieces of plant equipment is 
considered by the holdup model in Aspen HYSYS. The heat loss 
model influences the holdup by contributing an extra term to the 
energy balance equation.
Energy Balances
Heat is lost (or gained) from the holdup fluid through the wall 
and insulation to the surroundings. 
There are several underlying assumptions that are considered 
during a heat loss calculation:
 Figure 1.71-20
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Th• There is heat capacity associated with the wall (metal) 
and insulation housing the fluid.
• There is thermal conductivity associated with the wall 
and insulation housing the fluid.
• The temperature across the wall and insulation is 
assumed to be constant (lumped parameter analysis).
• You can now have different heat transfer coefficients on 
the inside of a vessel for the vapour and the liquid. The 
heat transfer coefficient between the holdup and the wall 
is no longer assumed to be same for the vapour and 
liquid.
• The calculation uses convective heat transfer on the 
inside and outside of the vessel.
• The calculations assume that the temperature does not 
vary along the height of the vessel, and there is a 
temperature gradient through the thickness of the wall 
and insulation.
A balance can be performed across the wall:
The balance across the insulation is:
where:  
A = the heat transfer area
x = the thickness
Cp = the heat capacity
T = the temperature
k = the thermal conductivity
h = the heat transfer coefficient
As shown, both the insulation and wall can store heat. The heat 
loss term that is accounted for in the energy balance around the 
holdup is . If Tfluid is greater than Twall, the 
heat is lost to the surroundings. If Tfluid is less than Twall, the heat 
(1.14)
(1.15)
td
d AxwallCpwallTwall[ ] h fluid wall,( )A Tfluid Twall–( )
kins
xins
--------A Twall Tins–( )–=
td
d AxinsCpins
Tins
2
---------⎝ ⎠
⎛ ⎞ kins
xins
--------A Twall Tins–( ) h ins surr,( )A Tins Tsurr–( )+=
h fluid wall,( )A Tfluid Twall–( )1-21
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1-22 Holdup Model
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This gained from the surroundings.
Heat Loss Parameters
The heat loss parameters can be specified for most unit 
operations in the Heat Loss page under 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 
• Detailed
Simple Model
The Simple model allows you to either specify the heat loss 
directly or have the heat loss calculated from specified values:
• Overall U value
• Ambient Temperature
The heat transfer area, A, and the fluid temperature, Tf, are 
calculated by Aspen HYSYS. The heat loss is calculated using:
The wall temperature and insulation temperature are 
considered separately. At any given time, the wall is 
assumed to be at the wall temperature, and the insulation is 
at the insulation temperature. The temperature at the wall/
insulation interface is calculated using a pseudo steady state 
assumption for viewing purpose.
(1.16)Q = UA(Tf Tamb )–1-22
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ThDetailed Model
The Detailed model allows you to specify more detailed heat 
transfer parameters. There are three radio buttons in the Heat 
Loss Parameters group as described in the table below:
1.3.5 Chemical Reactions
Chemical reactions that occur in plant equipment are considered 
by the holdup model in Aspen HYSYS. Reaction sets can be 
specified in the Results page of the Reactions tab.
The holdup model is able to calculate the chemical equilibria and 
reactions that occur in the holdup. In a holdup, chemical 
reactions are modeled by one of four mechanisms:
• Reactions handled inside thermophysical property 
packages
• Extent of reaction model
• Kinetic model
• Equilibrium model
Radio Button Description
Temperature 
Profile
Displays the temperatures of the:
• fluid
• wall
• insulation
• surroundings
Conduction Displays the conductive properties of the wall and 
insulation. The following properties can be specified by you:
• Conductivity of material
• Thickness of material
• Heat capacity of material
• Density of material
Equation (1.14) and (1.15) demonstrate how the 
parameters are used by the heat loss model. 
Convection Displays the convective heat transfer coefficients for heat 
transfer within the holdup and heat transfer occurring from 
the outside the holdup to the surroundings.
For more information on 
how reaction sets can be 
created and used within 
the simulation, see 
Chapter 5 - Reactions 
in the Aspen HYSYS 
Simulation Basis 
guide.1-23
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1-24 Holdup Model
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Th1.3.6 Related Calculations
There are calculations which are not handled by the holdup 
model itself, but can impact the holdup calculations. The 
following calculations require information and are solved in 
conjunction with the holdup model as described in the following 
table:
1.3.7 Advanced Holdup 
Properties
Located on each Holdup page found on the Dynamics tab of the 
unit operation property view there is an Advanced button. This 
button accesses the Holdup property view that provides more 
detailed information about the holdup of that unit operation.
Calculations Description
Vessel Level 
Calculations
The vessel level can be calculated from the vessel 
geometry, the molar holdup and the density for each 
liquid phase.
Vessel Pressure The vessel pressure is a function of the vessel volume 
and the stream conditions of the feed, product, and the 
holdup. The pressure in the holdup is calculated using 
a volume balance equation. Holdup pressures are 
calculated simultaneously across the flowsheet.
Tray Hydraulics Tray Hydraulics determines the rate from which liquid 
leaves the tray, and hence, the holdup and the 
pressure drop across the tray. The Francis Weir 
equation is used to determine the liquid flow based on 
the liquid level in the tray and the tray geometry.
Right-click anywhere in the property view to bring up the 
object inspect menu. Selecting the Open Page command 
displays the information on the Holdup page in a separate 
property view.1-24
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ThGeneral Tab
This tab provides the same information as shown in the Holdup 
page of the Dynamics tab. The accumulation, moles, and 
volume of the holdup appear on this tab. The holdup pressure 
also appears on this tab.
Select the Active Phase Flip Check checkbox to enable Aspen 
HYSYS to check if there is a phase flip between Liquid 1 (light 
liquid) and Liquid 2 (heavy liquid) during simulation and 
generate a warning message whenever the phase flip occur. If 
the checkbox is clear, Aspen HYSYS generates a warning only on 
the first time the phase flip occur.
 Figure 1.81-25
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1-26 Holdup Model
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ThNozzles Tab
The Nozzles tab displays the same information as shown in the 
Nozzles page of the Ratings tab. The nozzle diameters and 
elevations for each stream attached to the holdup appear on this 
tab. This section also displays the holdup elevation which is 
essentially equal to the base elevation of the piece of equipment 
relative to the ground. Changes to nozzle parameters can either 
be made in this tab or in the Nozzles page of the Ratings tab.
The Nozzles tab requires Aspen HYSYS Dynamics license.
 Figure 1.9
Refer to Section 1.6 - 
HYSYS Dynamics for 
more information.1-26
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Dynamic Theory 1-27
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ThEfficiencies Tab
The nozzle efficiencies can be specified in this tab. In Aspen 
HYSYS, you can indirectly specify the amount of mixing that 
occurs between the feed phases and existing holdup using feed, 
recycle and product efficiencies.
A flash efficiency, , is the fraction of feed stream that 
participates in the rigorous flash. If the efficiency is specified as 
100, the entire stream participates in the flash; if the efficiency 
is 0, the entire stream bypasses the flash and is mixed with the 
product stream.
The Efficiencies tab requires Aspen HYSYS Dynamics license.
 Figure 1.10
Nozzle Efficiency Description
Feed Nozzle 
Efficiency
The efficiencies of each phase for each feed 
stream into the holdup can be specified in these 
cells. These efficiencies are not used by the 
holdup model if there is flow reversal in the feed 
streams.
Refer to Section 1.6 - 
HYSYS Dynamics for 
more information.
For more information 
regarding feed, product, 
and recycle efficiencies, 
see Section 1.3.3 - 
Non-Equilibrium Flash.
η
1-27
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1-28 Holdup Model
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ThProperties Tab
The following fluid properties for each phase in the holdup 
appear on the Properties tab:
• Temperature
• Pressure
• Flow
• Molar Fraction of the specific phase in the holdup
• Enthalpy
• Density
• Molecular Weight
Product Nozzle 
Efficiency
Product nozzle efficiencies are used only when 
there is flow reversal in the product streams. In 
this situation, the product nozzles act as effective 
feed nozzles.
Recycle Efficiency Essentially, the recycle stream represents the 
material already existing in the holdup. Recycle 
efficiencies represent how much of the material in 
the holdup participates in the flash.
 Figure 1.11
Nozzle Efficiency Description1-28
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Dynamic Theory 1-29
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ThCompositions Tab
The compositional molar fractions of each phase in the holdup 
displays in the Compositions tab.
1.4 Pressure Flow Solver
Aspen HYSYS offers an advanced method of calculating the 
pressure and flow profile of a simulation case in Dynamics 
mode. Almost every unit operation in the flowsheet can be 
considered a holdup or carrier of material (pressure) and 
energy. A network of pressure nodes can therefore be conceived 
across the entire simulation case. The P-F solver considers the 
integration of pressure flow balances in the flowsheet. There are 
two basic equations which define most of the pressure flow 
network and these equations only contain pressure and flow as 
variables:
• Resistance equations. Which define flow between 
pressure holdups.
• Volume balance equations. Which define the material 
balance at pressure holdups.
The pressure flow balances both require information from and 
provide information to the holdup model. While the holdup 
model calculates the accumulation of material, energy, and 
composition in the holdup, the pressure flow solver equations 
 Figure 1.121-29
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1-30 Pressure Flow Solver
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Thdetermine the pressure of the holdup and flow rates around it. 
The holdup model brings the actual feed and product stream 
properties to holdup conditions for the volume balance 
equations using a rigorous or approximate flash. The pressure 
flow solver returns information essential to the holdup model 
calculations: the pressure of the holdup or the flow rates of 
streams around the holdup.
1.4.1 Simultaneous Solution in 
Pressure Flow Balances
All material streams within Aspen HYSYS can be solved for 
pressure and flow. All unit operations can be solved for 
pressure. As an example, consider the following flowsheet. 
There are 26 variables to solve for in the PF matrix. Twelve 
material streams contribute 24 variables to the flowsheet. The 2 
vessels, V-100 and V-101, contribute 1 variable each. The valve 
and tee operations are not considered nodes. These unit 
operations define a pressure flow relation between the inlet and 
exit streams, but rarely are they modeled with any inventory. 
A pressure-flow matrix is setup which solves the variables 
required. The matrix consists of: Volume balance equations, 
Resistance equations and Pressure-Flow specifications input by 
 Figure 1.131-30
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Thyou. The number of pressure flow specifications that need to be 
provided are discussed in Degrees of Freedom Analysis 
section.
1.4.2 Basic Pressure Flow 
Equations
The equations that are discussed in this section define the basis 
of the pressure flow network.
Volume Balance
For equipment with holdup, an underlying principle is that the 
physical volume of the vessel, and thus, the volume of material 
in the vessel at any time remains constant. Therefore, during 
calculations in dynamics, the change in volume of the material 
inside the vessel is zero:
where:  
V = volume of the vessel
t = time
flow = mass flowrate
h = holdup
P = vessel pressure
T = vessel temperature
As such, a vessel pressure node equation is essentially a 
(1.17)
(1.18)
V Constant f flow h P T, , ,( )= =
dV
dt
------ 0=1-31
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1-32 Pressure Flow Solver
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Thvolumetric flow balance and can be expressed as follows:
In the volume balance equation, pressure and flow are the only 
two variables to be solved in the matrix. All other values in the 
equation are updated after the matrix solves. Each vessel 
holdup contributes at least one volume balance equation to the 
pressure-flow matrix. When sufficient pressure-flow 
specifications are provided by you, any unknown(s) can be 
solved whether it be a vessel pressure or one of its flowrates.
The volume balance equation allows you to observe pressure 
effects in the vapour holdup due to disturbances in the feed. 
Consider a separator whose feed flow suddenly increases. 
Assume that the exit streams from the separator are specified 
by you and are thus, constant. The vessel pressure would 
increase for two reasons:
• Because the material of the exit streams remain 
constant, an increase in the vapour feed flow would 
increase the vapour holdup. An increase in the vapour 
holdup means that a larger amount of material is 
compressed into the same vapour volume resulting in a 
vessel pressure increase.
• The increase in the liquid level causes the vapour holdup 
to occupy a smaller volume within the vessel, causing the 
vessel pressure to rise.
Resistance Equations
Flows exiting from a holdup are calculated from a volume 
balance equation, specified by you, or calculated from a 
resistance equation. In general, the resistance equation 
calculates flowrates from the pressure differences of the 
surrounding nodes. Aspen HYSYS contains unit operations such 
as valves and heat exchangers which calculate flowrates using 
resistance equations. The resistance equations are modeled 
Volume change due to pressure + Volume change due to 
flows + Volume change due to temperature + 
Volume change due to other factors = 0
(1.19)1-32
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Dynamic Theory 1-33
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Thafter turbulent flow equations and have the form:
where:  
Flow = mass flow rate
k = conductance, which is a constant representing the 
reciprocal of resistance to flow
 = frictional pressure loss, which is the pressure drop 
across the unit operation without static head 
contributions
Equation (1.20) is a simplified form of the basic valve 
operation equation which uses the valve flow coefficient Cv. The 
mass flowrate through the valve is a function of the valve flow 
coefficient and the frictional pressure drop across the valve:
where:  
Flow = mass flowrate
Cv = conductance, which is a constant representing the 
reciprocal of resistance to flow
P1 = upstream pressure
P2 = downstream pressure
As shown, a resistance equation relates the pressures of two 
nodes and the flow that exists between the nodes. The following 
unit operations have a resistance equation associated with 
them. 
(1.20)
(1.21)
Unit Operation Resistance Term
Valve With a pressure flow specification, you can specify 
conductance, Cv, on the Specs page of the 
Dynamics tab.
Pump The heat flow and pump work define the pressure 
flow relation of the pump. These parameters can 
be specified and/or calculated on the Specs page 
of the Dynamics tab.
Flow k ΔP=
ΔP
Flow f Cv P1 P2,,( )=
For a more detailed 
description on the 
individual unit operations 
and the resistance 
equations associated with 
them, see the appropriate 
unit operation section in 
the Aspen HYSYS 
Operations Guide.1-33
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1-34 Pressure Flow Solver
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Th1.4.3 Pressure Flow 
Specifications
In Dynamics mode, you can specify the pressure and/or flow of 
a material stream in a flowsheet. The pressure flow 
specifications are made in the Dynamics tab of the Material 
Stream property view. 
To satisfy the degrees of freedom of the pressure-flow matrix, 
you must input a certain number of pressure-flow specifications. 
The volume balance equations, resistance equations, and 
pressure-flow relation equations make up a large number of 
equations in the pressure-flow matrix. However, you should be 
aware of the specifications that are needed before the matrix 
solves.
Degrees of Freedom Analysis
In almost all cases, a flowsheet being modeled dynamically 
using pressure-flow requires one pressure-flow specification per 
flowsheet boundary stream. A flowsheet boundary is one that 
crosses the model boundary and is attached to only one unit 
operation. Examples of such streams are the model’s feed and 
product streams. All other specifications for the flowsheet are 
Compressor/
Expander
The heat flow and compressor work define the 
pressure flow relation of the compressor. These 
parameters can be specified and/or calculated on 
the Specs page of the Dynamics tab.
Heater/Cooler/Heat 
Exchanger/Air 
Cooler/LNG
With a pressure flow specification, you can specify 
the k-value on the Specs page of the Dynamics 
tab.
Tray Sections, Weir 
Equation
The Weir equation determines liquid flow rate from 
the tray as a function of liquid level in the tray. 
Tray geometry can be specified on the Sizing page 
of the Ratings tab.
Tray Sections, K-
Value
The K-value is used to determine vapour flow 
exiting from the tray as a function of the pressure 
difference between trays. K-values can either be 
calculated or specified on the Specs page of the 
Dynamics tab.
Unit Operation Resistance Term
For more information on 
specifying Pressure-Flow 
specifications for a 
material stream, see 
Chapter 12 - Streams 
in the Aspen HYSYS 
Operations Guide.1-34
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Thhandled when each unit operation is sized using the 
conductance or valve flow coefficient. 
The following example confirms the “one P-F specification per 
flowsheet boundary stream” rule. In Figure 1.14, since there 
are four flowsheet boundary streams, you are required to make 
four pressure-flow specifications for the pressure flow matrix to 
solve. 
Notice that the pressure flow specifications do not necessarily 
have to be set for each flowsheet boundary stream. 
Specifications can be made for internal flowsheet streams as 
long as there is one P-F specification per flowsheet boundary 
stream.
In the flowsheet shown above, there are eight streams and one 
vessel holdup. To fully define the pressure flow matrix, the 
pressure and flow for each material stream and the pressure of 
each holdup must be solved for. In short, two variables are 
required for each material stream and one variable is required 
for each holdup: 
 Figure 1.14
8 material streams x 2 + 1 vessel holdup x 1 = 17 
pressure-flow variables
(1.22)1-35
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1-36 Pressure Flow Solver
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ThThe accumulation or amount of holdup is solved using material 
balances in the holdup model. Although the holdup is not solved 
by the pressure-flow matrix, it is used by the volume balance 
equation to calculate the vessel pressure of the holdup which is 
a variable in the matrix.
The pressure and flow of material streams are named Pstream name 
and Fstream name, respectively. The pressure of the holdup is 
named PH. 
There are a number of equations which describe the relationship 
between the pressures and flows in this network. They are as 
shown below:
Pressure-Flow Equation Description # of Eqns
Separator
Volume Balance equation  
The volume balance relates PH with F2, F3 and F5.
1
General Pressure 
relation
If the static head contribution option in the integrator is not 
activated, this general pressure relation is observed.
3
Valves
Resistance equations
This is the general form of the valve resistance equation. 
The actual equations vary according to inlet stream 
conditions.
3
General Flow relations
Since the valves are usually not specified with holdup, this 
relation is observed.
3
Mixer
General Pressure 
relation
The equalize option is recommended for the operation of 
the mixer in Dynamics mode. If this option is activated, this 
general pressure relation is observed.
2
dPH
dt
--------- f P T holdup flows,, ,( )=
PH P2 P3 P5= = =
F2 KVLV100 P1 P2–
F4 KVLV101 P3 P4–
F8 KVLV102 P7 P8–=
=
=
F1 F2
F3 F4
F7 F8=
=
=
P5 P6 P7= =1-36
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Dynamic Theory 1-37
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ThWith 17 variables to solve for in the network and 13 available 
equations, the degrees of freedom for this network is four. 
Therefore, four variables need to be specified to define this 
system. This is the same number of flowsheet boundary 
streams.
Pressure-Flow Specification 
Guidelines
The previous section outlined the number of pressure-flow 
specifications that are required by the flowsheet in order for the 
degrees of freedom to be satisfied. This section presents 
possible PF specifications that can be made for the inlet and exit 
streams of stand alone operations. 
The purpose of this section is to demonstrate the range of 
specifications that can be made for different unit operations in 
Aspen HYSYS. It is hoped that this section provides insight as to 
what should and should not be specified for each unit operation.
Valve
Rating information for the valve operation including the valve 
type and Cv values can be input on the Sizing page in the 
Ratings tab.
The dynamic valve can either be specified as having a set 
pressure drop or a pressure flow relation. This option is set on 
the Specs page of the Dynamics tab in the valve property view.
• For a pressure drop specification on the valve: one 
pressure spec and one flow spec is required for the inlet 
and exit streams.
General Flow relation
Since the mixer is usually not specified with holdup, this 
relation is observed.
1
Total Number of Pressure Flow Equations 13
Pressure-Flow Equation Description # of Eqns
F7 F5 F6+=1-37
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1-38 Pressure Flow Solver
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Th• For a pressure-flow specification on the valve: two 
pressures are specified or one pressure and one flow.
Pressure and level control can be achieved in a separator using 
valves on the vapour and liquid streams, respectively. It is best 
to use a pressure specification downstream of each valve. The 
percent openings on each valve can then be used to control the 
flow through each valve with a PID controller.
Heat Exchanger/Cooler/Heater
The dynamic heat exchanger can be specified as having a set 
pressure drop or a Overall K-Value (pressure-flow) relation. This 
option is set on the Specs page of the Dynamics tab in the heat 
exchanger property view:
• For a pressure drop specification on either the tube side 
or shell side: one pressure spec and one flow spec is 
recommended.
• For a K-value spec on either the tube or shell side: two 
pressures can be specified or one pressure and one flow.
K-values can be calculated using the Calculate K button on the 
Specs page of the Dynamics tab in the operation’s property 
view.
Heater and cooler operations are much like heat exchangers. 
However, they only have a single K-value on their process side.
The P-F spec option for conductance-type unit operations 
should be used as much as possible since it is much more 
realistic in determining pressure flow relations in an actual 
plant. The pressure drop option is provided to ease the 
transition between steady state and Dynamics mode. The 
pressure drop option can help more difficult simulations run 
since the initial exit stream conditions of the valve can be 
easily calculated using the pressure drop option.
The heat exchange operations, like the valve, should use the 
P-F spec option as much as possible to simulate actual 
pressure flow relations in the plant.1-38
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ThSeparators
Rating information including the volume of the vessel, boot 
capacity, and nozzle location can be input on the Sizing and 
Nozzles pages in the Ratings tab.
A separator with no valves attached to the inlet and exit streams 
requires at most one pressure specification. The other two 
streams are specified with flows. A more realistic way to run the 
separator is to attach valves to the inlet and exit streams of the 
vessel. The boundary streams of the separator with valves 
should be specified with pressure.
Condenser/Reboiler
Rating information for the condenser and reboiler including the 
vessel volume, boot capacity, and nozzle location can be input 
on the Sizing and Nozzles pages of the vessel’s Ratings tab.
It is highly recommended that the proper equipment be added 
to the reflux stream (e.g., pumps, valve, etc.). In all cases, level 
control for the condenser should be used to ensure a proper 
liquid level.
The Partial Condenser has three exit streams: 
• overhead vapour stream
• reflux stream
• distillate stream 
All three exit streams must be specified when attached to the 
main tray section. One pressure specification is recommended 
for the vapour stream. The other two exit streams must be 
specified with flow rates. Another option is to specify a Reflux 
Flow/Total Liq Flow value on the Specs page in the Dynamics 
tab. In this case, only one flow spec is required on either the 
reflux or distillate stream.
• The Fully-Refluxed Condenser has two exit streams: the 
overhead vapour stream and the reflux stream. 
One pressure and flow specification is required for the 
two exit streams.1-39
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1-40 Pressure Flow Solver
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Th• A Fully-Condensed Condenser has two exit streams: the 
reflux stream and the distillate stream. 
There are several possible configurations of pressure flow 
specifications for this type of condenser. A flow 
specification can be used for the reflux stream and a 
pressure flow spec can be used for the distillate stream. 
Two flow specifications can be used, however it is 
suggested that a vessel pressure controller be setup with 
the condenser duty as the operating variable.
• The Reboiler has two exit streams: the boilup vapour 
stream and the bottoms liquid stream. 
Only one exit stream can be specified. If a pressure 
constraint is specified elsewhere in the column, this exit 
stream must be specified with a flow rate.
Separation Columns
For all separation columns, the tray section parameters 
including the tray diameter, weir length, weir height, and tray 
spacing can be specified on the Sizing page in the Ratings tab of 
the Main TS property view.
The basic Absorber column has two inlet and two exit streams. 
When used alone, the absorber has four boundary streams and 
therefore requires four pressure-flow specifications. A pressure 
specification is always required for the liquid product stream 
leaving the bottom of the column. A second pressure 
specification should be added to the vapour product of the 
column, with the two feed streams having flow specifications.
The basic Refluxed absorber column has a single inlet and two 
or three exit streams, depending on the condenser 
configuration. When used alone, the refluxed ratios has three or 
four boundary streams (depending on the condenser) and 
requires four or five pressure-flow specifications; generally two 
pressure and three flow specifications. A pressure specification 
is always required for the liquid product stream leaving the 
bottom of the column.
The Reboiled Absorber column has a single inlet and two exit 
streams. When used alone, the reboiled absorber has three 
boundary streams and therefore requires three pressure-flow 
specifications; one pressure and two flow specifications. A 1-40
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Dynamic Theory 1-41
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Thpressure specification is always required for the vapour product 
leaving the column.
The basic Distillation column has one inlet and two or three exit 
streams, depending on the condenser configuration. When used 
alone, the distillation column has three or four boundary 
streams, but requires four or five pressure-flow specifications; 
generally one pressure and three or four flow specifications. The 
extra pressure flow specification is required due to the reflux 
stream, and is discussed in Section  - Column-Specific 
Operations from the Aspen HYSYS Operations Guide. 
Compressor/Expander/Pump
Rating information for the dynamic compressor, expander, and 
pump operations can be input on the Curves and Inertia pages 
in the Ratings tab.
In general, two specifications should be selected in the 
Dynamics Specifications group in the Specs page of the 
Dynamics tab in order for these unit operations to fully solve. 
You should be aware of specifications which causes 
complications or singularity in the pressure flow matrix. Some 
examples of such cases are:
• The Pressure rise checkbox should be cleared if the 
inlet and exit stream pressures are specified.
• The Speed checkbox should be cleared if the Use 
Characteristic Curves checkbox is cleared.
The compressor, expander and pump operations have one inlet 
stream and one exit stream. Two pressures are specified for the 
inlet and exit streams or one pressure and one flow are 
specified.
Mixer/Tee
The dynamic mixer and tee operations are very similar. It is 
recommended that the mixer be specified with the Equalize All 
option in Dynamics mode. It is also recommended that the 
dynamic tee not use the dynamic splits as specifications. These 1-41
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1-42 Dynamic Operations: General 
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Thoptions are set on the Specs page of the Dynamics tab in their 
respective operation property views.
By specifying the dynamic mixer and tee as recommended, the 
pressure of the surrounding streams of the unit operation are 
equal if static head contributions are not considered. This is a 
realistic situation since the pressures of the streams entering 
and exiting a mixer or tee must be the same. With the 
recommended specifications, flow to and from the tee is 
determined by pressures and resistance through the flowsheet. 
This is more realistic than using the split fractions which can also 
cause complications with regard to flow reversal.
A number of streams can enter or exit a mixer or tee. For stand 
alone operations, one stream must be specified with pressure. 
The other inlet/exit streams are specified with flow.
1.5 Dynamic Operations: 
General Guidelines
This section outlines some guidelines or steps that you follow in 
order to create and run a simulation case in Dynamics mode.
It is possible to create a case directly in Dynamics mode. Unit 
operations can be added just as easily in Dynamics mode as in 
steady state. The integrator should be run after every few 
additions of a unit operation to initialize exit stream conditions 
for the added unit operations.
It is also possible for you to build a dynamics case by first 
creating the case in Steady State mode. You can make the 
transition to Dynamics mode with some modifications to the 
flowsheet topology and stream specifications. Section 1.5.2 - 
Moving from Steady State to Dynamics outlines some 
general steps you can take to create a dynamics case from 
Steady State mode. 
The Dynamics Assistant can be used to quickly modify the 
steady state flowsheet so that it has a correct set of pressure 1-42
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Dynamic Theory 1-43
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Thflow specifications. It is important to note, however, that not all 
the modifications suggested by the assistant are always suited 
to your case.
It is suggested when you are first learning dynamics, that you 
build simple cases in Steady State mode so that the transition to 
Dynamics mode is relatively easy. Once the transition from 
Steady State to Dynamics mode is made, other unit operations 
can easily be added to better model the process system. If you 
are more experienced, you can adopt different and more 
efficient ways to create a dynamics case.
1.5.1 Differences between 
Dynamic & Steady State 
Mode
It is apparent that the specifications required by the unit 
operations in Dynamics mode are not the same as the Steady 
State mode. This section outlines the main differences between 
the two modes in regards to specifying unit operations.
Steady State
The Steady State mode uses modular operations which are 
combined with a non-sequential algorithm. Information is 
processed as soon as it is supplied. The results of any 
calculation are automatically propagated throughout the 
flowsheet, both forwards and backwards.
Material, energy, and composition balances are considered at 
the same time. Pressure, flow, temperature, and composition 
specifications are considered equally. For example, a column’s 
overhead flow rate specification is replaced by a composition 
specification in the condenser. The column can solve with either 
specification.1-43
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1-44 Dynamic Operations: General 
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ThDynamics
Material, energy and composition balances in Dynamics mode 
are not considered at the same time. Material or pressure-flow 
balances are solved for at every time step. Energy and 
composition balances are defaulted to solve less frequently. 
Pressure and flow are calculated simultaneously in a pressure-
flow matrix. Energy and composition balances are solved in a 
modular sequential fashion.
Because the pressure flow solver exclusively considers pressure-
flow balances in the network, P-F specifications are separate 
from temperature and composition specifications. P-F 
specifications are input using the “one P-F specification per 
flowsheet boundary stream” rule. Temperature and composition 
specifications should be input on every boundary feed stream 
entering the flowsheet. Temperature and composition are then 
calculated sequentially for each downstream unit operation and 
material stream using the holdup model.
Unlike in Steady State mode, information is not processed 
immediately after being input. The integrator should be run 
after the addition of any unit operation to the flowsheet. Once 
the integrator is run, stream conditions for the exit streams of 
the added unit operation is calculated.
1.5.2 Moving from Steady 
State to Dynamics
You should be aware that flow in the plant occurs because of 
resistance and driving forces. Before a transition from steady 
state to dynamics mode occurs, the simulation flowsheet should 
be set up so that a realistic pressure difference is accounted for 
across the plant. 1-44
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ThThe following table indicates some basic steps you can take to 
set up a case in Steady State mode and then switch to 
Dynamics mode.
Step Description
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 Dynamics mode. These unit 
operations include the separator operation and tray 
sections in a column operation. 
Add unit operations, such as valves, heat exchangers, and 
pumps, which define a pressure flow relation to these 
streams. It is also possible to specify a flow specification 
on this stream instead of using an operation to define the 
flow rate.
Equipment 
Sizing
Size all the unit operations in the simulation using actual 
plant equipment or predefined sizing techniques. Sizing of 
trays in columns can be accomplished using the Tray 
Sizing utility available from the Utilities page. Vessels 
should be sized to accommodate actual plant flowrates 
and pressures while maintaining acceptable residence 
times. 
General Equipment Sizing Rules
Vessels (Separators, Condensers, Reboilers) should be 
sized for 5 - 15 minutes of liquid holdup time. Sizing and 
Costing calculations are also performed using the Vessel 
Sizing utility in the Sizing page of the Rating tab.
Valves should be sized using typical flowrates. The valve 
should be sized with a 50% valve opening and a pressure 
drop between 15 and 30 kPa.
Column Tray Sizing Rules
Tray Sizing can be accomplished for separation columns 
using the Tray Sizing utility in the Utilities page. Any use 
of utilities should be restricted to Steady State mode. The 
trays are sized according to the existing flow rates and the 
desired residence times in the tray. Important variables 
include: 
• Tray diameter
• Weir length
• Weir height
• Tray spacing1-45
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1-46 Dynamic Operations: General 
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ThAdjusting 
Column 
Pressure
In steady state, the pressure profile of the column is user 
specified. In dynamics, it is calculated using dynamic 
hydraulic calculations. If the steady state pressure profile 
is very different from the calculated pressure drop, there 
can be large upsets in flow in the column when the 
integrator is run. 
A reasonable estimate of the column’s pressure profile can 
be calculated using the Tray Sizing utility. This utility 
provides a  value in the Results tab. The 
column pressure profile can be calculated using this value, 
the  value, and a desired pressure 
specification anywhere on the column.
You can change the  value to achieve a 
desired pressure profile across the column. This can easily 
be done by modifying the Weir height in the Ratings tab in 
the Tray Sizing utility. Reducing the weir height lowers the 
static head contributions and lowers the  
value. 
In Dynamics mode, the Nozzle Pressure Flow K-factors 
(found on the Dynamics tab of the Main TS property view) 
can also be adjusted to better model the pressure drop 
across the column.
Logical 
Operations
Some logical operations from the steady state are 
ignored. The Adjust operation can be replaced by PID 
Controllers. The recycle operation is redundant in 
Dynamics mode.
Adding Control 
Operations
Identify key control loops that exist within the plant. 
Implementing control schemes increases the realism and 
stability of the model. Disturbances in the plant can be 
modeled using the Transfer Function operation. The 
Events Scheduler can be used to model automated 
shutdowns and startups.
Enter Aspen 
HYSYS 
Dynamic 
Environment
Click on the Dynamic Mode icon to switch from Steady 
State mode to Dynamics mode. 
Step Description
MaxΔP Tray⁄
MaxΔP Tray⁄
MaxΔP Tray⁄
MaxΔP Tray⁄
Dynamic Mode icon1-46
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ThAdding 
Pressure-Flow 
Specifications
Specify one pressure-flow specification for each flowsheet 
boundary stream.
Controllers play a large part in stabilizing the PF Solver. 
Precautions
• Pay special attention to equipment with fixed 
pressure drops. Any fixed pressure drop 
specifications in equipment can yield unrealistic 
results, such as flow occurring in the direction of 
increasing pressure. Remember to check for fixed 
pressure drops in the reboiler and condenser of 
columns.
• Be cautious of Heaters/Coolers with fixed duties. This 
can cause problems if the flow in the heater/cooler 
happens to fall to zero. It is recommended to use a 
controller, or a Spreadsheet function, or a 
temperature specification to control the temperature 
of a stream.
• Feed and product streams entering and exiting tray 
sections should be at the same pressure as the tray 
section itself. Any large pressure differences between 
a feed or product stream and its corresponding tray 
section can result in large amounts of material 
moving into or out of the column.
It is necessary to isolate and converge single pieces of 
equipment in the plant using the Ignored feature for each 
unit operation if there is an especially large number of unit 
operations in the flowsheet.
Run the Integrator after any unit operation is added in 
Dynamics mode. Unlike the steady state environment, the 
exit streams of unit operations in Dynamics mode are not 
calculated until the Integrator is run. The Integrator 
should be run long enough to obtain reasonable values for 
the exit streams of the new operations.
Step Description
For more information 
regarding pressure-flow 
specifications for 
individual unit operations, 
refer to Section 1.4.3 - 
Pressure Flow 
Specifications.
For more information 
regarding the 
implementation of 
controllers in Aspen 
HYSYS, see Section 
3.5.4 - Setting Up a 
Control Strategy.1-47
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1-48 Dynamic Operations: General 
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ThTrouble 
Shooting
Error messages appear once the integrator is run.
Dynamics Assistant can be also be used to trouble shoot 
specification problems.
Too Many Specifications/Not Enough Specifications 
• The “Too many specifications” message indicates 
that Aspen HYSYS has detected too many 
specifications. The Equation Summary property view 
that appears with the message can provide help 
indicating the specification that is most likely not 
required. Click the Full Analysis button (or 
Partitioned Analysis button, if it is made available). 
At this point, Aspen HYSYS examines possible 
problem areas with the simulation case. Clicking the 
Extra Specs tab reveals the variable(s) most likely 
not required by Aspen HYSYS.
• The “Not enough specifications” message indicates 
that the simulator has detected too few 
specifications. The Extra Specs tab in the Equation 
Summary property view indicates possible variables 
that are missing from the simulation case. The 
Dynamics Assistant can aid in identifying which P-F 
specifications should be added or deleted from a 
dynamic simulation case.
Singular Problem
• This message indicates that not all of the equations 
in the P-F solver matrix are independent of one 
another. This occurs when one or more equations are 
redundant. For instance, if a valve operation is using 
a pressure drop specification, the inlet and exit 
streams cannot both be specified with pressure. The 
pressure drop equation becomes redundant. It is 
useful to overspecify a singular problem. Aspen 
HYSYS might be able to identify the redundant 
pressure flow specification and allow the case to 
solve.
The Pressure Flow Solver Failed to Converge
• This message indicates that one or more pressure-
flow specifications are unreasonable. This message 
can also appear if there are sudden large upsets to 
the simulation case. It is helpful to enter the 
Equation Summary property view to identify problem 
areas in the flowsheet. Click the Full Analysis button 
(or Partitioned Analysis button, if it is made 
available). By clicking the Update Sorted List button 
in the Unconverged tab, Aspen HYSYS shows the 
type of equation, location, and scaled error 
associated with the unconverged nodes in the 
flowsheet.
• Pay special attention to the unit operations with the 
largest errors in the Uncoverged tab. Check the 
vessel volumes of the uncoverged unit operations 
and ensure they are sized with reasonable residence 
times. Check the size of the valves attached to the 
unconverged unit operations.
Step Description
For more information 
regarding the Equations 
Summary property view, 
see Chapter 2 - 
Dynamic Tools.1-48
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Dynamic Theory 1-49
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Th1.6 Aspen HYSYS 
Dynamics
Aspen HYSYS Dynamics contains Fidelity (an extension of Aspen 
HYSYS) that provides advanced dynamic features to your 
simulation. Aspen HYSYS Dynamics allows you to put together 
very detailed models for operator training work or detailed 
dynamic studies. The capabilities exclusive to Aspen HYSYS 
Dynamics 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 starting up and shutdown of rotating equipment.
To use the Aspen HYSYS Dynamics features, you must 
purchase a Aspen HYSYS Dynamics license. If you do not 
have a Aspen HYSYS Dynamics license, or it is not activated, 
you are not able to see or access these features.1-49
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1-50 Aspen HYSYS Dynamics
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ThTo activate the features available with the Aspen HYSYS 
Dynamics license:
1. On the Aspen HYSYS desktop, select Tools | Preferences 
to access the Session Preferences property view.
2. Click on the Simulation tab, and select the Licensing page 
as shown in the figure below.
3. Open the Aspen HYSYS Fidelity drop-down list and select 
one of the following behavior option to access the Aspen 
HYSYS Dynamics license:
• No default behavior. This option prompts you to select 
a behavior option when changing a variable value that 
requires the license.
• Ask every time. This option prompts you to check the 
license when changing a variable value that requires the 
license.
• Validate immediately. This option causes the license in 
question to be validated as soon as any value that 
requires this license is changed. If the license is not 
available, an error message appears and the value is not 
changed. This option is recommended for users with 
standalone keys.
• Check out now and validate later. This option causes 
the license to be checked out immediately, but it is 
validated at a later stage (for example, the Aspen HYSYS 
Dynamics license is validated when you try to start the 
integrator). 
 Figure 1.151-50
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Dynamic Theory 1-51
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ThThis option is useful because you does not see any slow-
down due to network problems. Aspen HYSYS is 
assumed, however, that the license is there and you can 
make changes resulting in the case being stopped when 
the validation is complete. This option is recommended 
for users with network based security.
• Check out when in the relevant mode. This option 
checks the licenses in a run-time usage mode 
(essentially pay-per-use), so the Aspen HYSYS Dynamics 
license is only checked out while the case is in dynamics 
mode, and the Aspen HYSYS Dynamics license is 
returned while the case is in steady state mode. This 
option is recommended for users with token based 
security.
• Don’t check out. This option means that licenses are 
not checked out. It is recommended for users that do not 
have the licenses available. Input requiring the license is 
then ignored by Aspen HYSYS.
1.6.1 Detailed Heat Model
The Detailed Heat model is located on the Heat Loss page of the 
Rating tab for most of the unit operations that allow detailed 
heat model calculation. Refer to the figure below to see how the 
options are displayed:
There are two values that are common to each of the three radio 
buttons found in the Detailed Heat Loss Model: the Overall Heat 
 Figure 1.161-51
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1-52 Aspen HYSYS Dynamics
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ThLoss value and the Overall Heat Transfer Area.
The other parameters that appear by selecting one of the three 
radio buttons are described in the table below.
The governing equations relating heat loss from the vessel and 
the Detailed heat loss parameters shown here are discussed in 
Section 1.3.4 - Heat Loss Model.
Radio Button Description
Temperature 
Profile
Displays the temperatures of the various fluids, walls, 
insulation and surroundings.
Notice the parameters that appear on this page varies 
between different unit operations.
Conduction Displays the conductive properties of the wall and 
insulation. The following properties can be specified by 
you:
• Conductivity of material
• Thickness of material
• Heat capacity of material
• Density of material
Equation (1.14) and (1.15) demonstrate how the 
parameters are used by the heat loss model.
The heat transfer area is calculated from the vessel 
geometry. The rest of the heat transfer parameters are 
modified. The insulation thickness can be specified as 
zero to model vessels without insulation. The metal 
wall must be specified with a finite thickness.
Convection Displays the convective heat transfer coefficients for 
heat transfer within the holdup and heat transfer 
occurring from the outside the holdup to the 
surroundings.
Both the inside and outside heat transfer coefficients 
are modified from their default values.1-52
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Th1.6.2 Nozzles
The Nozzles page of the Rating tab contains information 
regarding the elevation and diameter of the nozzles. 
The elevations of each nozzle attached to the part of the 
equipment appear relative to two reference points:
• Inlet connection. Point from which the stream is 
entering the unit operation. The inlet connection 
indicates the height of the nozzle entering the unit 
operation.
• Outlet connection. Point where the stream comes out 
from a piece of equipment. The outlet connection 
indicates the height of the nozzle coming out from the 
unit operation.
You can view these elevations of each nozzle on the PFD by 
pressing:
• SHIFT I. Shows all the inlet connection point associated 
with each nozzle. 
• SHIFT O. Shows all the outlet connection point.
 Figure 1.17
You must have the Aspen HYSYS Dynamics license in order 
to activate the Nozzle page on the Rating tab.1-53
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1-54 Aspen HYSYS Dynamics
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ThThe inlet and outlet connection point are all relative to ground 
level, which is set by default at 0. All the elevation points can be 
user-specified but default settings are supplied by Aspen HYSYS 
as a starting point. You can adjust the elevations for each nozzle 
on the Nozzle page of the Rating tab as desired. To turn off the 
elevation display on the PFD, press SHIFT N and the name of 
each stream reappears.
The following is a summary of the Nozzle Parameters section:
The nozzle diameter and elevation as well as the levels of the 
phases inside the vessel determine what flows out through the 
nozzle. If the nozzle opening is exposed to several different 
phases, then the amount of each phase that flows out through 
the nozzle depends on how much of the nozzle opening is 
exposed to that particular phase. 
The nozzle elevation refers to the center of the nozzle opening, 
and not the bottom or top of it. For practical purposes, Aspen 
HYSYS moves nozzles located at the extreme bottom or top of 
the vessel very slightly. This minor adjustment is not displayed 
and does not impact the static head contributions. The 
adjustment is mostly done so that users do not have to consider 
The elevation of the nozzle is displayed as  on the 
PFD when the condition is inapplicable. For example, the 
final product stream will have an  inlet elevation 
and the initial feed stream will have an  outlet 
elevation.
Nozzle Parameter Description
Diameter The nozzle diameter appears and can be modified 
in this cell.
Elevation (Base) Elevation of the nozzle above the base (the bottom 
of the piece of equipment).
Elevation (Ground) Elevation of the nozzle above the ground.
Elevation (% of 
Height)
Elevation of the nozzle as a percentage of the 
height of the vessel.
Base Elevation 
relative to Ground
How far the piece of equipment is from the 
ground.
Diameter Allows you to specify the Diameter of the vessel.
Height Allows you to specify the Height of the vessel.1-54
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Dynamic Theory 1-55
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Thnozzle diameters carefully when setting elevations. The 
adjustment makes the calculated values more realistic.
Nozzle elevations can impact the pressure profile if static head 
contributions are enabled through the Integrator property view. 
Static head contributions come from levels inside vessels as well 
as elevation differences between pieces of equipment connected 
together. For example, consider a vessel filled with liquid. The 
pressure shown for the vessel is the pressure right at the top of 
the vessel. For product outlet streams the pressure is typically 
equal to the vessel pressure plus an appropriate static head 
contribution that depends on the nozzle elevation and the 
contents of the vessel. An outlet at the bottom of the vessel can 
have a higher pressure than an outlet located closer to the top 
of the vessel.
It is important to note that the pressure displayed for a stream 
is a point value. The stream pressure is taken immediately at 
the exit point of the operation. 
 Figure 1.18
If a valve is attached to the outlet 
stream of a vessel and the valve is 
at a much lower elevation than 
the vessel, then the pressure in 
the outlet stream does not show 
the static head contribution. The 
outlet stream sampling point is 
located at the vessel outlet and 
not at the lower valve inlet.
Height 
difference
Pressure taken 
at this point of 
the stream.1-55
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1-56 Aspen HYSYS Dynamics
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ThMain Tray Section
The Nozzle page for the Main Tray Section in the column 
environment is setup different from all the other unit operations.
The information is broken down into three groups: 
• Tray by tray for internal nozzle
• Feed nozzles
• Product nozzles 
The following information is available within these groups.
 Figure 1.19
Object Description
Traysection 
Elevation Relative to 
Ground
Specify the height of the tray section above the 
ground. The height is measured as the distance 
between the ground to the bottom tray.
Holdup RG The height of each tray relative to the ground. The 
values are dependent on the tray spacing and the 
value entered in the Tray section Elevation 
Relative to Ground cell.
VToAbove The elevation for vapour leaving the tray relative 
to the tray. By default this value is the tray 
spacing.
LToBelow The elevation for liquid leaving the tray relative to 
the tray. By default this value is zero.1-56
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ThThe discussion in the previous section about nozzle elevations, 
diameters and phases apply to tray section as well. Each stage 
is treated as if it was a vessel in itself. However, the liquid on a 
tray is typically an emulsion. Hence if there are two liquid 
phases in the tray, then both liquid phases flow through the 
nozzle located at the bottom of the tray.
1.6.3 Control Valve Actuator
The Actuator page, located on the Dynamics tab of the Valve 
unit operation, allows you to model valve dynamics in the valve 
operation. This page also contains information regarding the 
dynamic parameters of the valve and the per cent open 
positions of the actuator and the valve.
A control valve in Aspen HYSYS consists of a valve and an 
Elevation RH This value can be specified for both the feed and 
product nozzles. This gives the elevation of the 
nozzles relative to the height of the column.
Diameter Specifies the diameter of the nozzle for the Feed 
nozzles, Product nozzles, VToAbove and LtoBelow 
sections.
 Figure 1.20
Object Description1-57
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1-58 Aspen HYSYS Dynamics
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Thactuator. They are defined as follows:
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 behavior in actual control 
valves.
Valve Modes
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
Instantaneous Mode
In this mode, the actuator moves instantaneously to the desired 
actuator position defined by the controller. The equation defining 
the relationship is:
Valve component Description
Actuator An actuator is a device which applies the force required 
to cause movement in the valve.
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.
(1.23)Act% ActDesired%=1-58
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ThFirst Order Mode
A first order lag can be modeled in the response of the actuator 
position to changes in the desired actuator position. The 
movement of the actuator is defined by the solution of the 
following differential equation:
The actuator time constant, , in Equation (1.24) can be 
specified in the Actuator Time Constant cell.
Linear Mode
The actuator can be modeled to move to the desired actuator 
position at a constant rate. The actuator moves according to the 
following equation (if the desired actuator position is above the 
current actuator position):
The linear rate can be specified in the Actuator Linear Rate cell. 
Typical stroke times (closure rates) are as follows:
• Electric-Hydraulic Actuators: approximately 12 
inches/minute
• Piston Actuators (Motor Driven): under 70 inches/
minute
(1.24)
(1.25)
τd Act%( )
dt
--------------------- Act%+ ActDesired%=
τ
Act% Actuator Linear Rate( )Δt Acto% 
until Act%
+
ActDesired%
=
=
1-59
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1-60 Aspen HYSYS Dynamics
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ThValve Stickiness
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 behavior of the valve percent opening as a function 
of the actuator position is shown as follows:
The valve stickiness time constant is specified in the Valve 
Stickiness Time Constant cell. The offset can be specified in the 
Valve Position section. If the valve stickiness time constant is 
left empty, the time constant value is assumed to be zero.
If the Valve has Worn Trim checkbox has been selected, a 
0.1% offset is added to the right hand side of Equation (1.26). 
This offset disallows the valve percent opening to fully close.
K Value Damping Factor
On this same page, the user has the option of specifying the 
damping factor to modulate the iterative convergence of the 
computed valve conductance parameter K in the valve flow 
equation when convergence on this parameter is not achieved. 
This iterative computation is a modified successive substitution 
method, where
where:
q = damping or relaxation factor
F(K)  = value for  computed directly from the valve flow 
equation
Kn = value of   for the current iteration
Kn+1 = value of   at the end of the current iteration
(1.26)
(1.27)
τsticky
d Valve%( )
dt
---------------------------- Valve%+ Act%+ Offset=
Kn 1+ 1 q–( )F Kn( ) qKn+=1-60
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ThThe default value for q is 0.95 and the possible choices are in 
the range 0 to 1.
Fail Modes
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. 
Fail modes can be specified by selecting the corresponding radio 
button in the Positions group. The valve can be modeled to fail 
by selecting the Actuator has failed checkbox.
Fail Open Mode
In the event that the signal from the controller is cut off from 
the valve, the valve becomes wide open. In Aspen HYSYS, if the 
Fail Open radio button in selected, the signal received by the 
valve is modified by the valve as follows:
Notice that if ActDesired% (from controller) becomes zero in the 
event of a signal failure, the actuator becomes fully open. The 
fact that the signal from the controller is modified (by the valve 
operation using Equation (1.28)) has implications on the 
direction of the controller. If the Fail Open mode is selected for 
the valve, reverse acting controllers need to be toggled as 
direct-acting and direct-acting controllers need to be toggled as 
reverse acting.
Fail Shut Mode
In the event that the signal from the controller is cut off from 
the valve, the valve becomes fully closed. Aspen HYSYS does 
not modify the signal from the controller as with the Fail Open 
mode. If the signal from the controller becomes zero, so can the 
ActDesired% value. Since the signal from the controller is not 
modified by the valve, the controller’s direction does not have to 
ActDesired% = 100% - ActDesired% (from controller) (1.28)1-61
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1-62 Aspen HYSYS Dynamics
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Thbe changed.
Fail Hold Mode
In the event that the signal from the controller is cut off from 
the valve, the valve fails in its current position.
Positions Group
Various valve position parameters for the actuator and the valve 
appear in the Positions group:
1.6.4 Inertia
The inertia modeling parameters and the frictional or power loss 
associated with the rotating equipment in the Pump, 
Compressor, and Expander can be specified on the Inertia page 
Valve Position Parameter Definition
Minimum The minimum position the actuator or valve 
can physically achieve. Leaky valves can be 
modeled by specifying a non-zero value for 
the minimum valve position.
Maximum The maximum position the actuator or valve 
can physically achieve.
Current The actual position of the actuator or valve in 
time.
Desired The desired actuator position set by a PID 
Controller operation or imported from a 
Spreadsheet operation.
Offset The Offset defined in Equation (1.26) can be 
specified in this cell.1-62
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Thof the Rating tab for these unit operations.
Not all of the energy supplied to a dynamic pump or compressor 
is transferred to the fluid. Likewise, not all the energy provided 
by an expander translates to kinetic energy. There are frictional 
losses associated with the moving parts of these unit operations. 
Power is also required to accelerate the rotating impeller and 
shaft. In general, the total power or duty supplied to or provided 
from a pump, compressor, or expander has three parts:
• Rate of energy imparted to or provided by the fluid.
• Rate of energy required to accelerate or decelerate the 
rotational speed of the shaft.
• Rate of energy lost due to mechanical friction loss.
The rate of energy supplied to the fluid can be observed in the 
enthalpy change between the inlet and exit streams. For a pump 
or compressor, this is:
where:
F = molar flow rate 
h2 = molar enthalpy of the exit stream
h1 = molar enthalpy of the inlet stream
 Figure 1.21
(1.29)
The Size Inertia button option is only available in the Pump unit 
operation.
Rate of energy imparted to the fluid F MW( ) h2 h1–( )=1-63
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ThMW = molecular weight
The fluid power can also be calculated from the relationship:
where:
cf = correction factor
Fm = mass flow rate
H = head (adiabatic or polytropic depending on selection)
η = efficiency (adiabatic or polytropic depending on selection)
g = acceleration due to gravity
(1.30)
(1.31)
Rate of energy imparted to the fluid 
g cf( ) Fm( ) H( )
η
----------------------------------   (compressor, pump)=
Rate of energy imparted to the fluid cf Fm( ) H( ) η( )  (expander, turbine)=1-64
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ThInertial Modeling Parameters
The rate of energy required to accelerate the speed of a pump or 
compressor is a function of the rotational inertia of the impeller 
and the rotational speed. The rotational inertia, I, is calculated 
as follows: 
where:
M = mass of the impeller and rotating shaft
R = radius of gyration
The mass and radius of gyration can be specified in the Inertia 
page. If you have a total inertia value just back calculate the 
mass given any assumed radius. It is only the inertia which is 
used in the dynamic simulation. 
If you have inertia for different parts of the train rotating at 
different speeds, you need to pro-rate these to a common speed 
basis:
where:
Ib = rotational inertia with respect to some base speed
Iι = rotational inertia of the ith component on the rotating 
shaft (with respect to the speed at this gear ratio)
ωb = base nominal speed
ωi = nominal speed of the ith component on the rotating shaft 
(at this gear ratio)
GRi = gear ratio of the ith component
(1.32)
(1.33)
I MR2=
Ib
1
ωb
2
------ Ii∑ ωi
2( )
IiGRi
2
∑=
=
1-65
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1-66 Aspen HYSYS Dynamics
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ThThe power required to accelerate the equipment, PI, can be 
calculated using:
where:
  = rotation speed (radians/unit time)
The rotational power, which can be viewed on the Performance 
tab is related to torque, τ, via the following relationship:
You can also calculate the inertia by sizing method. The sizing 
method uses the following equation to calculation the inertia:
where:
I = inertia
P = design power
N = design speed
(1.34)
The rotational speed above and for friction loss is in radian 
units not revolutions. When the rotating equipment slows 
down, the stored rotational energy is either given back to 
the fluid or lost to friction.
(1.35)
(1.36)
The sizing method for calculating inertia is only available in 
the Pump unit operation.
PI I ω dωdt
------=
ω
τ
Pi
ω
----
Idωdt
------=
=
I 0.03768 P
N3
-----⎝ ⎠
⎛ ⎞ 0.9556
=
1-66
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ThTo calculate the inertia, specify the design power and design 
speed in the appropriate fields in the Design page and click the 
Size Inertia button on the Inertia page.
Friction Loss
The rate of energy lost from mechanical inefficiencies depends 
on the frictional power loss factor, ffric, which can be specified in 
the Friction Loss group. The frictional work, Pf, can be calculated 
as follows:
This relationship is the basis that the torque due to mechanical 
bearing loss varies directly with the speed, which is generally 
accurate down to about 20% of normal speed1. Below this it has 
been suggested the mechanical bearing loss torque is constant.
Although compressors normally startup and shutdown under 
flowing conditions, pumps are often started and stopped with no 
flow. In the situation where the fluid power is zero, the rate of 
acceleration /deceleration is actually a first order response with 
time constant as follows:
The user can also enter the time constant directly, and the 
friction loss factor is calculated from the inertia. 
The friction loss factor for a compressor should be obtained 
directly from the manufacturer, or it could be tuned to give the 
appropriate power loss at design. You can view this power loss 
on the Performance tab.
(1.37)
(1.38)
The approach to tuning the friction loss factor for rundown 
rate may produce an inaccurate power loss at design 
conditions because there is a short-circuit power loss of the 
fluid at zero load, which is not currently taken into account1. 
Pf ffricω ω=
τc I ffric⁄=1-67
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1-68 Aspen HYSYS Dynamics
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ThEnergy Balance Relationship
The complete energy balance for these unit operations is 
therefore given by Equation (1.39) and Equation (1.40).
For compressors and pumps:
For expanders and turbines:
The external Work, W, is positive normally meaning energy into 
the system for Equation (1.39), but energy out for Equation 
(1.40). The absolute speeds are only required in the rare 
circumstance of negative flow and rotation.
1.6.5 Static Head
For any unit operations with holdup, Aspen HYSYS calculates the 
static head considering the equipment holdup, the geometry, 
and the elevation of any attached nozzles. To use this static 
head calculation, you need to enable the calculations on the 
(1.39)
(1.40)
W
g cf( ) Fm( ) H( )
η
---------------------------------- I ω dωdt
------ ffricω ω+ +=
W g cf( ) Fm( ) H( ) η( ) I ω dωdt
------– ffricω ω–=1-68
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ThGeneral Options group on the Options tab of the Integrator.
The two checkboxes used to manipulate the static head are 
described below: 
The Static Head page, found on the Dynamics tab of the Main 
Tray Section in the column environment, allows you to choose 
the calculation method used to calculate the static head for this 
operation. There are four options given on this page:
• Use global option in integrator (full on or off)
• No contributions for this operation
• Internal levels contribution only (partial)
 Figure 1.22
Checkbox Description
Enable static 
head 
contributions
When you select this checkbox, Aspen HYSYS includes 
the effects of static head in the calculations.
Enable Implicit 
Static Head
Vessels can optionally be solved using implicit static 
head calculations for the pressure contributions 
associated with the levels inside the vessel rather than 
using explicit static head calculations. This option 
provides increased stability in applications where these 
static head contributions play a crucial role.
The options in the integrator are for global settings, while 
options in the column environment are for local settings.1-69
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Th• Levels and feed nozzle elevation differences (full)
In the column operation, the contribution of the internal levels is 
always calculated by Aspen HYSYS. Whether this calculation is 
used or not depends on the situation. 
For tray sections:
• If No contributions for this operation is selected, the 
static contributions will not be included for tray sections 
whether the integrator setting is selected or not.
• If Internal level contributions only or Levels and 
feed nozzle elevation differences is selected, the 
static contributions will be included for tray sections 
whether the integrator setting is selected or not.
• If Use global option in integrator (full on or off) is 
selected and the integrator setting is not selected, the 
static contributions will not be included for tray sections.
• If Use global option in integrator (full on or off) and 
the integrator setting are both selected, the static 
contributions will be included for tray sections.
For streams:
• If No contributions for this operation is selected, 
none of the static contributions will be included whether 
the integrator setting is selected or not.
• If Internal level contributions only is selected, only 
the level contributions will be included whether the 
integrator setting is selected or not.
• If Levels and feed nozzle elevation differences is 
selected, the full contributions will be included whether 
the integrator setting is selected or not.
• If Use global option in integrator (full on or off) and 
the integrator setting are both selected, the full 
contributions will be included.
• If Use global option in integrator (full on or off) is 
selected and the integrator setting is not selected, none 
of the static contributions will be included.1-70
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Th1.6.6  Design
The Design page, on the Rating tab of the Pump unit operation, 
lets you specify the following variables in the Design Flow group:
• Typical operating capacity, this parameter is used to aid 
in starting pumps up, which can have vapour in the line 
(for example, due to a reverse flow).
• Design power, this parameter is used to calculate the 
inertia value.
• Design speed, this parameter is also used to calculate 
the inertia value. 
The existence of vapour can cause difficulty when the pump 
starts up. Hence, if the flow is less than a certain fraction of this 
typical operating capacity, then the density is compensated to 
help start the pump up. The Typical operating capacity value 
enables Aspen HYSYS to decide when it is reasonable to 
compensate the density.
1.7 References
 1 Marks, Lionel S., Marks Mechanical Engineers Handbook, 3rd ed, 
p.1848. McGraw Hill Book Co., 1930.
 Figure 1.231-71
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1-72 References
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Th1-72
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Dynamic Tools 2-1
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Th2  Dynamic Toolsw.cadfamily.com    EMa
e document is for study 2.1  Introduction................................................................................... 3
2.2  Dynamics Assistant........................................................................ 4
2.2.1  General Tab............................................................................. 7
2.2.2  Streams Tab ............................................................................ 8
2.2.3  Pressure Flow Specs Tab ......................................................... 14
2.2.4  Unknown Sizing Tab................................................................ 15
2.2.5  Tray Sections Tab ................................................................... 18
2.2.6  Other Tab.............................................................................. 21
2.2.7  User Items Tab ...................................................................... 24
2.3  Equation Summary Property View................................................ 28
2.3.1  Summary Tab ........................................................................ 28
2.3.2  General Equations Tab ............................................................ 29
2.3.3  Unconverged Tab.................................................................... 30
2.3.4  Extra Variables Tab................................................................. 31
2.3.5  Extra Specifications Tab .......................................................... 32
2.3.6  Specified Equations Tab........................................................... 33
2.3.7  General Variables Tab ............................................................. 33
2.3.8  Specification Variables Tab....................................................... 34
2.3.9  Internal Specification Equations Tab.......................................... 35
2.3.10  Internal Specifications Variables Tab........................................ 36
2.3.11  Simultaneous Equations Tab................................................... 37
2.4  Integrator .................................................................................... 38
2.4.1  General Tab........................................................................... 38
2.4.2  Execution Tab ........................................................................ 39
2.4.3  Options Tab ........................................................................... 42
2.4.4  Heat Loss Tab ........................................................................ 442-1
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2-2 Dynamic Tools 
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The document is for study 2.5  Event Scheduler............................................................................45
2.5.1  Theory ..................................................................................45
2.5.2  Event Scheduler Property View .................................................47
2.5.3  Sequence Property View ..........................................................53
2.5.4  Event Property View ................................................................57
2.5.5  Analyzing a Schedule...............................................................70
2.5.6  Running a Schedule.................................................................71
2.6  Control Manager ...........................................................................72
2.7  Dynamic Initialization...................................................................73
2.7.1  Cold Initialization Tab ..............................................................76
2.7.2  Notes Page.............................................................................852-2
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Dynamic Tools 2-3
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Th2.1 Introduction
Modeling a process in dynamics is a complex endeavour. From 
the perspective of defining the model, you must consider 
parameters such as vessel holdups, valve sizing, and use of 
pressure flow specifications. 
To help simplify this process, Aspen HYSYS has several dynamic 
tools and they are described in the following table:
Tool Description
Dynamics 
Assistant
The Assistant provides a tool for easily converting old 
Aspen HYSYS dynamic cases to pressure flow 
dynamics. It provides general assistance to users who 
are learning how to create dynamic cases. It prepares 
steady state cases for dynamic simulation by ensuring 
that all the correct information is specified, thus 
avoiding over or under specified or singular problems.
View Equations 
Tool
Provides another means of analyzing cases for 
dynamic simulation. This tool provides a summary of 
the equations and variables used by the simulation 
when running in dynamics. By analyzing the case, it is 
possible to determine if there are required or 
redundant pressure flow specifications. In some 
instances, cases which are running in dynamics fails to 
converge, in this case, the View Equations tool can be 
used to help determine what part of the simulation is 
causing problems.
Integrator Allows you to control some of the integration 
parameters which are used by Aspen HYSYS. Simple 
parameters such as the time step or the integration 
stop time or advanced parameters such as the 
execution rates of the different balances can be set 
from this tool. Once a case is running in dynamics, the 
current simulation time and the real time factor can be 
viewed.
Event Scheduler Aspen HYSYS can perform predetermined actions at 
given times in the simulation; warn you by playing a 
sound when the temperature of a stream reaches a 
certain point, stop the integration once a condenser 
level stabilizes, or increase a feed rate after the 
simulation has run for a given time period.2-3
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Th2.2 Dynamics Assistant
The Dynamics Assistant provides a quick method for ensuring 
that a correct set of pressure flow specifications is used. The 
Assistant can be used when initially preparing your case for 
dynamics, or when opening an old Aspen HYSYS 1.x dynamic 
case. 
The Assistant recommends a set of specifications which is 
reasonable and guarantees that the case is not over specified, 
under specified, or singular. It has an option of doing a quick 
examination for potential problems that can occur while moving 
from steady state to dynamics as well as before running the 
case in dynamics. 
In the case of a simple separator, Aspen HYSYS adds pressure 
flow specifications as shown in the figure below.
The Assistant makes recommendations for specifying your 
model in Dynamics mode. You do not have to follow all the 
suggestions. It is recommended that you are aware of the 
effects of each change you make.
 Figure 2.1
Original 
Configuration
Configuration and 
Pressure Flow Specifications 
after using Dynamics Assistant2-4
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ThHowever, in more complicated models such as the one shown in 
the figure below, the Dynamics Assistant recommends the 
insertion of valves in some terminal streams. 
Although the Pressure Flow specifications added by the Assistant 
are adequate for starting a case in dynamics, detailed dynamic 
modeling can require more advanced modifications.
In cases where unit operations such as separators are directly 
connected through multiple streams, the flow cannot always be 
determined. As a temporary fix the Assistant adds a flow 
specification. However, you should add the missing unit 
operations (e.g., pumps, valves, etc.) to define the pressure-
flow relation between the vessel unit operations.
 Figure 2.2
A Flow Specification warning from the Dynamics Assistant 
usually indicates that your model is missing some 
equipment.
Original Configuration Configuration and Pressure 
Flow Specifications after using 
Dynamics Assistant2-5
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ThIn addition to ensuring that the correct Pressure Flow 
specifications are used for your dynamic case, the Assistant 
sizes all necessary equipment that has not yet been sized. 
The parameters sized are: 
• vessel volumes
• valve Cvs
• k values (for equipment such as heaters, coolers, and 
heat exchangers)
The assistant sizes required unit operations based on the flow 
conditions and specified residence times. The assistant also 
checks the Tray Section pressure profile for both steady state 
and dynamic model to ensure a smooth dynamics start. It also 
ensures that the tray section and attached stream have the 
same pressure.
Although the assistant ensures that your case runs in dynamics, 
it is not intended that the changes made are sufficient for your 
case to line out. It is still your responsibility to ensure that an 
adequate control scheme is added to the case and that your 
model is properly rated (i.e., existing vessels are adequately 
sized).
 Figure 2.3
Configuration of two units without 
Pressure Flow relationships when 
using the Dynamics Assistant.
Configuration of two units without 
Pressure Flow relationships when 
preparing for detailed dynamic 
modeling.2-6
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ThThe Dynamics Assistant can be opened using any of the four 
following methods:
• Using the hot key combination of CTRL Y
• Selecting Tools | Dynamics Assistant from the menu 
bar
• Clicking the Dynamics Assistant icon in the toolbar.
• Clicking the Dynamics Assistant button from the 
Equation Summary property view.
Each view of the Dynamics Assistant view has two buttons:
• Analyze Again re-evaluates the simulation case. This is 
particularly useful if you make changes to the case that 
affect the setup of the simulation for dynamics.
• Make Changes makes all the enabled changes within 
the Assistant.
2.2.1 General Tab
The General tab contains a summary of the changes which 
Aspen HYSYS recommends for dynamic simulation. Each item in 
the list has either a green checkmark or a red ‘x’ located to the 
right of the item indicating if the change is made. The 
checkmark indicates that the change is made while the ‘x’ 
indicates that the change is not made. 
The Preferences button opens the Assistant Preferences 
property view, which allows you to change the way Dynamics 
Assistant runs. This view displays three options:
• The Set stream pressure and flow specifications in 
the background option allows the Assistant to activate 
and deactivate stream pressure and flow specifications 
as it sees fit.
• The Perform checks before running dynamics option 
allows the Dynamics Assistant to check for any missing 
specification which can cause potential problems in 
dynamic simulation before you switch to Dynamics mode 
or run the Integrator. 
The Dynamics Assistant button is only available on the 
Equation Summary property view after the Full Analysis 
button is clicked and if there are problems with the case.
Dynamics Assistant icon2-7
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2-8 Dynamics Assistant
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Th• The Save steady state case on switch to dynamics 
option allows Aspen HYSYS to automatically save Steady 
State cases before they are transferred to Dynamics.
2.2.2 Streams Tab
The Streams tab consists of the following pages: 
• Pressure Specs
• Flow Specs
• Uninitialized
• Insert Valves
• Int. Flow Spec
Pressure Specs Page
The Pressure Specs page, lists all the streams that have 
pressure specifications either added or removed. 
The list of streams in the Remove pressure specifications in 
these streams group corresponds to the streams which currently 
have pressure specifications that the Assistant suggests you 
remove. Streams contained in the Set pressure specifications in 
these streams group are those where a pressure specification is 
recommended.
 Figure 2.42-8
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ThAs a rule, pressure specs should be added to boundary streams 
and not internal streams whose pressure should be determined 
by the surrounding equipment. If the Assistant plans to insert a 
valve, the stream in question is shown as not needing a 
pressure specification because the new stream at the other end 
of the valve receives the pressure specification.
If you do not want Aspen HYSYS to change the pressure 
specification for a stream listed in either group, clear the OK 
checkbox for the given stream. This prevents Aspen HYSYS from 
making the pressure specification change for the stream. To 
view a stream in either list, double-click on the stream to open 
its property view.
Flow Specs Page
The Flow Specs page, lists all the streams which have flow 
specifications either added or removed. The list of streams in 
the Remove flow specifications in these streams group 
corresponds to the streams that currently have flow 
specifications which are not recommended, and are removed. 
Streams contained in the Set flow specifications in these 
streams group are those which require a flow specification as a 
If you do not understand the reasoning behind the Dynamics 
Assistant’s recommendations, click the Tell me why button 
for a brief explanation.2-9
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2-10 Dynamics Assistant
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Thtemporary measure.
Flow rates should be determined by pressure differences or 
equipment such as pumps and compressors. For simplicity, the 
Dynamics Assistant can add flow specifications to feed streams 
instead of adding additional equipment.
If you do not want Aspen HYSYS to change the flow specification 
for a stream listed in either group, clear the OK checkbox for 
the given stream. This prevents Aspen HYSYS from making the 
flow specification change for the stream. To view a stream in 
either list, double-click on the stream to open its property view.
Uninitialized Page
The Uninitialized page contains the list of streams which are not 
completely defined. For Aspen HYSYS to initialize any streams 
listed on this page, you must be in Dynamics mode; uninitialized 
 Figure 2.52-10
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Thstreams is not initialized when in Steady State mode. 
Cases saved in Aspen HYSYS 1.2 dynamics lack stream phase 
information. This information can be replaced by flashing the 
streams in question. Switching back to steady state, solving, 
then returning to Dynamics mode also fixes these streams. 
Streams with no values receives initial estimates for 
temperature and pressure of 25°C and 101.33 kPa.
 Figure 2.62-11
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2-12 Dynamics Assistant
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ThInsert Valves Page
The Insert Valves page lists the valves which Aspen HYSYS 
inserts to ensure the pressure flow specifications are not 
singular (i.e., the pressure-flow matrix is unsolvable). Aspen 
HYSYS attempts to attach valves to boundary streams that are 
connected to unit operations without pressure flow (flow 
proportional to pressure difference) capability. 
The outlet stream for the valve is automatically assigned a 
pressure specification. Both the outlet stream and the valve 
added by Aspen HYSYS are named according to the original 
boundary stream; a valve attached to boundary stream 4 is 
named VLV-4 and the outlet stream 4-1. 
 Figure 2.7
A situation can arise where the Dynamics Assistant 
recommends the addition of a valve on a stream where you 
do not want to add one. In this case, you should ensure that 
those streams receive either a flow or pressure specification.2-12
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ThInternal Flow Specs Page
The Internal Flow Specs page lists the internal streams which 
require a flow specification. This is used primarily where 
separators are directly connected to each other by two or more 
streams. The flow specification which is added is sufficient to 
start a case in dynamics, however, it is highly recommended 
that a unit operation with pressure flow relationships (as such 
an operation is probably missing) be placed between such 
separators. 
Examples of unit operations with pressure flow relationships 
include:
• valve
• compressor
• pump
• heater
Once such a unit operation is placed between the separators, 
the flow specification can be removed.
 Figure 2.82-13
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2-14 Dynamics Assistant
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Th2.2.3 Pressure Flow Specs Tab
The Pressure Flow Specs tab consists of two pages: PF versus 
DP and LNG.
PF versus DP Page
The PF versus DP page lists the unit operations which currently 
have a specified pressure drop as the dynamic specification. The 
pressure drop option should not be used because this is 
physically unrealistic. Material flow is driven by pressure 
differences as well as resistances and stops when the pressures 
have been equalized. 
A fixed pressure drop specification does not allow for this 
process. The k values are calculated based on the initial 
specified pressure drop. 
If you do not want Aspen HYSYS to change the pressure drop 
specification to a pressure flow specification, simply clear the 
OK checkbox for the unit operation. Double-clicking on the 
name of the unit operation opens its property view.
 Figure 2.92-14
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ThLNG Page
The LNG page indicates which LNG exchangers are currently 
specified with a pressure drop specification or which LNGs are 
missing k values (depending on the dynamic rating method 
chosen). Pressure drop specifications should be changed to 
either pressure flow equation specifications (k values) or 
suitable pressure drop correlations.
2.2.4 Unknown Sizing Tab
The Unknown Sizing tab consists of the following pages: 
• Valves
• Volumes
• K values
Valves Page
The Valves page lists the valves that are not sized. The current 
conditions for the valve are listed and the calculated valve Cv 
based on the pressure drop and percent opening of the valve, 
both of which are changed directly on the page. 
 Figure 2.102-15
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ThBy default, any valve Cv values which are calculated to be less 
than 50 are defaulted to 50.
It is possible to change any of the sizing data for the valve. The 
Cv value is updated based on any changes that are made. If you 
change Cv, the new value is added to the valve when the Make 
Changes button is clicked.
If you do not want Aspen HYSYS to size a valve, clear the OK 
checkbox for the valve.
 Figure 2.112-16
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ThVolumes Page
The Volumes page lists the unit operations which have unknown 
volumes. 
Units which need volumes include: 
• Separators (regular, 3 phase and tanks)
• Condensers
• Reboilers
• Reactors
• Heat exchangers
• Coolers
• Air coolers
• Heaters
The unknown volumes are calculated based on the volumetric 
feed flow rate and the specified residence time. However, if the 
quick sizing feature is applicable, its volume and corresponding 
residence time is given.
If you do not want Aspen HYSYS to implement the calculated 
volume for a unit operation, clear the OK checkbox for the unit 
operation.
 Figure 2.12
If you do not like the suggested volume, either specify a new 
value or have it calculated by changing the volumetric flow 
rate and the residence time.2-17
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2-18 Dynamics Assistant
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ThK Values Page
The k values page lists unit operations for which the k value is 
unknown. The information required to calculate the k value is 
listed along with the current calculated k value.
Unit which can have ‘k’ values include:
• Heaters
• Coolers
• Heat exchangers
• Air coolers
• Valves
If you do not want Aspen HYSYS to calculate a k value, clear the 
OK checkbox for the unit operation.
2.2.5 Tray Sections Tab
The Tray Sections tab consists of the following pages: 
• SS Pressures
• Dry Hole Losses
• Stream Connection
 Figure 2.13
For more information on 
k values or other 
pressure flow 
parameters, see 
Chapter 1.4.2 - Basic 
Pressure Flow 
Equations.2-18
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ThSS Pressures Page
The SS Pressure page identifies tray sections where the total 
steady state pressure drop seems to be inconsistent with the 
total pressure drop calculated according to the dynamics rating 
model.  
By default, the Assistant provides tray section parameters based 
on its own internal sizing method. However, you can choose to 
use the Create sizing utility for selected sections button which 
brings up the Tray Sizing utility.
The results from the utility can be exported to the tray sections. 
You can also invoke the quick size feature for selected sections 
and then repeat the analysis to check if that resolves problems.
If the dry hole pressure loss is larger than the suggested 
maximum, a diameter is suggested. If you use this value and 
calculate k values based on it, the dry hole pressure losses 
should be more realistic. The Dynamics Assistant makes 
changes to the diameter such that the new pressure flow k 
values for the vapour gives reasonable pressure drops. The 
steady state pressure profile is updated to the suggested value.
 Figure 2.14
The SS Pressure page is active in Steady State mode only.2-19
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2-20 Dynamics Assistant
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ThYou can view the tray section property view by double-clicking 
on the tray section name.
Dry Hole Losses Page
The Dry hole losses page displays the tray sections where the 
dry hole pressure loss is very high. By default the assistant fixes 
this by changing the diameter of the tray section and calculating 
new k values based on it. 
The Sizing utility for selected sections button opens the tray 
sizing utility view which can be used to minimize the pressure 
losses.
 Figure 2.152-20
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ThStream Connections Page
The Stream connections page list the streams whose pressure 
do not match with those of the trays where they connect. For a 
ratings model you should supply any missing valves or pumps 
that are needed in the steady state model. 
2.2.6 Other Tab
The Other tab consists of the following pages: 
• Misc Specs
• Component Splitter
 Figure 2.16
The Stream Connections page is not active when static heads 
are enabled or in Dynamics mode.2-21
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2-22 Dynamics Assistant
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ThMisc Specs Page
The Misc page displays the dynamic specifications which are not 
based on sizing equipment, adding valves, or adding pressure 
and flow specs.
Some examples of the types of changes that are made are listed 
below:
• Pump delta P specs removed
• Pump power spec activated
• Mixer equal pressure option active
• Compressor power spec activated
• Expander power spec activated
 Figure 2.172-22
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ThComponent Splitter Page
The Component Splitter page display all the component splitters 
in your case. 
They do not follow pressure flow principles because they have 
fixed flow split fractions. Therefore, the Assistant can not 
properly cope with them.
One option is to use a pressure specification on each stream 
connected to the component splitter. If the component splitter is 
not connected to other operations, then only one stream needs 
a flow specification. 
The component splitter also offers an “equal pressures” option. 
With this option the specifications needed are the same as those 
you would need for a mixer or a tee, except that you would need 
one fewer specification for each overhead stream because of the 
split fractions.
Where a component splitter is not connected to other operations 
all feed flow rates are set to specifications with the two product 
 Figure 2.18
The material balance can be considered as providing an extra 
specification and for each overhead stream the split 
fractions provide an extra specification as well.2-23
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2-24 Dynamics Assistant
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Thflow rates not specified. Connected component splitters need 
your intervention.
2.2.7 User Items Tab
The User Items tab consists of the following pages: 
• Multiple Connections
• Conflicts
• Unit Operations
• Ignored Opers
• Flow direction
Multiple Connections Page
The Multiple Connections page lists any streams which are used 
in multiple connections. 
Streams which are attached as feeds or products to multiple unit 
operations are not allowed in dynamics since they represent 
physically impossible situations. You must make any changes 
listed here, the Dynamics Assistant does not correct any 
multiple connection errors. It is recommended to delete these 
streams and make proper reconnections.
 Figure 2.192-24
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ThConflicts Page
The Conflicts page lists any streams that have their flows 
directly controlled by controllers. 
It also lists streams that do not contain flow specifications or 
have specifications that the Assistant recommends you remove. 
Although this is allowed in Aspen HYSYS, it is recommended that 
a valve be placed in the stream to control the flow. If it is 
desired to directly control the flow of the stream another 
specification needs to be disabled.
This method can cause singular solution in the Flowsheet. The 
Assistant fixes this problem by moving the controller OP to a 
valve when one exists or is added. 
 Figure 2.20
You can open the property view for a given unit operation by 
double-clicking on its name.2-25
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2-26 Dynamics Assistant
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ThUnit Operations Page
The Unit Operations page lists any unit operations in the case 
that are not supported in dynamics. These unit operations 
should be either deleted, replaced with a suitable unit operation 
which is supported in dynamics, or disconnected from the active 
flowsheet and ignored.
Ignored Opers Page
The Ignored Opers page lists any unit operations which are 
currently ignored. 
 Figure 2.21
 Figure 2.222-26
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ThYou need to fix any problems listed on this page as the 
Dynamics Assistant cannot fix them. If an ignored unit operation 
should not be ignored, clear the Ignored checkbox for the unit 
operation.
If you want a unit operation not to be included in the dynamic 
simulation, disconnect it from the active flowsheet.
Flow Direction Page
The Flow direction page lists the valves where the flow is going 
from low to high pressure or where the pressure drop is zero for 
non zero flow. 
If the case is in Steady State mode, this alerts you to unrealistic 
setups that you should investigate before switching to Dynamics 
mode. Most likely the case is missing important pieces such as 
pumps or valves.
You can open the property view for a given unit operation by 
double-clicking on its name.
 Figure 2.23
The Flow direction page is only active when static heads are 
disabled.2-27
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2-28 Equation Summary Property View
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Th2.3 Equation Summary 
Property View
The Equation Summary property view is accessed by selecting 
Simulation | Equation View Summary from the menu bar.
2.3.1 Summary Tab
The Summary tab contains information that can help you find 
specification problems in your case.
The General Results group contains a summary of the number of 
equations and variables that are in the case. The group on the 
right, contains information regarding the general status of the 
case. If there are problems with the specifications, some basic 
information is provided. 
If you open Equation Summary property view from the menu 
bar, the Summary tab initially contains a single button. Clicking 
the Full Analysis button causes Aspen HYSYS to analyze the 
pressure-flow parameters in order to determine if there are 
enough specifications for the problem. 
If Aspen HYSYS determines a problem, the Dynamics Assistant 
button becomes visible and the Partitioned Analysis button 
 Figure 2.242-28
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Thbecome visible. The Full Analysis and Partitioned Analysis use 
different methods of analyzing pressure-flow parameters. 
In addition, should Aspen HYSYS detect any problems the 
Unconverged tab is replaced with a Extra Vars or Extra Specs 
tab, depending on the nature of the specification problem.
2.3.2 General Equations Tab
The General Eqns tab contains a list of all the equations that are 
used by the integrator. The number of equations corresponds to 
the value given in the Number of Equations cell on the Summary 
tab.
Double-clicking on the Equations cell for any given equation 
opens the property view for that equation. Double-clicking on 
the Owner cell opens the unit operation or stream that the 
equation is attached to.
 Figure 2.252-29
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2-30 Equation Summary Property View
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Th2.3.3 Unconverged Tab
If an error occurs while the integrator is running (PF solver failed 
to converge) you can view the equations that have solved up to 
the point of failure. 
When troubleshooting your simulation, it is best to begin at the 
top of the equation list. These equations contain serious errors 
and therefore give the greatest insight into correcting the 
simulation.
 Figure 2.26
Clicking the Update Sorted List button causes Aspen HYSYS to 
reveal the type of equations, location, and scaled error 
associated with the unconverged nodes in the flowsheet.2-30
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Th2.3.4 Extra Variables Tab
When the Full Analysis or Partitioned Analysis buttons are 
clicked, and Aspen HYSYS determines that not enough 
specifications were known, then the Extra Vars tab is added to 
the Equation Summary property view.
This property view shows possible variables that are missing 
from the case. 
Double-clicking on the Variable cell opens the Variable property 
view. Double-clicking on the Owner cell opens the unit operation 
or stream that requires the missing variable.
 Figure 2.272-31
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2-32 Equation Summary Property View
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Th2.3.5 Extra Specifications Tab
When the Full Analysis or Partitioned Analysis buttons are 
clicked, and Aspen HYSYS determines that too many 
specifications were known, then the Extra Specs tab is added to 
the Equation Summary property view.  
 Figure 2.28
Double-clicking on the Equations cell opens the property 
view for the equation. 
Double-clicking on the Owner cell opens the unit operation 
or stream that contains the extra specification.2-32
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Th2.3.6 Specified Equations Tab
The Spec Eqns tab contains a list of the specified equations in 
the case.
The Type cell displays the type of specified equation such as: 
Pressure Balance Equation or Flow Balance Equation.
2.3.7 General Variables Tab
The General Vars tab contains the list of variables being used by 
the integrator for dynamic simulation. 
 Figure 2.29
 Figure 2.30
This number 
corresponds to 
the number 
displayed in 
the Number of 
Variables cell 
on the 
Summary tab. 2-33
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2-34 Equation Summary Property View
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Th2.3.8 Specification Variables 
Tab
The Specs Vars tab contains the list of specified variables being 
used by the integrator for dynamic simulation. 
 Figure 2.31
This number 
corresponds 
to the 
number 
displayed in 
the User Spec 
Vars cell on 
the Summary 
tab. 2-34
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Th2.3.9 Internal Specification 
Equations Tab
The InterSpecEqns tab contains the list of Internal Specification 
Equations that are used by the integrator for dynamic 
simulation. 
 Figure 2.32
Double-clicking on the Equations cell opens the property 
view for the equation, while double-clicking on the Owner 
cell opens the unit operation or stream associated with the 
equation.2-35
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2-36 Equation Summary Property View
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Th2.3.10 Internal Specifications 
Variables Tab
The InterSpecVars tab contains the list of Internal Specification 
Variables that are used by the integrator for dynamic simulation. 
 Figure 2.33
Double-clicking on the Variables cell opens the variable 
property view, while double-clicking on the Owner cell opens 
the unit operation or stream from which the variable is 
taken.2-36
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Th2.3.11 Simultaneous 
Equations Tab
The Simultaneous Equations (SimulEqns) tab lists the equations 
that are solved simultaneously by the integrator.  
The Type cell displays the type of specified equation such as: 
Pressure Balance Equation or Flow Balance Equation.
 Figure 2.34
Double-clicking on the Equations cell opens the property 
view for the equation, while double-clicking on the Owner 
cell opens the unit operation or stream associated with the 
equation.2-37
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2-38 Integrator
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Th2.4 Integrator
The Integrator is used when running a case in Dynamics mode. 
You can access the Integrator property view by one of the 
following methods:
• Select Integrator command from the Simulation 
menu.
• Press CTRL I.
Aspen HYSYS solves all equations using the fully Implicit Euler 
integration method. On the Integrator property view, various 
integration parameters can be specified.
2.4.1 General Tab
The General tab has three groups which contain the time 
parameters for the integrator. The Integration Control group has 
the controls for whether the integration is Automatic or Manual. 
Manual integration lets you specify the number of time steps 
which Aspen HYSYS executes. It is usually used when you are 
trouble shooting or debugging your case because it allows you 
to move from one time step to the next.
 Figure 2.352-38
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ThOnce the integrator has executed the number of time steps, the 
integrator does not stop, but remains in a holding mode. If 
additional time steps are entered, the integrator continues 
integration for the given number of time steps.
2.4.2 Execution Tab
The Calculation Execution Rates group contains parameters that 
indicates the frequency at which the different balance equations 
are solved.
The default values for Pressure-Flow equations, Control and 
Logic Ops, Energy Calculations, and Composition and Flash are 
1, 2, 2, and 10 respectively. A value of 2 for the Energy 
Calculations means that an energy balance is performed every 2 
time steps. 
 Figure 2.36
Do not change any of these numbers unless you have good 
reason for doing so.2-39
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2-40 Integrator
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ThThe following is a table describing the setting for each of the 
four execution rates.
You can apply the execution rate values setting to all operations 
by selecting the Use these default periods for all operations 
checkbox. Clear the checkbox if you do not want to apply the 
execution rate values entered in this group.
You have the option of specifying the composition and energy 
balance execution rates per integration time step for individual 
dynamic unit operations.
To specify individual execution rates for different unit 
operations, it is necessary to add a new Dynamic Equipment 
Ops tab to the Workbook. 
1. Clear the Use these default periods for all operations 
checkbox in the Calculation Execution Rates group of the 
Integrator property view. 
2. Select the Workbook. Select Workbook from the menu bar, 
and then select Setup. This opens the Setup property view.
Field Description
Pressure Flow 
Solver
Since pressure and flow can change rapidly, their 
calculations are solved at the highest frequency and 
should be left at its default, 1.
Control and 
Logical Ops
The default number should always be sufficient, but 
you can reduce this number for special cases. (E.g., 
When you need rapid control responses or to mimic 
equipment where sample data can only be obtained at 
a low frequency.)
Energy 
Calculations
The energy calculation interpolates between the flash 
calculations. The value should be lower than that of the 
composition and flash calculations.
Composition and 
Flash 
Calculations
The composition and flash calculations number is the 
most important on this page. If you reduce this 
number the flashes will be performed more frequently. 
This can slow down the calculation speed of Aspen 
HYSYS, but it may result in more accurate 
compositions and results in some cases. This number 
can be reduced in cases where the phase change in an 
individual vessel is being studied and a high degree of 
accuracy is required with regard to the phase 
composition.2-40
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Th3. Click the Add button in the Workbook Tabs group. The New 
Object Type property view automatically opens. 
4. Scroll down the list and select the Dynamic Equipment Op 
item. Click the OK button. You are returned to the Setup 
property view. 
5. Close the Setup property view, and go to the Workbook 
property view. The Dynamic Equipment Op variable set is 
added as a Workbook tab.
On the Dynamic Equipment Ops tab, if you select or clear the 
any one of the Use integrator periods checkbox, all other Use 
integrator periods checkboxes in the flowsheet will also be 
selected or cleared. 
 Figure 2.37
 Figure 2.382-41
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2-42 Integrator
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ThYou must clear all the Use integrator periods checkbox, 
before you can specify individual execution rates for the 
different unit operations.
2.4.3 Options Tab
The Options tab contains the advanced parameters that are 
used in dynamics.
The General Options group contains the following parameters:
 Figure 2.39
Parameter Description
Enable static head 
contributions
Aspen HYSYS calculates the static head considering 
the equipment hold up, the geometry, and the 
elevation of any attached nozzles.
This option requires the Aspen HYSYS Dynamics 
feature. 
Enable implicit 
static head 
calculation
If this options is active, it will implicitly solve for 
static head contributions resulting from level changes 
inside the vessels. This can increase the stability in 
cases where the levels in a vessel are closely tied 
with the flow rates and the liquid height in the vessel 
can change rapidly. Use this option only if you are 
experiencing stability problems related to the above 
setup.
The Remove Fidelity 
button enables you to 
reset all variables 
(affected by the Aspen 
HYSYS Dynamics 
features) to their default 
values.
Refer to Section 1.6 - 
HYSYS Dynamics for 
more information.2-42
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ThEnable heat losses 
calculations as 
configured
When checkbox is selected, heat loss model settings 
for unit operations (such as vessel and tray sections) 
are accounted for. If checkbox is cleared, all heat 
losses are zero irrespective of individual heat loss 
settings.
Singularity 
pressure flow 
analysis before 
running
When checkbox is selected, Aspen HYSYS warns you 
of a possible singular solution matrix before starting 
integration.
For larger cases it is recommended that this 
checkbox be cleared to increase the overall start up 
speed. For cases where a singular solution is not 
considered to be a problem, this option can be 
disabled which increases the overall speed.
Rigorous non 
equilibrium mixed 
properties
It is recommended that this option remains active. 
Deactivating this option provides a slight speed 
increase when nozzle efficiencies are not 100%, 
although instabilities can occur. 
Skip flashes under 
acceptable 
conditions
It is recommended that this option remain inactive. 
Activating this option tells Aspen HYSYS to skip 
flashes calculations under acceptable conditions 
(e.g., valves with zero pressure drop or mixers/tees 
with only one effective feed). This provides a slight 
speed increase, although instabilities can occur.
Simultaneously 
solve heat transfer 
eqns with IOFlash
When the checkbox is selected and IOFlash is the 
flash algorithm selected in the Basis environment, 
then Aspen HYSYS tries to solve heat transfer 
equations (from heat exchangers) simultaneously 
with the flash, and potentially make the dynamics 
run faster.
If you encounter inconsistencies with your heat 
transfer equipment, clear this checkbox.
Access Fidelity 
license options
When selected, the Aspen HYSYS Dynamics features 
are activated. Once the features are used this 
checkbox cannot be cleared.
This option gives access to nozzle properties and 
other advanced features.
Model choking of 
liquid inside the 
valve
When selected, this activates the model liquid 
choking option for all the valves.
Use implicit check 
valve model
It is recommended to use this option if you are 
experiencing a delay in the opening and closing of 
check valves.
Parameter Description2-43
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2-44 Integrator
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Th2.4.4 Heat Loss Tab
The Heat loss tab allows you to specify the ambient temperature 
Truncate large 
volume integration 
errors
If there are large upsets or sudden severe changes in 
the system, it can result in a volume balance error, 
where the volume of the material shown does not 
match the physical volume that it occupies. If the 
error is large, you can enable the Truncate Large 
Volume Error option and Aspen HYSYS truncates 
the error and restore material inventory. 
However, truncating the error can cause a bump in 
the model and will also violate the overall material 
balance. For small errors (and if this option is turned 
off), Aspen HYSYS will slowly correct the error 
naturally over time. It is recommended that this 
option be turned off for depressuring utilities.
Reduced recycle 
efficiency for small 
timesteps
For smaller integration step sizes (where the 
composition time step ends up being less than 5 
seconds), you can enable the Small dt Reduced 
Recycle Efficiency option and Aspen HYSYS 
reduces the flash efficiency of material inside 
vessels. This option improves stability of the system, 
but in some cases can produce undesirable results. 
For example, the phases in a vessel may no longer 
be in equilibrium and can be at different 
temperatures. If you are reducing the integration 
step size or lower the composition period, you can 
turn this option off if you experience problems or 
unexpected results.
Close component 
material and 
energy balance
Enables Aspen HYSYS to perform careful calculation 
on component and energy balances to avoid 
imbalances.
Parameter Description2-44
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Ththat is used in the heat loss equations.
2.5 Event Scheduler
Using the Event Scheduler, it is possible to have Aspen HYSYS 
perform given tasks at predetermined times once a simulation is 
running in dynamics.
The tasks can be triggered by a pre-determined simulation time 
or elapsed time, a logical expression becoming true, or a 
variable stabilizing to within a given tolerance for a set amount 
of time.
2.5.1 Theory
The Event Scheduler property view as shown in Figure 2.42 
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 (a null 
 Figure 2.402-45
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2-46 Event Scheduler
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Thcondition is always true). 
The event may have zero or more enabled actions. Normal 
sequential stepping of the events occurs, although forward and 
backward branching may also be configured.
It is recommended that each schedule be treated and 
implemented as a separate autonomous set of sequences. 
Interaction between sequences in different schedules is not 
possible. Within each schedule, the user must configure one or 
more sequences. Each active sequence in the schedule is 
executed once per time step, in the order they appear. For an 
active sequence (that has been started and is not holding), the 
sequence status would be Waiting and there would be one 
current (or active) event. The current event means that this 
particular event is having its condition evaluated to determine 
any actions to take and which event in the sequence takes 
control next (becoming the next current event). You can think of 
the event's condition as being a logical True or False evaluation 
and the actions are manipulations on the simulation case. When 
the current event’s condition is met, its action(s) (if any) are 
 Figure 2.41
Schedule Manager
Schedule 1 Schedule 2 Schedule 3 Schedule N
Sequence A
Condition
Action List
Action 1
Action M
Event 1
Event X
Sequence Z2-46
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Thexecuted and by default control passes to the next event in the 
sequence. Normally (without Multi-Events turned on) only one 
event is evaluated and allowed to execute each time step. If 
branching has been configured for the current event, control 
may not be passed to the next step in the sequence. Instead 
control will be passed to the specified Jump To Event if the 
branching requirement is met. The branching requirement is 
nothing more than a condition that determines if or when the 
branching takes place. Both forward and backwards branching is 
permitted. If one sequence Starts, Stops, Pauses, or Resumes 
another sequence, this takes place on the current time step if 
the effected sequence is lower in the Schedules list; otherwise 
the effect is relevant on the next time step.
It is usually recommended to keep as many as possible of a 
series of steps within a single sequence. Rather than starting 
and stopping related sequences within a schedule, the use of 
branching can be easier to follow. Figure 2.46 shows the 
Sequence property view's Jump When column which allows you 
to see any branching within the sequence.
The Event Scheduler is an ideal way of implementing a flow 
chart representation of sequence logic control systems, which 
are typically implemented in Programmable Logic Controllers in 
the real plant.
2.5.2 Event Scheduler 
Property View
You can access the Event Scheduler by selecting Event 2-47
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2-48 Event Scheduler
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ThScheduler from the Simulation menu.  
You can also access the Event Scheduler by using the hot key 
CTRL E.
 Figure 2.42
The tree browser contains a list of the schedules in a 
case. Each schedule is comprised of sequences, 
which in turn are made up of events. 
Sequences group for a particular Schedule.
Sequence Control group
The Event Scheduler property view is now a combination of 
property views as shown in the figure below. 
The right pane of the property view now shows the Schedule, 
Sequence, Event, or specific Action information as you 
navigate through and click on the tree browser on the left 
pane. 
You can also view this information in separate property 
views as described in Section 2.5.3 - Sequence Property 
View and Section 2.5.4 - Event Property View.2-48
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ThThe buttons in the Schedule Options group are used to manage 
the schedules for the current case. 
The Legend group displays the status icons, which are shown as 
the tree icons in the tree browser, and indicate the current state 
of the schedule, sequence, event, or action. 
Button Description
Add Allows you to add new schedules to the case.
Delete Allows you to delete the selected schedule. This is only active 
when a schedule exists in the case.
Copy Allows you to make a copy of 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 re-order the schedules. This is only active when at 
least two schedules exist in the case.
 Figure 2.43
The following states are valid:
• Schedule: Fully Specified and 
Incomplete
• Sequence: Complete, Holding, Inactive, 
Incomplete and Waiting.
• Event: Complete, Fully Specified, 
Holding, Inactive, Incomplete, Running, 
Time Elapsed and Waiting.
• Action: Fully Specified and Incomplete.2-49
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ThThe Event Scheduler property view also consists of the following 
checkboxes:
When you add a new schedule, Aspen HYSYS opens a Schedule 
Sequences group. You can manage the sequences in the 
schedule on the Event Scheduler property view.
Checkbox Description
Smart Tree When this checkbox is selected and the Integrator is 
running, the tree browser expands and shows the current 
event of the last selected sequence or schedule.
Trace 
Messages
Allows you to enable or disable the tracing of Event 
Scheduler messages.The Aspen HYSYS Desktop Trace 
Window displays useful status and execution messages 
while the sequence is executing.
Trace Dump and Trace Message actions will always 
appear in the Trace Window.
Multi Events When the checkbox is selected, this option allows you to 
execute multiple contiguous events, with conditions of 
True, in a single time step.
The events will not continue after any Jump To Event 
occurs since a backwards jump could result in an infinite 
loop.
 Figure 2.442-50
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ThThe Schedule Sequences group consists of the following 
columns:
The buttons in the Schedule Sequences group are used to 
manage the events in a sequence. 
Field Description
Sequence Allows you to modify the name of the sequence without 
opening the Sequence property view. 
Double-clicking on any cell opens a relevant separate 
dedicated property view.
Run Mode Displays the sequence mode.
Status Displays the status of the sequence. 
Event Displays the number of the event currently being 
executed. 
Waiting For Displays the condition the sequence is waiting for before 
executing. 
Pending 
Actions
Displays the name of the action list associated with the 
event being executed. 
Button Description
View Allows you to view the selected sequence. The Sequence 
property view appears. This is only active when a sequence 
exists in the schedule.
Add Allows you to add a 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 of the selected sequence. This 
is only active when a sequence exists in the schedule.
Import Allows you to import a saved sequence from the disk. 
Schedules have the extension *.seq.
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 re-order the sequences. This is only active 
when at least two sequences exist in the schedule.
Refer to the section on 
the Settings Tab for 
more information.2-51
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2-52 Event Scheduler
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ThThe Sequence Control group is only available when you have 
added a sequence to the schedule.
The Schedule Name field allows you to modify the name of the 
schedule. 
You can view the status of the selected schedule by clicking on 
the Status Panel button although the individual schedule 
property view is not used by most users. The individual schedule 
property view also allows you to control the schedule by using 
the Start, Stop, Resume, and Hold buttons.
Button Description
Start Allows you to start the selected sequence(s).
Stop Allows you to stop the selected sequence(s).
Resume Allows you to resume paused sequence(s).
Hold Allows you to pause the selected sequence(s).
Force Allows you to execute the selected sequences current 
event. The actions will be executed.
Skip Allows you to skip the selected sequences current event. No 
actions will be executed.
 Figure 2.452-52
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Th2.5.3 Sequence Property View
When you select a particular sequence from the tree browser of 
the Event Scheduler property view, the Sequence property view 
appears.  
There are two tabs on the Sequence property view:
• Schedule of Events
• Settings
 Figure 2.46
You can also click the View button from the Schedule 
Sequences group of the Event Scheduler property view to 
view the Sequence property view.
The Schedule of Events tab is the best place to review your 
sequence branching by perusing the Jump When and Jump 
To columns.
The status bar indicates the current status of 
the sequence, which is also shown as the icon 
on the Event Scheduler tree browser. 
You can change the 
sequence name.2-53
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ThSchedule of Events Tab
The Schedule of Events tab shows the list of events for the 
schedule. The following columns are available on the Schedule 
of Events tab. 
Column Description
# Displays the number of the event on this row.
Event Displays the name of the event.
Double-clicking on any cell opens a relevant separate 
dedicated property view.
Specified Indicates whether the event is fully defined.
Condition Displays the condition name of the event.
Action List Displays the name of the events action list. 
Jump When Displays whether the event jumps over any events and if so 
under what condition. The following options are available 
from the drop-down list:
• Never. The default value is Never. It is configured on 
the Branching & Time Out behavior tab of the Event 
property view. The event cannot jump to another 
event.
• Always. The event always jumps to another event 
once it is executed.
• True. The event jumps to another event once the 
logic on the Condition tab is true.
• Timeout. The event jumps to another event if the 
timeout condition is met.
• False. The event jumps to another event once the 
logic on the Condition tab is False.
Jump To Displays the event to jump to. In this way a single event or 
a group of events can be skipped under certain 
circumstances.2-54
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ThThe following buttons are available on the Schedule of Events 
tab:
Settings Tab
On the Settings tab, you can specify the sequence’s universal 
settings, status window reporting, and the event conditions 
default timeout behavior. 
Button Description
View Allows you to view the selected event. This is only active 
when an event exists in the sequence.
Add Allows you to add a new event to the sequence. 
Delete Allows you to delete the selected event. This is only active 
when an event exists in the sequence.
Copy Allows you to make a copy of the selected event. This is 
only active when an event exists in the sequence.
Sort Allows you to re-order the events. This is only active when 
at least two events exist in the sequence.
Analyze This is only active when a sequence is incomplete. 
 Figure 2.47
Refer to Section 2.5.5 - 
Analyzing a Schedule 
for more information.2-55
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ThThe Universal Settings group consists of the following options:
You can also specify the default time out behavior, which applies 
to all events unless a specific event overrides this behavior. 
• Hold. Pauses the execution of the sequence when an 
event in the sequence times out.
• Play Audio File. Plays a specified audio file when an 
event in the sequence times out. You can also test the 
sound by clicking on the Test Sound button after you 
have selected the audio file. 
• Stop Integrator. The Integrator is stopped when an 
event in the sequence times out.
Options Description
Run Mode You can select the sequence mode from the drop-
down list. There are two options:
• One Shot. Executes all its Events in order 
then changes its status to Complete.
• Continuous. Returns to the first Event after 
the last Event has executed in a continuous 
loop.
You can also set the Run Mode from the Event 
Scheduler property view as shown in Figure 2.42.
Unit Set for Logical 
Expressions
Allows you to select a unit set from the drop-down 
list. These are the available unit sets that are on 
the Variables tab of the Session Preferences 
property view. 
Synchronize All Time 
Sensitive Conditions
When you select the checkbox, this ensures 
execution of a particular event at an exact 
simulation time. The Logic Condition Wait for Time 
or Elapsed Time may internally adjust time step to 
integrate precisely to a time sensitive event 
condition.
True Event does NOT 
step if 
JumpWhen=Timeout
It is NOT recommended to select this checkbox. It 
is used for backward compatibility of models built 
in older Aspen HYSYS versions. In older Aspen 
HYSYS versions, if the user specified Branching 
Jump When Timeout, a potential problem 
occurred. When the events condition evaluated 
True and remained True, the actions repeated 
multiple times and the event would not naturally 
step to the next event in the sequence. Instead 
the event would wait for the timeout timer to 
elapse and then jump to the requested event. 
Aspen HYSYS cases built in older versions will load 
into the new Aspen HYSYS version with this 
checkbox on.
Refer to Section 12.3.1 - 
Units Page in the Aspen 
HYSYS User Guide for 
more information.
Refer to Section 2.5.4 - 
Event Property View on 
the Branching & Time 
Out Behaviour Tab for 
more information.2-56
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Th2.5.4 Event Property View
When you select a particular event from the tree browser of the 
Event Scheduler property view, the Event property view 
appears. 
You do not have to select any of the Default TimeOut 
checkboxes, in which case the default behavior is to stop the 
sequence upon Timeout. 
If an events Branching Jump When Timeout is set, this will 
complement any of the above behavior.
 Figure 2.48
You can also click the View button from the Schedule of 
Events tab of the Sequence property view to view the Event 
property view.
You can change the event name and 
condition name using these fields.
This status is also shown on the 
tree browser of the Event 
Scheduler property view.2-57
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2-58 Event Scheduler
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ThThe Event Options group consists of two buttons.
There are three tabs on the Event property view:
• Condition
• Action List
• Branching & Time Out behavior
Condition Tab
The Condition tab as shown in Figure 2.48 shows the four 
possible conditions located within the Wait For group. The 
condition the user selects determines:
• When the actions associated with this event are 
executed.
• When the sequence proceeds to the next event.
A True or 1 condition is required for the above.
The Logic to Evaluate True
The Wait For Logic to Evaluate True requires variables whose 
Tags become part of the Logical ‘If’ Expression. 
 Figure 2.49
The Force and Skip buttons affect the current event which 
may not be the event you are viewing.
 Figure 2.50
Allows you to bypass 
the current event of 
the selected sequence.
Allows you to execute 
the current event of 
the selected 
sequence.
Shows the timer which counts up to 
the user specified True For time.
The True For field is optional. The 
logical expression must continuously 
evaluate True for the amount of time 
you have entered in this field.2-58
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ThVariables are selected by clicking the Add Variable button, which 
opens the Variable Navigator. If no expression is provided, the 
condition always and immediately evaluates True. 
Help is available by clicking the Expression Help button. This is 
the same logic that is used by the Spreadsheet. 
The Syntax State reports whether or not the expression is valid 
syntax and returns Incomplete, Bad Syntax, or Complete (not 
the evaluation of the expression). 
An Elapsed Amount of Time
The actions execute after the user specified amount of time has 
elapsed.
 Figure 2.51
 Figure 2.52
 Figure 2.53
Allows you to specify the amount of 
time you want to wait before the 
actions are executed.
Shows the timer which counts up to 
the user specified elapsed time.2-59
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2-60 Event Scheduler
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ThA Specific Simulation Time
Actions execute only at the specific simulation time. This option 
is not commonly used since it is based upon the sequence 
always starting at the same absolute simulation time. Wait For 
An Elapsed Amount of Time is the preferred option. 
A Variable to Stabilize
The actions execute after a variable has stabilized.
For the variable to stabilize the following information is required:
• basis for calculation
• bandwidth about the basis (percent or absolute)
• period that the variable must be within the bandwidth
The basis for calculation is a selected simulation variable from 
your case. The percentage bandwidth is specified as a percent 
delta from the average value over the stabilization period.
 Figure 2.54
 Figure 2.55
The bandwidth as entered is plus or minus so for a ±2kPa 
bandwidth, for example, the variable must remain within a 
range of 4kPa.
Allows you to specify the exact time 
you want the actions to execute at.2-60
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ThThe condition evaluates True when the selected variable value is 
within the given bandwidth for the specified length of time.
You can view what the variable movement has been over the 
most recent stabilization period. The Variable Stability field 
shows the variable movement in either engineering units or 
percent depending on the type of tolerance you selected.
Action List Tab
The user may define zero or more actions which execute once 
the Condition is met. The user might specify no actions, if the 
user wants to simply perform a branching operation. 
The name and type of the action selected in the List Of Actions 
For This Event group can be changed in the Individual Action 
Specification group. 
The Name field allows you to change the action name while the 
Type drop-down list provides the available action types. After an 
Variable stability is not available until the event has been the 
current event of the sequence for at least the stabilization 
period.
 Figure 2.562-61
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2-62 Event Scheduler
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Thaction type is selected, the Configuration group changes to show 
the information required for the action type.
Specify Variable
The Specify Variable action requires an Object and a Value. 
To select an Object, click the Select Target button which opens 
the Variable Navigator. The current value and units of the 
selected object are shown. If the Increment Only checkbox is 
selected the Value cell is added to the Current Value cell at the 
time of execution otherwise the Value cell replaces the Current 
Value cell.
Start/Stop/Hold/Resume Sequence
The Start, Stop, Hold, and Resume Sequence actions all require 
that a Sequence is selected from the drop-down list.  
 Figure 2.57
 Figure 2.582-62
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ThThe View button displays the selected sequence. It is not 
possible for a Sequence to Start or Resume itself, but it can Stop 
itself and put itself in Hold Mode.
Play Sound
The Play Sound action requires a *.wav audio file, which can be 
selected by clicking the Select Audio File button. When a file is 
selected, the Test sound button is enabled and the audio file can 
be played.
Trace Dump
The Trace Dump action requires the selection of a source 
variable by clicking the Select Source button.
The available Sequences are only those that are part of the 
parent Schedule. You cannot control sequences of other 
schedules.
 Figure 2.59
 Figure 2.602-63
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2-64 Event Scheduler
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ThThe default format for the variable is shown, and can be 
changed by clicking the button at the end of the Format cell. 
When an object is selected, the Trace Dump button is 
activated, allowing you to dump the current variable information 
to the Trace Window.
Play Script Action
The Play Script action requires a Aspen HYSYS script file that 
can be selected by clicking the Select File button. When a file is 
selected, the Play Script button is enabled and the script can be 
played.
Save Snapshot
The Save Snapshot action requires a snapshot file that can be 
selected by clicking the Select Snapshot File button. 
 Figure 2.61
 Figure 2.62
For more information, 
refer to Taking a 
Snapshot in the Event 
Scheduler section from 
Section 11.12.3 - 
External Snapshots in 
the Aspen HYSYS User 
Guide.2-64
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ThSend DDE Command
The Send DDE Command action requires inputs of Service, 
Topic, and Command.
For example, when execution occurs the Excel Sheet1, MyMacro 
runs using the Dynamic Data Exchange protocol. When the 
three inputs are specified, the Execute button becomes enabled 
and the action can be tested. 
In addition, the Transaction Time Out checkbox can be 
selected which then requires an input time. The command 
execution by the external service (i.e., Excel) returns control to 
Aspen HYSYS after the input time expires otherwise it waits until 
the external command completes its execution. When the 
Transaction Time Out checkbox is selected, the Time Out 
Notification checkbox becomes enabled and selecting this 
checkbox notifies you if a time out occurs.
Stop Integrator
The Stop Integrator action stops the integrator.
 Figure 2.63
 Figure 2.64
Service Excel
Topic Sheet1
Command MyMacro2-65
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2-66 Event Scheduler
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ThRamp Controller
The Ramp Controller action requires a Controller, a Target Set 
Point, and a Ramp Duration as inputs. 
The controller is selected by clicking the Select Target button 
that browses all the controllers in the case. After selecting a 
controller, the current set point is shown for reference and the 
View button is enabled. 
Click the View button to open the Controller property view. 
Optionally, the Increment Only checkbox can be selected 
which increments the set point rather than setting it. Upon 
execution the controller is switched to Ramp mode, the target 
set point and duration are specified and the Ramp Controller 
message is sent.
Set Controller Mode
The Set Controller Mode action requires a Controller and a New 
Mode as inputs.
 Figure 2.65
 Figure 2.662-66
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ThThe controller is selected by clicking the Select Target button 
that browses all the controllers in the case. After selecting a 
controller the current mode is shown and the View button is 
enabled.
Trace Message
The Trace Message action traces a message to the Aspen HYSYS 
Trace Window. The user specifies the text of the message in the 
Message field.
 Figure 2.672-67
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2-68 Event Scheduler
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ThBranching & Time Out behavior Tab
The information contained on this tab is not necessary for an 
event to be completely defined. Branching is optional and 
permits sequence flow control other than the default sequential 
forward stepping.
The Jump When drop-down list contains five conditions, which 
are described in the table below:
 Figure 2.68
Condition Description
Never This is the default. The Event cannot jump to another 
Event.
Always The Event always jumps to another Event regardless of the 
condition evaluation. If however, the condition is True, the 
actions will execute first.
True The Event jumps to another Event after the Condition is 
met. The actions execute first.
Timeout The Event jumps to another Event after the timeout period.
False The Event jumps to another event once the logic on the 
Condition tab is False.2-68
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ThIf a Jump When condition other than Never is selected, an Event 
must be selected in the Jump To field. The Jump To field 
contains a drop-down list with all the Events in the Sequence. 
Events can only jump to other Events in the same Sequence. 
Backwards jumping is permitted.
You can also specify a Time Out for the event so that the 
sequence does not get stuck and remain on the event 
indefinitely. If you want to configure Time Out behavior, you 
must select the Event Logic Condition Time Out After 
checkbox and enter the time out period. The Count Up field 
shows the alarm counter such that when it reaches the Time Out 
time, the event times out.
Selecting the Event Specific behavior checkbox enables the 
behavior group options. The three checkboxes in the group can 
be selected in any combination from all three to none. These 
selections override the default behavior defined in the parent 
sequence.
If you have the Jump When not set to Timeout and you have 
not selected any of the three checkboxes in the behavior 
group, the sequence becomes inactive (stops). You can only 
re-start the sequence.
If you have the Jump When not set to Timeout and you have 
any of the three checkboxes selected in the behavior group, 
you have to resume the sequence or start the Integrator if 
necessary.   You must manually either Force or Skip the 
Timeout Event.
If you have the Jump When set to Timeout and you have 
selected zero to all checkboxes in the behavior group, the 
sequence remains active. Then, the sequence proceeds with 
the Jump To Event. You may have to resume the sequence or 
start the Integrator.2-69
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2-70 Event Scheduler
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Th2.5.5 Analyzing a Schedule
Aspen HYSYS has a built in tool for analyzing Schedules which 
are not yet fully defined. From the Sequence property view, it is 
possible to analyze an incomplete sequence by clicking the 
Analyze button. When the Analyze button is clicked, Aspen 
HYSYS opens the Analysis of Sequence property view, which is 
Modal. 
This property view displays a matrix with a list of all the Events 
that are not fully specified. Selecting an Event and double-
clicking in the Event column or clicking the Analyze Event button 
opens the Analysis of Event property view.
The Analysis of Event property view provides feedback about the 
 Figure 2.69
The Analyze button is only active when the Sequence is 
incomplete.
 Figure 2.702-70
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ThEvent. The contents of the Required Specifications group 
changes depending on the Event Wait For condition. Fields that 
are  can be specified at this point or the View Event 
button can be clicked, opening the Event property view where 
the specifications can be made. Actions that are not fully 
specified are displayed in the Incomplete Actions matrix. 
Double-clicking the Event cell in the matrix opens the Event on 
the Action List tab with the Action selected.
When you click the View Event button from the Analysis of 
Event property view, a Modal Event property view appears. You 
can click the Pin button to make the property view non-Modal, 
the active property view returns to the Analysis of Event 
property view which is also Modal.
2.5.6 Running a Schedule
After a Schedule is fully defined, each Sequence is in an Inactive 
state. Sequences can be run from the Event Scheduler property 
view as shown in Figure 2.42.
To activate a Sequence, select it from the tree browser and click 
the Start button. Alternatively, you can select the desired 
Schedule from the tree browser and then select the sequence 
from the list in the Schedule Sequences group. The status 
changes from Inactive to Waiting. The integrator has to be 
running for evaluations to occur.
At every time step, any Sequences that are in the Waiting state 
have their current Event’s Wait For Condition evaluated. When a 
current event evaluates True, the associated Action List items 
are executed, and provided no Branching behavior is specified, 
the next Event in the list becomes the current event and 
evaluations continue.
When you edit an event while the schedule is running, the 
sequence will either continue running or be put in the Hold 
state if further information is required before everything is 
sufficiently satisfied. 
Click the Resume button to continue the sequence after your 
edits. 2-71
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2-72 Control Manager
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ThIf the Sequence RunMode is One Shot and the last event in the 
list executes, the status changes from Waiting to Complete and 
the Sequence is reset.
2.6 Control Manager
The Control Manager is a summary of the PID Controllers and 
MPC Controllers contained within the current simulation. There 
are three tabs available within this property view: PIDs, MPCs, 
and User Variables.
PIDs Tab
This tab provides a summary of the PID Controllers within the 
current simulation. There are three modes displayed:
• Controller mode. Allows you to set the controller to 
automatic, manual, or off.
• Aspen HYSYS mode. Allows you to toggle the Aspen 
HYSYS mode between Internal and External.
• Sp mode. Displays whether the SP is set to local or 
remote.
 Figure 2.712-72
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ThThere are also three key variables displayed: set point (Sp), 
process variable (Pv), and operating target object (Op). 
MPC Tab
This tab provides a summary of the MPC Controllers within the 
current simulation. There are two modes that appear:
• Controller mode. Allows you to set the controller to 
automatic, manual, or off.
• Aspen HYSYS mode. Allows you toggle the Aspen 
HYSYS mode between Internal and External.
There are also three key variables displayed: set point (Sp), 
process variable (Pv), and operating target object (Op).
User Variable Tab
The User Variables tab allows you to create and implement 
variables in the Aspen HYSYS simulation case. 
2.7 Dynamic Initialization
Dynamic Initialization allows you to initialize user-specified unit 
operations and/or streams, all at once or in different areas, to a 
consistent operating state in dynamics modeling. Typically you 
can initialize the selected objects to cold, empty, depressurized, 
Refer to Chapter 5 - 
User Variables in the 
Aspen HYSYS 
Customization Guide 
for more information.2-73
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2-74 Dynamic Initialization
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Thand off-line.  
To access Dynamic Initialization, select Dynamic Initialization 
from the Flowsheet menu. The Dynamic Initialization Manager 
appears.  
The Dynamic Initialization Manager organizes multiple dynamic 
initialization areas. A dynamic Initialization area is a grouping of 
flowsheet objects with specific initialization stream(s), and cold 
Correct set up of the cold initialization configuration is not 
only necessary for obtaining accurate results, but it can also 
prevent movements in the model, and pressure-flow non-
convergence error.
It is recommended that you run the Integrator a few time 
steps on any unit operations and streams before dynamic 
initialization can have an effect on those operations. After 
performing dynamic initialization, if the simulation 
progresses through numerous steps without any errors or 
significant movement in the process, the initialization is 
successful.
If significant movement within the model is detected after 
initialization, set the Integrator in Manual mode and 
initialize the case again to locate the source for movement. 
In most cases, controls and logic operations, boundary 
streams, isolation valves, and pressure/flow specifications 
are the common sources for movements and errors.
 Figure 2.72
You can also access the Dynamic Initialization Manager by 
right-clicking on a selected object or a blank area on the 
PFD, in the object inspect menu select the Dynamic 
Initialization menu, and select View Dynamic Init Manager 
command.2-74
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Dynamic Tools 2-75
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Thinitialization conditions.
There are four buttons in the Dynamic Initialization Manager:
• Add. Adds a new dynamic initialization area to the 
simulation case and opens a blank Dynamic Initialization 
property view.
• View. Opens the property view of the dynamic 
initialization area selected from the drop-down list.
• Delete. Delete an existing dynamic initialization area 
selected from the drop-down list.
• Cancel. Exit the Dynamic Initialization Manager property 
view.2-75
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2-76 Dynamic Initialization
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ThThe Dynamic Initialization property view contains two tabs:
• Cold Initialization
• Notes
2.7.1 Cold Initialization Tab
The Cold Initialization tab allows you to enforce certain cold 
initialization conditions for the selected streams and unit 
operations. All objects of this Dynamic Initialization Area using 
the same Fluid Package must be initialized to the same 
conditions as the Initialization Stream. You can apply similar 
initialization parameters to successive dynamic simulations to 
eliminate the need to either return to the steady state 
environment or manually specify each individual simulation 
object. Dynamic Initialization saves time during multiple 
dynamic simulation runs.
There are three pages on the Cold Initialization tab: Objects, 
Fluid Packages, and Configuration pages. These pages contain 
the initialization parameters and are described in the following 
sections.
 Figure 2.73
Refer to Fluid Packages 
Page for more 
information on fluid 
package and initialization 
stream.2-76
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ThObjects Page
The Objects page displays a list of objects that have been 
selected for this dynamic initialization area. In the Selected 
Objects table, the Name, and Type for each selected object is 
shown.
The objects listed in the Selected Objects table are to be cold 
initialized using the data specified on the Fluid Packages and 
Configuration pages. 
• You can add more objects to the list by clicking the Add 
Object button. 
• You can remove objects from the Selected Objects 
table, by clicking the checkbox in the Remove column, 
or by selecting the row containing the object and clicking 
the Remove Object button. 
• You can remove all of the objects from the Selected 
Objects table by click the Remove All button.
 Figure 2.74
You can rename 
the dynamic 
initialization area 
in the Name field.2-77
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2-78 Dynamic Initialization
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ThYou can also add object(s) to the Selected Objects table 
directly from the PFD by performing the following:
1. Select the object(s) in the PFD.
To select multiple objects on the PFD, hold down SHIFT and 
left-click the mouse button (or click and drag the cursor) 
over the objects you want to add to the list.
2. Right-click on one of the selected objects or right-click on a 
blank area of the PFD. 
The Object Inspect menu appears.
3. Select Dynamic Initialization from the Object Inspect 
menu. Three sub-menus appear:
• Add Objects to Default. Adds the selected object(s) 
directly to the Selected Objects table on the Objects 
Page.
• Remove Objects from Default. Removes the selected 
object(s) directly from the Selected Objects table on the 
Objects Page.
• View Dynamic Init Manager. Opens the Dynamic 
Initialization Manager.
 Figure 2.75
The Add Objects to Default and Remove Objects from Default 
functions operate on the last Dynamic Initialization Area 
which was selected in the Dynamic Initialization Manager.2-78
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ThFluid Packages Page
The Fluid Packages page allows you to specify the parameters of 
an initialization stream for a specific fluid package. The holdups 
as well as the streams for the selected objects on the Objects 
page are initialized to be identical to the values of the 
initialization stream.
 Figure 2.76
Object Description
Fluid Package 
drop-down list
Displays a list of fluid packages for the simulation case.
Each fluid package used by the selected objects must 
have an initialization stream.
Stream drop-
down list
Displays all the streams that use the selected fluid 
package. The selected stream is the initialization 
stream for the selected fluid package.
The initialization fluid must be vapour thus the stream 
must have a vapour fraction of 1 (gas phase only).
Stream Data Set 
radio buttons
Allows you to display the condition or composition for 
the selected stream in the Stream Properties table.
View Stream 
button
Opens the property view for the selected stream.
Stream 
Properties table 
• Condition. Displays the temperature, pressure, 
and the vapour fraction for the selected stream.
• Composition. Displays the mole fractions of each 
component in the selected stream.
Initialize Objects 
button
Allows you to cold initialize all the selected objects on 
the Objects page to the parameters specified for the 
selected initialization stream.2-79
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2-80 Dynamic Initialization
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ThConfiguration Page
The Configuration page allows you to activate the initialization 
settings for specific equipment, logical operations, and stream 
specifications for this dynamic initialization area. 
You can also use a macro to programmatically manipulate the 
initialization settings for any objects, particularly control logics 
in spreadsheet operations. You can specify the name of the 
macro in the User Variable Macro group. Once you click the 
Initialize Objects button, the macro is executed.
To prevent heat transfer in a cold plant, the temperature of 
the initialization stream must be set to the Default Ambient 
Temperature, which is specified on the Heat Loss tab in the 
Integrator property view. 
Similarly, any heat transfer unit operations (e.g., heat 
exchangers) that have two or more fluid packages used by 
their associated streams should be initialized to the same 
temperature as the Default Ambient Temperature.
 Figure 2.77
Aspen HYSYS provides default settings (in the Configuration 
page) for a standard cold initialization process. 
Refer to Chapter 2.4.4 - 
Heat Loss Tab for more 
information.
Refer to Boolean Logical 
Operations, 
Spreadsheets, and 
Event Schedules 
sections for more 
information.2-80
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ThThe following operations and specifications are listed in the Unit 
Operation Setting group:
Any unchecked operations or specifications listed in the Unit 
Operation Setting group will have no effect when you 
initialize objects.
Checkbox Description
Controller 
Initialization
You can select this checkbox to initialize all controllers. 
By default, the controller mode is set to Manual, and 
the output is set to 0%. 
You can specify a desired OP value other than the 
default (0%). For more information, refer to the 
following sections in the Aspen HYSYS Operations 
Guide:
• Section 5.4.2 - Split Range Controller 
(Initialization Page)
• Section 5.4.3 - Ratio Controller 
(Initialization Page)
• Section 5.4.4 - PID Controller 
(Initialization Page)
To restore the default Controller Initialization setting, 
delete the Cold Init OP value in the Initialization page 
(i.e., the Cold Init OP field is  on the 
Initialization page).
Reset Digital 
Point
There are three options available from the drop-down 
list:
• No. Ignores the initialization option.
• OFF. Sets the output state to OFF.
• ON. Sets the output state to ON.
If the particular Digital Point(s) is in Manual mode, the 
output state will remain when the integrator is started. 
You can also specify the OP state in the Parameters 
tab of the Digital Point operation. 
Reset Selector Allows you to initialize the Selector output to 0.
Once the Integrator starts, the Selector will only 
maintain its output at the desired value if the Selector 
is in Hand Select mode or if the operations feeding the 
Selector have been properly initialized.
You can specify a user-defined OP value for each 
Selector. 
Reset Transfer 
Function
Allows you to initialize the transfer function with the OP 
and PV equal to their minimum.
Transfer function block should have a user-specified PV 
to ensure that its output remains constant.
You can specify a user-defined OP value for each 
Transfer Function Block. 
Reset Energy 
Streams
Allows you to set the duty and any duty fluid flows to 
0.
Refer to Spreadsheets 
section for more 
information on Controller, 
Digital Point, Selector, 
and Transfer Function.
Refer to Section 5.5.3 - 
Parameters Tab in the 
Aspen HYSYS 
Operations Guide for 
more information.
Refer to Selection Mode 
Page section from 
Section 5.8.3 - 
Parameters Tab in the 
Aspen HYSYS 
Operations Guide for 
more information.
Refer to Configuration 
Page section from 
Section 5.12.3 - 
Parameters Tab in the 
Aspen HYSYS 
Operations Guide for 
more information.2-81
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2-82 Dynamic Initialization
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ThReset Stream 
Flow Specs
Allows you to reset all stream flow specifications to 0.
Stream flow specifications in Aspen HYSYS dynamics 
are not a natural way for modeling real life behavior 
since flow usually occurs due to pressure gradients. 
You should use the flow specifications with caution. It 
is recommended to set the flow specifications to 0 for 
cold initialization. It should be noted that when there is 
a zero flow specification into a closed valve or flow 
path, an absolute zero pressure is produced, and the 
corrected pressure message appears in the trace 
window.
Reset Stream 
Pressure Specs
You can select this checkbox to set the pressure equal 
to that of the initialization stream (Atmospheric 
pressure is commonly used).
If you insert isolation valves on boundary streams that 
have a different pressure than the initialization stream, 
then you do not need to select this checkbox. Most 
users would prefer to preserve pressure specifications.
Reset Valves There are three options available from the drop-down 
list:
• No. Ignores this initialization option.
• Fail Position. The valve is set to its failed 
position as specified on the Actuator page in the 
Dynamics tab of the Valve. 
Refer to Section 6.7 - Valve in the Aspen 
HYSYS Operations Guide for more information.
• Closed. Indicates a 0% opening.
If a valve is used as a pressure drop device, most likely 
it should not be set to Closed. Instead the actuator 
minimum valve position could be set to 100%, so that 
any non-open signal from the dynamic initialization will 
be rejected.
Pump Off Allows you to set the duty, delta P, head, capacity, and 
speed to 0 for all the pumps. 
This option also turns off the pump by clearing the On 
checkbox in the Pump property view. 
Air Cooler Fans 
Off
Allows you to turn off all the air cooler fans.
Reset 
Compressor/
Expander
Allows you to set the speed, head, pressure drop, and 
duty to 0 for all compressors and expanders.
Reset Heater/
Cooler specs
Allows you to set the duty to 0 for all heaters and 
coolers. The product temperatures are set to be the 
same as the temperature of the initialization stream.
All dynamic unit operations will have their holdups initialized 
to that of the Initialization Stream.
Checkbox Description
Refer to Special 
Behaviours and Usage 
for Unit Operations 
during Dynamic 
Initialization for 
additional information on 
valve behavior during 
dynamic initialization.
Refer to Section 9.3.2 - 
Pump Property View in 
the Aspen HYSYS 
Operations Guide for 
more information.2-82
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ThSpecial behaviors and Usage for Unit 
Operations during Dynamic 
Initialization
The following sections describe the special behaviors and usage 
for some of the unit operations during dynamic initialization:
Valves
When an isolation valve is initialized to be closed, but its 
downstream equipment is not in this initialization area, the 
valves downstream holdup is still initialized to that of the 
initialization stream. As a result, when a valve is cold initialized 
without including any of its downstream equipment, the valve 
will be closed but its downstream holdup (especially for non-
zero volume) will be set to the upstream initialization contents.
Boolean Logical Operations
Boolean logical operations and any complex logic modeling could 
be done on a separate flowsheet with the output connections 
going through a spreadsheet. This way, the flowsheet or 
spreadsheet can be ignored or turned off and you can manually 
line up any inconsistent logic. The cause-and-effect matrix 
operation is not automatically initialized in any way.
Spreadsheets
The Spreadsheet operation is often used in control modeling. 
The Integrator must be started on a new model before 
performing the dynamic initialization. If you are having 
difficulties converging the pressure-flow solver or you just 
want to quickly cold initialize your new model, just set the 
Integrator in manual mode, and start the Integrator. You do 
not need to take any steps. Now you can cold initialize.2-83
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2-84 Dynamic Initialization
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ThHowever in dynamic initialization, a spreadsheet cell should NOT 
be used for the start of control or logic propagation. Instead, 
Controller, Selector, Transfer Function, or Digital Control Points 
could be used as the operator accessible points for turning on or 
off the plant. For instance, the OP state of a Digital Point can be 
imported into a spreadsheet cell and then operated upon before 
it is sent out to the On/Off switch of a pump. As long as the 
spreadsheets do not contain any time dependent behavior, they 
should initialize with the inputs coming from one of the four 
mentioned logical operations.
Event Schedules
It is recommended that all the schedules in the Event Scheduler 
be reset manually to their OFF state before running the case in 
dynamic initialization.
Static Head Contributions
There may be small movements in the cold initialized model if 
static head contributions are turned on and the unit operations 
have non-zero elevations entered. These should be minimal 
since all streams and holdup should contain light gas.2-84
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Th2.7.2 Notes Page
The Notes page provides a text editor that allows you to record 
any comments or information for future reference. You can use 
the Notes page to describe the portion of the simulation case 
that is initialized in this Dynamic Initialization Area.
 Figure 2.782-85
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2-86 Dynamic Initialization
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Th2-86
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Control Theory 3-1
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Th3  Control Theory3-1
3.1  Introduction................................................................................... 2
3.2  Process Dynamics .......................................................................... 3
3.2.1  Characteristic Parameters of the Process System .......................... 3
3.3  Basic Control................................................................................ 10
3.3.1  Available Control Operations .................................................... 16
3.4  Advanced Control......................................................................... 32
3.4.1  Model Predictive Control.......................................................... 32
3.5  General Guidelines....................................................................... 39
3.5.1  Effect of Characteristic Process Parameters on Control................. 39
3.5.2  Choosing the Correct Controller................................................ 41
3.5.3  Choosing Controller Tuning Parameters ..................................... 42
3.5.4  Setting Up a Control Strategy .................................................. 48
3.6  References................................................................................... 58
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3-2 Introduction
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Th3.1 Introduction
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 operating ability of the plant.
While 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 is suitable for implementation
• stress the system with disturbances as desired to test for 
plant performance
Even after a plant has started operation, process engineers can 
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 
analyze performance off-line on a dynamic simulator, instead of 
disturbing the actual process.
Three topics are covered in this chapter:
• The characteristic parameters of a process are discussed 
in the Section 3.2 - Process Dynamics. 
• The control strategies available in Aspen HYSYS are 
discussed in the Section 3.3 - Basic Control and the 
Section 3.4 - Advanced Control. 
• The Section 3.5 - General Guidelines outlines some 
steps you can follow to implement a control strategy in 
Aspen HYSYS. Included in this section are several 
techniques that can be used to determine possible initial 
tuning values for the controller operations.3-2
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Control Theory 3-3
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Th3.2 Process Dynamics
As a precursor to understanding the concepts of process control, 
the dynamic characteristics of the process are discussed. The 
task of designing a control scheme is best carried out if there is 
a good understanding of the process system being studied. A 
process’ response to 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.
As detailed in Chapter 1 - Dynamic Theory, 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.
3.2.1 Characteristic 
Parameters of the Process 
System
It is easiest to define a chemical process system using the 
general conservation principle which states that:
Rate of accumulation = Input - Output + Internal 
Generation
(3.1)3-3
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3-4 Process Dynamics
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ThTo describe some characteristic parameters of a chemical 
process system, the general conservation principle is applied to 
a flow relation first order liquid level system:
The conservation of material in the tank is expressed as follows:
where:  
H = liquid height in the tank
A = cross-sectional area of the tank
Fi = inlet flow rate
Fo = exit flow rate
There is a non-linear relationship describing the flow out of the 
bottom of the tank, Fo, and the liquid height in the tank, H. 
However, to express Equation (3.2) 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 where the flow does not vary considerably over time. 
The exit flow, Fo, can be expressed in terms of the linearity 
constant, R (the valve resistance):
 Figure 3.1
(3.2)
(3.3)
AdHdt
------- Fi Fo–=
Fo
H
R
---=3-4
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Control Theory 3-5
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ThEquation (3.2) can therefore be expressed as:
Equation (3.5) is a general first-order differential which can be 
expressed in terms of two characteristic parameters: the steady 
state gain, K, and the time constant, :
where:  
y(t) = output of the system
u(t)= input to the system
K = steady state gain 
 = time constant of the system
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, , 
and the steady state gain, K, can be expressed as:
(3.4)
(3.5)
(3.6)
(3.7)
AdHdt
------- Fi
H
R
---   –=
RAdHdt
------- H+ RFi=
τ
τdydt
----- y t( )+ Ku t( )=
τ
τ
τ AR  and  K R ==3-5
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3-6 Process Dynamics
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ThWhen a step function of magnitude M is applied to the general 
first-order system, the output response, y(t), is as follows:
As shown, the output, y(t), attains 63.2% of its final steady 
state value in one time constant. The output’s response in 
equation form is:
or in terms of the first-order tank example:
The following is a list of characteristic parameters that can be 
defined in terms of the first-order response illustrated in the 
previous example.
 Figure 3.2
(3.8)
(3.9)
y t( ) MK 1 e
t–
τ
----
–=
H t( ) MR 1 e
t
AR
-------–
–=3-6
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Control Theory 3-7
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ThProcess Gain
The process gain is defined as the ratio of the change/deviation 
in the process output to the change/deviation in the process 
input. The change in the process input is defined in Equation 
(3.6) as u(t). The change in the process output is defined as 
y(t). The first term in Equation (3.6) is transient and becomes 
zero at steady state. Therefore, the gain can be calculated as 
shown in the equation below. 
where:  
ySSnew = new steady state y
uSSnew = new steady state U
For this liquid level example, the steady state gain, K, is the 
valve resistance, R. Therefore, a step change in the flow into the 
tank of magnitude M results in a change in liquid level, H(t), in 
the tank equal to MR.
Time Constant
The time constant, , defines the speed of the response. The 
response of the system always follow the profile shown in 
Figure 3.2. After  time units, the response y(t) equals 
0.632MK or 63.2% of the ultimate gain. This is always true for 
first-order systems without time delays. For this liquid level 
example, the time constant is the product of the area of the 
tank, A, and the resistance of the exit valve, R.
(3.10)Steady state gain
ySSnew ySS–
uSSnew uSS–
------------------------------ K= =
τ
τ
3-7
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3-8 Process Dynamics
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ThCapacity
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 of the system and the time constant of the response:
In the liquid level example, the capacity is the cross sectional 
area of the tank. Since the capacity of a system is proportional 
to the time constant, , 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:
(3.11)
The time constant is the same as the hold up or residence 
time.
(3.12)
C τ
R
--=
τ
Attenuation 1  Response Amplitude out of the system
Disturbance Amplitude into the system
---------------------------------------------------------------------------------------------–
Attenuation =   1 Amplitude Ratio–
=
3-8
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Control Theory 3-9
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ThThe input, u(t), to the first order system can be changed from a 
step function to a sinusoidal function: 
where:  
 = frequency of the input response
M = amplitude of the input function
The response of the system y(t) becomes:
where:  
After the transient term becomes negligible (the first bracketed 
term), an ultimate periodic response remains (the second term). 
The response amplitude of the system is therefore:
Since the disturbance amplitude into the system is M, the 
amplitude ratio is:
(3.13)
(3.14)
(3.15)
(3.16)
u t( ) M ωt( )sin=
ω
y t( ) MK ωτ
ωτ( )2 1+
-----------------------e t τ/– 1
ωτ( )2 1+
--------------------------- ωt φ+( )sin+=
φ ωτ–( )1–tan=
y t( ) MK
ωτ( )2 1+
---------------------------=
AR K
ωτ( )2 1+
---------------------------=3-9
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3-10 Basic Control
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ThDead 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 be 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 is easily modeled using the Transfer 
Function block operation.
3.3 Basic Control
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 PID 
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 can be realized using interacting PID Controller 
operations. There is a feed forward controller built into the PID.
Instrumentation dynamics can also be modeled in Aspen HYSYS, 
increasing the fidelity of the simulation with real valve 
dynamics. Final control elements can be modeled with 
hysteresis. The valve response to controller input can be 
modeled as instantaneous, linear, or first order. Dead time, lags, 
leads, whether they originate from disturbances or within the 
process control loop can be modeled effectively using the 
Transfer function operation.
One has cascade control when the output of one controller is 
used as the set point of another controller.3-10
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Control Theory 3-11
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ThTerminology
Before reviewing the major control operations that are available 
in Aspen HYSYS, it is useful to be familiar with the following 
terms.
Disturbances
A disturbance upsets the process system and causes the output 
variables to move from their desired set points. 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
An open loop response from a process is determined by varying 
the input to a system and measuring the output’s response. The 
open loop response to a first-order system from a step input is 
shown in Figure 3.2. In open loop control, the controller sets 
the input to the process without any knowledge of the output 
variable that closes the loop in feedback control schemes. 
 Figure 3.33-11
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3-12 Basic Control
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ThA 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.
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 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.
Feedback control attempts to maintain the output variable, PV, 
at a user-defined set point, SP. There are some basic steps that 
are carried out by the controller to achieve this task:
1. Measure the output variable, PV.
2. Compare the measured value, PV, with the desired set point 
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.
 Figure 3.43-12
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Control Theory 3-13
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Th3. Supply the error, E(t), to the general control equation. The 
value of the desired percent opening of the control valve, 
OP%, is calculated.
4. The value of OP% is passed to the final control element 
which determines the input to the process, U(t).
5. The entire procedure is repeated. 
The general control equation for a PID controller is given by: 
where:  
OP(t) = controller output at time t
E(t) = error at time t
Kc = proportional gain of the controller
Ti = integral (reset) time of the controller
Td = derivative (rate) time of the controller
 Figure 3.5
(3.17)OP t( ) KcE t( )
Kc
Ti
----- E t( ) KcTd
dE t( )
dt
-------------+∫+=3-13
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3-14 Basic Control
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ThDirect and Reverse Acting
The input to the feedback controller is called the error or the 
difference between the output process variable and the set 
point. The error is defined 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 
defined as reverse acting.
where:  
SP(t) = set point
PV(t) = measured output process variable
If the PV rises above the SP, the OP, or input to the process, 
decreases. If the 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:
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 valve and hence the flow of heat.
(3.18)
(3.19)
E t( ) SP t( ) PV t( )–=
E t( ) PV t( ) SP t( )–=3-14
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Control Theory 3-15
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ThStability
The stability of a system is a very important aspect to consider 
when designing control schemes. Most systems can 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 responds in one of three ways:
• The response proceeds to new steady state and 
stabilizes.
• The response oscillates continuously with a constant 
amplitude.
• The response grows continuously and never reaches 
steady state conditions.
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 grow unbounded. 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.
 Figure 3.63-15
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3-16 Basic Control
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ThFor example, a pure integrating process, such as a tank with a 
pumped (constant) exit flow, is non-self-regulating since a 
bounded increase in the flow input to the system from steady 
state results in the response (liquid height) to increase 
continuously.
A prerequisite for closed loop control is that the closed loop 
response is stabilizable. 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 of Kc cannot increase 
indefinitely. The response remains stable up to a certain value of 
Kc. Increasing Kc beyond the stability limit can 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
3.3.1 Available Control 
Operations
Modeling Hardware Elements
The plant can be simulated more accurately by modeling the 
hardware elements of the control loop. Non-linearities can be 
modeled in the Valve operation in the Actuator page of the 
Dynamics tab.3-16
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Control Theory 3-17
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ThSensors
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 Min and Max 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.
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 can be used. The 
following is a list of the available dynamic modes for the valve 
operation:
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 can be 
defined more easily by expressing flow as a percentage, Cv (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 %Cv as a function of the 
actual valve percent opening.
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 modeled in the response of the actuator 
position to changes in the desired OP%. The actuator time constant 
can be specified in the Parameters field.
Similarly, 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 is specified in the Parameters field. In 
effect, a second order lag can be modeled between the valve 
position and the desired OP%.
Linear The actuator can be modeled to move to the desire OP% at a 
constant rate. This rate is specified in the Parameters field.
For more information 
about the dynamic valve 
operation, see Section 
6.7 - Valve in the Aspen 
HYSYS Operations 
Guide.3-17
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3-18 Basic Control
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ThThere 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.
The valve characteristics are shown graphically as follows:
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.
 Figure 3.7
CV% Valve Opening %=
CV% Valve Opening %( )0.5=
CV% Valve Opening %( )3=
0 20 40 60 80 100
0
20
40
60
80
100
Quick Opening
Linear
Equal 
Percentage
% Valve Position
% Cv
CONTROL VALVE FLOW CHARACTERISTICS3-18
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Control Theory 3-19
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ThFeedback Control
Digital On/Off
Digital On/Off control is one of the 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 set point, the heating element turns 
on. When the temperature rises above the set point, 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 set point. 
The opposite is true when the OP is ON option is set to “PV > 
Threshold”.
(3.20)
(3.21)
For more information on 
the Digital Point 
operation in Aspen 
HYSYS, see Section 5.5 
- Digital Point in the 
Aspen HYSYS 
Operations Guide.
OP 0% for PV > SP and OP 100% for PV < SP ==
OP 0% for PV < SP and OP 100% for PV > SP ==3-19
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3-20 Basic Control
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ThOne main characteristic of the On/Off controller is that the PV 
always cycle about the set point.
The cycling frequency depends on the dynamics of the process. 
Those systems with a large capacity (large time constant) cycles 
less frequently. The On/Off controller is an appropriate controller 
if the deviation from the set point is within an acceptable range 
and the cycling does not destabilize the rest of the process.
 Figure 3.83-20
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Control Theory 3-21
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ThProportional Control (P-only)
Unlike On/Off control, proportional control can damp out 
oscillations from disturbances and stop the cycling of the 
process variable. P-only control is implemented in Aspen HYSYS 
by setting the values of Td and Ti to infinity (or in Aspen HYSYS 
to ) in the PID Controller operation. With P-only 
control, oscillations that occur in the process variable due to 
disturbances or changes in the set point dampen out the 
quickest (have the smallest natural period) among all other 
simple feedback control schemes. The output of the proportional 
control is defined as: 
The value of the bias, OPss, is calculated when the controller is 
switched to Automatic mode. The set point is defaulted to equal 
the current PV. In effect, the error becomes zero and OPss is 
then set to the value of OP(t) at that time.
A sustained offset between the process variable and the set 
point is always present in this sort of control scheme. The error 
becomes zero only if:
• the bias, OPss, equals the operating variable, OP
• Kc becomes infinitely large
However, Kc cannot practically become infinitely large. The 
magnitude of Kc is restricted by the stability of the closed loop 
system.
(3.22)
In general, a higher controller gain gives tighter control. 
However, the value of Kc cannot increase indefinitely. The 
response remains stable up to a certain value of Kc. 
Increasing Kc beyond the stability limit causes the closed 
loop response to become unstable.
OP t( ) OPss KcE t( )+=3-21
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3-22 Basic Control
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ThThe following shows the effect of the magnitude of Kc on the 
closed loop response of a first order system to a unit step 
change in set point.
Proportional 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.
 Figure 3.93-22
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Control Theory 3-23
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ThProportional Integral Control (PI)
Unlike P-only control, proportional + integral control can 
dampen out oscillations and return the process variable to the 
set point. 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 takes longer to line out 
(dampen) the process variable than P-only control. The output 
of the proportional controller + integral controller is defined as:
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. 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:
(3.23)
 Figure 3.10
OP t( ) KcE t( )
Kc
Ti
----- E t( )∫+=3-23
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3-24 Basic Control
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ThThe integral time, Ti, is defined as the amount of time required 
for the controller output to move an amount equivalent to the 
error. Because the relationship between Ti and the control action 
is reciprocal, increasing Ti results in less integral action, while 
decreasing Ti results in greater integral action. The integral time 
should be decreased (increased integral action) just enough to 
return the process variable to the set point. Any more action 
only serves to lengthen the response time. 
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 can cause oscillations in the system 
and cause the control system to become unstable. The larger 
the integral action the more likely it becomes 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: 
Td 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 is ramped, as well. For the same 
ramped input the derivative only response is 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 
(3.24)OP t( ) KcE t( )
Kc
Ti
----- E t( ) KcTd
dE t( )
dt
-------------+∫+=3-24
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Control Theory 3-25
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Thinfinity in either the set point or the measured process variable 
is not physically possible, signals which have short rise times 
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 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 PID 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 PID 
control which shortens the controller’s response to be 
comparable to the response of a P-only controller. 
However, if a controller has a noisy input which cannot be 
filtered or minimized in the process, PID control is not a suitable 
control scheme.
 Figure 3.113-25
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3-26 Basic Control
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ThCascade Control
Cascade control is a technique that implements a secondary 
feed back control loop within a primary feedback control loop. 
Cascade control can be used when there are significant 
disturbances to the manipulated variable of the primary loop. 
A secondary loop is created to control the manipulated variable 
of the primary loop. The primary loop then manipulates the set 
point of the secondary controller.
Consider an example where the main process variable could 
better be controlled using a cascade control scheme. The Btms 
stream of a distillation column is being heated in a reboiler 
whose energy source is steam. 
As shown above, the objective is to regulate T, the temperature 
of the Btms stream. A possible feedback control scheme is to 
control the Btms stream temperature using the steam valve 
opening. This traditional control setup works well if the steam 
valve opening corresponds exactly with the flow of steam to the 
reboiler. 
 Figure 3.123-26
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Control Theory 3-27
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ThHowever, disturbances occurring in the steam supply header can 
alter the flow of steam significantly even at a constant valve 
opening. The only way this type of disturbance can be detected 
by the controller is with a change in the Btms temperature, T. If 
there is a large lag associated with the reboiler heating system 
and if disturbances to the steam header flow occur frequently, 
there is the possibility that the Btms temperature, T, never 
settles to the desired set point.
Cascade control dampens disturbances to the inlet flow of steam 
by using a feedback controller within a feedback controller. The 
primary (or Master) controller measures the variable to be 
controlled, the exit temperature T, and determines the required 
steam flow requirement in a feedback loop. 
The steam flow requirement becomes the set point of the 
secondary feedback controller. The secondary control loop is set 
up as a regular feedback controller. It measures the flow rate of 
steam as the Process Variable Source. 
 Figure 3.133-27
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3-28 Basic Control
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ThThe Output Target Object is the steam control valve opening 
which needs to be sized as usual. Because the Output Target 
Object of the primary loop is the secondary loop’s set point, the 
primary loop does not have a control valve. It is the secondary 
controller that physically interacts with the system by adjusting 
the steam valve.
Cascade control can be successfully implemented if the following 
occur:
• Disturbances affect the input to the primary feedback 
controller.
• Disturbances are measurable by the secondary control 
loop.
• The response period of the primary loop is more than 3 
times greater than the response period of the secondary 
loop.
Starting up a Cascade System1
To put a cascade system into operation in Aspen HYSYS:
1. Place the primary control in manual. This breaks the cascade 
and allows the secondary controller to be tuned.
2. Tune the secondary controller as if it were the only control 
loop present.
3. Return the secondary controller to the remote set point by 
placing the primary controller in automatic.
4. Now tune the primary loop normally. If the system begins to 
oscillate when the primary controller is put in automatic, 
reduce the primary controller’s gain.
Feedforward Control
Feedforward control can 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 3-28
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Control Theory 3-29
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Thto a process. However, feedforward control is useful if there are 
measured disturbances which affect the output of the process.
With feedback control, the controller requires information about 
the controlled process variable, PV, and the set point, SP, in 
order to determine the value of OP%, the desired valve percent 
opening of the input to the process. To determine the value of 
OP%, the feedforward controller requires information from two 
variables: the set point of the process variable, SP, and the 
disturbance affecting the process. 
The current implementation contains a lead/lag transfer 
function.
 Figure 3.143-29
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3-30 Basic Control
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ThConsider an example of a liquid stream being heated in a steam 
heat exchanger.
It is desired to control the Exit stream temperature, T2, at a 
certain set point, SP, using the Steam flow as the manipulated 
variable. However, the process suffers from frequent changes in 
the Feed temperature, T1. 
To determine the value of OP%, the values of SP and T1 are 
required by the controller. At steady state, the overall energy 
balance relates the steam flow to the disturbance of the process, 
T1, and the temperature of stream Exit, T2:
where:  
Fs = steam flow
 = heat of condensation for steam
F = flow of stream Exit
Cp = specific heat of stream Exit
 Figure 3.15
(3.25)Fsλ FCp T2 T1–( )– 0=
λ
3-30
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Control Theory 3-31
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ThFrom 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 
(3.25) becomes:
To calculate the feedforward controller output, a linear relation 
is assumed to exist between the steam flow and the valve 
opening of the steam valve. Therefore, the final form of the 
feedforward controller equation is:
To successfully implement a feedforward control system, 
consider the following:
• It cannot be implemented if the disturbance is not 
measurable. If unexpected disturbances enter the 
process when pure feedforward control is used, no 
corrective action is taken and the errors build up in the 
system.
• A fairly accurate model of the system is required.
• 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.
The process variable to be controlled is not measured using 
feedforward control. There is no way of confirming that the 
process variable is attenuating disturbances or maintaining a 
desired set point.
Considering that an accurate model of the process is not usually 
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 is probably an offset between the actual 
controlled variable and its desired set point. 
(3.26)
(3.27)
Fs
Cp
λ
------F SP T1–( )=
OP t( )
Cp
λ
-----F SP T1–( )steam valve span
100%
---------------------------------------------=3-31
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3-32 Advanced Control
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ThTherefore, 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. If the 
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.
3.4 Advanced Control
3.4.1 Model Predictive Control
Model Predictive Control (MPC) refers to a class of algorithms 
that compute a sequence of manipulated variable adjustments 
to optimize the future behavior of a plant. A typical MPC has the 
following capabilities:
• Handles multi-variable systems with process interactions.
• Encapsulates the behavior of multiple Single Input Single 
Output (SISO) controllers and de-couplers.
• Uses a process model, i.e., a first order model or a step 
response data is required.
• Incorporates the features of feedforward control, i.e., 
must be a measured disturbance by taking in 
consideration the model disturbances in its predictions.
• Poses as an optimization problem and is therefore 
capable of meeting the control objectives by optimizing 
the control effort, and at the same time is capable of 
handling constraints.
MPC technology was originally developed to meet the specialized 
control needs of power plants and petroleum refineries, but it 
can now be found in a wide variety of application areas 
including:
• chemicals
• food processing
• automotive
• aerospace
• metallurgy
• pulp and paper3-32
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Control Theory 3-33
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ThMPC vs. PID
When the MPC controller is compared to the ubiquitous PID 
controller, some differences are readily apparent. First, it is 
essential that there exist a model of the process to use an MPC 
controller. 
Like most advanced controllers, the model of the process is used 
to form predictions of the process outputs based on present and 
past values of the input and outputs. This prediction is then 
used in an optimization problem in which the output is chosen 
so that the process reaches or maintains its set point at some 
projected time in the future.
MPC Theory
Currently most model predictive control techniques like Dynamic 
Matrix Control (DMC) and Model Algorithmic Control (MAC) are 
based on optimization of a quadratic objective function involving 
the error between the set point and the predicted outputs. In 
these cases, a discrete impulse response model can be used to 
derive the objective function.
Let a0, a1, a2,...,aT represent the value of the unit step response 
function obtained from a typical open loop process, as shown in 
the figure below:
 Figure 3.163-33
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3-34 Advanced Control
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ThFrom the figure above, you can define ai = 0 for . Consider a 
step response resulting from a change ( ) in the input. Let  
be the actual output,  the predicted value of the output 
variable and  the value of the manipulated variable at the nth 
sampling interval. If there is no modeling error and no 
disturbances to them.
Denoting , the convolution model of the single 
step response function, see figure below, is given as follows:  
The control horizon U is the number of control actions (or 
control moves) that are calculated to affect the predicted 
outputs over the prediction horizon V, (i.e., over the next V 
sampling periods). Similarly, the discrete model can be written 
as follows:
(3.28)
(3.29)
 Figure 3.17
(3.30)
i 0≤
Δm cn
ĉn
mn
ĉn cn=
Δmi mi mi 1––=
ĉn 1+ c0 aiΔmn 1 i–+
i 1=
T
∑+=
ĉn 1+ c0 himn 1 i–+
i 1=
T
∑+=3-34
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Control Theory 3-35
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ThWhere hi is the impulse response coefficient. Shifting the model 
back one time step, Equation (3.30) can be written as follows:
Subtracting Equation (3.31) from Equation (3.30), a 
recursive form of the model expressed in incremental change 
can be obtained:
To provide corrections for the influence of model errors and 
unmeasured load changes during the previous time step, a 
corrected prediction is used in the model. This corrected 
value is obtained by comparing the actual value of  with  
and then shifting the correction forward, as follows:
Substituting the corrected prediction in Equation (3.32) results 
in the following recursive form:
The Equation (3.34) can be extended to incorporate 
predictions for a number of future time steps allowing the 
model-based control system to anticipate where the process is 
heading. 
(3.31)
(3.32)
(3.33)
(3.34)
ĉn c0 himn i–
i 1=
T
∑+=
Δm
ĉn 1+ ĉn hiΔmn 1 i–+
i 1=
T
∑+=
cn 1+
*
cn ĉn
cn 1+
* ĉn 1+– cn ĉn–=
cn 1+
* cn hiΔmn 1 i–+
i 1=
T
∑+=3-35
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3-36 Advanced Control
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ThA new design parameter called prediction horizonV is 
introduced, which influences control system performance, and is 
expressed in terms of incremental changes in the manipulated 
variable:
where:  
j = 1, 2, ..., V
Suppose that arbitrary sequence of U input changes are made, 
then using a prediction horizon of V sampling periods, Equation 
(3.35) can be expressed in a vector-matrix form as:
where:  
Using the predicted behavior of the process (see Equation 
(3.36)) over the prediction horizon, a controller in model 
(3.35)
(3.36)
(3.37)
(3.38)
(3.39)
cn j+
* cn j 1–+
* hiΔmn j i–+
i 1=
T
∑+=
cn 1+
*
cn 2+
*
cn 3+
*
.
.
.
cn V+
*
a1 0 0 … 0
a2 a1 0  0
a3 a2 a1  0
.    .
.    .
.    .
aV aV 1– aV 2– … aV U– 1+
Δmn
Δmn 1+
Δmn 2+
.
.
.
Δmn U 1–+
cn P1+
cn P2+
cn P3+
.
.
.
cn PV+
+=
ai hj
j 1=
i
∑=
Pi Sj
j 1=
i
∑= for i 1 2 … V,,,=
Si hi Δmn j i–+
i j 1+=
T
∑= for j 1 2 … V,,,=3-36
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Control Theory 3-37
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Thpredictive control can be designed. The control objective is to 
compute the manipulated variables to ensure that the predicted 
response has certain desirable characteristics (i.e., to have the 
corrected predictions  approach the set point as closely as 
possible). 
One sampling period after the application of the current control 
action, the predicted response is compared with the actual 
response. Using the corrective feedback action for any errors 
between actual and predicted responses, the entire sequence of 
calculation is then repeated at each sampling instant. 
Denoting the set point trajectory, (in other words, the desired 
values of the set point V time steps into the future), as 
, Equation (3.36) can be written as:
where:
 = the  triangular matrix 
 = the  vector of future control moves. 
 and  = the closed loop and open loop predictions, 
respectively, and are defined as follows:
For a perfect match between the predicted output trajectory of 
the closed loop system and the desired trajectory, then  
and Equation (3.40) becomes:
(3.40)
(3.41)
cn j+
*
rn j+ j 1 2 … V,,,=,
Ê A Δm– E'ˆ+=
A V U×
Δm U 1×
Ê E'ˆ
Ê
rn 1+ cn 1+
*–
rn 2+ cn 2+
*–
.
.
.
rn V+ cn V+–
= E'ˆ
En P1–
En P2–
.
.
.
En PV–
=
Ê 0=
Δm A( ) 1– E'ˆ=3-37
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3-38 Advanced Control
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ThThe best solution can be obtained by minimizing the 
performance index:
Here, the optimal solution for an over-determined system  
turns out to be the least squares solution and is given by 
where:
 = pseudo-inverse matrix 
 = matrix of feedback gains (with dimensions )
One of the shortcomings of Equation (3.42) is that it can result 
in excessively large changes in the manipulated variable, when 
 is either poorly defined or singular. One way to overcome 
this problem is by modifying the performance index by 
penalizing movements of the manipulated variable.
where:
 and  are positive-definite weighting matrices for 
predicted errors and control moves, respectively. These 
matrices allows you to specify different penalties to be 
placed on the predicted errors resulting in a better 
tuned controller. 
The resulting control law that minimizes  is 
The weighting matrices  and  contains a potentially large 
number of design parameters. It is usually sufficient to select 
 and  (  is the identity matrix and  is a scalar 
(3.42)
(3.43)
(3.44)
(3.45)
J Δm[ ] ÊTÊ=
U V<( )
Δm ATA( )
1– ATE'ˆ KcE'ˆ= =
ATA( )
1– AT
Kc V U×
ATA
J Δm[ ] ÊTΓuÊ ΔmTΓyΔm+=
Γu Γy
J
Δm ATΓuA Γy+( )
1– ATΓuE'ˆ KcE'ˆ= =
Γu Γy
Γu I= Γy f I= I f3-38
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Control Theory 3-39
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Thdesign parameter). Large values of  penalize the magnitude of 
 more, thus giving less vigorous control. When , the 
controller gains are very sensitive to , largely because of the 
poor definition of , and  must be made small.
3.5 General Guidelines
3.5.1 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 
outline of 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 of a system and the period of the 
disturbances to the system. From Terminology section, 
attenuation is defined as:
The time constant, , 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. 
(3.46)
f
Δm f 0=
U
ATA U
Attenuation = 1 K
ωτ( )2 1+
---------------------------–
τ
3-39
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3-40 General Guidelines
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ThThe frequency of incoming disturbances affects the system’s 
ability to attenuate these disturbances. High-frequency 
disturbances are more easily attenuated than low-frequency 
disturbances.
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 to 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.
With capacity-dominated processes (with little or no dead 
time), proportional-only control can achieve much better 
disturbance rejection. The system itself is able to attenuate 
disturbances in the frequency range that the controller 
cannot. High frequency disturbances can be handled by the 
system. Low frequency disturbances are handled best with 
the controller.3-40
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Control Theory 3-41
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Th3.5.2 Choosing the Correct 
Controller
You should consider what type of performance criteria is 
required for the set point 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 set point. 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 is very robust. It 
is less likely to become unstable.
The following is a flowchart that outlines a method for selecting 
a feedback controller2.
 Figure 3.183-41
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3-42 General Guidelines
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ThIn general, 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 is appropriate. 
3.5.3 Choosing Controller 
Tuning Parameters
The following is a list of general tuning parameters appropriate 
for various processes3. The suggested controller settings are 
optimized for a quarter decay ratio error criterion. 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 help you obtain 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.
Generally, proportional control can be considered the principal 
component of controller equation. 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.3-42
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Control Theory 3-43
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ThFlow Control
Flow in a pipe is typically a fast responding process. The dead 
time and capacity associated with a length of pipe is generally 
small. It is therefore not unusual for the process to be limited by 
the final control element (valve) dynamics. You can easily 
incorporate valve dynamics in the Aspen HYSYS model by 
modifying the valve parameters in the Actuator page of the 
Dynamics tab.
Tuning a flow loop for PI control is a relatively easy task. For the 
flow measurement to track the set point closely, the gain, Kc, 
should be set between 0.4 and 0.65 and the integral time, Ti, 
should be set between 0.05 and 0.25 minutes. 
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 can cause the 
control loop to become unstable at different operating 
conditions. 
Therefore, the highest process gain should be used to tune the 
controller. If a stability limit is reached, the gain should be 
decreased, but the integral action should not. Since flow 
measurement is naturally noisy, derivative action is not 
recommended.
Liquid Pressure Control
Like the flow loop process, the liquid pressure loop is typically 
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 and Integral-only 
control, depending on your performance requirements. Like flow 
control, the highest process gain should be used to tune the 
controller. Typically, the process gain for pressure is smaller than 
the flow process gain. The controller gain, Kc, should be set 
between 0.5 and 2 and the integral time, Ti, should be set 
between 0.1 and 0.25 minutes.3-43
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3-44 General Guidelines
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ThLiquid 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 of level 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, OPss, 
be specified as 50% for P-only control. 
This ensures that the control valve is wide open for a level of 
75% and completely shut when the level is 25% for a set point 
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, Ti, should be set between 1 and 5 minutes.
Common sense dictates that the manipulated variable for level 
control should be the stream with the most direct impact on the 
level. For example, in a column with a reflux ratio of 100, there 
are 101 units of vapour entering the condenser and 100 units of 
reflux leaving the reflux drum for every unit of distillate leaving. 
The reflux flow or vapour boilup is used to control the level of 
the reflux drum. If the distillate flow is used, it would only take a 
change of slightly more than 1% in the either the reflux or 
vapour flow to cause the controller to saturate the distillate 
valve.3-44
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Control Theory 3-45
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ThGas 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.
The pressure loop can easily be tuned for PI control. The 
controller gain, Kc, should be set between 2 and 10 and the 
integral time, Ti, should be set between 2 and 10 minutes.
Like liquid level control, it is necessary to determine what affects 
pressure the most. For example, on a column with a partial 
condenser, you can determine whether removing the vapour 
stream affects pressure more than condensing the reflux. If the 
column contains noncondensables, these components can affect 
the pressure considerably. In this situation, the vent flow, 
however small, should be used for pressure control.
Temperature Control
Temperature dynamic responses are generally slow, so PID 
control is used. Typically, the controller gain, Kc, should be set 
between 2 and 10, the integral time, Ti, should set between 2 
and 10 minutes, and the derivative time Td, should be set 
between 0 and 5 minutes.
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 an amplitude 
that neither falls nor rises over time. At a system’s stability 
margins, there are two important system parameters, the 
ultimate period and the ultimate gain, which allow the 
calculation of the proportional, integral, and derivative gains.3-45
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3-46 General Guidelines
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ThATV 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:
1. 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, then move valve -
2h%.
4. When the process variable (PV) crosses the set point, move 
the valve position +2h%.
5. Continue until a limit cycle is established.
6. Record the amplitude of the response, A. Make sure to 
express A as a fraction of the PV span.
 Figure 3.193-46
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Control Theory 3-47
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Th7. The tuning parameters are calculated as follows:
Ziegler-Nichols Tuning Technique
The Ziegler-Nichols4 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 process variable, PV.
3. The tuning parameters are calculated as follows:
Tuning Parameter Equation
Ultimate Gain
Ultimate Period
Controller Gain
Controller Integral Time
Tuning Parameter Equation
Ultimate Gain
Ultimate Period
Ku
4h
πa
------=
Pu Period taken from limit cycle=
Kc
Ku
3.2
------=
Ti 2.2Pu=
Ku Controller gain that produces limit cycle=
Pu Period taken from limit cycle=3-47
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3-48 General Guidelines
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Th3.5.4 Setting Up a Control 
Strategy
This section outlines how to create a control strategy in Aspen 
HYSYS. First follow the guidelines outlined in Section 1.5.2 - 
Moving from Steady State to Dynamics to setup a stable 
dynamic case. In many cases, an effective control strategy 
serves to stabilize the model.
You can install controllers in the simulation case either in steady 
state or Dynamics 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.
In the following sections you will:
• select the controlled variables in the plant
• select controller structures for each controlled variable
• set final control elements
• set up the data book and strip charts
• set up the controller faceplates
• set up the integrator
• fine tune the controllers
Selecting 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. 
Controller Gain
Controller Integral Time
Tuning Parameter Equation
Kc
Ku
2.2
------=
Ti Pu 1.2⁄=3-48
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Control Theory 3-49
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ThIn Aspen HYSYS, there are a number a variables that 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 pressure-flow 
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. 
These specifications should be reserved 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, you can attach 
the variable to the Transfer Function operation. A variety of 
different forcing functions and disturbances can be modeled in 
this manner.
The behavior of controlled variables are determined by the type 
of controller and the tuning parameters associated with the 
controller. Typically, the number of control valves in a plant 
dictate the possible number of controlled variables. There are 
more variables to control in Dynamics mode than in Steady 
State mode. 
For example, a two-product column in Steady State mode 
requires two steady state specifications. The simulator then 
manipulates the other variables in the column to satisfy the 
provided specifications and the column material and energy 
balances. 
The same column in Dynamics mode requires five specifications. 
The three new specifications correspond to the inventory or 
integrating specifications that were not fixed in steady state. 
The inventory variables include the condenser level, the reboiler 
For more information on 
setting pressure-flow 
specifications in a 
dynamic simulation case, 
see Chapter 1 - 
Dynamic Theory.3-49
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Thlevel, and the column pressure.
Selecting Controller Structures for 
Each Controlled Variable
Select appropriate controller structures for each controlled 
variable in the simulation case. 
The controller operations can be added in either steady state or 
Dynamics mode. However, controllers have no effect on the 
simulation in Steady State mode. You must specify the following 
information to fully define the PID Controller operation.
Connections Tab
In the Connections tab, you can specify/select the variable 
information entering and exiting the controller.
Process Variable (PV)
The process variable can be specified by clicking the Select PV 
button. The controller measures the process variable in an 
attempt to maintain it at a specified set point, SP.
Operating Variable (OP)
The operating variable, OP, can be specified 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 can be specified as a physical 
valve in the plant, a material stream, or an energy stream.
A good controller strategy includes the control of both 
integrating variables and steady state variables. By 
maintaining the integrating variables at specified set points, 
controllers add stability to the plant. Other controllers 
maintain the desired steady state design specifications such 
as product composition and throughput.
Some general guidelines 
in selecting appropriate 
controllers can be found 
in Section 3.5.2 - 
Choosing the Correct 
Controller.3-50
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ThOperating 
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 can be 
modeled if a physical valve is selected.
Material Stream If a material stream is selected 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:
Aspen HYSYS varies the flow specification of the 
material stream according to the calculated controller 
output, OP. (Therefore, a non-realistic situation can 
arise in the dynamic case since material flow is not 
dependent on the surrounding conditions.)
Energy Stream If an energy stream is selected as an operating 
variable, you can select a Direct Q or a Utility Fluid 
Duty Source by clicking the Control Valve button. 
If the Direct Q option is selected, 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:
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.
Flow OP %( )
100
----------------- Flowmax Flowmin–( ) Flowmin+=
Energy Flow OP %( )
100
----------------- Flowmax Flowmin–( ) Flowmin+=3-51
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ThParameters Tab
The direction of the controller, the controller’s PV range, and the 
tuning parameters can be specified in the Parameters tab. 
A controller’s direction (whether it is direct or reverse acting) 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 that the sensor is to measure.
Tuning parameters are specified in the tuning field. 
Final Control Elements
Set the range on the control valve at roughly twice the steady 
state flow you are controlling. This can be achieved by sizing the 
valve with a pressure drop between and 15 and 30 kPa with a 
valve percent opening of 50%. 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 property view 
of the material or energy stream. This property view is displayed 
by clicking on the Control Valve button in the PID Controller 
property view.
The final 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 of final 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 can possibly become more stable.
For more information 
about whether a 
controller is direct or 
reverse acting, refer to 
Terminology section.
For more information 
about the choice of tuning 
parameters for each 
controller, see Section 
3.5.3 - Choosing 
Controller Tuning 
Parameters.
For more information 
about the 
characterization of final 
control elements in 
Aspen HYSYS, see 
Modeling Hardware 
Elements section from 
Section 3.3.1 - 
Available Control 
Operations.3-52
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Control Theory 3-53
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ThSetting up the Databook & Strip 
Charts
Set up strip charts for your model. Enter the Databook property 
view. Select the desired variables that are to be included in the 
strip chart in the Variables tab. 
From the Strip Charts tab, add a new strip chart by clicking the 
Add button and activate the variables to be displayed on the 
strip chart. No more than six variables should be selected for 
each strip chart to keep it readable.
Click on the Strip Chart button in the View group to see the strip 
chart. Size as desired and then right-click on the strip chart. 
Select Graph Control command from the Object Inspect menu. 
 Figure 3.20
 Figure 3.213-53
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3-54 General Guidelines
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ThThere are six tabs, where you can manipulate the strip chart 
display features, set the numerical ranges of the strip chart for 
each variable, the nature of the lines for each variable, and 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.
Setting up the Controller Faceplates
Click on the Face Plate button in the PID Controller property 
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.
 Figure 3.22
Clicking the Face Plate button opens 
the Face Plate property view.
You can edit the 
set point or mode 
directly from the 
Face Plate.3-54
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Control Theory 3-55
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ThSetting up the Integrator
The integration step size can be modified in the Integrator 
property view located in the Simulation menu. If desired, 
change the integration step size to a smaller interval. The 
default integration time step is 0.5 seconds. 
Changing the step size causes the model to run slowly, but 
during the initial switch from steady state to Dynamics mode, 
the smaller step sizes allow the system to initialize better and 
enable 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 might be 
specified if the process can maintain stability.
 Figure 3.233-55
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3-56 General Guidelines
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ThFine Tuning of Controllers
Before the Integrator runs, each controller should be turned off 
and then put back in manual mode. This initializes the 
controllers. Placing the controllers in manual defaults the set 
point 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 once the Integrator runs. The transition 
of most unit operations from steady state to Dynamics mode is 
smooth. 
However, controller tuning is critical if the plant simulation is to 
remain stable. Dynamic columns, for example, are not open 
loop stable like many of the unit operations in Aspen HYSYS. 
Any large disturbances in the column can result in simulation 
instability.
After the Integrator is running:
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 closed loop performance is obtained. 3-56
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Control Theory 3-57
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Th5. 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 are tested to reduce the energy requirements of 
unit operations.
Stability
It is 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 (3.47).
To control the process, the controller must interact with the 
actual process. This is achieved by using the effective gain, Keff, 
which is essentially the controller gain with units. The effective 
gain is defined as:
(3.47)
(3.48)
The stability of the closed loop response is not only 
dependent on the controller gain, Kc, but also on the PV 
range parameters provided and the maximum flow allowed 
by the control valve. Decreasing the PV range increases the 
effective gain, Keff, and therefore decreases the stability of 
the overall closed loop response. Decreasing the final control 
element’s flow range decreases the effective gain, Keff, and 
therefore increases the stability of the closed loop response.
Kc
OP% PV Range×
error
--------------------------------------------=
Keff
Kc Flowmax Flowmin–( )
PV Range
------------------------------------------------------------=3-57
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3-58 References
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ThIt is therefore possible to achieve tight control in a plant and to 
have the simulation case become unstable due to modifications 
in the PV range or Cv 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 of the 
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.
3.6 References
 1 Svrcek, Bill. A Real Time Approach to Process Controls First Edition 
(1997) p. 91
 2 Svrcek, Bill. A Real Time Approach to Process Controls First Edition 
(1997) p. 70
 3 Svrcek, Bill. A Real Time Approach to Process Controls First Edition 
(1997) p. 105-123
 4 Ogunnaike, B.A. and W.H. Ray. Process Dynamics, Modelling, and 
Control Oxford University Press, New York (1994) p. 531
 5 Seborg, D. E., T. F. Edgar and D. A. Mellichamp. Process Dynamics 
and Control John Wiley & Sons, Toronto (1989) p. 649-667
The process gain has units which are reciprocal to the 
effective gain.3-58
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Index
www.c
The doA
Accumulation 1-14
Advanced Holdup
componsitions tab 1-29
efficiencies tab 1-27
general tab 1-25
nozzles tab 1-26
properties tab 1-28
Advanced Holdup Properties 1-24
Ambient 2-44
ATV Tuning 3-46
B
Balance
 See Material Balance, Component Bal-
ance and Energy Balance
C
Capacity 3-8, 3-39
Cascade Control 3-26
See also PID Controller
Cold Initialization 2-76
adding objects 2-77
configuring objects 2-79
removing objects 2-77
selecting fluid package 2-79
Component Balance 1-9
Control Strategy 3-48
Controller
available control operations 3-16
choosing correct 3-41
selecting variables 3-50
tuning 3-42
Controller Theory
capacity 3-8
dead time 3-10
process gain 3-7
terminology 3-11
time constant 3-7
Cv
See Valve Flow Coefficient (Cv)
D
Databook 3-53
DDE 2-65
Dead Time 3-10, 3-40
Degrees of Freedom 1-34
Distributed Models 1-6
Dynamic Assistant
general tab 2-7
other specs tab 2-21
pressure flow specs tab 2-14
streams tab 2-8
unknown sizing tab 2-15
user items tab 2-24
Dynamic Initialization 2-73
See also Cold Initialization
Dynamic Simulation
control strategy 3-48
converting steady state models 1-44
degrees of freedom 1-34
differences from steady state 1-43
general concepts 1-6
linear 1-7
non-linear 1-7
theory 1-3, 1-6
Dynamics Assistant 2-4
E
Efficiencies 1-16
Energy Balance 1-10, 1-20
Equation Summary
property view 2-28
Event Scheduler 2-45
F
Face Plate 3-54
Feedback Control 3-12, 3-19
direct acting 3-14
reverse acting 3-14
Feedforward Control 3-28
Flash
non-equilibrium 1-15
Flow Control 3-43
Flowsheet Menu
dynamic initialization 2-73
G
Gas Pressure Control 3-45
H
Heat Loss Model 1-20
detailed 1-23
parameters 1-22
simple 1-22
Holdup Model 1-13I-1
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I-2 
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The doadvantages of 1-13
assumptions 1-14
I
Implicit Euler Method 1-11
Integration Strategy 1-12
Integrator 2-38, 3-55
K
k Values 1-32
L
Liquid Level Control 3-44
Liquid Pressure Control 3-43
Lumped Models 1-6
M
Material Balance 1-7
N
Nozzles 1-18
O
Open Loop Control 3-11
Operations
general guidelines 1-42
Ordinary Differential Equations 1-7
P
Partial Differential Equations 1-6
PID Controller
ATV tuning 3-46
choosing correct controller 3-41
See also Controller Theory
tuning 3-42, 3-56
Ziegler-Nichols tuning 3-47
Pressure Flow
failed convergence 1-48
model 1-31
specifications 1-34, 1-37
volume balance 1-31
Pressure Flow Solver 1-29
simultaneous solution 1-30
Process Dynamic 3-3
Process Gain 3-7
Proportional Control 3-21
Proportional Integral Control (PI) 3-23
Proportional Integral Derivative Control (PID) 
3-24
R
Reactions 1-23
Resistance Equation 1-33
S
Scripts 2-64
Sensors 3-17
Singular Problem 1-48
Solution
Implicit Euler 1-11
Stability 3-15, 3-57
Static Head Contributions 1-19
T
Temperature Control 3-45
Time Constant 3-7
Tuning Methods 3-45
V
Valve
fail-safe function 1-61
modes 1-58
positions 1-62
stickiness 1-58, 1-60
Valve Dynamics 3-17
Valve Flow Coefficient (Cv) 1-33
Valve Type 3-17
Volume Balance 1-31
Z
Ziegler-Nichols Tuning Technique 3-47I-2
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