Fundamentals of Process Plant Equipment Control Manual

Fundamentals of Process Plant Equipment Control 25-28 June Petroleum Training Centre

Ron Frend

Fundamentals of Process Plant & Equipment Control ©Ron Frend 2006

1

FUNDAMENTALS AND HYDRAULICS

8

PFD - Process Flow Diagram The Process Flow Diagram - PFD, a schematic illustration of the system

9 9

P&ID - Piping and Instrumentation Diagram

11

P&ID / PFD Symbols General Instrument or Function Symbols -1 Instrument or Function Symbols - 2 General Instrument or Function Symbols 3 General Instrument or Function Symbols – 4

13 13 14 15 16

FIRST LAW OF THERMODYNAMICS Thermodynamics Example 1:

17 17 23

HYDRAULICS & FLUID FLOW

24

Pressure & Head

25

Bernoulli’s Theorem How pressure and velocity interact

33 33

Liquid Flow Flow Units Restriction Flow Sensors

35 35 36

Two Phase & Multiphase Flow

38

Reynolds Number

41

SOME NOTES FOR THE METRIC PIPE FRICTION CHART SHOWN BELOW

43

FRICTION LOSS FOR METRIC PIPE, VALVES AND FITTINGS

44

PUMPS & COMPRESSORS

45

Centrifugal pump designs

45

Pump Affinity Laws

47

Performance Curves

49

Compressors and Expanders CENTRIFUGAL COMPRESSORS

52 54

HEAT TRANSFER AND REACTION ENGINEERING

58

Thermal Conductivity

59

Conduction & Convection

61


2

Conduction: Examples of conduction: Convection: Example of convection:

61 61 61 61

Insulation Heat transfer coefficients and calculation Heat exchangers, type and sizing Steam Reboilers Condensers and sub-cooling Introduction to energy recovery

62 63 65 69 70 73

An Introduction to Pinch Technology What is Pinch Technology? Basic Concepts of Pinch Analysis Steps of Pinch Analysis

76 76 79 79

Catalysts and Reaction Engineering Chemical Reactions Reaction Kinetics

92 93 95

Crude Distillation

96

Catalytic Cracking Introduction FCC Process Configuration Main Characteristics Equipment in FCC Feedstock & Yield Conclusion

99 99 100 100 101 101 101

Catalysis

103

CATALYSIS AND DISTILLATION

104

Distillation and Other Separation Processes

105

Distillation basics ATMOSPHERIC DISTILLATION Feeds and Products for Atmospheric Distillation Feed Preheat Exchanger Train Atmospheric Crude Fractionator Trends and Variations in Atmospheric Unit Design Phase behavior and vapour/liquid equilibrium Gas/Liquid separation

106 106 106 107 111 112 114 117

Industrial uses of Fractional Distillation 128 Trays: function, pressure drop, efficiency, flooding, operations, and damage 129 Tower Capacity: Equipment and Column Sizing Pressure Drop Column Height

133 133 136 137

Absorption & Adsorption Separation > Absorption

138 138


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Separation > Adsorption Solid Liquid Separation Introduction

MODULE 5 – PROCESS CONTROL & ECONOMICS

138 138 139

141

Process Control Basics Measured Variables

142 142

Process Control Systems Why Control? Control Objectives Techniques of Control

142 142 143 144

Process Economics Refinery Economics Crude Slate Product Slate 22.1. Control of distillation columns.

148 148 148 151 153


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FIGURE 1 FIRST LAW OF THERMODYNAMICS ......................................................................18 FIGURE 2 CONTROL VOLUME CONCEPTS...............................................................................19 FIGURE 3 OPEN SYSTEM CONTROL VOLUMES .....................................................................20 FIGURE 4 MULTIPLE CONTROL VOLUMES IN SAME SYSTEM..........................................21 FIGURE 5 STATIC HEAD................................................................................................................25 FIGURE 6 HEAD EXAMPLE ............................................................................................................27 FIGURE 7 HEAD EXAMPLE 2 ........................................................................................................29 FIGURE 8 SUCTION HEAD EXAMPLE ........................................................................................31 FIGURE 9 THREE DIFFERENT TYPES OF RESTRICTIONS COMMONLY ARE USED TO CONVERT FLOW RATE TO A PRESSURE DIFFERENCE, P1 - P2.............................37 FIGURE 10 MULTIPHASE FLOW ..................................................................................................38 FIGURE 11 THREE DIFFERENT TYPES OF OBSTRUCTION FLOW METERS .................40 FIGURE 12 PIPE FRICTION HEAD LOSS NOMOGRAPH ......................................................43 FIGURE 13 FRICTION LOSS FOR FITTINGS...........................................................................44 FIGURE 14 CENTRIFUGAL PUMP PERFORMANCE CURVE .................................................49 FIGURE 15 PUMP CURVES SHOWING SPEED & DIAMETER.............................................50 FIGURE 16 COMPRESSOR TYPES...............................................................................................52 FIGURE 17 COMPRESSOR SELECTION NOMOGRAPH ........................................................53 FIGURE 18 CENTRIFUGAL COMPRESSOR FLOW RANGE ..................................................55 FIGURE 19 COMPRESSOR CURVES...........................................................................................55 FIGURE 20 TUBE & SHELL HEAT EXCHANGER ........................................................................66 FIGURE 21 PLATE TYPE HEAT EXCHANGER ...........................................................................67 FIGURE 22 PARALLEL FLOW HEAT EXCHANGER....................................................................67 FIGURE 23 COUNTER FLOW HEAT EXCHANGER ....................................................................68 FIGURE 24 CROSS FLOW HEAT EXCHANGER..........................................................................68 FIGURE 25 A STEAM REBOILER .................................................................................................69 FIGURE 26 REBOILER SCHEMATIC ...........................................................................................69 FIGURE 27 CONDENSER ...............................................................................................................70 FIGURE 28 CHANGE OF SECTION - CHANGE IN PRESSURE ...........................................72 FIGURE 29 A SIMPLE FLOW SCHEME WITH T-H PROFILE ...............................................77 FIGURE 30 IMPROVED FLOW SCHEME WITH T-H PROFILE.............................................77 FIGURE 31 GRAPHIC REPRESENTATION OF TRADITIONAL AND PINCH DESIGN APPROACHES ...........................................................................................................................78 FIGURE 32 STEPS OF PINCH ANALYSIS..................................................................................80 FIGURE 33 HEAT TRANSFER EQUATION.................................................................................82 FIGURE 34 TEMPERATURE-ENTHALPY RELATIONS USED TO CONSTRUCT COMPOSITE CURVES ............................................................................................................83 FIGURE 35 COMBINED COMPOSITE CURVES .......................................................................84 FIGURE 36 GRAND COMPOSITE CURVE..................................................................................85 FIGURE 37 HEN AREA MIN ESTIMATION FROM COMPOSITE CURVES........................86 FIGURE 38 ENERGY-CAPITAL COST TRADE OFF (OPTIMUM DTMIN) ..........................87 FIGURE 39 TYPICAL GRID DIAGRAM .......................................................................................89 FIGURE 40 FLUID CATALYTIC CRACKING ............................................................................100 FIGURE 41 EARLY BATCH FRACTIONATION ........................................................................106 FIGURE 42 DESALTING - SINGLE STAGE.............................................................................108 FIGURE 43 DESALTING - 2 STAGE .........................................................................................108 FIGURE 44 CRUDE UNIT FURNACE .........................................................................................109 FIGURE 45 TEMPERATURE-COMPOSITION DIAGRAM FOR AMMONIA-BUTANE AT 20.7 BAR..................................................................................................................................115 FIGURE 46 T-X-Y DIAGRAM FOR AMMONIA-BUTANE AT 20.7 BAR ...........................116 FIGURE 47 T-X-Y DIAGRAM FOR AMMONIA-BUTANE AT 4, 10, AND 20.7 BAR ....116 FIGURE 48 PARTICLE DIAMETERS OF TYPICAL CONTAMINANTS..............................118 FIGURE 49 COALESCER CUT-AWAY VIEW ...........................................................................120 FIGURE 50 AEROSOL SIZES ......................................................................................................121 FIGURE 51 COALESCER EFFICIENCY CHANGE VS. GAS FLOW RATE ........................122 FIGURE 52 LIQUID AEROSOL SEPARATION EFFICIENCY TEST SCHEMATIC ..........123


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FIGURE 53 EFFECT OF CHEMICAL TREATMENT ON COALESCER PERFORMANCE.125 FIGURE 54 SCHEMATIC OF PALL LG COALESCER TEST STAND ..................................126 FIGURE 55 FIELD TEST RESULTS OF GAS STREAMS IN REFINERIES AND GAS PROCESSING PLANTS.........................................................................................................127 FIGURE 56 TYPICAL DISTILLATION TOWERS IN OIL REFINERIES.............................128 FIGURE 57 VALVE TRAYS (PHOTOS COURTESY OF PAUL PHILLIPS).........................129 FIGURE 58 VAPOUR & LIQUID FLOW ACROSS COLUMN/TRAY .................................130 FIGURE 59 LIQUID DISTRIBUTORS - GRAVITY (LEFT), SPRAY (RIGHT)(PHOTOS COURTESY OF PAUL PHILLIPS).......................................................................................131 FIGURE 60 TRAY PACKINGS ......................................................................................................132 FIGURE 61 STRUCTURED PACKING (PHOTO COURTESY OF PAUL PHILLIPS)........132 FIGURE 62 TYPICAL GRAVITY SEPARATION SYSTEM ......................................................139 FIGURE 63 FEEDBACK CONTROL LOOP ................................................................................145 FIGURE 64 LARGE MAGNITUDE DISTURBANCE.................................................................146 FIGURE 65 TIME DELAY ..............................................................................................................147 FIGURE 66 DISTILLATION COLUMN WITH SIX SINGLE-LOOP CONTROL SYSTEMS. ....................................................................................................................................................155 FIGURE 67 DISTILLATION COLUMN WITH SINGLE-LOOP AND CASCADE CONTROL SYSTEMS .................................................................................................................................156

TABLES TABLE TABLE TABLE TABLE

1THERMAL CONDUCTIVITY PROPERTIES 2 TYPICAL STREAM DATA 3 TYPES OF LIQUID/GAS SEPARATORS 4 COMPARISON OF THE DOP AND LASE


60 81 121 123

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Consultants Profile

•

•

•

Ronald Frend B.Sc. M.Vib.Inst M.ThermographicInst. o Shell Tankers (UK) Ltd o 1970 – 1984 o Marine Engineer Certified Chief Engineer Petroleum Development (Oman) o 1984 – 1989 o Rotating Equipment Specialist – Vibration Analysis o Head of Maintenance Planning o Head of Surface Maintenance (North Oman) Private Consultant o 1989 – present o Petro-Chemical, o Manufacturing, o Shipping, o Process


7

Fundamentals And Hydraulics Basics • • •

Process equipment and flow diagrams P&IDs Mass and energy balances


8

PFD - Process Flow Diagram

The Process Flow Diagram - PFD, a schematic illustration of the system A Process Flow Diagram - PFD - (or System Flow Diagram - SFD) shows the relationships between the major components in the system. PFD also tabulate process design values for the components in different operating modes, typical minimum, normal and maximum. A PFD does not show minor components, piping systems, piping ratings and designations. A PFD should include: • • • • • • •

Process Piping Major equipment symbols, names and identification numbers Control, valves and valves that affect operation of the system Interconnection with other systems Major bypass and recirculation lines System ratings and operational values as minimum, normal and maximum flow, temperature and pressure Composition of fluids

This figure depicts a small and simplified PFD:


9

System Flow Diagrams should not include: • • • • • • • •

pipe class pipe line numbers minor bypass lines isolation and shutoff valves maintenance vents and drains relief and safety valve code class information seismic class information


10

P&ID - Piping and Instrumentation Diagram A Piping and Instrumentation Diagram - P&ID, is a schematic illustration of functional relationship of piping, instrumentation and system equipment components P&ID shows all of piping including the physical sequence of branches, reducers, valves, equipment, instrumentation and control interlocks. The P&ID are used to operate the process system. A P&ID should include: • • • • • • • • • • • • • • • • •

Instrumentation and designations Mechanical equipment with names and numbers All valves and their identifications Process piping, sizes and identification Miscellaneous - vents, drains, special fittings, sampling lines, reducers, increasers and swagers Permanent start-up and flush lines Flow directions Interconnections references Control inputs and outputs, interlocks Interfaces for class changes Seismic category Quality level Annunciation inputs Computer control system input Vendor and contractor interfaces Identification of components and subsystems delivered by others Intended physical sequence of the equipment

This figure depicts a very small and simplified P&ID:


11

A P&ID should not include: • • • • • • • •

Instrument root valves control relays manual switches equipment rating or capacity primary instrument tubing and valves pressure temperature and flow data elbow, tees and similar standard fittings extensive explanatory notes


12

P&ID / PFD Symbols

General Instrument or Function Symbols -1


13

Instrument or Function Symbols - 2


14

General Instrument or Function Symbols 3


15

General Instrument or Function Symbols – 4


16

FIRST LAW OF THERMODYNAMICS

Thermodynamics The First Law of Thermodynamics states: Energy can neither be created nor destroyed, only altered in form. For any system, energy transfer is associated with mass and energy crossing the control boundary, external work and/or heat crossing the boundary, and the change of stored energy within the control volume. The mass flow of fluid is associated with the kinetic, potential, internal, and "flow" energies that affect the overall energy balance of the system. The exchange of external work and/or heat complete the energy balance. The First Law of Thermodynamics is referred to as the Conservation of Energy principle, meaning that energy can neither be created nor destroyed, but rather transformed into various forms as the fluid within the control volume is being studied. The energy balance spoken of here is maintained within the system being studied. The system is a region in space (control volume) through which the fluid passes. The various energies associated with the fluid are then observed as they cross the boundaries of the system and the balance is made. A system may be one of three types: isolated, closed, or open. The open system, the most general of the three, indicates that mass, heat, and external work are allowed to cross the control boundary. The balance is expressed in words as: all energies into the system are equal to all energies leaving the system plus the change in storage of energies within the system. Remember that energy in thermodynamic systems is composed of • • • • •

kinetic energy (KE), potential energy (PE), internal energy (U), and flow energy (PL); as well as heat and work processes.

For most industrial plant applications that most frequently use cycles, there is no change in storage (i.e. heat exchangers do not swell while in operation). In equation form, the balance appears as indicated in the heat balance figure below:


17

Figure 1 First Law of Thermodynamics


18

Heat and/or work can be directed into or out of the control volume. But, for convenience and as a standard convention, the net energy exchange is presented here with the net heat exchange assumed to be into the system and the net work assumed to be out of the system. If no mass crosses the boundary, but work and/or heat do, then the system is referred to as a "closed" system. If mass, work and heat do not cross the boundary (that is, the only energy exchanges taking place are within the system), then the system is referred to as an isolated system. Isolated and closed systems are nothing more than specialized cases of the open system. In this text, the open system approach to the First Law of Thermodynamics will be emphasized because it is more general. Also, almost all practical applications of the first law require an open system analysis. An understanding of the control volume concept is essential in analyzing a thermodynamic problem or constructing an energy balance. Two basic approaches exist in studying Thermodynamics: •

the control mass approach and the

•

control volume approach.

The former is referred to as the LeGrange approach and the latter as the Eulerian approach. In the control mass concept, a "clump" of fluid is studied with its associated energies. The analyzer "rides" with the clump wherever it goes, keeping a balance of all energies affecting the clump.

Figure 2 Control volume concepts


19

The control volume approach is one in which a fixed region in space is established with specified control boundaries, as shown above. The energies that cross the boundary of this control volume, including those with the mass crossing the boundary, are then studied and the balance performed. The control volume approach is usually used today in analyzing thermodynamic systems. It is more convenient and requires much less work in keeping track of the energy balances.

Figure 3 Open system control volumes


20

Figure 4 Multiple Control Volumes in Same System

The forms of energy that may cross the control volume boundary include those associated with the mass (m) crossing the boundary. Mass in motion has potential (PE), kinetic (KE), and internal energy (U). In addition, since the flow is normally supplied with some driving power (a pump for example), there is another form of energy associated with the fluid caused by its pressure. This form of energy is referred to as flow energy (Pn-work). The thermodynamic terms thus representing the various forms of energy crossing the control boundary with the mass are given as m (u + Pn + ke + pe). In open system analysis, the u and Pn terms occur so frequently that another property, enthalpy, has been defined as h = u + Pn. This results in the above expression being written as m (h + ke + pe). In addition to the mass and its energies, externally applied work (W), usually designated as shaft work, is another form of energy that may cross the system boundary. In order to complete and satisfy the conservation of energy relationship, energy that is caused by neither mass nor shaft work is classified as heat energy (Q). Then we can describe the relationship in equation form as follows.

m(hin+pein-kein) Q = m(hout-peout+keout) +W


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where:


22

Example 1 illustrates the use of the control volume concept while solving a first law problem involving most of the energy terms mentioned previously.

Example 1: Open System Control Volume The enthalpies of steam entering and leaving a steam turbine are 1349 Btu/lbm and 1100 Btu/lbm, respectively. The estimated heat loss is 5 Btu/lbm of steam. The flow enters the turbine at 164 ft/sec at a point 6.5 ft above the discharge and leaves the turbine at 262 ft/sec. Determine the work of the turbine Where”


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Hydraulics & Fluid Flow Hydraulics and Fluid Flow • Pressure and head • Bernoulli’s theorem and its field applications • Flow of liquids • Reynolds number and pressure drop in pipes • Two-phase and multi-phase flow • Pumps and compressors • Mixing and mixers


24

Pressure & Head It turns out that head is a very convenient term in the pumping business. Pressure is not as convenient a term because the amount of pressure that the pump will deliver depends upon the weight (specific gravity) of the liquid being pumped and the specific gravity changes with the fluid temperature and concentration. Each litre of liquid has weight, so we can easily calculate the kilograms per minute being pumped. Head or height is measure in meters so if we multiply these two together we get kilogram meters per minute which converts directly to work at the rate of 610 kgm/min = 1 kilowatt. If you are more comfortable with metric horsepower units, you should know that 735.5 watts makes one metric horsepower If you will refer to the Figure below you should get a clear picture of what is meant by static head. Please note that we always measure from the centreline of the pump to the highest liquid level

Figure 5 Static Head

To calculate head accurately we must calculate the total head on both the suction and discharge sides of the pump. In addition to the static head we will learn that there is a head caused by resistance in the piping, fittings and valves called friction head, and an additional head caused by any pressure that might be acting on the liquid in the tanks, including atmospheric pressure. This head is called "surface pressure head". Once we know all of these heads it gets simple. We subtract the suction head from the discharge head and the head remaining will be the amount of head that the pump must be able to generate at its rated flow. Here is how it looks in a formula:


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System head = total discharge head - total suction head or H = hd - hs The total discharge head is made from three separate heads: hd = hsd + hpd + hfd • • • •

hd = hsd = hpd = hfd =

total discharge head discharge static head discharge surface pressure head discharge friction head

The total suction head also consists of three separate heads hs = hss + hps - hfs • • • •

hs = hss = hps = hfs =

total suction head suction static head suction surface pressure head suction friction head

As we make these calculations you must be sure that all your calculations are made in either "meters of liquid gauge" or "meters of liquid absolute". In case you have forgotten "absolute" means that you have added atmospheric pressure (head) to the gauge reading. Normally head readings are made in gauge readings and we switch to the absolute readings only when we want to calculate the net positive suction head available (NPSHA) to find out if our pump is going to cavitate. We use the absolute term for these calculations because we are often calculating a vacuum or using negative numbers We will begin by making some actual calculations. You will not have to look up the friction numbers because I am going to give them to you, but you can find them in a number of publications and these charts: • •

Piping friction losses, metric, Valves and fittings losses, metric,

The next illustration (Figure #2) shows that the discharge head is still measured to the liquid level, but you will note that it is now below the maximum height of the piping. Although the pump must deliver enough head to get up to the maximum piping height it will not have to continue to deliver this head when the pump is running because of the "siphon effect". There is of course a maximum siphon effect. It is derived from the formula to convert pressure to head:


26

Since atmospheric pressure at seal level is one bar we get a maximum siphon distance of 10.2 meters

Figure 6 Head example

We will begin with the total suction head calculation •

•

The suction head is negative because the liquid level in the suction tank is below the centreline of the pump: o hss = -2 meters The suction tank is open so the suction surface pressure equals atmospheric pressure : o hps = 0 meters gauge

In these examples you will not be calculating the suction friction head. When you learn how you will find that there are two ways to do it •

•

You would look at the charts and add up the K factors for the various fittings and valves in the piping. You would then multiply these K factors by the velocity head that is shown for each of the pipe sizes and capacities. This final number would be added to the friction loss in the piping for the total friction head. Or, you can look at a chart that shows the equivalent length of pipe for each of the fittings and add this number to the length of the piping in the system to determine the total friction loss.

For this example, I will tell you the total friction head on the suction side of the pump is: hfs = 1.5 meters at rated flow •

The total suction head is going to be a gauge value because atmosphere was given as 0, o hs = hss + hps - hfs = - 2 + 0 - 1.5 = - 3.5 meters of liquid gauge at rated flow


27

•

The total discharge head calculation is similar o The static discharge head is: hsd = 40 meters o The discharge tank is also open to atmospheric pressure, so: hpd = 0 feet, gauge o I will give you the discharge friction head as: hfd = 7 meters at rated flow o The total discharge head is: hd = hsd + hpd + hfd = 40 + 0 +7 = 47 meters of liquid gauge at rated flow

The total system head calculation becomes: Head = hd - hs = 47 - (-3.5) = 50.5 meters of liquid at rated flow Our next example involves a few more calculations, but you should be able to handle them without any trouble. If we were pumping from a vented suction tank to an open tank at the end of the discharge piping we would not have to consider vacuum and absolute pressures. In this example we will be pumping from a vacuum receiver that is very similar to the hotwell we find in many condenser applications Again, to make the calculations you will need some pipe friction numbers that are available from charts: • •

Piping friction losses, metric, Valves and fittings losses, metric,

I will give you the friction numbers for the following examples. Specifications: • • • • • • • •

Transferring 300 m3/hr weak acid from the vacuum receiver to the storage tank Specific Gravity = 0.98 Viscosity = equal to water Piping = all 150 mm Schedule 40 steel pipe Discharge piping rises 15 meters vertically above the pump centreline and then runs 135 meters horizontally. There is one 90° elbow in this line Suction piping has 1.5 meters of pipe, one gate valve, and one 90° elbow, all of which are 150 mm in diameter. The minimum level in the vacuum receiver is 2 meters above the pump centreline. The pressure on top of the liquid in the vacuum receiver is 500 mm of mercury, vacuum.


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Figure 7 Head example 2

To calculate suction surface pressure use the following formula:

Now that you have all of the necessary information we will begin by dividing the system into two different sections using the pump as the dividing line. Total suction head calculation •

•

•

The suction side of the system shows a minimum static head of 2 meters above suction centreline. Therefore, the static suction head is o hss = 2 meters Using the first conversion formula, the suction surface pressure is

The suction friction head fs equals the sum of all the friction losses in the suction line. If you referenced the metric pipe friction loss table you would learn that the friction loss in 150 mm. pipe at 300 m3/hr is 9 meters per 100 meters of pipe.

In 1.5 meters of pipe, friction loss = 15/100 x 9 = 0.14 meters

Fitting

Equivalent length of straight pipe

150 mm normal bend elbow

3.4 meters

150 mm Gate valve

2.1 meters


29

In a real life pumping application there would be other valves and fittings that experience friction losses. You might find: • • • • • •

Check valves Foot valves Strainers Sudden enlargements Shut off valves Entrance and exit losses

The loss in the suction fittings becomes: In 5.5 meters of pipe friction loss = 55 / 100 x 9 = 0.50 meters The total friction loss on the suction side is: hfs = 0.14 + 0.50 = 0.64 meters at 300 m3/hr The total suction head then becomes: hs = hss + hps - hfs = 2 - 7.14 - 0.64 = - 5.78 meters gauge at 300 m3/hr Now we will look at the total discharge head calculation • • •

Static discharge head = hsd = 15 meters Discharge surface pressure = hpd = 0 meters gauge Discharge friction head = hfd = sum of the following losses :

Friction loss in 150 mm pipe at 300 m3/hr, from the chart is 9 meters per hundred feet of pipe. • •

In 150 meters of pipe the friction loss = 150/100 x 9 = 13.5 meters Friction loss in 150 mm. Elbow:= 3.4/100 x 9 = 0.31 meters 1

The discharge friction head is the sum of the above losses, that is: hfd = 13.5 + .31 = 13.81 meters at 300 m3/hr The total discharge head then becomes: hd = hsd + hpd + hfd = 15 + 0 + 13.81 = 28.81 meters at 300 m3/hr. Total system head calculation: H = hd - hs = 28.81 - (-5.78) = 34.59 meters at 300 m3/hr Our next example will be the same as the one we just finished except that there is an additional 3 meters of pipe and another 90° flanged elbow in the vertical leg. The total suction head will be the same as in the previous example. Take a look at the figure below


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Figure 8 Suction head example

Nothing has changed on the suction side of the pump so the total suction head will remain the same: hs = - 5.78 meters at 300 m3/hr Total discharge head calculation •

• • •

The static discharge head (hsd) will change from 15 meters to 12 meters since the highest liquid surface in the discharge is now only 12 meters above the pump centreline. This value is based on the assumption that the vertical leg in the discharge tank is full of liquid and that as this liquid falls it will tend to pull the liquid up and over the loop in the pipeline. This arrangement is called a siphon leg. The discharge surface pressure is unchanged: hpd = 0 meters The additional 3 meters of pipe and the additional elbow will increase the friction loss in the discharge pipe. In 3 meters of pipe the friction loss = 3 / 100 x 9 = 0.27 meters The friction loss in the additional elbow = 3.4 / 100 x 9 = 0.31 meters The friction head will then increase as follows: hfd = 0.27 + 0.31 = 0.58 at 300 m3/hr. The total discharge head becomes: hd = hsd + hpd + hfd = 12 + 13.81+ 0 + 0.58 = 26.39 meters at 300 m3/hr


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Total system head calculation Head = hd - hs = 26.39 - (-5.78) = 32.17 meters at 300 m3/hr.


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Bernoulli’s Theorem Bernoulli's theorem: in fluid dynamics, relation among the pressure, velocity, and elevation in a moving fluid (liquid or gas), the compressibility and viscosity (internal friction) of which are negligible and the flow of which is steady, or laminar. First derived (1738) by the Swiss mathematician Daniel Bernoulli, the theorem states, in effect, that the total mechanical energy of the flowing fluid, comprising the energy associated with fluid pressure, the gravitational potential energy of elevation, and the kinetic energy of fluid motion, remains constant. Bernoulli's theorem is the principle of energy conservation for ideal fluids in steady, or streamline, flow. Bernoulli's theorem implies, therefore, that if the fluid flows horizontally so that no change in gravitational potential energy occurs, then a decrease in fluid pressure is associated with an increase in fluid velocity. If the fluid is flowing through a horizontal pipe of varying cross-sectional area, for example, the fluid speeds up in constricted areas so that the pressure the fluid exerts is least where the cross section is smallest. This phenomenon is sometimes called the Venturi effect, after the Italian scientist G.B. Venturi (1746-1822), who first noted the effects of constricted channels on fluid flow. Bernoulli's theorem is the basis for many engineering applications, such as aircraftwing design. The air flowing over the upper curved surface of an aircraft wing moves faster than the air beneath the wing, so that the pressure underneath is greater than that on the top of the wing, causing lift.

How pressure and velocity interact static pressure + dynamic pressure = total pressure = constant static pressure + 1/2 x density x velocity2 = total pressure = constant

General Concept: The Bernoulli effect is simply a result of the conservation of energy. The work done on a fluid (a fluid is a liquid or a gas), the pressure times the volume, is equal to the change in kinetic energy of the fluid. General Facts: Where there is slow flow in a fluid, you will find increased pressure. Where there is increased flow in a fluid, you will find decreased pressure. In a real flow, friction plays a large role - a lot of times you must have a large pressure drop (decrease in pressure) just to overcome friction. This is the case in your house. Most water pipes have small diameters (large friction), hence the need for "water pressure" - it is the energy from that pressure drop that goes to friction. In a real flow i.e. around an immersed body, friction plays a large role – most of the time when the ship is in service you have a large pressure drop (decrease in pressure) just to overcome friction. For example, if you have a water pipe with a small diameter (large friction), hence the need for "water pressure" – it is the energy from that pressure drop that goes to friction.

Example: the showerhead


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A showerhead (if you have a fancy one) has a number of different operation modes. If you go for the "massage" mode, you are moving a little water fast. For the "lite shower," you are moving a lot of water slowly. It takes the same amount of energy to move a little water fast as it does to move a lot of water slowly. This is the amount of energy you have due to your "water pressure".

Example When a liquid runs freely through a pipe of a constant area (B), to which three ascension pipes (D,E,F) are connected, the static pressure will decrease along the dashed line towards the outlet (Fig.1), The pressure decreases as result of friction loss in the horizontal pipe.

Fig. 1 In (Fig.2) the area has been changed in two places, with a thinner pipe at section (G) and a thicker pipe at section (H). The following occurs: Section (G) The resultant constriction causes the liquid to move at a higher speed, increasing the dynamic pressure, with the result that the static pressure in pipe (D) falls below the dashed line. Section (H) In section (H), which has a much larger area, the static pressure rises above the dashed line, the speed of the liquid having decreased due to the larger area, with the result that the dynamic pressure will be decreased.

Fig. 2


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Liquid Flow Flow Units The units used to describe the flow measured can be of several types, depending on how the specific process needs the information. The most common descriptions are the following: 1. Volume flow rate Expressed as a volume delivered per unit time. Typical units are gallons/min, m3/hr, ft3hr. 2. 2. Flow velocity Expressed as the distance the liquid travels in the carrier per unit time. Typical units are m/min, ft/min. This is related to the volume flow rate by

where V = flow velocity •

Q = volume flow rate

•

A = cross-sectional area of flow carrier (pipe, and so on)

3. Mass or weight flow rate Expressed as mass or weight flowing per unit time. Typical units are kg/hr, Ib/hr. This is related to the volume flow rate by F = pQ (5.36) where F = mass or weight flow rate •

p = mass density or weight density

•

Q = volume flow rate

EXAMPLE Water is pumped through a 1.5-in diameter pipe with a flow velocity of 2.5 ft per second. Find the volume flow rate and weight flow rate. The weight density is 62.4 Ib/ft3. Solution The flow velocity is given as 2.5 ft/s, so the volume flow rate can be found from Equation (5.35), Q = VA. The area is given by A = pd24 where the diameter d = (1.5 in)(l/12 ft/in) = 0.125 ft so that A = (3.14)(0.125)2/4 = 1.0122 ft2. Then, the volume flow rate is •

Q = (2.5ft/s)(0.0122ft2)(60s/min)

•

Q = 1.8 ft3/min

The weight flow rate is found from Equation (5.36): F = (62.4 lb/ft3)(1.8 ft3/min) F = 112 lb/min


35

Restriction Flow Sensors One of the most common methods of measuring the flow of liquids in pipes is by introducing a restriction in the pipe and measuring the pressure drop that results across the restriction. When such a restriction is placed in the pipe, the velocity of the fluid through the restriction increases, and the pressure in the restriction decreases. We find that there is a relationship between the pressure drop and the rate of flow such that as the flow increases, the pressure drops. In particular, one can find an equation of the form

where Q = volume flow rate K = a constant for the pipe and liquid type Dp = drop in pressure across the restriction The constant K depends on many factors, including the type of liquid, size of pipe, velocity of flow, temperature, and so on. The type of restriction employed also will change the value of the constant used in this equation. The flow rate is linearly dependent not on the pressure drop, but on the square root. Thus, if the pressure drop in a pipe increased by a factor of 2 when the flow rate was increased, the flow rate will have increased only by a favor of 1.4 (the square root of 2). Certain standard types of restrictions are employed in exploiting the pressure-drop method of measuring flow. Figure 5.37 shows the three most common methods. It is interesting to note that having converted flow information to pressure, we now employ one of the methods of measuring pressure, often by conversion to displacement, which is measured by a displacement sensor before finally getting a signal that will be used in the processcontrol loop. The most common method of measuring the pressure drop is to use a differential pressure sensor similar to that shown in Figure 5.33. These are often described by the name DP cell.


36

Figure 9 Three different types of restrictions commonly are used to convert flow rate to a pressure difference, p1 - p2

EXAMPLE Flow is to be controlled from 20 to 150 gal/min. The flow is measured using an orifice plate system such as that shown in Figure 5.37c. The orifice plate is described by Equation (5.37) with K = 119.5 (gal/minypsi1/2. A bellows measures the pressure with an LVDT so that the output is 1.8 V/psi. Find the range of voltages that result from the given flow range. Solution From Equation (5.37), we find the pressures that result from the given flow: Dp = (Q/K)2 For 20 gal/min Dp = (20/119.5)2 = 0.0280 psi and for 150 gal/min Dp = (150/119.5)2 = 1.5756 psi Because there are 1.8 V/psi, the voltage range is easily found. For 20 gal/min, V = 0.0280(1.8) = 0.0504 V For 150 gal/min, V = 1.5756(1.8) = 2.836 V


37

Two Phase & Multiphase Flow Multiphase flow means that a mixture of two or more substances (gases, liquids or solid particles) flow together without being dissolved in each other. Such flows are ubiquitous - e.g. in the atmosphere, in the food industry, in cooling systems, in the process industry and in petroleum reservoirs.

Figure 10 Multiphase flow

Multiphase flow is simultaneous flow of gas, liquid and/or solid phases in different combinations. Multiphase flow, particularly two-phase flows, are probably the most common flow cases in nature, examples being the flow of blood in our body and drift of clouds in the sky. The bubbles rising in a glass of soda provide another good example of multiphase flow. Multiphase flows are also important in most industrial applications, such as energy conversion, paper manufacturing, food manufacturing, bio-technological and medical applications. In industry, controlling multiphase flow is essential for efficiency, quality and profitability and often has a decisive effect on environmental aspects. In oil and gas production, multiphase flow often occurs in wells and pipelines because the wells produce gas and oil simultaneously. This is called two-phase flow. In addition to gas and oil, water is also often produced at the same time. This is called three-phase flow. In the North Sea, the oil companies previously often built large production platforms standing on the sea floor, equipped with process facilities separating gas. oil and water. The gas was sent to market in one pipeline, while the oil was shipped directly or sent to shore in another pipeline. Today this is usually too expensive. Instead the operators often choose subsea developments where the untreated well stream is sent directly from a subsea template to an existing platform or to shore in one multiphase pipeline. In cases where a subsea development with multiphase transport is feasible, billions may be saved by dispensing with costly platforms. Multiphase transport implies many new challenges. Under unfavourable conditions, oil and water may be expelled from the pipeline in large batches (slug flow) which can disturb the receiving facilities. Oil and water may form thick emulsions which give high pressure losses and reduced production. Wax and hydrates (an ice-like substance) may precipitate and block the pipe. Unfavourable water chemistry may lead to fatal corrosion attacks piercing the pipe. Before commissioning a field, it is important to be able to predict possible production problems and to predict flow patterns and pressure losses as accurately as possible so that pipelines and process plants may be designed optimally.


38

Obstruction Flow Sensor Another type of flow sensor operates by the effect of flow on an obstruction placed in the flow stream. In a rotameter, the obstruction is a float that rises in a vertical tapered column. The lifting force and thus the distance to which the float rises in the column is proportional to the flow rate. The lifting force is produced by the differential pressure that exists across the float, because it is a restriction in the flow. This type of sensor is used for both liquids and gases. A moving vane flow meter has a vane target immersed in the flow region, which is rotated out of the flow as the flow velocity increases. The angle of the vane is a measure of the flow rate. If the rotating vane shaft is attached to an angle-measuring sensor, the flow rate can be measured for use in a process-control application. A turbine type of flow meter is composed of a freely spinning turbine blade assembly in the flow path. The rate of rotation of the turbine is proportional to the flow rate. If the turbine is attached to a tachometer, a convenient electrical signal can be produced. In all of these methods of flow measurement, it is necessary to present a substantial obstruction into the flow path to measure the flow. For this reason, these devices are used only when an obstruction does not cause any unwanted reaction on the flow system.

Magnetic Flow Meter It can be shown that if charged particles move across a magnetic field, a potential is established across the flow, perpendicular to the magnetic field. Thus, if the flowing liquid is also a conductor (even if not necessarily a good conductor) of electricity, the flow can be measured by allowing the liquid to flow through a magnetic field and measuring the transverse potential produced. The pipe section in which this measurement is made must be insulated and a nonconductor itself, or the potential produced will be cancelled by currents in the pipe. A diagram of this type of flow meter is presented in Figure 5.39. This type of sensor produces an electrical signal directly and is convenient for process-control applications involving conducting fluid flow.


39

Figure 11 Three different types of obstruction flow meters


40

Reynolds Number

Flow performance can be affected by a dimensionless unit called the Reynolds Number. It is defined as the ratio of the liquid's inertial forces to its drag forces.

Laminar and turbulent flow are most common in flow regimes or in liquid flow measurement operations but there is also transitional flow. If we want to calculate the Reynolds number , we can use the following equation R = 3160 x Q x Gt D x µ where:


41

R = Reynolds number Q = liquid's flow rate, gpm Gt = liquid's specific gravity D = inside pipe diameter, in. µ = liquid's viscosity, cp • • •

When the Reynolds number is less than 2000, flow will be described as laminar When the Reynolds number is greater than 4000, flow will be described as turbulent When the Reynolds number is in the range of 2000 to 4000 the flow is considered transitional.

Viscosity can be a major factor that affects the value of the Reynolds number. For e.g You may find highly viscous hydraulic oils may exhibit laminar flow in most conditions while things like water will be turbulent


42

SOME NOTES FOR THE METRIC PIPE FRICTION CHART SHOWN BELOW • • • • • •

The chart is calculated for fresh water at 15°C. Use actual bores rather than nominal pipe size. For stainless steel pipe multiply the numbers by 1.1. For steel pipe multiply the numbers by 1.3 For cast iron pipe multiply the numbers by 1.7 The losses are calculated for a fluid viscosity similar to fresh water.

Figure 12 Pipe friction head loss nomograph


43

FRICTION LOSS FOR METRIC PIPE, VALVES AND FITTINGS

Figure 13 Friction Loss for fittings


44

Pumps & Compressors Centrifugal pump designs The overwhelming majority of contractor pumps use centrifugal force to move water. Centrifugal force is defined as the action that causes something, in this case water, to move away from its center of rotation. All centrifugal pumps use an impeller and volute to create the partial vacuum and discharge pressure necessary to move water through the casing. The impeller and volute form the heart of the pump and help determine its flow, pressure and solid handling capability. An impeller is a rotating disk with a set of vanes coupled to the engine/motor shaft that produces centrifugal force within the pump casing. A volute is the stationary housing (in which the impeller rotates) that collects, discharges and recirculates water entering the pump. A diffuser is used on high pressure pumps and is similar to a volute but more compact in design. Many types of material can be used in their manufacture but cast iron is most commonly used for construction applications. In order for a centrifugal pump, or self priming, pump to attain its initial prime the casing must first be manually primed or filled with water. Afterwards, unless it is run dry or drained, a sufficient amount of water should remain in the pump to ensure quick priming the next time it is needed.

As the impeller churns the water (see figure above), it purges air from the casing creating an area of low pressure, or partial vacuum, at the eye (centre) of the impeller. The weight of the atmosphere on the external body of water pushes water rapidly through the hose and pump casing toward the eye of the impeller. Centrifugal force created by the rotating impeller pushes water away from the eye, where pressure is lowest, to the vane tips where the pressure is highest. The velocity of the rotating vanes pressurizes the water forced through the volute and discharges it from the pump. Water passing through the pump brings with it solids and other abrasive material that will gradually wear down the impeller or volute. This wear can increase the distance between the impeller and the volute resulting in decreased flows, decreased heads and longer priming times. Periodic inspection and maintenance is necessary to keep pumps running like new.


45

Another key component of the pump is its mechanical seal. This spring loaded component consists of two faces, one stationary and another rotating, and is located on the engine shaft between the impeller and the rear casing (see figure below). It is designed to prevent water from seeping into and damaging the engine. Pumps designed for work in harsh environments require a seal that is more abrasion resistant than pumps designed for regular household use.

Typically seals are cooled by water as it passes through the pump. If the pump is dry or has insufficient water for priming it could damage the mechanical seal. Oillubricated an occasionally grease-lubricated seals are available on some pumps that provide positive lubrication in the event that the pump is run without water. The seal is a common wear part that should be periodically inspected.


46

Pump Affinity Laws There are occasions when you might want to permanently change the amount of fluid you are pumping, or change the discharge head of a centrifugal pump. There are four ways you could do this: •

Regulate the discharge of the pump.

•

Change the speed of the pump.

•

Change the diameter of the impeller.

•

Buy a new pump

Of the four methods the middle two are the only sensible ones. In the following paragraphs we will learn what happens when we change either the pump speed or impeller diameter and as you would guess other characteristics of the pump are going to change along with speed or diameter. To determine what is going to happen you begin by taking the new speed or impeller diameter and divide it by the old speed or impeller diameter. Since changing either one will have approximately the same affect I will be referring to only the speed in this part of the discussion. As an example:

The capacity, or amount of fluid you are pumping, varies directly with this number. Example: 100 Gallons per minute x 2 = 200 Gallons per minute Or in metric, 50 Cubic meters per hour x 0,5 = 25 Cubic meters per hour The head varies by the square of the number. Example : a 50 foot head x 4 (22) = 200 foot head Or in metric, a 20 meter head x 0,25 ( 0,52) = 5 meter head The horsepower required changes by the cube of the number. Example : a 9 Horsepower motor was required to drive the pump at 1750 rpm.. How much is required now that you are going to 3500 rpm? We would get: 9 x 8 (23) = 72 Horse power is now required. Likewise if a 12 kilowatt motor were required at 3000 rpm. and you decreased the speed to 1500 the new kilowatts required would be: 12 x 0,125 (0.53) = 1,5 kilowatts required for the lower rpm.


47

The following relationships are not exact, but they give you an idea of how speed and impeller diameter affects other pump functions. The net positive suction head required by the pump manufacturer (npshr) varies by the square of the number. Example : A 3 meter NPSHR x 4 (22) = 12 meter N.P.S.H.R. Or: 10 foot NPSHR x 0.25 ( 0.52) = 2.5 foot N.P.S.H.R. The amount of shaft run out ( deflection) varies by the square of the number As an example : If you put a dial indicator on the shaft and noticed that the total run out at 1750 rpm. was 0.005 inches then at 3500 rpm the run out would be 0.005" x 4 (22), or 0.020 inches. Likewise if you had 0,07 mm. run out at 2900 rpm. and you slowed that shaft down to 1450 rpm the run out would decrease to 0,07 mm x 0,25 ( 0.52) or 0,018 mm. The amount of friction loss in the piping varies by about 90% of the square of the number. Fittings and accessories varies by almost the square of the number. As an example : If the system head loss was calculated or measured at 65 meters at 1450 rpm., the loss at 2900 rpm. would be : 65 meters x 4 (22) = 260 x 0.9 = 234 Meters If you had a 195 foot loss at 3500 rpm the loss at 1750 rpm. would be : 195 x 0.25 (0.52) = 48.75 0.9 = 43.87 feet of head loss. The wear rate of the components varies by the cube also Example : At 1750 rpm. the impeller material is wearing at the rate of 0.020 inches per month. At 3500 rpm the rate would increase to: 0.020 " x 8 (23) or 0.160 inches per month. Likewise a decrease in speed would decrease the wear rate eight times as much. We started this section by stating that a change in impeller speed or a change in impeller diameter has approximately the same affect. This is true only if you decrease the impeller diameter to a maximum of 10% . As you cut down the impeller diameter the housing is not coming down in size so the affinity laws do not remain accurate below this 10% maximum number. The affinity laws remain accurate for speed changes and this is important to remember when we convert from jam packing to a balanced mechanical seal. We sometimes experience an increase in motor speed rather than a drop in amperage during these conversions and the affinity laws will help you to predict the final outcome of the change.


48

Performance Curves

Figure 14 Centrifugal Pump Performance Curve

Please look at the above illustration. You will note that we have plotted the head of the pump against its capacity. The head of a pump is read in feet or meters. The capacity units will be either gallons per minute, liters per minute, or cubic meters per hour. According to the above illustration this pump will pump a 40 capacity to about a 110 head, or a 70 capacity to approximately a 85 head (you can substitute either metric or imperial units as you see fit) The maximum head of this pump is 115 units. This is called the maximum shutoff head of the pump. Also note that the best efficiency point (BEP) of this impeller is between 80% and 85% of the shutoff head. This 80% to 85% is typical of centrifugal pumps, but if you want to know the exact best efficiency point you must refer to the manufacturers pump curve. Ideally a pump would run at its best efficiency point all of the time, but we seldom hit ideal conditions. As you move away from the BEP the shaft will deflect and the pump will experience some vibration. You will have to check with your pump manufacturer to see how far you can safely deviate from the BEP (a maximum of 10% either side is typical) Now look at the following illustration:


49

Figure 15 Pump curves showing speed & diameter

Note that we have added some additional curves to the original illustration. These curves show what happens when you change the diameter of the impeller. Impeller diameter is measured in either inches or millimeters. If we wanted to pump at the best efficiency point with a 11.5 impeller we would have to pump a capacity of 50 to a 75 head. The bottom half of the illustration shows the power consumption at various capacities and impeller diameters. We have labeled the power consumption horsepower, but in the metric system it would be called kilowatts Each of the lines represents an impeller diameter. The top line would be for the 13 impeller the second for the 12.5 etc. If we were pumping a capacity of 70 with a 13 impeller it would take about 35 horsepower. A capacity of 60 with the 12 impeller would take about 20 horsepower. Most pump curves would show you the percent of efficiency at the best efficiency point . The number varies with impeller design and numbers from 60% to 80% are normal. When you will look at an actual pump curve you should have no trouble reading the various heads and corresponding capacities for the different size impellers. You will note however, that the curve will usually show an additional piece of information and that is NPSHR which stands for net positive suction head required to prevent the pump from cavitating. Depending upon the pump curve you might find a 10 foot (3.0 meter) NPSH required head at a capacity of 480 Gallons per minute (110 cubic meters per hour) if you were using a 13 inch (330 mm.) diameter impeller.


50

You should keep in mind that the manufacture assumed you were pumping 20°C (68°F) fresh water and the N.P.S.H. Required was tested using this assumption. If you are pumping water at a different temperature or if you are pumping a different fluid, you are going to have to add the vapour pressure of that product to the N.P.S.H. Required. The rule is that Net Positive Suction Head Available minus the Vapour Pressure of the product you are pumping (converted to head) must be equal to or greater than Net Positive Suction Head Required by the manufacturer.

Suppose we wanted to pump some liquid Butane at 32 degrees Fahrenheit (0 degrees Centigrade) with this pump. If we look at the curve for Butane on a vapour pressure chart similar to the one shown in the charts and graphs section of this web site you will note that Butane at 32°F needs at least 15 psi (1,0 Bar) to stay in a liquid state. To convert this pressure to head we use the standard formula :

In other words Butane at this temperature would not vapourize as long as I had the above absolute heads available at the suction side of the pump.


51

Compressors and Expanders Depending on application, compressors are manufactured as positive displacement, dynamic, or thermal type. Positive displacement types fall in two basic categories: reciprocating and rotary

Figure 16 Compressor Types

The reciprocating compressor consists of one or more cylinders each with a piston or plunger that moves back and forte displacing a positive volume with each stroke. The diaphragm compressor uses a hydraulically pulsed flexible diaphragm to displace the gas. Rotary compressors cover lobe-type, screw-type, vane-type, and liquid ring type, each having a casing with one or more rotating elements that either mesh with each other such as lobes or screws, or that displaces a fixed volume with each rotation. The dynamic types include radial-flow (centrifugal), axial flow and mixed flow machines. They are rotary continuous-flow compressors in which the rotating element (impeller or bladed rotor s accelerates the gas as it passes through the element, converting the velocity head into static pressure, partially in the rotating element and partially in stationary diffusers or blades. Ejectors are "thermal" compressors that use a high velocity gas or steam jet to entrain the inflowing gas then convert the velocity of the mixture to pressure in a diffuser. The Figure below covers normal range of operation for compressors of the commercially available types Fig (4) summarizes the difference between reciprocating and centrifugal compressors.


52

Figure 17 Compressor selection nomograph


53

CENTRIFUGAL COMPRESSORS This section is intended to supply information sufficiently accurate lo determine whether a centrifugal compressor should be considered for a specific job. The secondary objective is to present information for evaluating compressor performance. The compressor selection nomograph gives an approximate idea of the flow range that a centrifugal compressor will handle. A multi-wheel (multistage) centrifugal compressor is normally considered for inlet volumes between 500 and 200,000 inlet acfm. A singlewheel (single stage) compressor would normally have application between 100 and 150,000 inlet acfm. A multi-wheel compressor can be thought of as a series of single wheel compressors contained in a single casing. Historically, centrifugal compressors used to operate at speeds of 3,000 rpm no higher, a limiting factor being impeller stress considerations as well as velocity limitation of 0.8 to 0.85 Mach number at the impeller tip and eye. Recent advances in machine design have resulted in production of some units running at speeds in excess of 40,000. Centrifugal compressors are usually driven by electric motors, steam or gas turbines (with or without speed-increasing gears), or turbo-expanders. There is an overlap of centrifugal and reciprocating compressors on the low end of the flow range. At the higher end of the flow range an overlap with the axial compressor exists. The extent of this overlap depends on a number of things. Before a technical decision could be reached as to the type of compressor that would be installed, the service, operational requirements and economics would have to be considered.


54

Figure 18 Centrifugal Compressor flow range

The operating characteristics must be determined before an evaluation of compressor suitability for the application can be made. Fig. (29) gives a rough comparison of the characteristics of the axial, centrifugal, and reciprocating compressor. The centrifugal compressor approximates the constant head - variable volume machine, while the reciprocating is a constant volume-variable head machine. The axial compressor, which is a low head, high flow machine, falls somewhere in between. A compressor is a part of the system, and its performance is dictated by the system resistance. The desired system capability or objective must be determined before a compressor can be selected. With variable speed, the centrifugal compressor can deliver constant capacity at variable pressure, variable capacity at constant pressure, or a combination variable capacity and variable pressure.

Figure 19 Compressor Curves


55

Basically the performance of the centrifugal compressor, at speeds other than design, is such that the capacity will vary directly as the speed, the head developed as the square of the speed, and the required horsepower as the cube of the speed. As the speed deviates from the design speed, the error of these rules, known as the affinity laws, or fan laws, increases. The fan laws only apply to single stages or multi-stages with vent low compression ratios or very low Mach numbers. Fan Laws:


56

By varying speed, the centrifugal compressor will meet any load and pressure condition demanded by the process system within the operating limits of the compressor and the driver. It normally accomplishes this as efficiently as possible, since only the head required by the process is developed by the compressor. This compares to the essentially constant head developed by the constant speed compressor.

Page 57

Heat Transfer And Reaction Engineering Heat Transfer • Thermal conductivity • Conduction and convection • Insulation • Heat transfer coefficients and calculation • Heat exchangers, type and sizing • Steam reboilers • Condensers and sub-cooling • Introduction to energy recovery

Page 58

Thermal Conductivity In physics, thermal conductivity, k, is the intensive property of a material that indicates its ability to conduct heat. It is defined as the quantity of heat, Q, transmitted in time t through a thickness L, in a direction normal to a surface of area A, due to a temperature difference ∆T, under steady state conditions and when the heat transfer is dependent only on the temperature gradient. thermal conductivity = heat flow rate × distance / (area × temperature difference)

Examples In metals, thermal conductivity approximately tracks electrical conductivity, as the freely moving valence electrons transfer not only electric current but also heat. However, this correlation does not apply to some materials, as shown in the table below, where highly electrically conductive silver is shown to be less thermally conductive than diamond, which is an electrical semiconductor. Thermal conductivity is not a simple property, and depends intimately on structure and temperature. For instance, pure, crystalline substances also exhibit highly variable thermal conductivities along different crystal axes. One particularly notable example is sapphire, for which the CRC Handbook reports a thermal conductivity perpendicular to the c-axis of 2.6 W·m-1·K-1 at 373 K, and 6000 W·m-1·K-1 at 35 K for an angle of 36 degrees to the c-axis. Air and other gases are generally good insulators, in the absence of convection. Therefore, many insulating materials function simply by having a large number of gas-filled pockets which prevent large-scale convection. Examples of these include polystyrene (styrofoam) and silica aerogel. Thermal conductivity is clearly an important quantity for construction and related fields. However, materials used in such trades are rarely subjected to chemical purity standards. Several construction materials' k values are listed below. These should be considered approximate due to the uncertainties related to material definitions. The following table is meant as a small sample of data to illustrate the thermal conductivity of various types of substances. For more complete listings of measured k-values, see the references.

Page 59

Table 1Thermal conductivity properties

Thermal conductivity (W·m-1·K-1)

Temperature (K)

Notes

Diamond

1,000a

273

type I diamond

Silver

429a

300

Highest electrical conductivity of any metal

Iron, pure

80.2a

300

Stainless Steel

14a

273

Limestone

1.3b

Ice

2.2a

Soil

0.2-1.1c

Oak

0.16a

298

Rubber

0.16a

303

Polystyrene

0.033a

98-298

Nitrogen

0.026a

300

(92%)

Air

(100 kPa) 0.0262a

Silica aerogel 0.003a

273

300 98-298

For general scientific use, thermal conductance is the quantity of heat that passes in unit time through a plate of particular area and thickness when its opposite faces differ in temperature by one degree. For a plate of thermal conductivity k, area A and thickness L this is kA/L, measured in W·K-1. This matches the relationship between electrical conductivity (A·m-1·V-1) and electrical conductance (A·V-1). There is also a measure known as heat transfer coefficient: the quantity of heat that passes in unit time through unit area of a plate of particular thickness when its opposite faces differ in temperature by one degree. The reciprocal is thermal insulance. In summary: • •

thermal conductance = kA/L, measured in W·K-1 -1 o thermal resistance = L/kA, measured in K·W -1 -2 heat transfer coefficient = k/L, measured in W·K ·m 2 -1 o thermal insulance = L/k, measured in K·m ·W .

The heat transfer coefficient is also known as thermal admittance, but this term has other meanings.

Page 60

Conduction & Convection Conduction: In metals, the dominant method of conduction is through the movement of electrons. This method of conduction does not operate in non-metals because there are no free electrons (other than graphite). When a metal is heated, the electrons closest to the heat source vibrate more rapidly. Electrons then collide with these atoms and gain more kinetic energy (movement energy). The electrons therefore move around faster and collide with other free electrons which then gain more kinetic energy. Kinetic energy is therefore transferred between the electrons and through the metal from the point closest to the heat source towards points futher away. The electrons all travel very short distances but are very fast moving therefore conduction of heat happens very quickly. In metals and in insulators, there is conduction of heat due to the vibration of the atoms. As atoms closest to the heat source absorb heat/thermal energy, they make their neighbouring atoms vibrate more rapidly which then in turn make their neighbouring atoms vibrate more. Examples of conduction: The wire gauzes used on tripods are metal therefore they are good heat conductors. Gauzes on cookers are also metal so that heat is conducted quickly and food is cooked fast. Poor thermal conductors (insulators) are used for saucepan handles so that they don't heat up and can still be handled. Metals are used for the containers which heat liquids e.g. pans, kettles on hobs Air is a poor conductor therefore materials that trap air are used for insulation in lofts and hot water cylinders. Convection: The cool particles gain kinetic energy when they are heated from the source and expands as it heats up. The particles become less dense than the surrounding cold air therefore it rises and displace the cool air. Cool particles are more dense therefore they fall and move towards the heat source to take the place of the warm particles. They then heat up and rise while other particles cool down and fall. Example of convection: Convection is used in fridges to cool it down. Heat is carried away, therefore the back of fridges are always warm. Land & sea breezes are due to convection. Atmospheric winds. Hot water systems.

Page 61

Insulation Insulation is any material used to reduce or “slow down” or “resist” the flow of energy. There are several different types of insulators: • • •

Thermal insulators reduce the flow of heat. Electrical insulators reduce the flow of electricity. Acoustical insulators reduce the flow of sound.

A material may insulate well in more than one way. Some materials, such as diamond, are superb insulators in one way (electrical), but extremely poor insulators in another way (thermal). A purified synthetic diamond conducts heat even better than copper, and has the highest thermal conductivity of any known solid at room temperature. Thus it is the worst thermal insulator known that's solid at room temperature. Heat is the internal kinetic, vibrational energy that all materials contain (except at absolute zero). Heat spontaneously flows from a high temperature region to a low temperature region, and the greatest heat flow occurs through the path of least resistance. The proximity of a high temperature region to a low temperature region constitutes a temperature gradient. Thermal insulation maintains a thermal gradient by reducing the flow of heat across the temperature gradient. Insulation exists in most large appliances, for example, in ovens, refrigerators, freezers, and water heaters. In some cases, the insulation serves to prevent heat loss to the environment. In other cases, it serves to prevent heat gain from the environment.

Page 62

Heat transfer coefficients and calculation The heat transfer coefficient is used as a fudge factor in calculating heat transfer in thermodynamics. The heat transfer coefficent is often calculated from the Nusselt number (a dimensionless number). Below is an example where it is used to find the heat lost from a hot tube to the surrounding area.

where • • • •

Q = power input or heat lost h = overall heat transfer coefficient A = outside surface area of tubing ∆T = difference in temperature between tubing surface and surrounding area

There are different heat transfer relations for different liquids, flow regimes, and thermodynamic conditions. A common example pertinent to many of the necessary power plant efficiency and thermal hydraulic calculations is the Dittus-Boelter heat transfer corelation, valid for water in a circular pipe with Reynolds numbers between 100 000 and 120 000 and Prandtl numbers between 0.7 and 120. An example is shown below where it is used to calculate the heat transfer from a tubing wall to water.

where • • •

Hdb = => Dittus-Boelter correlation kw = thermal conductivity of water Nu = Nusselt number

•

Pr = Prandtl number =

• • • • • •

Re = Reynolds number = DH = hydraulic diameter = mass flow rate µ = water viscosity Cp = heat capacity at constant pressure A = cross-sectional area of flow

The heat transfer coefficient has SI units in watts per meter squared-kelvin. Often it can be estimated by dividing the thermal conductivity by a length scale. Heat transfer coefficients add inversely, like resistances. It can be thought of as a thermal resistance. Shown below is an addition of heat transfer coefficients where one is estimated as a thermal conductivity divided by a length scale.

where • • •

Q = power input h = heat transfer coefficient t = tubing thickness

Page 63

• • •

k = thermal conductivity of metal tube A = cross-sectional area of flow ∆T = difference in temperature between outer wall of tubing and sample water.

Page 64

Heat exchangers, type and sizing A heat exchanger is a component that allows the transfer of heat from one fluid (liquid or gas) to another fluid. Reasons for heat transfer include the following: 1. To heat a cooler fluid by means of a hotter fluid 2. To reduce the temperature of a hot fluid by means of a cooler fluid 3. To boil a liquid by means of a hotter fluid 4. To condense a gaseous fluid by means of a cooler fluid 5. To boil a liquid while condensing a hotter gaseous fluid Regardless of the function the heat exchanger fulfills, in order to transfer heat the fluids involved must be at different temperatures and they must come into thermal contact. Heat can flow only from the hotter to the cooler fluid. In a heat exchanger there is no direct contact between the two fluids. The heat is transferred from the hot fluid to the metal isolating the two fluids and then to the cooler fluid.

Types of Heat Exchanger Construction Although heat exchangers come in every shape and size imaginable, the construction of most heat exchangers falls into one of two categories: tube and shell, or plate. As in all mechanical devices, each type has its advantages and disadvantages.

Tube and Shell The most basic and the most common type of heat exchanger construction is the tube and shell, as shown in below. This type of heat exchanger consists of a set of tubes in a container called a shell. The fluid flowing inside the tubes is called the tube side fluid and the fluid flowing on the outside of the tubes is the shell side fluid. At the ends of the tubes, the tube side fluid is separated from the shell side fluid by the tube sheet(s). The tubes are rolled and press-fitted or welded into the tube sheet to provide a leak tight seal. In systems where the two fluids are at vastly different pressures, the higher pressure fluid is typically directed through the tubes and the lower pressure fluid is circulated on the shell side. This is due to economy, because the heat exchanger tubes can be made to withstand higher pressures than the shell of the heat exchanger for a much lower cost. The support plates shown also act as baffles to direct the flow of fluid within the shell back and forth across the tubes.

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Figure 20 Tube & Shell Heat Exchanger

Plate A plate type heat exchanger, as illustrated below, consists of plates instead of tubes to separate the hot and cold fluids. The hot and cold fluids alternate between each of the plates. Baffles direct the flow of fluid between plates. Because each of the plates has a very large surface area, the plates provide each of the fluids with an extremely large heat transfer area. Therefore a plate type heat exchanger, as compared to a similarly sized tube and shell heat exchanger, is capable of transferring much more heat. This is due to the larger area the plates provide over tubes. Due to the high heat transfer efficiency of the plates, plate type heat exchangers are usually very small when compared to a tube and shell type heat exchanger with the same heat transfer capacity. Plate type heat exchangers are not widely used because of the inability to reliably seal the large gaskets between each of the plates. Because of this problem, plate type heat exchangers have only been used in small, low pressure applications such as on oil coolers for engines. However, new improvements in gasket design and overall heat exchanger design have allowed some large scale applications of the plate type heat exchanger. As older facilities are upgraded or newly designed facilities are built, large plate type heat exchangers are replacing tube and shell heat exchangers and becoming more common.

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Figure 21 Plate Type Heat Exchanger

Because heat exchangers come in so many shapes, sizes, makes, and models, they are categorized according to common characteristics. One common characteristic that can be used to categorize them is the direction of flow the two fluids have relative to each other. The three categories are parallel flow, counter flow and cross flow. Parallel flow, as illustrated in below, exists when both the tube side fluid and the shell side fluid flow in the same direction. In this case, the two fluids enter the heat exchanger from the same end with a large temperature difference. As the fluids transfer heat, hotter to cooler, the temperatures of the two fluids approach each other. Note that the hottest cold-fluid temperature is always less than the coldest hot-fluid temperature

Figure 22 Parallel Flow Heat Exchanger

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Counter flow, as illustrated later, exists when the two fluids flow in opposite directions. Each of the fluids enters the heat exchanger at opposite ends. Because the cooler fluid exits the counter flow heat exchanger at the end where the hot fluid enters the heat exchanger, the cooler fluid will approach the inlet temperature of the hot fluid. Counter flow heat exchangers are the most efficient of the three types. In contrast to the parallel flow heat exchanger, the counter flow heat exchanger can have the hottest cold- fluid temperature greater than the coldest hot-fluid temperature.

Figure 23 Counter Flow Heat Exchanger Cross flow, as illustrated below, exists when one fluid flows perpendicular to the second fluid; that is, one fluid flows through tubes and the second fluid passes around the tubes at 90º angle. Cross flow heat exchangers are usually found in applications where one of the fluids changes state (2-phase flow). An example is a steam system's condenser, in which the steam exiting the turbine enters the condenser shell side, and the cool water flowing in the tubes absorbs the heat from the steam, condensing it into water. Large volumes of vapour may be condensed using this type of heat exchanger flow.

Figure 24 Cross Flow Heat Exchanger

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Steam Reboilers A reboiler is a special kind of heat exchanger used to put heat into a distillation column Steam may also be used to evapourate (or vapourise) a liquid, in a type of shell and tube heat exchanger known as a reboiler. These are used in the petroleum industry to vapourise a fraction of the bottom product from a distillation column. These tend to be horizontal, with vapourisation in the shell and condensation in the tubes

Figure 25 A Steam reboiler

Figure 26 Reboiler schematic

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Condensers and sub-cooling The steam condenser, shown below, is a major component of the steam cycle in power generation facilities. It is a closed space into which the steam exits the turbine and is forced to give up its latent heat of vapourization. It is a necessary component of the steam cycle for two reasons. One, it converts the used steam back into water for return to the steam generator or boiler as feedwater. This lowers the operational cost of the plant by allowing the clean and treated condensate to be reused, and it is far easier to pump a liquid than steam. Two, it increases the cycle's efficiency by allowing the cycle to operate with the largest possible delta- T and delta-P between the source (boiler) and the heat sink (condenser). Because condensation is taking place, the term latent heat of condensation is used instead of latent heat of vapourization. The steam's latent heat of condensation is passed to the water flowing through the tubes of the condenser. After the steam condenses, the saturated liquid continues to transfer heat to the cooling water as it falls to the bottom of the condenser, or hotwell. This is called subcooling, and a certain amount is desirable. A few degrees subcooling prevents condensate pump cavitation. The difference between the saturation temperature for the existing condenser vacuum and the temperature of the condensate is termed condensate depression. This is expressed as a number of degrees condensate depression or degrees subcooled. Excessive condensate depression decreases the operating efficiency of the plant because the subcooled condensate must be reheated in the boiler, which in turn requires more heat from the reactor, fossil fuel, or other heat source

Figure 27 Condenser

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There are different condenser designs, but the most common, at least in the large power generation facilities, is the straight-through, single-pass condenser illustrated above. This condenser design provides cooling water flow through straight tubes from the inlet water box on one end, to the outlet water box on the other end. The cooling water flows once through the condenser and is termed a single pass. The separation between the water box areas and the steam condensing area is accomplished by a tube sheet to which the cooling water tubes are attached. The cooling water tubes are supported within the condenser by the tube support sheets. Condensers normally have a series of baffles that redirect the steam to minimize direct impingement on the cooling water tubes. The bottom area of the condenser is the hotwell. This is where the condensate collects and the condensate pump takes its suction. If non-condensable gasses are allowed to build up in the condenser, vacuum will decrease and the saturation temperature at which the steam will condense increases. Non-condensable gasses also blanket the tubes of the condenser, thus reducing the heat transfer surface area of the condenser. This surface area can also be reduced if the condensate level is allowed to rise over the lower tubes of the condenser. A reduction in the heat transfer surface has the same effect as a reduction in cooling water flow. If the condenser is operating near its design capacity, a reduction in the effective surface area results in difficulty maintaining condenser vacuum. The temperature and flow rate of the cooling water through the condenser controls the temperature of the condensate. This in turn controls the saturation pressure (vacuum) of the condenser. To prevent the condensate level from rising to the lower tubes of the condenser, a hotwell level control system may be employed. Varying the flow of the condensate pumps is one method used to accomplish hotwell level control. A level sensing network controls the condensate pump speed or pump discharge flow control valve position. Another method employs an overflow system that spills water from the hotwell when a high level is reached. Condenser vacuum should be maintained as close to 29 inches Hg as practical. This allows maximum expansion of the steam, and therefore, the maximum work. If the condenser were perfectly air-tight (no air or non-condensable gasses present in the exhaust steam), it would be necessary only to condense the steam and remove the condensate to create and maintain a vacuum. The sudden reduction in steam volume, as it condenses, would maintain the vacuum. Pumping the water from the condenser as fast as it is formed would maintain the vacuum. It is, however, impossible to prevent the entrance of air and other noncondensable gasses into the condenser. In addition, some method must exist to initially cause a vacuum to exist in the condenser. This necessitates the use of an air ejector or vacuum pump to establish and help maintain condenser vacuum. Air ejectors are essentially jet pumps or eductors, as illustrated in Figure 10 below. In operation, the jet pump has two types of fluids. They are the high pressure fluid that flows through the nozzle, and the fluid being pumped which flows around the nozzle into the throat of the diffuser. The high velocity fluid enters the diffuser where its molecules strike other molecules. These molecules are in turn carried along with the high velocity fluid out of the diffuser creating a low pressure area around the mouth of the nozzle. This process is called entrainment. The low pressure area will draw more fluid from around the nozzle into the throat of the diffuser. As the fluid moves down the diffuser, the increasing area converts the velocity back to pressure. Use of steam at a pressure between 200 psi and 300 psi as the high pressure fluid enables a single stage air ejector to draw a vacuum of about 26 inches Hg.

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Figure 28 Change of section - change in pressure

Normally, air ejectors consist of two suction stages. The first stage suction is located on top of the condenser, while the second stage suction comes from the diffuser of the first stage. The exhaust steam from the second stage must be condensed. This is normally accomplished by an air ejector condenser that is cooled by condensate. The air ejector condenser also preheats the condensate returning to the boiler. Two-stage air ejectors are capable of drawing vacuums to 29 inches Hg. A vacuum pump may be any type of motor-driven air compressor. Its suction is attached to the condenser, and it discharges to the atmosphere. A common type uses rotating vanes in an elliptical housing. Single-stage, rotary-vane units are used for vacuums to 28 inches Hg. Two stage units can draw vacuums to 29.7 inches Hg. The vacuum pump has an advantage over the air ejector in that it requires no source of steam for its operation. They are normally used as the initial source of vacuum for condenser start-up.

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Introduction to energy recovery

Energy recovery. A process of converting used oil into usable energy, e.g., burned to recover energy, heat building, or incinerator At various stages in the refining process, useful energy carriers may be lost. The most important (energy) sources are the recovery of combustible products for useful applications, which would have been flared otherwise, as well as the recovery of hydrogen from different flue and process gas streams. The latter will reduce the need for additional hydrogen makeup; an energy-intensive and expensive process.

Flare gas recovery (or zero flaring) is a strategy evolving from the need to improve environmental performance. Generally, conventional flaring practice has been to operate at some flow greater than the manufacturer’s minimum flow rate to avoid damage to the flare (Miles, 2001). Typically, flared gas consists of background flaring (including planned intermittent and planned continuous flaring) and upset-blowdown flaring. In offshore flaring, background flaring can be as much as 50% of all flared gases. In refineries, background flaring will generally be less than 50%, depending on practices in the individual refinery. Recent discussions on emissions from flaring from the California Bay area refineries has highlighted the issue from an environmental perspective (Ezerksy, 2002).7 The report highlighted the higher emissions compared to previous assumptions of the Air Quality District, due to larger volumes of flared gases. The report also demonstrated the differences among various refineries, and plants within the refineries. Reduction of flaring will not only result in reduced air pollutant emissions, but also in increased energy-efficiency replacing fuels, as well as less negative publicity around flaring. Reduction of flaring can be achieved by improved recovery systems, including installing recovery compressors. New compressors and liquid-seals have been installed, and the two flare gas recovery systems have reduced flaring to near-zero levels (Fisher and Brennan, 2002). A plantwide assessment of the Equilon refinery in Martinez (now fully owned by Shell) highlighted the potential for flare gas recovery. The refinery will install new recovery compressors to reduce flaring. No specific costs were available for the flare gas recovery project, as it is part of a large package of measures for the refinery. The overall project has projected annual savings of $52 million and a payback period of 2 years (US DOEOIT, 2002). However, emissions can be further reduced by improved process control equipment and new flaring technology. Development of gas-recovery systems, development of new ignition systems with low-pilot-gas consumption or elimination of pilots altogether with the use of new ballistic ignition systems can reduce the amount of flared gas considerably. Development and demonstration of new ignition systems without a pilot may result in increased energy efficiency and reduced emissions.

Hydrogen Management and Recovery. Hydrogen is used in the refinery in processes such as hydrocrackers and desulphurisation using hydrotreaters. The production of hydrogen is an energy-intensive process using natural gas-fueled reformers. However, these processes and other processes generate gases that may contain a certain amount of hydrogen not used in the processes, or generated as byproduct of distillation of conversion processes. In addition, different processes have varying quality (purity) demands for the hydrogen feed. Reducing the need for hydrogen make-up will reduce energy use in the reformer and reduce the need for purchased natural gas. Natural

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gas is an expensive energy input in the refinery process, and lately associated with large fluctuation in prices. The major technology developments in the hydrogen management within the refinery are hydrogen process integration (or hydrogen cascading) and hydrogen recovery technology (Zagoria and Huycke, 2003). Revamping and retrofitting existing hydrogen networks can increase hydrogen capacity between 3% and 30% (Ratan and Vales, 2002).

Hydrogen integration at refineries is a new and important application of pinch analysis. Most hydrogen systems in refineries feature limited integration and pure hydrogen flows are sent from the reformers to the different processes in the refinery. But as the use of hydrogen is increasing the value of hydrogen is more and more appreciated. Using the approach of composition curves used in pinch analysis the production and uses of hydrogen of a refinery can be made visible. This allows us to identify the best matches between different hydrogen sources and uses based on quality of the hydrogen streams. It allows the user to select the appropriate and most cost-effective technology for hydrogen purification. A recent improvement of the analysis technology also accounts for gas pressure, to reduce compression energy needs (Hallale, 2001). The analysis method accounts also for costs of piping, besides the costs for generation, fuel use and compression power needs. It can be used for new and retrofit studies. The BP refinery at Carson, in a project with the California Energy Commission, has executed a Hydrogen Pinch analysis of the large refinery. Total potential savings of $4.5 million on operating costs were identified, but the refinery decided to realize a more cost effective package saving $3.9 million per year. As part of the plant-wide assessment of the Equilon (Shell) refinery at Martinez, an analysis of the hydrogen network has been included (US DOE-OIT, 2002). This has resulted in the identification of large energy savings. Further development and application of the analysis method at Californian refineries, especially as the need for hydrogen is increasing due to reduced future sulfur content of diesel and other fuels, may result in reduced energy needs at all refineries with hydrogen needs (all, except San Joaquin Refining in Bakersfield) (Khorram and Swaty, 2002). One refinery identified savings of $6 million/year in hydrogen savings without capital projects (Zagoria and Huycke, 2003).

Hydrogen recovery is an important technology development area to improve the efficiency of hydrogen recovery, reduce the costs of hydrogen recovery and increase the purity of the resulting hydrogen flow. Hydrogen can be recovered indirectly by routing low-purity hydrogen streams to the hydrogen plant (Zagoria and Huycke, 2003) or can be recovered from off gases by routing it to the existing purifier of the hydrogen plant or by installing additional purifiers to treat the off gases and vent gases. The cost savings of recovered hydrogen are around 50% of the costs of hydrogen production (Zagoria and Huycke, 2003). Membranes are an attractive technology for hydrogen recovery. If the content of recoverable products is higher than 2-5% (or preferably 10%), recovery may make economic sense (Baker et al., 2000). New membrane applications for the refinery and chemical industry are under development. Membranes for hydrogen recovery from ammonia plants have first been demonstrated about 20 years ago (Baker et al., 2000), and are used in various state-of-the-art plant designs. Refinery off gas flows have a different composition, making different membranes necessary for optimal recovery. Membrane plants have been demonstrated for recovery of hydrogen from hydrocracker off gases. Various suppliers offer membrane technologies for hydrogen recovery in the refining industry, including Air Liquide, Air Products and UOP. The hydrogen content has to be at least 25% for economic recovery of the hydrogen, with a recovery yield of 85-95% and a purity of 95%.

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Membrane technology generally represents the lowest cost option for low product rates, but not necessarily for high flow rates (Zagoria and Huycke, 2003). For high-flow rates PSA technology is often the conventional technology of choice. Development of low-cost and efficient membranes is an area of research interest to improve cost-effectiveness of hydrogen recovery, and enable the recovery of hydrogen from gas streams with lower concentrations.

Heat Recovery. Heat is recovered and re-used throughout the refinery. Next to efficient integration of heat flows throughout the refinery, the efficient operation of heat exchangers is a major area of interest. In a complex refinery most processes occur under high temperature and pressure conditions; the management and optimization of heat transfer among processes is therefore key to increasing overall energy efficiency. Fouling, a deposit buildup in units and piping that impede heat transfer, requires the combustion of additional fuel. For example, the processing of many heavy crude oils in the U.S. increases the likelihood of localized coke deposits in the heating furnaces, thereby reducing furnace efficiency and creating potential equipment failure. An estimate by the Office of Industrial Technology at the U.S. Department of Energy noted that the cost penalty for fouling could be as much as $2 billion annually in material and energy costs. The problem of fouling is expected to increase with the trend towards processing heavier crudes. Fouling is the effect of several process variables and heat exchanger design. Fouling may follow the combination of different mechanisms (Bott, 2001). Several methods of investigation have been underway to attempt to reduce fouling including the use of sensors to detect early fouling, physical and chemical methods to create high temperature coatings (without equipment modification), the use of ultrasound, as well as the improved long term design and operation of facilities. The U.S. Department of Energy initially funded preliminary research into this area, but funding has been discontinued (Huangfu, 2000; Bott, 2000). Initial analysis on fouling effects of a 100,000 bbl/day crude distillation unit found an additional heating load of 12.3 kBtu/barrel (13.0 MJ/barrel) processes (Panchal and Huangfu, 2000). Reducing this additional heating load could results in significant energy savings. This technology is still in the conceptual and basic research stage and therefore it is difficult to assess capital costs at this time. Argonne National Laboratory (ANL) has been the lead in working with the refining industry in the area. Progress so far has included: a basic understanding of fouling mechanisms developed (for example, the presence of iron sulfide in crude oil and its link to fouling), the development of a threshold fouling model by ANL, the testing of prototype fouling detection units, the development of a Heat Exchanger Design Handbook (1999 Edition) incorporating ANL’s petroleum fouling threshold model, and the preparation of a guideline document on Heat Exchanger Fouling in the Crude Oil Distillation Unit (Panchal, 2000). Besides ANL, several other groups have worked in the area of fouling reduction. Outside the U.S., groups in Europe and Canada have worked on fouling. While the issue of fouling is now on the radar screen of plant managers (there is a biannual Fouling Mitigation conference held by the American Institute for Chemical Engineers), a stronger commitment by the refining industry would be needed to advance this technology to the next stage of development. Some sources believe that the future development of this area is expected to be in the area of Condition-Based Maintenance of Heat-Transfer Equipment that will be based on Knowledge-Based and Monitoring –Based Mitigation of Fouling/Corrosion (Panchal, 2000, see also section on process control systems). Furthermore, developments in heat exchanger design and process intensification may also contribute to reducing the problem of fouling.

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An Introduction to Pinch Technology While oil prices continue to climb, energy conservation remains the prime concern for many process industries. The challenge every process engineer is faced with is to seek answers to questions related to their process energy patterns. A few of the frequently asked questions are: 1. Are the existing processes as energy efficient as they should be? 2. How can new projects be evaluated with respect to their energy requirements? 3. What changes can be made to increase the energy efficiency without incurring any cost? 4. What investments can be made to improve energy efficiency? 5. What is the most appropriate utility mix for the process? 6. How to put energy efficiency and other targets like reducing emissions, increasing plant capacities, improve product qualities etc, into a one coherent strategic plan for the overall site?

What is Pinch Technology?

Meaning of the term "Pinch Technology" The term "Pinch Technology" was introduced by Linnhoff and Vredeveld to represent a new set of thermodynamically based methods that guarantee minimum energy levels in design of heat exchanger networks. Over the last two decades it has emerged as an unconventional development in process design and energy conservation. The term ‘Pinch Analysis’ is often used to represent the application of the tools and algorithms of Pinch Technology for studying industrial processes. Developments of TM TM TM rigorous software programs like PinchExpress , SuperTarget , Aspen Pinch have proved to be very useful in pinch analysis of complex industrial processes with speed and efficiency.

Basis of Pinch Analysis Pinch technology presents a simple methodology for systematically analysing chemical processes and the surrounding utility systems with the help of the First and Second Laws of Thermodynamics. The First Law of Thermodynamics provides the energy equation for calculating the enthalpy changes (dH) in the streams passing through a heat exchanger. The Second Law determines the direction of heat flow. That is, heat energy may only flow in the direction of hot to cold. This prohibits ‘temperature crossovers’ of the hot and cold stream profiles through the exchanger unit. In a heat exchanger unit neither a hot stream can be cooled below cold stream supply temperature nor a cold stream can be heated to a temperature more than the supply temperature of hot stream. In practice the hot stream can only be cooled to a temperature defined by the ‘temperature approach’ of the heat exchanger. The temperature approach is the minimum allowable temperature difference (DTmin) in the stream temperature profiles, for the heat exchanger unit. The temperature level at which DTmin is observed in the process is referred to as "pinch point" or "pinch condition". The pinch defines the minimum driving force allowed in the exchanger unit.

Objectives of Pinch Analysis Pinch Analysis is used to identify energy cost and heat exchanger network (HEN) capital cost targets for a process and recognizing the pinch point. The procedure first predicts, ahead of design, the minimum requirements of external energy, network area, and the number of units for a given process at the pinch point. Next a heat exchanger network design that satisfies these targets is synthesized. Finally the network is optimized by comparing energy cost and the capital cost of the network so that the total annual cost is minimized. Thus, the prime objective of pinch analysis is to achieve financial savings by better process heat integration (maximizing process-to-process heat recovery and reducing the external utility loads). The concept of process heat integration is illustrated in the example discussed below.

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A Simple Example of Process Integration by Pinch Analysis Consider the following simple process [Figure 29] where feed stream to a reactor is heated before inlet to a reactor and the product stream is to be cooled. The heating and cooling are done by use of steam (Heat Exchanger -1) and cooling water (Heat Exchanger-2), respectively. The Temperature (T) vs. Enthalpy (H) plot for the feed and product streams depicts the hot (Steam) and cold (CW) utility loads when there is no vertical overlap of the hot and cold stream profiles.

Figure 29 A Simple Flow Scheme with T-H profile An alternative, improved scheme is shown below where the addition of a new ‘Heat Exchanger–3’ recovers product heat (X) to preheat the feed. The steam and cooling water requirements also get reduced by the same amount (X). The amount of heat recovered (X) depends on the ‘minimum approach temperature’ allowed for the new exchanger. The minimum temperature approach between the two curves on the vertical axis is DTmin and the point where this occurs is defined as the "pinch". o From the T-H plot, the X amount corresponds to a DTmin value of 20 C. Increasing the DTmin value leads to higher utility requirements and lower area requirements..

Figure 30 Improved Flow Scheme with T-H profile

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Development of the Pinch Technology Approach When the process involves single hot and cold streams (as in above example) it is easy to design an optimum heat recovery exchanger network intuitively by heuristic methods. In any industrial set-up the number of streams is so large that the traditional design approach has been found to be limiting in the design of a good network. With the development of pinch technology in the late 1980’s, not only optimal network design was made possible, but also considerable process improvements could be discovered. Both the traditional and pinch approaches are depicted below

Figure 31 Graphic Representation of Traditional and Pinch Design Approaches

Traditional Design Approach: First, the core of the process is designed with fixed flow rates and temperatures yielding the heat and mass balance for the process. Then the design of a heat recovery system is completed. Next, the remaining duties are satisfied by the use of the utility system. Each of these exercises is performed independently of the others.

Pinch Technology Approach: Process integration using pinch technology offers a novel approach to generate targets for minimum energy consumption before heat recovery network design. Heat recovery and utility system constraints are then considered in the design of the core process. Interactions between the heat recovery and utility systems are also considered. The pinch design can reveal opportunities to modify the core process to improve heat integration. The pinch approach is unique because it treats all processes with multiple streams as a single, integrated system. This method helps to optimize the heat transfer equipment during the design of the equipment.

Areas of Applications of Pinch Technology Pinch originated in the petrochemical sector and is now being applied to solve a wide range of problems in mainstream chemical engineering. Wherever heating and cooling of process materials takes places there is a potential opportunity. Thus initial applications of the technology were found in projects relating to energy saving in industries as diverse as iron and steel, food and drink, textiles, paper and cardboard, cement, base chemicals, oil, and petrochemicals. Early emphasis on energy conservation led to the misconception that conservation is the main area of application for pinch technology. The technology, when applied with imagination, can affect reactor design, separator design, and the overall process optimization in any plant. It has been applied to processing problems that go far beyond energy conservation. It has been employed to solve problems as diverse as improving effluent quality, reducing emissions, increasing product yield, debottlenecking, increasing throughput, and improving the flexibility and safety of the processes.

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Basic Concepts of Pinch Analysis Most industrial processes involve transfer of heat either from one process stream to another process stream (interchanging) or from a utility stream to a process stream. In the present energy crisis scenario all over the world, the target in any industrial process design is to maximize the process-toprocess heat recovery and to minimize the utility (energy) requirements. To meet the goal of maximum energy recovery or minimum energy requirement (MER) an appropriate heat exchanger network (HEN) is required. The design of such a network is not an easy task considering the fact that most processes involve a large number of process and utility streams. As explained in the previous section, the traditional design approach has resulted in networks with high capital and utility costs. With the advent of pinch analysis concepts, the network design has become very systematic and methodical. A summary of the key concepts, their significance, and the nomenclature used in pinch analysis is given below: Combined (Hot and Cold ) Composite Curves: Used to predict targets for Minimum energy (both hot and cold utility) required, • Minimum network area required, and • Minimum number of exchanger units required. • DTmin and Pinch Point: The DTmin value determines how closely the hot and cold composite curves can be ‘pinched’ (or squeezed) without violating the Second Law of Thermodynamics (none of the heat exchangers can have a temperature crossover). • Grand Composite Curve: Used to select appropriate levels of utilities (maximize cheaper utilities) to meet over all energy requirements. • Energy and Capital Cost Targeting: Used to calculate total annual cost of utilities and capital cost of heat exchanger network. • Total Cost Targeting: Used to determine the optimum level of heat recovery or the optimum DTmin value, by balancing energy and capital costs. Using this method, it is possible to obtain an accurate estimate (within 10 - 15%) of overall heat recovery system costs without having to design the system. The essence of the pinch approach is the speed of economic evaluation. • Plus/Minus and Appropriate Placement Principles: The "Plus/Minus" Principle provides guidance regarding how a process can be modified in order to reduce associated utility needs and costs. The Appropriate Placement Principles provide insights for proper integration of key equipments like distillation columns, evaporators, furnaces, heat engines, heat pumps, etc. in order to reduce the utility requirements of the combined system. • Total Site Analysis: This concept enables the analysis of the energy usage for an entire plant site that consists of several processes served by a central utility system.

Steps of Pinch Analysis In any Pinch Analysis problem, whether a new project or a retrofit situation, a well-defined stepwise procedure is followed. It should be noted that these steps are not necessarily performed on a oncethrough basis, independent of one another. Additional activities such as re-simulation and data modification occur as the analysis proceeds and some iteration between the various steps is always required.

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Figure 32 Steps of Pinch Analysis

Identification of the Hot, Cold and Utility Streams in the Process ‘Hot Streams’ are those that must be cooled or are available to be cooled. e.g. product cooling before storage ‘Cold Streams’ are those that must be heated e.g. feed preheat before a reactor. ‘Utility Streams’ are used to heat or cool process streams, when heat exchange between process streams is not practical or economic. A number of different hot utilities (steam, hot water, flue gas, etc.) and cold utilities (cooling water, air, refrigerant, etc.) are used in industry. The identification of streams needs to be done with care as sometimes, despite undergoing changes in temperature, the stream is not available for heat exchange. For example, when a gas stream is compressed the stream temperature rises because of the conversion of mechanical energy into heat and not by any fluid to fluid heat exchange. Hence such a stream may not be available to take part in any heat exchange. In the context of pinch analysis, this stream may or may not be considered to be a process stream.

2. Thermal Data Extraction for Process & Utility Streams For each hot, cold and utility stream identified, the following thermal data is extracted from the process material and heat balance flow sheet: o

•

Supply temperature (TS C) : the temperature at which the stream is available.

•

Target temperature (TT C) : the temperature the stream must be taken to.

•

Heat capacity flow rate (CP kW/ C) : the product of flow rate (m) in kg/sec and specific heat 0 (Cp kJ/kg C).

o

o

CP = m x Cp •

Enthalpy Change (dH) associated with a stream passing through the exchanger is given by the First Law of Thermodynamics:

First Law energy equation: d H = Q ± W In a heat exchanger, no mechanical work is being performed: W = 0 (zero)

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The above equation simplifies to: d H = Q, where Q represents the heat supply or demand associated with the stream. It is given by the relationship: Q= CP x (TS - TT). Enthalpy Change, dH = CP x (TS - TT) ** Here the specific heat values have been assumed to be temperature independent within the operating range. The stream data and their potential effect on the conclusions of a pinch analysis should be considered during all steps of the analysis. Any erroneous or incorrect data can lead to false conclusions. In order to avoid mistakes, the data extraction is based on certain qualified principles. For details on principles of data extraction, check out Link-2 at the end of the article. The data extracted is presented in the below.

Table 2 Typical Stream Data STREAM NUMBER

STREAM NAME

SUPPLY TEMP. °C

TARGET TEMP. °C

HEAT CAP. kW /°C

ENTH. FLOW. kW

1

FEED

60

205

20

2900

2

REAC.OUT

270

160

18

1980

3

PRODUCT

220

70

35

5250

4

RECYCLE

160

210

50

2500

CHANGE

Selection of Initial DTmin value The design of any heat transfer equipment must always adhere to the Second Law of Thermodynamics that prohibits any temperature crossover between the hot and the cold stream i.e. a minimum heat transfer driving force must always be allowed for a feasible heat transfer design. Thus the temperature of the hot and cold streams at any point in the exchanger must always have a minimum temperature difference (DTmin). This DTmin value represents the bottleneck in the heat recovery. In mathematical terms, at any point in the exchanger Hot stream Temp. ( TH ) - ( TC ) Cold stream Temp. >= DTmin

The value of DTmin is determined by the overall heat transfer coefficients (U) and the geometry of the heat exchanger. In a network design, the type of heat exchanger to be used at the pinch will determine the practical Dtmin for the network. For example, an initial selection for the Dtmin value for shell and 0 tubes may be 3-5 C (at best) while compact exchangers such as plate and frame often allow for an 0 initial selection of 2-3 C. The heat transfer equation, which relates Q, U, A and LMTD (Log Mean Temperature Difference) is depicted in Figure 4.

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Figure 33 Heat Transfer Equation For a given value of heat transfer load (Q), if smaller values of DTmin are chosen, the area requirements rise. If a higher value of DTmin is selected the heat recovery in the exchanger decreases and demand for external utilities increases. Thus, the selection of DTmin value has implications for both capital and energy costs. This concept will become clearer with the help of composite curves and total cost targeting discussed later. Just as for a single heat exchanger, the choice of DTmin (or approach temperature) is vital in the design of a heat exchanger networks. To begin the process an initial DTmin value is chosen and pinch analysis is carried out. Typical DTmin values based on experience are available in literature for reference. A few values based on Linnoff March’s application experience are tabulated below for shell and tube heat exchangers. No

Industrial Sector

Experience DTmin Values

1

Oil Refining

20-40ºC

2

Petrochemical

10-20ºC

3

Chemical

10-20ºC

4

Low Processes

Temperature 3-5ºC

Construction of Composite Curves and Grand Composite Curve COMPOSITE CURVES: Temperature - Enthalpy (T - H) plots known as ‘Composite curves’ have been used for many years to set energy targets ahead of design. Composite curves consist of temperature (T) – enthalpy (H) profiles of heat availability in the process (the hot composite curve) and heat demands in the process (the cold composite curve) together in a graphical representation. In general any stream with a constant heat capacity (CP) value is represented on a T - H diagram by a straight line running from stream supply temperature to stream target temperature. When there are a number of hot and cold streams, the construction of hot and cold composite curves simply involves the addition of the enthalpy changes of the streams in the respective temperature intervals. An example of hot composite curve construction is shown in Figure 5(a) and (b). A complete hot or cold composite curve consists of a series of connected straight lines, each change in slope represents a change in overall hot stream heat capacity flow rate (CP).

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Figure 34 Temperature-Enthalpy Relations Used to Construct Composite Curves For heat exchange to occur from the hot stream to the cold stream, the hot stream cooling curve must lie above the cold stream-heating curve. Because of the ‘kinked’ nature of the composite curves (Figure 6), they approach each other most closely at one point defined as the minimum approach temperature (DTmin). DTmin can be measured directly from the T-H profiles as being the minimum vertical difference between the hot and cold curves. This point of minimum temperature difference represents a bottleneck in heat recovery and is commonly referred to as the "Pinch". Increasing the DTmin value results in shifting the of the curves horizontally apart resulting in lower process to process heat exchange and higher utility requirements. At a particular DTmin value, the overlap shows the maximum possible scope for heat recovery within the process. The hot end and cold end overshoots indicate minimum hot utility requirement (QHmin) and minimum cold utility requirement (QCmin), of the process for the chosen DTmin. Thus, the energy requirement for a process is supplied via process to process heat exchange and/or exchange with several utility levels (steam levels, refrigeration levels, hot oil circuit, furnace flue gas, etc.). Graphical constructions are not the most convenient means of determining energy needs. A numerical approach called the "Problem Table Algorithm" (PTA) was developed by Linnhoff & Flower (1978) as a means of determining the utility needs of a process and the location of the process pinch. The PTA lends itself to hand calculations of the energy targets. To summarize, the composite curves provide overall energy targets but do not clearly indicate how much energy must be supplied by different utility levels. The utility mix is determined by the Grand Composite Curve.

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Figure 35 Combined Composite Curves GRAND COMPOSITE CURVE (GCC): In selecting utilities to be used, determining utility temperatures, and deciding on utility requirements, the composite curves and PTA are not particularily useful. The introduction of a new tool, the Grand Composite Curve (GCC), was introduced in 1982 by Itoh, Shiroko and Umeda. The GCC (Figure 7) shows the variation of heat supply and demand within the process. Using this diagram the designer can find which utilities are to be used. The designer aims to maximize the use of the cheaper utility levels and minimize the use of the expensive utility levels. Low-pressure steam and cooling water are preferred instead of high-pressure steam and refrigeration, respectively. The information required for the construction of the GCC comes directly from the Problem Table Algorithm developed by Linnhoff & Flower (1978). The method involves shifting (along the temperature [Y] axis) of the hot composite curve down by ½ DTmin and that of cold composite curve up by ½ DTmin. The vertical axis on the shifted composite curves shows process interval temperature. In other words, the curves are shifted by subtracting part of the allowable temperature approach from the hot stream temperatures and adding the remaining part of the allowable temperature approach to the cold stream temperatures. The result is a scale based upon process temperature having an allowance for temperature approach (DTmin). The Grand Composite Curve is then constructed from the enthalpy (horizontal) differences between the shifted composite curves at different temperatures. On the GCC, the horizontal distance separating the curve from the vertical axis at the top of the temperature scale shows the overall hot utility consumption of the process.

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Figure 36 Grand Composite Curve Figure 7 shows that it is not necessary to supply the hot utility at the top temperature level. The GCC indicates that we can supply the hot utility over two temperature levels TH1 (HP steam) and TH2 (LP steam). Recall that, when placing utilities in the GCC, intervals, and not actual utility temperatures, should be used. The total minimum hot utility requirement remains the same: QHmin = H1 (HP steam) + H2 (LP steam). Similarly, QCmin = C1 (Refrigerant) +C2 (CW). The points TH2 and TC2 where the H2 and C2 levels touch the grand composite curve are called the "Utility Pinches." The shaded green pockets represent the process-to-process heat exchange. In summary, the grand composite curve is one of the most basic tools used in pinch analysis for the selection of the appropriate utility levels and for targeting of a given set of multiple utility levels. The targeting involves setting appropriate loads for the various utility levels by maximizing the least expensive utility loads and minimizing the loads on the most expensive utilities.

Estimation of Minimum Energy Cost Targets Once the DTmin is chosen, minimum hot and cold utility requirements can be evaluated from the composite curves. The GCC provides information regarding the utility levels selected to meet QHmin and QCmin requirements. If the unit cost of each utility is known, the total energy cost can be calculated using the energy equation given below.

Estimation of Heat Exchanger Network ( HEN ) Capital Cost Targets The capital cost of a heat exchanger network is dependent upon three factors: 1. the number of exchangers, 2. the overall network area,

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3. the distribution of area between the exchangers Pinch analysis enables targets for the overall heat transfer area and minimum number of units of a heat exchanger network (HEN) to be predicted prior to detailed design. It is assumed that the area is evenly distributed between the units. The area distribution cannot be predicted ahead of design. •

AREA TARGETING: The calculation of surface area for a single counter-current heat exchanger requires the knowledge of the temperatures of streams in and out (dTLM i.e. Log Mean Temperature Difference or LMTD), overall heat transfer coefficient (U-value), and total heat transferred (Q). The area is given by the relation Area = Q / [ U x dTLM ]

The composite curves can be divided into a set of adjoining enthalpy intervals such that within each interval, the hot and cold composite curves do not change slope. Here the heat exchange is assumed to be "vertical" (pure counter-current heat exchange). The hot streams in any enthalpy interval, at any point, exchanges heat with the cold streams at the temperature vertically below it. The total area of the HEN (Amin) is given by the formula in Figure below, where i denotes the ith enthalpy and interval j denotes the jth stream and dTLM denotes LMTD in the ith interval.

Figure 37 HEN AREA min Estimation from Composite Curves The actual HEN total area required is generally within 10% of the area target as calculated above. With inclusion of temperature correction factors area targeting can be extended to non counter-current heat exchange as well. NUMBER OF UNITS TARGETING: For the minimum number of heat exchanger units (Nmin) required for MER (minimum energy requirement or maximum energy recovery), the HEN can be evaluated prior to HEN design by using a simplified form of Euler’s graph theorem. In designing for the minimum energy requirement (MER), no heat transfer is allowed across the pinch and so a realistic target for the minimum number of units (NminMER) is the sum of the targets evaluated both above and below the pinch separately. NminMER=[Nh+Nc+Nu–1]AP +[Nh+Nc+Nu–1]BP Where : Nh = Number of hot streams Nc=Number of cold streams Nu = Number of utility streams AP / BP : Above / Below Pinch

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•

HEN TOTAL CAPITAL COST TARGETING: The targets for the minimum surface area (Amin) and the number of units (Nmin) can be combined together with the heat exchanger cost law to determine the targets for HEN capital cost (CHEN). The capital cost is annualized using an annualization factor that takes into account interest payments on borrowed capital. The equation used for calculating the total capital cost and exchanger cost law is given below.

For the Exchanger Cost Equation shown above, typical values for a carbon steel shell and tube exchnager would be a = 16,000, b = 3,200, and c = 0.7. The installed cost can be considered to be 3.5 times the purchased cost given by the Exchanger Cost Equation. •

Estimation of Optimum DTmin Value by Energy-Capital Trade Off

To arrive at an optimum DTmin value, the total annual cost (the sum of total annual energy and capital cost) is plotted at varying DTmin values (Figure 7). Three key observations can be made from Figure 9: An increase in DTmin values result in higher energy costs and lower capital costs. A decrease in DTmin values result in lower energy costs and higher capital costs. An optimum DTmin exists where the total annual cost of energy and capital costs is minimized. Thus, by systematically varying the temperature approach we can determine the optimum heat recovery level or the DTminOPTIMUM for the process.

Figure 38 Energy-Capital Cost Trade Off (Optimum DTmin)

Estimation of Practical Targets for HEN Design The heat exchanger network designed on the basis of the estimated optimum DTmin value is not 0 always the most appropriate design. A very small DTmin value, perhaps 8 C, can lead to a very complicated network design with a large total area due to low driving forces. The designer, in practice, 0 selects a higher value (15 C) and calculates the marginal increases in utility duties and area requirements. If the marginal cost increase is small, the higher value of DTmin is selected as the practical pinch point for the HEN design.

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Recognizing the significance of the pinch temperature allows energy targets to be realized by design of appropriate heat recovery network. So what is the significance of the pinch temperature? The pinch divides the process into two separate systems each of which is in enthalpy balance with the utility. The pinch point is unique for each process. Above the pinch, only the hot utility is required. Below the pinch, only the cold utility is required. Hence, for an optimum design, no heat should be transferred across the pinch. This is known as the key concept in Pinch Technology. To summarize, Pinch Technology gives three rules that form the basis for practical network design: No external heating below the Pinch. No external cooling above the Pinch. No heat transfer across the Pinch. Violation of any of the above rules results in higher energy requirements than the minimum requirements theoretically possible. Plus/Minus Principle: The overall energy needs of a process can be further reduced by introducing process changes (changes in the process heat and material balance). There are several parameters that could be changed such as reactor conversions, distillation column operating pressures and reflux ratios, feed vaporization pressures, or pump-around flow rates. The number of possible process changes is nearly infinite. By applying the pinch rules as discussed above, it is possible to identify changes in the appropriate process parameter that will have a favorable impact on energy consumption. This is called the "Plus/Minus Principle." Applying the pinch rules to study of composite curves provide us the following guidelines: •

Increase (+) in hot stream duty above the pinch.

•

Decrease (-) in cold stream duty above the pinch.

This will result in a reduced hot utility target, and any •

Decrease (-) in hot stream duty below the pinch.

•

Increase (+) in cold stream duty below the pinch

will result in a reduced cold utility target. These simple guidelines provide a definite reference for the adjustment of single heat duties such as vaporization of a recycle, pump-around condensing duty, and others. Often it is possible to change temperatures rather than the heat duties. The target should be to •

Shift hot streams from below the pinch to above and

•

Shift cold streams from above the pinch to below.

The process changes that can help achieve such stream shifts essentially involve changes in following operating parameters: •

reactor pressure/temperatures

•

distillation column temperatures, reflux ratios, feed conditions, pump around conditions, intermediate condensers

•

evaporator pressures

•

storage vessel temperatures

For example, if the pressure for a feed vaporizer is lowered, vaporization duty can shift from above to below the pinch. The leads to reduction in both hot and cold utilities. Appropriate Placement Principles: Apart from the changes in process parameters, proper integration of key equipment in process with respect to the pinch point should also be considered. The pinch concept of "Appropriate Placement" (integration of operations in such a way that there is reduction in the utility requirement of the combined system) is used for this purpose. Appropriate placement principles have been developed for distillation columns, evaporators, heat engines,

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furnaces, and heat pumps. For example, a single-effect evaporator having equal vaporization and condensation loads, should be placed such that both loads balance each other and the evaporator can be operated without any utility costs. This means that appropriate placement of the evaporator is on either side of the pinch and not across the pinch. In addition to the above pinch rules and principles, a large number of factors must also be considered during the design of heat recovery networks. The most important are operating cost, capital cost, safety, operability, future requirements, and plant operating integrity. Operating costs are dependent on hot and cold utility requirements as well as pumping and compressor costs. The capital cost of a network is dependent on a number of factors including the number of heat exchangers, heat transfer areas, materials of construction, piping, and the cost of supporting foundations and structures. With a little practice, the above principles enable the designer to quickly pan through 40-50 possible modifications and choose 3 or 4 that will lead to the best overall cost effects. The essence of the pinch approach is to explore the options of modifying the core process design, heat exchangers, and utility systems with the ultimate goal of reducing the energy and/or capital cost.

9. Design of Heat Exchanger Network The design of a new HEN is best executed using the "Pinch Design Method (PDM)". The systematic application of the PDM allows the design of a good network that achieves the energy targets within practical limits. The method incorporates two fundamentally important features: (1) it recognizes that the pinch region is the most constrained part of the problem (consequently it starts the design at the pinch and develops by moving away) and (2) it allows the designer to choose between match options. In effect, the design of network examines which "hot" streams can be matched to "cold" streams via heat recovery. This can be achieved by employing "tick off" heuristics to identify the heat loads on the pinch exchanger. Every match brings one stream to it target temperature. As the pinch divides the heat exchange system into two thermally independent regions, HENs for both above and below pinch regions are designed separately. When the heat recovery is maximized the remaining thermal needs must be supplied by hot utility. The graphical method of representing flow streams and heat recovery matches is called a ‘grid diagram’ (Figure 10).

Figure 39 Typical Grid Diagram All the cold (blue lines) and hot (red line) streams are represented by horizontal lines. The entrance and exit temperatures are shown at either end. The vertical line in the middle represents the pinch temperature. The circles represent heat exchangers. Unconnected circles represent exchangers using utility heating and cooling.

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The design of a network is based on certain guidelines like the "CP Inequality Rule", "Stream Splitting", "Driving Force Plot" and "Remaining Problem Analysis". Having made all the possible matches, the two designs above and below the pinch are then brought together and usually refined to further minimize the capital cost. After the network has been designed according to the pinch rules, it can be further subjected to energy optimization. Optimizing the network involves both topological and parametric changes of the initial design in order to minimize the total cost. Benefits and Applications of Pinch Technology One of the main advantages of Pinch Technology over conventional design methods is the ability to set energy and capital cost targets for an individual process or for an entire production site ahead of design. Therefore, in advance of identifying any projects, we know the scope for energy savings and investment requirements. General Process Improvements In addition to energy conservation studies, Pinch Technology enables process engineers to achieve the following general process improvements: Update or Modify Process Flow Diagrams (PFDs): Pinch quantifies the savings available by changing the process itself. It shows where process changes reduce the overall energy target, not just local energy consumption. Conduct Process Simulation Studies: Pinch replaces the old energy studies with information that can be easily updated using simulation. Such simulation studies can help avoid unnecessary capital costs by identifying energy savings with a smaller investment before the projects are implemented. Set Practical Targets: By taking into account practical constraints (difficult fluids, layout, safety, etc.), theoretical targets are modified so that they can be realistically achieved. Comparing practical with theoretical targets quantifies opportunities "lost" by constraints - a vital insight for long-term development. Debottlenecking: Pinch Analysis, when specifically applied to debottlenecking studies, can lead to the following benefits compared to a conventional revamp: •

Reduction in capital costs

•

Decrease in specific energy demand giving a more competitive production facility

For example, debottlenecking of distillation columns by Column Targeting can be used to identify less expensive alternatives to column retraying or installation of a new column. Determine Opportunities for Combined Heat and Power (CHP) Generation: A well-designed CHP system significantly reduces power costs. Pinch shows the best type of CHP system that matches the inherent thermodynamic opportunities on the site. Unnecessary investments and operating costs can be avoided by sizing plants to supply energy that takes heat recovery into consideration. Heat recovery should be optimized by Pinch Analysis before specifying CHP systems. Decide what to do with low-grade waste heat: Pinch shows, which waste heat streams, can be recovered and lends insight into the most effective means of recovery. Industrial Applications The application of Pinch Technology has resulted in significant improvements in the energy and capital efficiency of industrial facilities worldwide. It has been successfully applied in many different industries from petroleum and base chemicals to food and paper. Both continuous and batch processes have been successfully analyzed on an individual unit and site-wide basis. Pinch technology has been extensively used to capitalize on the mistakes of the past. It identifies the existence of built-in spare heat transfer areas and presents the designer with opportunities for cheap retrofits. In case of the design of new plants, Pinch Analysis has played a very important role and minimized capital costs. A Case Study: When Pennzoil was adding a residual catalytic cracking (RCC) unit, the gas plant associated with the RCC and an alkylation unit at its Atlas Refining facility in Shreveport, energy efficiency was one of their major considerations in engineering the refinery expansion. Electric Power Research Institute (EPRI) and Pennzoil's energy provider, SWEPCO, used pinch technology to carry

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out an optimization study of the new units and the utility systems that serve them rather than simply incorporating standard process packages provided by licensors. The pinch study identified opportunities for saving up to 23.7% of the process heating through improved heat integration. Net savings for Pennzoil were estimated at $13.7 million over 10 years.

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Catalysts and Reaction Engineering • • •

Chemical reactions Reaction kinetics Introduction to catalysis

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Chemical Reactions

Separation: heavy on the bottom, light on the top

Modern separation involves piping oil through hot furnaces. The resulting liquids and vapours are discharged into distillation towers, the tall, narrow columns that give refineries their distinctive skylines.

Inside the towers, the liquids and vapours separate into components or fractions according to weight and boiling point. The lightest fractions, including gasoline and liquid petroleum gas (LPG), vapourize and rise to the top of the tower, where they condense back to liquids. Medium weight liquids, including kerosene and diesel oil distillates, stay in the middle. Heavier liquids, called gas oils, separate lower down, while the heaviest fractions with the highest boiling points settle at the bottom. These tarlike fractions, called residuum, are literally the "bottom of the barrel." The fractions now are ready for piping to the next station or plant within the refinery. Some components require relatively little additional processing to become asphalt base or jet fuel. However, most molecules that are destined to become high-value products require much more processing.

Conversion: cracking and rearranging molecules to add value This is where refining's fanciest footwork takes place--where fractions from the distillation towers are transformed into streams (intermediate components) that eventually become finished products. This also is where a refinery makes money, because only through conversion can most low-value fractions become gasoline. The most widely used conversion method is called cracking because it uses heat and pressure to "crack" heavy hydrocarbon molecules into lighter ones. A cracking unit consists of one or more tall, thick-walled, bullet-shaped reactors and a network of furnaces, heat exchangers and other vessels.

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Fluid catalytic cracking, or "cat cracking," is the basic gasoline-making process. Using intense heat (about 1,000 degrees Fahrenheit), low pressure and a powdered catalyst (a substance that accelerates chemical reactions), the cat cracker can convert most relatively heavy fractions into smaller gasoline molecules. Hydrocracking applies the same principles but uses a different catalyst, slightly lower temperatures, much greater pressure and hydrogen to obtain chemical reactions. Although not all refineries employ hydrocracking, Chevron is an industry leader in using this technology to cost-effectively convert medium- to heavyweight gas oils into high-value streams. The company's patented hydrocracking process, which takes place in the Isocracker unit, produces mostly gasoline and jet fuel. Some refineries also have cokers, which use heat and moderate pressure to turn residuum into lighter products and a hard, coallike substance that is used as an industrial fuel. Cokers are among the more peculiar-looking refinery structures. They resemble a series of giant drums with metal derricks on top. Cracking and coking are not the only forms of conversion. Other refinery processes, instead of splitting molecules, rearrange them to add value. Alkylation, for example, makes gasoline components by combining some of the gaseous byproducts of cracking. The process, which essentially is cracking in reverse, takes place in a series of large, horizontal vessels and tall, skinny towers that loom above other refinery structures. Reforming uses heat, moderate pressure and catalysts to turn naphtha, a light, relatively low-value fraction, into high-octane gasoline components. Chevron's patented reforming process is called Rheniforming for the rheniumplatinum catalyst used.

Treatment: the finishing touch Back when the first refineries used to boil crude oil to get kerosene, they didn't have to worry about customer specifications or government standards. Today, however, a major portion of refining involves blending, purifying, fine-tuning and otherwise improving products to meet these requirements. To make gasoline, refinery technicians carefully combine a variety of streams from the processing units. Among the variables that determine the blend are octane level, vapour pressure ratings and special considerations, such as whether the gasoline will be used at high altitudes. Technicians also add performance additive and dyes that distinguish the various grades of fuel. Refining has come a long way since the oil boiling days of refning. By the time a gallon of gasoline is pumped into a car's tank, it contains more than 200 hydrocarbons and additives. All that changing of molecules pays off in a product that ensures smooth, high-performance driving.

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Reaction Kinetics A simple chemical reaction - the rearrangement of electrons and bonding partners - occurs between two small molecules. From understanding the kinetics of the reaction, and the equilibrium extent to which it can proceed, come applications: the network of reactions during combustion, the chain reactions that form polymers, the multiple steps in the synthesis of a complex pharmaceutical molecule, the specialized reactions of proteins and metabolism. Chemical kinetics is the chemical engineer's tool for understanding chemical change. A catalyst influences the reaction rate. Catalysts are sought for increasing production, improving the reaction conditions, and emphasizing a desired product among several possibilities. The challenge is to design the catalyst, to increase its effectiveness and stability, and to create methods to manufacture it. The chemical reactor should produce a desired product reliably, safely, and economically. In designing a reactor, the chemical engineer must consider how the chemical kinetics, often modified by catalysis, interacts with the transport phenomena in flowing materials. New microreactor designs are expanding the concept of what a reactor may do, how reactions may be conducted, and what is required to scale a process from laboratory to production.

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Crude Distillation Distillation is the first step in the processing of crude oil and it takes place in a tall steel tower called a fractionation column. The inside of the column is divided at intervals by horizontal trays. The column is kept very hot at the bottom (the column is insulated) but as different hydrocarbons boil at different temperatures, the temperature gradually reduces towards the top, so that each tray is a little cooler than the one below. The crude needs to be heated up before entering the fractionation column and this is done at first in a series of heat exchangers where heat is taken from other process streams which require cooling before being sent to rundown. Heat is also exchanged against condensing streams from the main column. Typically, the crude will be heated up in this way upto a temperature of 200 - 280 0C, before entering a furnace. As the raw crude oil arriving contains quite a bit of water and salt, it is normally sent for salt removing first, in a piece of equipment called a desalter. Upstream the desalter, the crude is mixed with a water stream, typically about 4 - 6% on feed. Intense mixing takes place over a mixing valve and (optionally) as static mixer. The desalter, a large liquid full vessel, uses an electric field to separate the crude from the water droplets. It operates best at 120 - 150 0C, hence it is conveniently placed somewhere in the middle of the preheat train.

Part of the salts contained in the crude oil, particularly magnesium chloride, are hydrolysable at temperatures above 120 0C. Upon hydrolysis, the chlorides get converted into hydrochloric acid, which will find its way to the distillation column's overhead where it will corrode the overhead condensers. A good performing desalter can remove about 90% of the salt in raw crude.

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Downstream the desalter, crude is further heated up with heat exchangers, and starts vapourising, which will increase the system pressure drop. At about 170 -200 0C, the crude will enter a 'pre-flashvessel', operating at about 2 - 5 barg, where the vapours are separated from the remaining liquid. Vapours are directly sent to the fractionation column, and by doing so, the hydraulic load on the remainder of the crude preheat train and furnace is reduced (smaller piping and pumps). Just upstream the preflash vessel, a small caustic stream is mixed with the crude, in order to neutralise any hydrochloric acid formed by hydrolysis. The sodium chloride formed will leave the fractionation column via the bottom residue stream. The dosing rate of caustic is adjusted based on chloride measurements in the overhead vessel (typically 10 - 20 ppm). At about 200 - 280 0C the crude enters the furnace where it is heated up further to about 330 370 0C. The furnace outlet stream is sent directly to the fractionation column. Here, it is separated into a number of fractions, each having a particular boiling range. At 350 0C, and about 1 barg, most of the fractions in the crude oil vapourise and rise up the column through perforations in the trays, losing heat as they rise. When each fraction reaches the tray where the temperature is just below its own boiling point, it condenses and changes back into liquid phase. A continuous liquid phase is flowing by gravity through 'downcomers' from tray to tray downwards. In this way, the different fractions are gradually separated from each other on the trays of the fractionation column. The heaviest fractions condense on the lower trays and the lighter fractions condense on the trays higher up in the column. At different elevations in the column, with special trays called draw-off trays, fractions can be drawn out on gravity through pipes, for further processing in the refinery. At top of the column, vapours leave through a pipe and are routed to an overhead condenser, typically cooled by air fin-fans. At the outlet of the overhead condensers, at temperature about 40 0C, a mixture of gas, and liquid naphtha exists, which is falling into an overhead accumulator. Gases are routed to a compressor for further recovery of LPG (C3/C4), while the liquids (gasoline) are pumped to a hydrotreater unit for sulfur removal. A fractionation column needs a flow of condensing liquid downwards in order to provide a driving force for separation between light and heavy fractions. At the top of the column this liquid flow is provided by pumping a stream back from the overhead accumulator into the column. Unfortunately, a lot of the heat provided by the furnace to vapourise hydrocarbons is lost against ambient air in the overhead fin-fan coolers. A clever way of preventing this heat lost of condensing hydrocarbons is done via the circulating refluxes of the column. In a circulating reflux, a hot side draw-off from the column is pumped through a series of heat exchangers (against crude for instance), where the stream is cooled down. The cool stream is sent back into the column at a higher elevation, where it is been brought in contact with hotter rising vapours. This provides an internal condensing mechanism inside the column, in a similar way as the top reflux does which is sent back from the overhead accumulator. The main objective of a circulating reflux therefore is to recover heat from condensing vapours. A fractionating column will have several (typically three) of such refluxes, each providing sufficient liquid flow down the corresponding section of the column. An additional advantage of having circulating refluxes is that it will reduce the vapour load when going upwards in the column. This provided the opportunity to have a smaller column diameter for top sections of the tower. Such a reduction in diameter is called a 'swage'.

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The lightest side draw-off from the fractionating column is a fraction called kerosene, boiling in the range 160 - 280 0C, which falls down through a pipe into a smaller column called 'sidestripper'. The purpose of the side stripper is to remove very light hydrocarbons by using steam injection or an external heater called 'reboiler'. The stripping steam rate, or reboiled duty is controlled such as to meet the flashpoint specification of the product. Similarly to the atmospheric column, the side stripper has fractionating trays for providing contact between vapour and liquid. The vapours produced from the top of the side stripper are routed back via pipe into the fractionating column. The second and third (optional) side draw-offs from the main fractionating column are gasoil fractions, boiling in the range 200 - 400 0C, which are ultimately used for blending the final diesel product. Similar as with the kerosene product, the gasoil fractions (light and heavy gasoil) are first sent to a side stripper before being routed to further treating units. At the bottom of the fractionation column a heavy, brown/black coloured fraction called residue is drawn off. In order to strip all light hydrocarbons from this fraction properly, the bottom section of the column is equipped with a set of stripping trays, which are operated by injecting some stripping steam (1 - 3% on bottom product) into the bottom of the column.

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Catalytic Cracking

Introduction Already in the 30's it was found that when heavy oil fractions are heated over clay type materials, cracking reactions occur, which lead to significant yields of lighter hydrocarbons. While the search was going on for suitable cracking catalysts based on natural clays, some companies concentrated their efforts on the development of synthetic catalyst. This resulted in the synthetic amorphous silica-alumina catalyst, which was commonly used until 1960, when it was slightly modified by incorporation of some crystalline material (zeolite catalyst). When the success of the Houdry fixed bed process was announced in the late 1930s, the companies that had developed the synthetic catalyst decided to try to develop a process using finely powdered catalyst. Subsequent work finally led to the development of the fluidised bed catalytic cracking (FCC) process, which has become the most important catalytic cracking process. Originally, the finely powdered catalyst was obtained by grinding the catalyst material, but nowadays, it is produced by spray-drying a slurry of silica gel and aluminium hydroxide in a stream of hot flue gases. Under the right conditions, the catalyst is obtained in the form of small spheres with particles in the range of 1-50 microns. When heavy oil fractions are passed in gas phase through a bed of powdered catalyst at a suitable velocity (0.1-0.7m/s), the catalyst and the gas form a system that behaves like liquid, i.e. it can flow from one vessel to another under the influence of a hydrostatic pressure. If the gas velocity is too low, the powder does not fluidise and it behaves like a solid. If velocity is too high, the powder will just be carried away with the gas. When the catalyst is properly fluidised, it can be continously transported from a reactor vessel, where the carcking reactions take place and where it is fluidised by the hydrocarbon vapour, to a regenerator vessel, where it is fluidised by the air and the products of combustion, and then back to the reactor. In this way the proces is truly continous. The first FCC unit went on stream in Standard Oil of New Jersey's refinery in Baton Rounge, Louisiana in May 1942. Since that time, many companies have developed their own FCC process and there are numerous varieties in unit configuration.

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Figure 40 Fluid catalytic cracking

FCC Process Configuration : Hot feed, together with some steam, is introduced at the bottom of the riser via special distribution nozzles. Here it meets a stream of hot regenerated catalyst from the regenerator flowing down the inclined regenerator standpipe. The oil is heated and vapourised by the hot catalyst and the cracking reactions commence. The vapour, initially formed by vapourisation and successively by cracking, carries the catalyst up the riser at 10-20 m/s in a dilute phase. At the outlet of the riser the catalyst and hydrocarbons are quickly separated in a special device. The catalyst (now partly deactivated by deposited coke) and the vapour then enter the reactor. The vapour passes overhead via cyclone separator for removal of entrained catalyst before it enters the fractionator and further downstream equipment for product separation. The catalyst then descends into the stripper where entrained hydrocarbons are removed by injection of steam, before it flows via the inclined stripper standpipe into the fluidised catalyst bed in the regenerator. Air is supplied to the regenerator by an air blower and distributed throughout the catalyst bed. The coke deposited is burnt off and the regenerated catalyst passes down the regenerator standpipe to the bottom of the riser, where it joins the fresh feed and the cycle recommences. The flue gas (the combustion products) leaving the regenerator catalyst bed entrains catalyst particles. In particular, it entrains "fines", a fine dust formed by mechanical rubbing of catalyst particles taking place in the catalyst bed. Before leaving the regenerator, the flue gas therefore passes through cyclone separators where the bulk of this entrained catalyst is collected and returned to the catalyst bed. Normally modern FCC is driven by an expansion turbine to mimimise energy consumption. In this expansion turbine, the current of flue gas at a pressure of about 2 barg drives a wheel by striking impellers fitted on this wheel. The power is then transferred to the air blower via a common shaft. This system is usually referred to as a "power recovery system". To reduce the wear caused by the impact of catalyst particles on the impellers (erosion), the flue gas must be virtually free of catalyst particles. The flue gas is therefore passed through a vessel containing a whole battery of small, highly efficient cyclone separators, where the remaining catalyst fines are collected for disposal. Before being disposed of via a stack, the flue gas is passed through a waste heat boiler, where its remaining heat is recovered by steam generation. In the version of the FCC process described here, the heat released by burning the coke in the regenerator is just sufficient to supply the heat required for the riser to heat up, vapourise and crack the hydrocarbon feed. The units where this balance occurs are called " heat balanced" units. Some feeds caused excessive amounts of coke to be deposited on the catalyst, i.e. much more than is required for burning in the regenerator and to have a "heat balanced" unit. In such cases, heat must be removed from the regenerator, e.g. by passing water through coils in the regenerator bed to generate steam. Some feeds cause so little coke to be deposited on the catalyst that heat has to be supplied to the system. This is done by preheating the hydrocarbon feed in a furnace before contacting it with the catalyst.

Main Characteristics

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• • •

•

•

A special device in the bottom of the riser to enhance contacting of catalyst and hydrocarbon feed. The cracking takes place during a short time (2-4 seconds) in a riser ("short-contact time riser") at high temperatures ( 500-540 0C at riser outlet). The catalyst used is so active that a special device for quick separation of catalyst and hydrocarbons at the outlet of the riser is required to avoid undesirable cracking after the mixture has left the riser. Since, no cracking in thereactor is required, the reactor no longer functions as a reactor; it merely serves as a holding vessel for cyclones. The regenerator takes place at 680-720 0C. With the use of special catalysts, all the carbon monoxide (CO) in the flue gas is combusted to carbon dioxide (CO2) in the regenerator. Modern FCC includes a power recovery system for driving the air blower.

Equipment in FCC • • • • • •

Large storage vessels for catalyst (fresh and equilibrium) Regenerator Reactor Main Fractionator Product Work Up section (several distillation columns in series Product treating facilities

Feedstock & Yield Before the introduction of residues, vacumn distillates were used as feedstock to load the Catalytic Cracker fully. These days, even residues are used to load the cracker. The term used for this type of configuation is Long Residue Catalytic Cracking Complex. The only modification or addition needed are a residue desalter and a bigger and more heat resistent reactor. The yield pattern of an FCC unit is typically as follows:

Product C3 & C4 Gasoline Heavy Gas Oil Coke

% wt on fresh feed 15 40-50 10 5

Conclusion

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The FCC Unit can a real margin improver for many refineries. It is able to convert the residues into high value products like LPG , Butylene, Propylene and Mogas together with Gasoil. The FCC is also a start for chemical production (poly propylene). Many FCC's have 2 modes: a Mogas mode and a Gasoil mode and FCC's can be adapted to cater for the 2 modes depending on favourabale economic conditions. The only disavantage of an FCC is that the products produced need to be treated (sulfur removal) to be on specification. Normally Residue FCCs act together with Residue Hydroconversion Processes and Hydrocrackers in order to minimise the product quality give away and get a yield pattern that better matches the market specifications. Via product blending, expensive treating steps can be avoided and the units prepare excellent feedstock for eachother: desulfurised residue or hydrowax is excellent FCC feed, while the FCC cycle oils are excellent Hydrocracker feed. In the near future, many refiners will phase the challenge how to desulfurise cat cracked gasoline without destroying its octane value. Catalytic destillation appears to be one of the most promising candidate processes for that purpose.

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Catalysis Catalysts and initiators start or promote chemical reactions that are used to produce organic chemicals, polymers and adhesives. A chemical catalyst is a substance that increases the rate at which a chemical reaction occurs; however, the catalyst itself does not undergo chemical change. An initiator is a chemical compound that helps start a chemical reaction such as polymerization. Unlike a catalyst, an initiator is usually consumed in the reaction. Substances such as organic peroxides are commonly used as initiators. According to some estimates, more than half of all petrochemical processes use catalysts and initiators. In heterogeneous catalysis, a chemical catalyst provides a surface on which reactants become adsorbed temporarily, and where chemical bonds in the reactants are weakened, allowing new bonds to be created. Because the bonds between the products and the catalyst are weaker, the products are released from the chemical catalyst. Continuous process catalysts (CPC) are used to process industrial chemicals such as solvents, plasticizers, monomers and intermediates. Catalytic solutions include a variety of specialized catalyst products. There are two basic types of catalysis: homogeneous catalysis, in which both the catalyst and reactants are in the same phase (for example, liquid or gas), and heterogeneous catalysis, in which the catalyst and reactants are in different phases (for example, solid catalysts and gaseous reactants). Metal catalysts and initiators are made from precious metals such as gold, iridium, osmium, palladium, platinum, rhodium, ruthenium and silver. They are used as heterogeneous catalysts for reactions such as hydrogenation and isomerization. Zeolites, minerals with a porous structure, can also be used as catalysts. Synthetic zeolites are the most important catalysts in petrochemical refineries. The proper selection of catalysts and initiators is an important consideration. For example, using rhodium or platinum as catalysts can produce different products depending on whether methane or ethane are used. ASTM International (formerly called the American Society for Testing and Materials (ASTM), maintains standards for catalysts such as ASTM D3766, standard terminology relating to catalysts and initiators. Some catalysts and initiators must be handled as hazardous materials. The National Fire Protection Association (NFPA) maintains NFPA 432, a standard which covers the catalyst organic peroxide.

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Catalysis And Distillation

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Distillation and Other Separation Processes • • • • • • •

Distillation basics Phase behavior and vapour/liquid equilibria Gas/Liquid separation Trays: function, pressure drop, efficiency, flooding, operations, and damage Bubble and dew points: calculation and application Foam: formation, detection, cause Packed v. trayed columns

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Distillation basics

ATMOSPHERIC DISTILLATION

The Essence of Atmospheric Distillation The purpose of atmospheric distillation is to recover light materials and fractionate into sharp light fractions. This is accomplished by distilling at atmospheric pressure with steam stripping for improved cuts. Atmospheric distillation is historically the oldest refining process and is the first step in crude oil processing.

The Development of Crude Distillation Oil refining was sufficiently advanced by the 1500s that feed preheat, reflux, reboiling, and temperature regulation were all utilized for distillation. But it was the idea of Samuel Kier, a pharmacist from Tarentum Pennsylvania, that oil sludge from his father’s salt wells could be therapeutic. A suggestion from J.G. Booth, a Philadelphia chemist, led to Kier using distillation for purification of his “Rock Oil” cure around 1846. A simple batch still to improve the color was devised by a whiskey distiller using first a one-barrel and then a five-barrel iron kettle. A second stage was added to remove hydrogen sulfide. Initially Kier sold the distilled oil as bottled medicine but soon Kier was distilling petroleum for illumination purposes. Kerosene was born. He called it Kier's Carbon Oil. The typical refinery was a vertical pot, many taken from coal tar distillation facilities that were displaced by entry of kerosene as an illuminant. The overhead from the pot was condensed in a water box with coils submerged in water that was pumped in and flowed out by gravity to a cooling pond. It was deemed a radical improvement when the still was laid on its side and made of thinner metal, thereby increasing heat transfer.

Figure 41 Early Batch Fractionation

Feeds and Products for Atmospheric Distillation The feeds and products are illustrated in the Refinery Schematic and on the figure of the atmospheric tower.

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Feed Preheat Exchanger Train Since crude oil is elevated from atmospheric liquid temperature to over 700°F, recovery of heat is of prime importance in crude distillation economics. The path of crude being preheated is typical. The crude is first used to absorb part of the overhead condensation load after which there is exchange with one or more of the liquid sidestreams withdrawn, beginning with the top sidestream. The crude desalting is placed within the feed preheat exchanger train. The point at which the crude is desalted is carefully selected. Normally it is at a temperature of 250°F to 300°F and is a function of the gravity of the crude, with lighter crudes (lighter than 40°API) being desalted at 250°F and those heavier than 30°API being desalted at 300°F. Care must be taken in the temperature and pressure balance of the heat exchanger train that water does not vapourize.

Crude Electrostatic Desalting Nearly all crudes contain "salts" the concentration of which is expressed as pounds of sodium chloride per thousand barrels of crude. Other chlorides such as magnesium chloride are also present. The salt is present in the emulsified water in the crude. Salt water pumped to the surface with the crude is settled in facilities in the field and is treated with heat and chemicals to break oil water emulsions. Nonetheless crude arrives at the battery limits of the crude unit with emulsified salt water. Salt can lead to deposits on heat exchangers and drastically reduce heat transfer. Formation of hydrogen chloride by hydrolysis can lead to corrosion. In addition to salt the water contains metals in various compounds that can deposit on various catalysts in the refinery. Water washing is employed to remove the dissolved salt and salt crystals plus the dissolved metals and dirt from the oil. The untreated oil is mixed with fresh wash water, demulsifiers are added and the streams are mixed and heated and subjected to additional mixing, followed by settling. The oil, salt water, demulsifier and wash water mixture is separated in an electrostatic settling drum which uses a high voltage electric field across the drum to promote coalescing of water into droplets which collect in a water layer at the bottom of the settler. The quantity of emulsified water in the crude is variable but the added wash water can be as much as 10% of the crude oil charge. About 90% of the water can be recovered. If that frequently is not sufficient and a second stage of desalting is employed, the water recovery is raised to 99%.

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Figure 42 Desalting - single stage

Figure 43 Desalting - 2 stage

The wash water also washes out sediment such as fine clay, rust and other solids. Incomplete settling of sediments in production tanks causes more solids to be moved with the product for removal at the refinery. Normal efficient water washing will remove over half of these suspended solids. It should be noted that crude desalting may be difficult to operate because of the variability of emulsions. In addition the wash water now must be treated for benzene recovery. After desalting, crude is exchange with hotter sidestream liquids further down in the tower.

Crude Unit Furnace and Overflash Generalizations about the furnace vary according to the crude and the refiner’s experience. The following are rough guidelines.

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Figure 44 Crude unit furnace

Crude Furnace Duty After further heat exchange in the train, to raise the temperature to about 550°F, the crude oil is directed to the furnace, whose tubular configuration led to the term pipe still. Crude unit furnaces can be fired with oil, refinery fuel gas or natural gas. They can be a box or cabin furnace, usually with horizontal tubes or a cylindrical vertical tube furnace. Heat flux in these furnaces is not excessively high at 10,000 btu/hr-ft2 and coking is ordinarily not a problem if the desalter unit is operating properly. Since the atmospheric tower does not have a reboiler, the heat content of the furnace supports the total vapour rate to the column plus additional duty called overflash. Heat removal in the tower is accomplished by condensation of vapour with liquid cooled in pumparounds. Depending on the crude slate, perhaps half or more of the crude is flashed. Heater Outlet and Transfer Line Heater outlet temperature is limited to approximately 750°F by thermal cracking of the feedstock, which impairs distillate product smoke points and color. Depending on the crude, this temperature may range from 700°F to 800°F. But cracking of paraffinic and naphthenic crudes occurs at approximately 650°F - 700°F. The outlet from the furnace is directed to the flash zone in the fractionation tower via a transfer line. The pressure drop in the transfer line between the furnace outlet and inlet to the tower is assumed to be 5 psid and the temperature loss roughly corresponds to 5°F - 10°F.

Overflash

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The furnace is normally operated to produce overflash. Overflash is defined as vapourization in excess of requirements for lifting all of the products taken overhead and withdrawn as side-streams. The purpose of overflash is to generate internal reflux in the wash trays between the flash zone and the bottom side-stream draw tray. Overflash vapours are condensed and wash the trays to prevent carryover and coking. Overflash is generally 3% to 5% (volume) of gross vapour from the flash zone which is essentially overhead and sidestream products. Overflash is also defined in terms of crude charge to tower and is 2% to 3% (volume) on that basis. Although not shown on the schematic, when processing high vapour pressure crudes, a flash drum is placed in the train after the desalter and before the furnace inlet control valves. Flashing off light gases and water lowers the feed vapour pressure to avoid flashing of the crude before the control valves, which leads to maldistribution in the furnace.

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Atmospheric Crude Fractionator

Flash Zone and Stripping Section The flash zone pressure is set as low as possible to maximize vapourization, minimize flash zone temperature, and reduce furnace duty while optimizing compression on the tower overhead vapour stream. Flash zone pressure is determined by overhead condensation pressure plus pressure drop in the tower. If the reflux drum operates at 5 psig, the pressure drop across the overhead cooling system is 5 psid, and the pressure drop through the tower is 5 psid, the flash zone will operate at 15 psig. The temperature in the flash zone is a function of the onset of crude cracking, the pressure of the flash zone, and the amount of stripping steam. With a maximum furnace outlet temperature limit of 750°F, and transfer line temperature loss of about 10°F, the maximum bulk temperature in the flash zone is about 740°F. There is a 4 tray stripping section below the flash zone. Stripping steam is added to lower the partial pressure of the hydrocarbon and increase vapourization at lower temperatures. Traditionally, the amount of stripping steam has been 10 lb of steam per barrel of atmospheric resid. However, stripping steam must be recovered as sour water and designs to minimize stripping steam use 5 lb of steam per barrel of atmospheric resid

Wash Section The wash section consists of 3 to 4 trays above the flash zone and below the bottom gas oil draw. The purpose of the wash section is to provide reflux to the vapours from the flash zone to wash resins and materials that may contaminate the products. The reflux is the condensed overflash vapour. Either sieve trays or grid are utilized. Overhead System, Number of Trays and Pressure Profile in Tower The overhead vapours from the tower are cooled and partially condensed by exchange with cold feed followed by condensation with air fin or water condensers. The vapours if any are directed from the overhead accumulator to the fuel system. Operating pressures have increased from the 1970s to be high enough to reduce noncondensing vapours to a minimum. The intent is to reduce compression on the overhead system. On the other hand, high operating pressures decrease vapourization, increase flash zone temperatures and furnace duty, and affect yields. Pressures in the reflux drum may vary according to the design and be as low as 0.5 psig to as high as 20 psig if the overhead vapour is totally condensed. This discussion will use a reflux drum pressure of 5 psig as a basis. The pressure drop across the overhead condensers is also variable but is on the order of 3 – 10 psid; this discussion will assume 5 psid. Pressure drops across trays average 0.1 psid/tray to 0.2 psid/tray. This discussion will assume a total of 32 trays above the flash zone including 4 wash trays, resulting in a 5 psid pressure drop in the tower from the flash zone to the top of the column. Assumed are 6 trays below the reflux to the naphtha draw, 6 trays below to the light distillate draw, 6 trays below to the heavy distillate draw, and 10 trays below to the gas oil draw. There are 4 wash trays below the gas oil draw. Trays may be either sieve trays or grid.

Sidestreams

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Liquid product from the overhead is straight run gasoline, a part of which is returned to the tower as reflux for the top section of the tower. There is a pumparound on the fractionator where liquid is taken from a draw tray and cooled and returned to the next tray down as subcooled reflux. This not only reduces the overhead condensing load but achieves uniform tower loadings by providing overflow at lower points in the tower since successive product withdrawal reduces liquid overflow. The cut point for each sidestream fraction is the final boiling point of the stream being withdrawn. However, the liquid has a lighter component tail that must be removed from the sidestream. Traditionally, the liquid product sidestream was directed to a stripper that used live steam for stripping. Emphasis on reducing steam stripping and sour water has led to the replacement of live steam injection with reboiled strippers in some instances.

Trends and Variations in Atmospheric Unit Design Briefly summarized are several approaches to crude unit design that represent current thinking about stream separation. Each offers distillation combinations to provide additional separations, additional capacity or energy conservation: •

A combined gasoline and naphtha stream is desulfurized in one hydrotreating unit before routed to a splitter to separate the streams for further processing.

•

Benzene is removed from the naphtha stream by a superfractionator after the naphtha hydrotreater and before the catalytic reformer. Benzene goes overhead and the naphtha bottoms is routed to the reformer.

•

Three-step fractionation permits a crude unit capacity expansion with production of LPG, isopentane, straight run and naphtha streams from the crude unit overhead.

•

The Technip Progressive Distillation process minimizes energy consumption.

•

Shell's Bulk CDU Integrated Processing provides for full integration of units in the refinery.

Combined Gasoline and Naphtha Stream One trend is the desulfurization of straight run gasoline and naphtha in one hydrotreater unit, followed by a precise fractionation of the desulfurized materials at a cut point of about 180°F. The overhead gasoline containing benzene and cyclohexane go to the isomerization unit where the benzene is hydrogenated to cyclohexane, cycloparaffins are isomerized to olefins, and normal paraffins are isomerized to iso-paraffins for octane improvement. This is usually performed in a two-reactor system. The splitter bottoms naphtha, which contains no benzene or cyclohexane, goes to reforming. This is desirable if there is no facility for extraction or hydrogenation of benzene to eliminate benzene from gasoline. Additionally, both gasoline and naphtha are also desulfurized and the atmospheric tower is simplified. However, a larger hydrotreater plus splitter are required.

Removal of Aromatics Before Reforming To reduce the amount of aromatics to the reformer and ultimately to the gasoline stock, benzene is removed from the naphtha stream by a superfractionator after the naphtha hydrotreater and before the catalytic reformer. Benzene goes overhead and the naphtha bottoms is routed to the reformer.

Three Step Fractionation

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Three-step fractionation permits a crude unit capacity expansion with production of LPG, isopentane, straight run gasoline, and naphtha streams from the crude unit overhead. A combined stream of straight run gasoline and naphtha will be taken overhead on the crude tower. The liquids are then fractionated in a splitter to make naphtha for reformer feed as the splitter bottoms. The splitter overhead goes to a stabilizer. The overhead from the stabilizer is LPG is directed to the saturates gas plant. The stabilizer bottoms goes to a deisopentanizer that produces isopentane overhead and straight run gasoline bottoms.

Progressive Distillation for Energy Reduction In this arrangement, crude is first desalted then prefractionated in a reboiled preflash tower with the bottoms being fractionated in a second preflash tower employing both a reboiler and steam. The two preflash overheads are then processed in a gas plant and precision fractionators to produce LPG, Light Naphtha, Medium Naphtha, and Heavy Naphtha. The bottoms from the second preflash goes to a conventional heater and atmospheric column arrangement that produces four distillate streams. Atmospheric tower bottoms go to conventional vacuum distillation for production of gas oils and vacuum resid. This is a grass roots design offered by Technip for minimization of energy consumption.

The Shell Bulk Crude Distillation (CDU) Process This process combines and integrates crude distillation with hydrodesulfurization and high vacuum separation. It also incorporates catalytic cracking, hydrocracking and visbreaking with the separation processes. The basic concept is first to separate naphtha / gasoline overhead and long resid bottoms from the distillate midfractions in a crude distillation. The bulk middle distillate fraction is desulfurized and cut into distillate streams in a superfractionator. Overhead naphtha / gasoline from the crude unit and streams from various other plants are gathered and processed in light hydrocarbon and light oil fractionators to make LPG, light gasoline and several naphtha fractions. Bottoms resid from the crude tower is processed in a high vacuum unit that is integrated with cracking processes. The flow sheet is quite complex and varies with each design.

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Phase behavior and vapour/liquid equilibrium

Single Phase, Non-Saturated Equilibrium States For the simplest cases, a two component mixture existing as either a superheated vapour or a subcooled liquid, Gibbs's phase rule shows that now three independent properties must be known to establish the intensive state of the system; i.e. temperature, pressure, and concentration. Using the Patel-Teja equation of state modified for a mixture there are as many as three numerical solutions at a given temperature, pressure, and composition, but only the solutions using the compressibility corresponding to the vapour (largest positive real Z) or liquid (smallest positive real Z) apply. Equations 2-8 through 2-22 and 2-39 through 242 provide an equal number of equations and unknowns.

Vapour-Liquid Equilibrium and Single Phase, Saturated Equilibrium States When two mixture phases are in equilibrium, Gibbs's phase rule requires only two independent, intensive properties to define the state. For example, if a mixture exists as a saturated liquid and vapour in equilibrium at a fixed temperature and pressure, Gibb's phase rule finds that the compositions in each phase are defined by the two independent properties of temperature and pressure. However, mathematically, the Patel-Teja equation of state has solutions at many vapour and liquid compositions for a given temperature and pressure. Equating the fugacity for each component in the liquid to its respective fugacity in the vapour provide the two additional equations that are needed to solve for the vapour phase composition and the liquid phase composition. (2-45) This equation allows the calculation of temperature composition diagrams and pressure composition diagrams for two component mixtures. The Figure below displays a temperature composition diagram constructed using the Patel-Teja equation of state and equation 2-45 for the ammonia-butane system at 20.7 bar (300 psi). The lower line represents the bubble point temperature line. For example, when a subcooled 50/50 liquid mixture of ammonia and butane is heated from 300K (point 1), the first bubble of vapour will form at just above 316 K (point 2) and the mixture will be in vapour-liquid equilibrium. Upon further heating, the overall composition of both the vapour and the mixture will remain 50/50, but the compositions of the liquid and vapour phases will vary (see point 3). Finally, as the last drop of liquid is evapourated, the 50/50 mixture is now a saturated vapour (point 4). Further heating will superheat the vapour (point 5). Point 4 lies on the dew point line so named because this is where the first drop of condensation would form if it was approached via cooling from point 5. Note at a temperature around 316 K, the equilibrium vapour and liquid phases are at the same composition (~0.82). At this point, known as an azeotrope, the azeotropic mixture boils at a constant temperature with constant vapour and liquid phase compositions (similar to a pure substance). It is important to note that there may be two compositions required when making vapour liquid equilibrium (VLE) calculations. At a saturated temperature and pressure in a mixture there are two cubic equations for the compressibility; one for the liquid and one for the vapour. The liquid compressibility cubic equation is obtained by evaluating equations 2-39 through 2-41 using liquid compositions. The vapour compressibility cubic equation is obtained by evaluating equations 2-39 through 2-41 using vapour compositions. The liquid's compressibility is still the smallest positive real root but of the liquid compressibility equation. The vapour's compressibility is the largest positive real root of the vapour compressibility equation.

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Figure 45 Temperature-Composition Diagram for Ammonia-Butane at 20.7 bar

Liquid-Liquid Equilibrium States As shown above, a two component mixture in a single, sub-cooled liquid phase needs three independent, intensive properties to define its equilibrium state. For example, in the ammonia-water mixture at 4 bar and 260 K, an equilibrium state exists at all compositions. For mixtures displaying only one liquid phase, the Patel-Teja equation of state has only one solution for a given set of three independent, intensive properties. However, some mixtures display liquid-liquid equilibrium (LLE) for which two liquid phases are in equilibrium with each other. In such mixtures, the phase rule requires only two independent, intensive properties. In other words, at a given temperature and pressure, the compositions of the two liquid phases are actually dependent properties. For example, in the ammonia-butane system at 5.29 bar and 273.15 K, a two liquid phase equilibrium state exists in which the composition of butane is 0.8526 mole fraction in one of the liquid phases and 0.106 mole fraction in the other (Wilding, 1996). The equilibrium condition for liquid-liquid equilibrium is: (2-46)

Vapour-Liquid-Liquid Equilibrium States When a mixture displaying liquid-liquid equilibrium is heated to its bubble point a third phase, vapour, comes into equilibrium with the two liquid phases. This point is known as the three phase flash point. According to Gibbs's phase rule, this state is dictated by only one independent, intensive property. The equilibrium condition for the three phase vapour-liquidliquid (VLLE) equilibrium state is: (2-47) Calculations produce three different cubic equations for the compressibility since there are three sets of compositions. The compressibility of each liquid is represented by the smallest positive real root of each liquid's compressibility cubic equation while the compressibility of the vapour is the largest positive real root of the vapour's cubic compressibility equation.

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Figure 46 T-x-y Diagram for Ammonia-Butane at 20.7 bar

Figure 47 T-x-y Diagram for Ammonia-Butane at 4, 10, and 20.7 bar

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Gas/Liquid separation Recent Development In Liquid/Gas Separation Technology Removing liquids and solids from a gas stream is very important in refining and gas processing applications. Effective removal of these contaminants can prevent costly problems and downtime with downstream equipment like compressors, turbines, and burners. In addition, hydrocarbons and solid contaminants can induce foaming in an amine contactor tower and can contribute to premature catalyst changeouts in catalytic processes. In compressors that use oil to lubricate cylinders, the lube oil often gets into the discharge gas causing contamination downstream. A thin film of hydrocarbon deposited on heat exchangers will thicken and coke, decreasing heat transfer efficiency, increasing energy consumption and creating a risk of hot spots and leaks. Several technologies are available to remove liquids and solids from gases. This paper will first provide selection criteria for the following gas/liquid separation technologies: •

gravity separators

•

centrifugal separators

•

filter vane separators

•

mist eliminator pads

•

liquid/gas coalescers

and then focus on the separation of fine aerosols from gases using liquid/gas coalescing technology. •

Removal Mechanisms

•

Liquid/Gas Separation Technologies

•

Formation of Fine Aerosols

•

Ratings/Sizing

•

Design and Its Impact on Sizing

•

Field Testing For Liquid/Gas Coalescers

•

Test Procedure

•

Field Test Results

•

Conclusions

Removal Mechanisms Before evaluating specific technologies, it is important to understand the mechanisms used to remove liquids and solids from gases. These can be divided into four different categories.2 The first and easiest to understand is gravity settling, which occurs when the weight of the droplets or particles (ie. the gravitation force) exceeds drag created by the flowing gas. A related and more efficient mechanism is centrifugal separation which occurs when the centrifugal force exceeds the drag created by the flowing gas. The centrifugal force can be several times greater than gravitational force.

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The third separation mechanism is called inertial impaction which occurs when a gas passes through a network, such as fibers and impingement barriers. In this case, the gas stream follows a tortuous path around these obstacles while the solid or liquid droplets tend to go in straighter paths, impacting these obstacles. Once this occurs, the droplet or particle loses velocity and/or coalesces, and eventually falls to the bottom of the vessel or remains trapped in the fiber medium. And finally, a fourth mechanism of separation occurs with very small aerosols (less than 0.1 µm). Called diffusional interception or Brownian Motion, this mechanism occurs when small aerosols collide with gas molecules. These collisions cause the aerosols to deviate from the fluid flow path around barriers increasing the likelihood of the aerosols striking a fiber surface and being removed.3 Throughout this paper, reference to droplet and particle sizes will be in the unit micron. One micron is 1/1000 of a millimeter or 39/1,000,000 of an inch. Figure 1 shows the size of various material in microns.

Figure 48 Particle Diameters of Typical Contaminants

Liquid/Gas Separation Technologies Gravity Separators In a gravity separator or knock-out drum, gravitational forces control separation. The lower the gas velocity and the larger the vessel size, the more efficient the liquid/gas separation. Because of the large vessel size required to achieve settling, gravity separators are rarely designed to remove droplets smaller than 300 microns.4 A knock-out drum is typically used for bulk separation or as a first stage scrubber. A knock-out drum is also useful when vessel internals are required to be kept to a minimum as in a relief system or in fouling service.5 Gravity separators are not recommended as the soul source of removal if high separation efficiency is required. Centrifugal Separators In centrifugal or cyclone separators, centrifugal forces can act on an aerosol at a force several times greater than gravity. Generally, cyclonic separators are used for removing aerosols greater than 100 µm in diameter and a properly sized cyclone can have a reasonable removal efficiency of aerosols as low as 10 µm. A cyclone’s removal efficiency is very low on mist particles smaller than 10µm.6 Both cyclones and knock-out drums are recommended for waxy or coking materials. Mist Eliminators

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The separation mechanism for mist eliminator pads is inertial impaction. Typically, mist eliminator pads, consisting of fibers or knitted meshes, can remove droplets down to 1-5 microns but the vessel containing them is relatively large because they must be operated at low velocities to prevent liquid reentrainment. Filter Vane Separators Vane separators are simply a series of baffles or plates within a vessel. The mechanism controlling separation again is inertial impaction. Vane separators are sensitive to mass velocity for removal efficiency, but generally can operate at higher velocities than mist eliminators, mainly because a more effective liquid drainage reduces liquid reentrainment. However, because of the relatively large paths between the plates constituting the tortuous network, vane separator can only remove relatively large droplet sizes (10 microns and above). Often, vane separators are used to retrofit mist eliminator pad vessels when gas velocity exceeds design velocity.7 Liquid/Gas Coalescers Liquid/gas coalescer cartridges combine features of both mist eliminator pads and vane separators, but are usually not specified for removing bulk liquids. In bulk liquid systems, a high efficiency coalescer is generally placed downstream of a knock-out drum or impingement separator. Gas flows through a very fine pack of bound fibrous material with a wrap on the outer surface to promote liquid drainage (See Figure 2 below). A coalescer cartridge can trap droplets down to 0.1 micron. When properly designed and sized, drainage of the coalesced droplets from the fibrous pack allows gas velocities much higher than in the case of mist eliminator pads and vane separators with no liquid reentrainment or increase in pressure drop across the assembly.

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Figure 49 Coalescer Cut-away View

Table 2 summarizes each of these technologies and provides guidelines for proper selection. As you can see, for systems containing very fine aerosols, under 5 µm, a coalescer should be selected. Removing very fine aerosols from gases results in major economic, reliability, and maintenance benefits in compressor systems.

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Table 3 Types of Liquid/Gas Separators Technology

Droplet Size Removed

Gravity Separator

Down to 300µm

Centrifugal Separator

Down to 8-10µm

Mist Eliminator Pad

Down to 10µm

Vane Separator

Down to 10µm

High Efficiency L/G Coalescer

Down to 0.1µm

Formation of Fine Aerosols There are several different ways that very fine liquid aerosols can get into a gas stream. • •

Condensation from a saturated vapour,

•

Atomization (spray effect through a flow restriction) and,

•

Liquid reentrainment.

Recent studies on aerosol size distribution in a natural gas stream have identified that significant quantities of droplets below 5 microns are the norm whenever choke valves and other restrictions are present9 or when vapours are at their dew points.10 The measurements shown in Figure 3 were performed to determine concentration ofliquid aerosols in natural gas stream sampled downstream of vane separators (combination of gravity separator and horizontal filter barrier and equivalent to a mist eliminator pad). Results show that in many cases, large quantities of aerosols can go through this type of separator because the droplets are too small to be trapped by these separation devices. As a result, a liquid/gas coalescer should be the technology of choice whenever high recovery rates are required to protect downstream equipment or to recover valuable liquid products.

Figure 50 Aerosol Sizes

Ratings/Sizing

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It is important to note that a coalescer is different from a filter in that it performs both filtration of fine solid particles and coalescence of liquid aerosols from a gas stream. The sizing and rating criteria for coalescers, as it pertains to liquids removal, is very critical to the ultimate performance of the coalescer. An undersized coalescer will result in continuous liquid reentrainment, very low liquid separation efficiency and will be vulnerable to any process changes. The critical nature of coalescer sizing is illustrated in Figure 4 which shows that coalescer performance can drop very rapidly once the coalescer is challenged by too much liquid (either because of high aerosol concentration in the gas stream or because of a high gas flowrate). This marks a dramatic departure from most other separation equipment whose performance gradually diminishes as it is pushed passed its rated maximum.

Figure 51 Coalescer Efficiency Change vs. Gas Flow Rate

Traditional means of coalescer performance validation is the DOP (dioctylphthalate) test.11 In this test, a monodispersed aerosol of 0.3 µm diameter is continuously generated by a condensation of DOP vapour under controlled conditions. When aerosol generation is stabilized (constant particle size and aerosol concentration), the concentration of DOP is measured upstream and downstream of the coalescer by a light scattering photometer. Results are expressed as a percent of DOP penetration at the flow rate used. Some major drawbacks of the DOP test include:1 1. The test is performed on a dry or unsaturated cartridge. A dry cartridge, in essence, acts like a sponge, absorbing any liquid which goes through it. What the DOP test does not measure is the coalescer’s ability to retain liquids when liquids saturate the coalescer medium and could be re-entrained downstream. 2. This leads to a second drawback; the pressure drop measured across the assembly is underestimated when compared with actual pressure drops across a saturated element. The saturated DP is approximately 2-4 times greater than the clean DP. 3. The test is performed under a partial vacuum where gas properties (density and viscosity) are very different from those prevailing at actual operating pressure. DOP test conditions tend to overstate the efficiency of the coalescer element. In order to avoid shortcomings of the DOP test, Pall has developed the Liquid Aerosol Separation Efficiency (LASE) test. This test was developed solely for the purpose of measuring coalescer performance in a compressed gas stream under conditions more similar to those found in a refinery or a gas processing plant. The system used for this test is schematically represented in Figure 5. 5:

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Figure 52 Liquid Aerosol Separation Efficiency Test Schematic

The LASE test differs from the DOP test in the following ways: 1. It gives a more accurate and meaningful measure of efficiency. The DOP efficiency essentially tells you what percent of 0.3 µm dioctylphthalate droplets will be removed by a dry coalescer; the LASE test tells you what ppmw of contaminants will be in the gas downstream of the coalescer. In other words, what the LASE test tells you is how much contaminant your downstream equipment will be exposed to. 2. The DOP uses monodispersed (ie. same sized) droplets of DOP, a liquid not commonly encountered in a gas processing or refinery gas streams; the LASE test uses a lube oil which has droplet sizes that range from 0.1-0.9 µm. 3. The LASE test more closely simulates process conditions, by being run on a saturated cartridge and being performed under positive pressure. Table 4 comparison of the DOP and LASE

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Design and Its Impact on Sizing The goal for improving coalescer design is to maximize efficiency while preventing liquid reentrainment. Reentrainment occurs when liquid droplets accumulated on a coalescer element are carried off by the exiting gas. This occurs when velocity of the exiting gas, or annular velocity exceeds the gravitational forces of the draining droplet. We earlier discussed the importance of correct coalescer sizing. In designing and sizing a coalescer, the following parameters must be taken into account: •

Gas velocity through the media,

•

Annular velocity of gas exiting the media,

•

Solid and liquid aerosol concentration in the inlet gas, and

•

Drainability of the coalescer

Each of these factors with the exception of the inlet aerosol concentration can be controlled. At a constant gas flow rate, media velocity can be controlled by either changing the coarseness of the medium’s pore structure or by increasing or decreasing the number of cartridges used. The coarser the medium, however, the less efficient the coalescer will be at removing liquid. At a constant gas flow rate, the exiting velocity of the gas can be controlled by increasing or decreasing the size of the vessel or the space between the cartridges. Drainage can be improved by either selecting low surface energy coalescer materials or by treating the coalescer medium with a chemical that lowers the surface energy of the medium to a value lower than the surface tension of the liquid to be coalesced.13 Having a low surface energy material prevents liquid from wetting the filter medium and accelerates drainage of liquids down along the medium’s fibers. The liquid coalesced on the fibrous material falls rapidly through the network of fibers without accumulating in the pores where it would otherwise be pushed through by the gas and be reentrained. Figure 6 shows the effect that a chemical treatment can have on a coalescer. It shows that the maximum flowrate of a chemically treated cartridge is more than twice that of a similar cartridge that is not treated.

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Figure 53 Effect of Chemical Treatment on Coalescer Performance

One can conclude from these design parameters that a large housing with a large number of cartridges that have very fine pores would easily eliminate any liquid problems you may encounter in a gas stream. Obviously, the costs associated with such a vessel is very high. As vessel size and cartridge quantity are reduced so is the probability of reentrainment and poorer removal efficiency. In addition, as the assembly size decreases, the pressure drop increases which can result in increased operating costs. So, an optimization is required. When evaluating a coalescer assembly, make sure that all of these parameters are taken into consideration when the assembly is sized. A coalescer is best used in conjunction with a knock-out drum or other impingement separator.

Field Testing For Liquid/Gas Coalescers Field testing a gas stream where liquids need to be removed can provide the following information: 1. the amount of liquid in the gas, 2. the ability to efficiently coalesce liquids, and 3. the amount of solid particulate matter present. As a result, accurate sampling becomes critical. It is very important to measure accurately gas flow rates through a test coalescer cartridge to determine the amount and the nature of the liquid present in the gas. For that purpose, a complete test kit has been designed to perform side stream liquid/gas coalescer testing. This test kit is shown in Figure 7. It includes: (1) a coalescer housing for one cartridge connected to an independent sump by a small ball valve; (2) an orifice flowmeter downstream of the coalescer housing that includes flanges, orifice plate and differential pressure gauge; (3) a needle valve to regulate the flow of gas through the coalescer housing; (4) two sample ports, upstream and downstream of the coalescer housing, to which two of the gas test kits can be hooked up simultaneously to analyze influent and effluent gas quality; and (5) two long flexible stainless steel hoses connecting the test kit to the main gas line and the discharge line.

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Figure 54 Schematic of Pall LG Coalescer Test Stand

Test Procedure Before going on-site for a field test, the plant is contacted to obtain system conditions (pressure, temperature, gas flow rate, type of gas and if possible liquid concentration in the gas stream). Based on this information, an orifice plate is selected to measure gas flow rates in the range indicated. The orifice is also selected to minimize pressure drop so that gas condensation and hydrate formation is not induced. After putting the side stream test kit on-line, the flow rate is adjusted below the critical flow rate, so as not to get reentrained. Once the coalescer cartridge is saturated, test membranes are inserted in the test jigs upstream and downstream of the coalescer housing, the sump is emptied of any liquid that may have been accumulated during the cartridge saturation period, and the actual test begins. t the end of the test, the volume of liquid accumulated in the sump is measured and collected in a sample bottle for subsequent lab analysis. Test membranes are also collected to determine the amount of solids suspended in the gas and for qualitative identification of the solid contaminants. Liquid aerosol concentration is determined from the amount of liquid coalesced and the quantity of gas sampled.

Field Test Results The results of field tests on 49 gas streams (natural gas, carbon dioxide, hydrogen and fuel gas) in both gas processing plants and refineries show that significant quantities of liquid are present in most gas streams. Figure 8 summarizes these results of tests. Of the 49 streams tested, over 85% (43 out of 49 tests) had liquids concentration greater than 1 ppmw. This concentration of liquid can result in significant rotating equipment problems and can contribute to poor process operations in an amine contacting unit.

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Figure 55 Field Test Results of Gas Streams in Refineries and Gas Processing Plants

Conclusions 4. Selecting gas/liquid separation technologies requires not only knowledge of the process conditions, but a knowledge of the characteristics of the liquid contaminants. Selection should be made based on droplet size, concentration, and whether the liquid has waxing or fouling tendencies. 5. Through an analysis of field data, it was shown that due to the presence of very fine liquid droplets (below 1 micron) in most gas processes, high efficiency liquid/gas coalescers should be recommended whenever high recovery rates are required to protect downstream equipment or to recover valuable liquids. 6. The sizing and design of a coalescer is of critical importance. Once a coalescer is challenged with too much liquid, either because of excessive aerosol concentrations or large gas flow rates, its efficiency will decrease rapidly. 7. The Liquid Aerosol Separation Efficiency (LASE) test is a meaningful performance test of liquid/gas coalescers, as it allows coalescer cartridges to be tested under conditions closely resembling actual operating conditions (saturated element, realistic pressure drops and gas properties (density, viscosity). 8. A surface treatment of the coalescer medium improved liquid drainage in the fibrous materials and decreased by 50% the number of cartridges required to handle a given flow. 9. Field testing has demonstrated that significant amounts of liquids are present in gas stream in refinery and gas processing plants.

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Industrial uses of Fractional Distillation Distillation is the most common form of separation technology in the chemical industry. In most chemical processes, the distillation is continuous steady state, where batch fractionation is not as economical. New feed is always being added to the distillation column and products are always being removed. Unless the process is disturbed due to changes in feed, heat, ambient temperature, or condensing, the amount of feed being added and the amount of product being removed are normally equal. This is known as continuous, steadystate fractional distillation. The most widely used industrial applications of continuous, steady-state fractional distillation are in petroleum refineries, petrochemical plants and natural gas processing plants.

Figure 56 Typical distillation towers in oil refineries

Industrial distillation is typically performed in large, vertical cylindrical columns known as "distillation towers" or "distillation columns" with diameters ranging from about 65 centimeters to 6 meters and heights ranging from about 6 meters to 60 meters or more. The distillation towers have liquid outlets at intervals up the column which allow for the withdrawal of different fractions or products having different boiling points or boiling ranges. The "lightest" products (those with the lowest boiling point) exit from the top of the columns and the "heaviest" products (those with the highest boiling point) exit from the bottom of the column. Large-scale industrial towers also use reflux to achieve more complete separation of products. Fractional distillation is also used in air separation, producing liquid oxygen, liquid nitrogen, and high purity argon. Distillation of chlorosilanes also enable the production of high-purity silicon for use as a semiconductor. In industrial uses, sometimes a packing material is used in the column instead of trays, especially when low pressure drops across the column are required, as when operating under vacuum. This packing material can either be random dumped packing (1-3" wide) or structured sheet metal. Typical manufacturers are Koch, Sulzer and other companies. Liquids tend to wet the surface of the packing and the vapours pass across this wetted surface, where mass transfer takes place. Unlike conventional tray distillation in which every tray represents a separate point of vapour liquid equilibrium, the vapour liquid equilibrium curve in a packed column is continuous. However, when modeling packed columns it is useful to compute a number of "theoretical stages" to denote the separation efficiency of the packed column with respect to more traditional trays. Differently shaped packings have different surface areas and void space between packings. Both of these factors affect packing performance.

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Trays: function, pressure drop, efficiency, flooding, operations, and damage

Trays and Plates

The terms "trays" and "plates" are used interchangeably. There are many types of tray designs, but the most common ones are : Bubble cap trays A bubble cap tray has a riser or chimney fitted over each hole, and a cap that covers the riser. The cap is mounted so that there is a space between riser and cap to allow the passage of vapour. Vapour rises through the chimney and is directed downward by the cap, finally discharging through slots in the cap, and finally bubbling through the liquid on the tray. Valve trays In valve trays, perforations are covered by liftable caps. Vapour flows lifts the caps, thus self creating a flow area for the passage of vapour. The lifting cap directs the vapour to flow horizontally into the liquid, thus providing better mixing than is possible in sieve trays.

Figure 57 Valve trays (photos courtesy of Paul Phillips)

Sieve trays Sieve trays are simply metal plates with holes in them. Vapour passes straight upward through the liquid on the plate. The arrangement, number and size of the holes are design parameters.

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Because of their efficiency, wide operating range, ease of maintenance and cost factors, sieve and valve trays have replaced the once highly thought of bubble cap trays in many applications.

Liquid and Vapour Flows in a Tray Column The next few figures show the direction of vapour and liquid flow across a tray, and across a column.

Figure 58 Vapour & Liquid Flow across Column/Tray Each tray has 2 conduits, one on each side, called ‘downcomers’. Liquid falls through the downcomers by gravity from one tray to the one below it. The flow across each plate is shown in the above diagram on the right. A weir on the tray ensures that there is always some liquid (holdup) on the tray and is designed such that the the holdup is at a suitable height, e.g. such that the bubble caps are covered by liquid. Being lighter, vapour flows up the column and is forced to pass through the liquid, via the openings on each tray. The area allowed for the passage of vapour on each tray is called the active tray area.

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The picture on the left is a photograph of a section of a pilot scale column equipped with bubble capped trays. The tops of the 4 bubble caps on the tray can just be seen. The downcomer in this case is a pipe, and is shown on the right. The frothing of the liquid on the active tray area is due to both passage of vapour from the tray below as well as boiling.

As the hotter vapour passes through the liquid on the tray above, it transfers heat to the liquid. In doing so, some of the vapour condenses adding to the liquid on the tray. The condensate, however, is richer in the less volatile components than is in the vapour. Additionally, because of the heat input from the vapour, the liquid on the tray boils, generating more vapour. This vapour, which moves up to the next tray in the column, is richer in the more volatile components. This continuous contacting between vapour and liquid occurs on each tray in the column and brings about the separation between low boiling point components and those with higher boiling points. Tray Designs A tray essentially acts as a mini-column, each accomplishing a fraction of the separation task. From this we can deduce that the more trays there are, the better the degree of separation and that overall separation efficiency will depend significantly on the design of the tray. Trays are designed to maximise vapour-liquid contact by considering the liquid distribution and vapour distribution on the tray. This is because better vapour-liquid contact means better separation at each tray, translating to better column performance. Less trays will be required to achieve the same degree of separation. Attendant benefits include less energy usage and lower construction costs.

Figure 59 Liquid distributors - Gravity (left), Spray (right)(photos courtesy of Paul Phillips)

Packings Page 131

There is a clear trend to improve separations by supplementing the use of trays by additions of packings. Packings are passive devices that are designed to increase the interfacial area for vapour-liquid contact. The following pictures show 3 different types of packings.

Figure 60 Tray Packings

These strangely shaped pieces are supposed to impart good vapour-liquid contact when a particular type is placed together in numbers, without causing excessive pressure-drop across a packed section. This is important because a high pressure drop would mean that more energy is required to drive the vapour up the distillation column.

Figure 61 Structured packing (photo courtesy of Paul Phillips)

Packings versus Trays A tray column that is facing throughput problems may be de-bottlenecked by replacing a section of trays with packings. This is because: •

packings provide extra inter-facial area for liquid-vapour contact

•

efficiency of separation is increased for the same column height

•

packed columns are shorter than trayed columns

Packed columns are called continuous-contact columns while trayed columns are called staged-contact columns because of the manner in which vapour and liquid are contacted.

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Tower Capacity: Factors, calculation, modification

Equipment and Column Sizing In order to have stable operation in a distillation column, the vapour and liquid flows must be managed. Requirements are: • • • • • •

vapour should flow only through the open regions of the tray between the downcomers liquid should flow only through the downcomers liquid should not weep through tray perforations liquid should not be carried up the column entrained in the vapour vapour should not be carried down the column in the liquid vapour should not bubble up through the downcomers

These requirements can be met if the column is properly sized and the tray layouts correctly determined.

Tray layout and column internal design is quite specialized, so final designs are usually done by specialists; however, it is common for preliminary designs to be done by ordinarily superhuman process engineers. These notes are intended to give you an overview of how this can be done, so that it won't be a complete mystery when you have to do it for your design project. Basically in order to get a preliminary sizing for you column, you need to obtain values for • • • •

the tray efficiency the column diameter the pressure drop the column height

Tray Construction & Hydraulics Three main types of trays are to be discussed: • • •

Bubble Cap Trays Sieve Trays Valve Trays

Typically, the liquid flow between trays is governed by a weir on each tray. The flow depends on the length of the weir and how high the liquid level on the tray is above the weir. The Francis weir equation is one example of how the flow off a tray may be modeled.

Tray Efficiency

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Ideally, tray efficiencies are determined by measurements of the performance of actual trays separating the materials of interest; however, this is usually not practical in the early phases of a design. Consequently, some form of estimation is required. Estimates can be based on theory or on data collected from other columns. The O'Connell correlation is based on data collected from actual columns. It is based on bubble cap trays and is conservative for sieve and valve trays. It correlates the overall efficiency of the column with the product of the feed viscosity and the relative volatility of the key component in the mixture. These properties should be determined at the arithmetic mean of the column top and bottom temperatures. A fit of the data has been determined:

This, or a similar data set, can be used to get preliminary estimates of efficiency numbers.

Column Diameter Column diameter is found based on the constraints imposed by flooding. The number of ideal stages isn't needed to find the diameter -- only the vapour and liquid loads. You do need the number of actual stages to get the column height. Before beginning a diameter calculation, you want to know the vapour and liquid rates throughout the column. You then do a diameter calculation for each point where the loading might be an extreme: the top and bottom trays; above and below feeds, sidedraws, or heat addition or removal; and any other places where you suspect peak loads. Once you've calculated these diameters, you select one to use for the column, then check it to make sure it will work. Some columns will have two sections with different diameters -consider this possibility if you end up with regions where the estimated diameter varies by 20% or more, but realize it will be more expensive than a column that is the same all the way up. One issue that ought to be considered is the validity of your design numbers. If you are following the "traditional" approach, you've probably designed your column for reflux rates in the range of 1.1 to 1.2 times the minimum. This may not give you a column that can handle "upsets" well, so you may want to design for a capacity slightly greater than that -- increasing the flows by about 20% might be wise. Flooding Downcomer flooding occurs when liquid backs up on a tray because the downcomer area is two small. This is not usually a problem. More worrisome is entrainment flooding, caused by too much liquid being carried up the column by the vapour stream. A number of correlations and techniques exist for calculating the flooding velocity; from this, the active area of the column is calculated so that the actual velocity can be kept to no more than 80-85% of flood; values down to 60% are sometimes used. A force balance can be made on droplets entrained by the vapour stream (which can lead to entrainment flooding). This balance yields an expression relating the vapour and liquid densities and a capacity factor (C, with velocity units) to the flooding velocity:

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Capacity Factors The capacity factor can be determined from theory (it depends on droplet diameter, drag coefficient, etc.), but is usually obtained from correlations based on experimental data from distillation tray tests. Depending on the correlation used, C may include the effects of surface tension, tendency to foam, and other parameters. A common correlation is one proposed by Fair in the late 50s - early 60s. The version for sieve trays is available in a wide range of sources (including Figure 21.28 of MSH). The correlation takes the form of a plot of a capacity factor (which must be corrected for surface tension) vs. a functional group based on the liquid to vapour mass ratio:

Enter the plot from the bottom with this number, and then read the capacity factor from the left. This capacity factor applies to nonfoaming systems and trays meeting certain hole and weir size restrictions. It will need to be corrected for surface tension:

where the surface tension is in dynes/cm. Other correlations for the capacity factor are also available. Several are based on more recent information, and may well be more accurate than the Fair plot; however, they also tend to be less broadly known and often require more a priori information on the system. You should use a correlation that is acceptable for your problem. Diameter Once you have the capacity factor, you can readily solve for the flooding velocity:

(this solution is for the Fair correlation, and adds the surface tension correction). We know that flow=velocity*area, so we can calculate the flow area from the known vapour flow rate and the desired velocity (a fraction of flood). This area needs to be increased to account for the downcomer area which is unavailable for mass transfer. The resulting tray area can then be used to calculate the column diameter. So, with everything lumped together, we have:

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The only "new" term is the ratio of downcomer area to tray area. This should probably never be less than 0.1, and probably seldom will be greater than 0.2. Trays probably aren't a good idea for columns less than about 1.5 ft in diameter (you can't work on them) -- these are normally packed. Packing is less desirable for large diameter columns (over about 5 ft in diameter). Pressure Drop There is a pressure gradient through the column -- otherwise the vapour wouldn't flow. This gradient is normally expressed in terms of a pressure drop per tray, usually on the order of 0.10 psi. The best source of pressure drop information is to measure the actual drop between trays, but this isn't always feasible at the beginning of a design. Detailed calculations are possible, but these depend so much on the actual tray specifications that final values are usually obtained from experts, but approximate methods can be used to get values to put in your design basis. There are two main components to the pressure drop: the "dry tray" drop caused by restrictions to vapour flow imposed by the holes and slots in the trays and the head of the liquid that the vapour must flow through.

Dry Tray Losses The dry tray head loss can be related to an orifice flow equation:

This equation determines the dry tray drop in inches of fluid (your text has a similar equation in SI units). The constant 0.186 takes care of the units and is appropriate for sieve trays. The orifice size coefficient Co depends on the tray configuration and will usually fall between 0.65 and 0.85. The hole velocity can be obtained by dividing the vapour flow rate by the total hole area of the tray. Liquid Losses The liquid head pressure drop includes the effects of surface tension and of the frothing on the tray. It is typically represented as the product of an aeration factor and the height of liquid on the tray:

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Correlations are available for the aeration factor (beta); a value of 0.6 is good for a wide variety of situations. The height of liquid on the tray is the sum of the weir height and the height of liquid over the weir. The total height can be calculated directly from the volume of liquid on the tray and its active area. Another approach is to back the height out of a version of the Francis weir equation (which relates flow off a tray to liquid height and weir length). One version, for a straight weir, in units of inches and gal/min is:

Realize that these equations depend on the size and shape of the weir. Column Height The height of a trayed column is calculated by multiplying the number of (actual) stages by the tray separation. Tray spacing can be determined as a cost optimum, but is usually set by mechanical factors. The most common tray spacing in 24 inches -- it allows enough space to work on the trays whenever the column is big enough around (>5 ft diameter) that workers must crawl inside. Smaller diameter columns may be able to get by with 18 inch tray spacings. In addition to the space occupied by the trays, height is needed at the top and bottom of the column. Space at the top -- typically an additional 5 to 10 ft -- is needed to allow for disengaging space. The bottom of the tower must be tall enough to serve as a liquid reservoir. Depending on your boss's feelings about keeping inventory in the column, you will probably design the base for about 5 minutes of holdup, so that the total material entering the base can be contained for at least 5 minutes before reaching the bottom tray. The total of height added to the top and bottom will usually amount to about 15% or so added to that required by the trays. You rarely will see a real tower that is more than about 175 ft. tall. Tall, skinny towers are not a good idea, so watch the height/diameter ratio. You generally want to keep it less than 20 or 30. If your tower ends up exceeding these values, you probably want to look at a redesign, maybe by reducing the tray spacing, or splitting the tower into two parts.

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Absorption & Adsorption

Separation > Absorption Absorption processes are employed to recover valuable light components such as propane/propylene and butane/butylene from the vapours that leave the top of crude-oil or process-unit fractionating columns within the refinery. These volatile gases are bubbled through an absorption fluid, such as kerosene or heavy naphtha, in equipment resembling a fractionating column. The light products dissolve in the oil while the dry gases—such as hydrogen, methane, ethane, and ethylene—pass through undissolved. Absorption is more effective under pressures of about 7 to 11 kilograms per square centimetre (100 to 150 pounds per square inch) than it is at atmospheric pressure. The enriched absorption fluid is heated and passed into a stripping column, where the light product vapours pass upward and are condensed for recovery as liquefied petroleum gas (LPG). The unvapourised absorption fluid passes from the base of the stripping column and is reused in the absorption tower. Absorption is generally used to separate a higher-boiling constituent from other components of a system of vapours and gases. The absorption medium is usually an oil in the range of gas oil. Absorption is widely employed in the recovery of natural gasoline from well gas and of vapours given off by storage tanks. Absorption also obtains light hydrocarbons from many refining processes (catalytic cracking, hydrocracking, coking etc.). The solvent oil may be heavy gasoline, kerosenes, or even heavier oils. The absorbed products are recovered by fractionating or steamstripping.

Separation > Adsorption Certain highly porous solid materials have the ability to select and adsorb specific types of molecules, thus separating them from other materials. Silica gel is used in this way to separate aromatics from other hydrocarbons, and activated charcoal is used to remove liquid components from gases. Adsorption is thus somewhat analogous to the process of absorption with an oil, although the principles are different. Layers of adsorbed material only a few molecules thick are formed on the extensive interior surface of the adsorbent; the interior surface may amount to several hectares per kilogram of material. Adsorption is employed for about the same purpose as absorption; in the processjust mentioned natural gasoline may be separated from natural gas by adsorption on charcoal. Adsorption is also used to remove undesirable colours from lubricating oils, usually employing activated clay. The use of molecular sieves in separating close boiling components deserves a special mention. Molecular sieves are a special form of adsorbent. Such sieves are produced by the dehydration of naturally occurring or synthetic zeolites (crystalline alkali-metal aluminosilicates). The dehydration leaves intercrystalline cavities that have pore openings of definite size, depending on the alkali metal of the zeolite. Under adsorptive conditions, normal paraffin molecules can enter the crystalline lattice and be selectively retained, whereas all other molecules are excluded. This principle is used commercially for the removal of normal paraffins from gasoline fuels, thus improving their combustion properties. The use of molecular sieves is also extensive in the manufacture of high-purity solvents.

Solid Liquid Separation

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Introduction Separation techniques concentrate contaminated solids through physical and chemical means. These processes seek to detach contaminants from their medium (i.e., the soil, sand, and/or binding material that contains them). Description:

Figure 62 Typical Gravity Separation System

The separation processes are used for removing contaminated concentrates from soils, to leave relatively uncontaminated fractions that can then be regarded as treated soil. Ex situ separation can be performed by many processes. Gravity separation and sieving/physical separation are two well-developed processes that have long been primary methods for treating municipal wastewaters. Magnetic separation, on the other hand, is a much newer separation process that is still being tested. Gravity Separation Gravity separation is a solid/liquid separation process, which relies on a density difference between the phases. Equipment size and effectiveness of gravity separation depends on the solids settling velocity, which is a function of the particles size, density difference, fluid viscosity, and particle concentration (hindered settling). Gravity separation is also used for removing immiscible oil phases, and for classification where particles of different sizes are separated. It is often preceded by coagulation and flocculation to increase particle size, thereby allowing removal of fine particles. Magnetic Separation Magnetic separation is used to extract slightly magnetic radioactive particles from host materials such as water, soil, or air. All uranium and plutonium compounds are slightly magnetic while most host materials are nonmagnetic. The process operates by passing contaminated fluid or slurry through a magnetized volume. The magnetized volume contains a magnetic matrix material such as steel wool that extracts the slightly magnetic contamination particles from the slurry. Sieving/Physical Separation Sieving and physical separation processes use different size sieves and screens to effectively concentrate contaminants into smaller volumes. Physical separation is based on

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the fact that most organic and inorganic contaminants tend to bind, either chemically or physically, to the fine (i.e., clay and silt) fraction of a soil. The clay and silt soil particles are, in turn, physically bound to the coarser sand and gravel particles by compaction and adhesion. Thus, separating the fine clay and silt particles from the coarser sand and gravel soil particles would effectively concentrate the contaminants into a smaller volume of soil that could then be further treated or disposed

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Module 5 – Process Control & Economics

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Process Control Basics

Measured Variables Definition: The physical quantity, property, or condition which is to be measured. Common measured variables are temperatures, pressure, rate of flow, thickness, speed, etc. 2. The pan of the process that is monitored to determine the actual condition of the controlled variable.

Process Control Systems This part of the course starts with an outline and overview of basic control concepts. Questions which process engineers routinely have to answer about process control include the following: •

I have this process. What should I control?

•

Where on the process do I put my control loops?

•

As I proceed with the design of a process, what aspects of control should I consider at which stages?

Most books with the words `process control' in the title do little to answer these questions. Classical linear control theory, which forms the basis of most books on control, is much concerned with how to design controllers and is less helpful on how to design complete control systems. Other problems with this classical approach, for most process engineers wishing to design control systems for real chemical processes, are the restriction of most of its methods to idealised process models, and the extensive use of rather specialised mathematics. Satisfactory answers to questions such as the above frequently require little conventional mathematics. What they do require, however, is a good understanding of what a process is intended to do and how it works. In this book we will approach process control from the standpoint of a chemical or process engineer, and address these questions and others like them. We will consider the process and its control system in the language of process engineering. We will use mathematics, as such, only when necessary, and the language of classical control engineering only when it is unavoidable, or will add very significantly to the process engineer's understanding.

Why Control? Chemical plants are intended to be operated under known and specified conditions. There are several reasons why this is so: Safety: Formal safety and environmental constraints must not be violated. Operability: Certain conditions are required by chemistry and physics for the desired reactions or other operations to take place. It must be possible for the plant to be arranged to achieve them. Economic:

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Plants are expensive and intended to make money. Final products must meet market requirements of purity, otherwise they will be unsaleable. Conversely the manufacture of an excessively pure product will involve unnecessary cost. A chemical plant might be thought of as a collection of tanks in which materials are heated, cooled and reacted, and of pipes through which they flow. Such a system will not, in general, naturally maintain itself in a state such that precisely the temperature required by a reaction is achieved, a pressure in excess of the safe limits of all vessels be avoided, or a flowrate just sufficient to achieve the economically optimum product composition arise.

Control Objectives Control systems in chemical plants have, as noted, three functions. •

Safety.

•

Operability, i.e. to ensure that particular flows and holdup are maintained at chosen values within operating ranges.

•

To control product quality, process energy consumption etc.

To a large extent these are quite separate objectives. Indeed, in the case of safety systems separate equipment is generally used. The aims of control for operability are secondary to those of strategic control for quality etc., which directly affect process profitability.

Control for Safety Concern for safety is paramount in designing a chemical plant and its control systems. Ideally a process design should be `intrinsically safe', that is, plant and equipment should be such so that any deviation, such as an increase in reactor pressure, will itself change operating conditions so that it is rapidly removed, for example by a fall in reaction rate. For many perturbations this type of responsive, passive safety system will not be possible and active systems will be required. These active safety systems must be robust and of high integrity. Current processes achieve this through simplicity. The ultimate safety system is in most cases the mechanical relief valve which simply vents the plant to atmosphere, possibly through a flare or scrubber. We will not discuss control for safety explicitly in this book. Generally speaking a complete and separate system is provided to handle emergency control action. The need for this, and its design requirements, are established in hazard and operability or hazop studies. These are typicaly carried out on the complete process with its `normal' control systems in place. A number of safety issues will be addressed in the course of developing the design of the control systems for normal operation, but it must be emphasised that our treatment of this vital issue will be relatively restricted.

Control for Operability The operator of a process quite simply has to •

know what it is doing

•

be able to make it do what he or she wants, rather than to follow its natural inclinations.

The issue of making a plant behave in this way is called operability.

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The majority of control loops in a plant control system are associated with operability. Specific flow rates have to be set, levels in vessels maintained and chosen operating temperatures for reactors and other equipment achieved.

Control for Profitability There is no point in building a plant which is totally safe and can be made to take up any (safe) conditions of flow, temperature etc., if the conditions under which it is operated do not produce the correct amount of product to the correct specification, thus allowing its operators to make a profit. The top level of process control, what we will refer to as the strategic control level is thus concerned with achieving the appropriate values principally of: •

Production rate,

•

Product quality, and

•

Energy economy.

Techniques of Control

Basic Concepts of Feedback Control The task of maintaining these required conditions falls to one or, more usually several, process control systems with which the plant will be equipped. The practical aspects of these will be discussed more fully in the following module. The underlying principle of most process control, however, is already understood by anyone who has grasped the operation of the domestic hot water thermostat: •

The quantity whose value is to be maintained or regulated, e.g. the temperature of the water in a cistern, is measured.

•

Comparison of the measured and required values provides an error, e.g. `too hot' or `too cold'.

•

On the basis of the error, a control algorithm decides what to do. Such an algorithm might be: If the temperature is too high then turn the heater off. If it is too low then turn the heater on.

•

The adjustment chosen by the control algorithm is applied to some adjustable variable, such as the power input to the water heater.

This summarises the basic operation of a feedback control system such as one would expect to find carrying out nearly all control operations on chemical plants, and indeed in most other circumstances where control is required. The diagram belows a feedback control loop.

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Figure 63 Feedback control loop

Notice that this extremely simple idea has a number of very convenient properties. The feedback control system seeks to bring the measured quantity to its required value or setpoint. The control system does not need to know why the measured value is not currently what is required, only that this is so. There are two possible causes of such a disparity: •

The system has been disturbed. This is the common situation for a chemical plant subject to all sorts of external upsets. However, the control system does not need to know what the source of the disturbance was.

•

The setpoint has been changed. In the absence of external disturbance, a change in setpoint will introduce an error. The control system will act until the measured quantity reaches its new setpoint.

A control system of this sort should also handle simultaneous changes in setpoint and disturbances.

Advantages of Feedback Control Not only does the feedback control system require no knowledge of the source or nature of disturbances, but it requires minimal detailed information about how the process itself works. Feedback control action is entirely empirical, so long as an adjustment is being made in the correct `sense', e.g. more heat means increasing temperature and vice versa, then the control system should remove the effect of an external disturbance. As we will see, it helps to know more than this, but the minimum information required to make a feedback control system work is whether the adjustment makes the measurement go up or down.

Disadvantages of Feedback Control The main disadvantage of feedback control is that the disturbance enters into the process and upsets it. It is after the process output is different from the setpoint that the controller takes some corrective actions. Although most processes allow some fluctuation of controlled variable within a certain range, there are two process conditions which can make the overall effectiveness of feedback control quite unsatisfactory. One of these is the occurrence of disturbances of a large magnitude that is strong enough to seriously affect or even damage

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the process. The other is the occurrence of a large amount of lag (time delay) within the process. These are discussed further below.

Large Magnitude Disturbance An example where the occurrence of disturbances of large magnitude that are strong enough to seriously effect the process, is temperature control of a catalyst reactor in which strong exothermic reaction takes place. The reaction heat is very high, therefore the reactant gas mixture is diluted by a inert gas to carry away most reaction heat, although the temperature of the reactor is maintained by feedback control of a coolant flowrate in coils inside. Assuming a large magnitude disturbance, the sudden large increase in the reactant concentration in the feed, enters the reactor, a sudden increase in the temperature is so large and so quick that the catalyst is burnt out before the control system senses the change and takes any actions. A diagram of this situation is shown below.

Figure 64 Large magnitude disturbance

Large Time Delay A simple example of a large time delay is the distillation column as outlined in the figure below. If we use feedback control to regulate the purity of the top product, when the feed composition changes (disturbance), the control system is not aware any takes no action until the effects of the disturbance travels and arrives at the sensor position at the top. When the controller takes the correction, the whole column may be far away from the designed conditions.

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Figure 65 Time delay

The question of importance of either occurrence is defined in economic terms. In either case, the principle concern is the existence of errors that have significant economic consequences in the overall process operation. In these cases, feedforward control can be used to deal with these disadvantages or inadequacies of feedback control.

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Process Economics

Refinery Economics The overall economics or viability of a refinery depends on the interaction of three key elements: the choice of crude oil used (crude slates), the complexity of the refining equipment (refinery configuration) and the desired type and quality of products produced (product slate). Refinery utilization rates and environmental considerations also influence refinery economics. Using more expensive crude oil (lighter, sweeter) requires less refinery upgrading but supplies of light, sweet crude oil are decreasing and the differential between heavier and more sour crudes is increasing. Using cheaper heavier crude oil means more investment in upgrading processes. Costs and payback periods for refinery processing units must be weighed against anticipated crude oil costs and the projected differential between light and heavy crude oil prices. Crude slates and refinery configurations must take into account the type of products that will ultimately be needed in the marketplace. The quality specifications of the final products are also increasingly important as environmental requirements become more stringent.

Crude Slate Different types of crude oil yield a different mix of products depending on the crude oil´s natural qualities. Crude oil types are typically differentiated by their density (measured as API gravity) and their sulphur content. Crude oil with a low API gravity is considered a heavy crude oil and typically has a higher sulphur content and a larger yield of lower-valued products. Therefore, the lower the API of a crude oil, the lower the value it has to a refiner as it will either require more processing or yield a higher percentage of lower-valued byproducts such as heavy fuel oil, which usually sells for less than crude oil. Crude oil with a high sulphur content is called a sour crude while sweet crude has a low sulphur content. Sulphur is an undesirable characteristic of petroleum products, particularly in transportation fuels. It can hinder the efficient operation of some emission control technologies and, when burned in a combustion engine, is released into the atmosphere where it can form sulphur dioxide. With increasingly restrictive sulphur limits on transportation fuels, sweet crude oil sells at a premium. Sour crude oil requires more severe processing to remove the sulphur. Refiners are generally willing to pay more for light, low sulphur crude oil. Most refineries in Western Canada and Ontario were designed to process the light sweet crude oil that is produced in Western Canada. Unlike leading refineries in the U.S., Canadian refineries in these regions have been slower to reconfigure their operations to process lower cost, less desirable crude oils, instead choosing to rely extensively on the abundant, domestically-produced, light, sweet crudes. As long as these lighter crudes were available, refining economics were insufficient to warrant new investment in heavy oil conversion capacity. However, with growing oil sands production and the declining production of conventional light sweet crudes, refineries in Western Canada and Ontario have started to make the investment required to process the increasing supply of heavier crudes. In 2003, Shell Canada completed the conversion of their Scotford refinery to use bitumen feedstock. In the fall of 2003, Consumer´s Co-operative Refineries Ltd completed a 35,000 bbls/day expansion of their refinery in Regina, Saskatchewan. This increased their heavy oil refining capacity to approximately 85,000 bbls/day. Petro-Canada has also announced plans to do a major refitting of their Edmonton refinery. Although this construction is not expected to increase their capacity, it will allow them to upgrade and refine oil sands feedstock. The

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$1.2 billion CDN project will significantly expand the existing coker at Edmonton allowing for approximately 53,000 bbls/day of bitumen upgrading. Similarly, Suncor is expected to do a feedstock conversion at its Sarnia refinery to run more lower value oil sands feedstock. Much of this investment by the large integrated oil companies (companies that are involved in both the production of crude oil and the manufacturing and distribution of petroleum products) is associated with ensuring a market for their growing oil sands production. In Western Canada and Ontario, almost 50% of the crude oil processed by refiners is conventional light, sweet crude oil and another 25% is high quality synthetic crude oil. Synthetic crude is a light crude oil that is derived by upgrading oil sands. Most of the remaining crude oil processed by these refineries is heavy, sour crude. The crude slate is expected to change significantly in the years ahead as refiners increase their capacity to process heavy crude oil and lower quality synthetic crudes. Refineries in Atlantic Canada and Quebec are dependent on imported crudes and tend to process a more diverse crude slate than their counterparts in Western Canada and Ontario. These refiners have the capacity to purchase crude oil produced almost anywhere in the world and therefore have incredible flexibility in their crude buying decisions. Approximately 1/3 of crude processed in Eastern Canada and Quebec is conventional, light sweet crude and another 1/3 is medium sulphur, heavy crude oil. The remaining 1/3 is a combination of sour light, sour heavy and very heavy crude oil. The crude slate in Eastern Canada is expected to remain much more static than that in Western Canada and Ontario, as these refiners are not constrained by the quality or volume of domestic crude production. Figure 3 illustrates the product yield for six typical types of crude oil processed in Canada. It includes both light and heavy as well as sweet and sour crude oils. A very light condensate (62 API) and a synthetic crude oil are also included. The chart compares the different output when each crude type is processed in a simple distillation refinery. The output is broken down into five main product groups: gasoline, propane and butane (C3/C4), Cat feed (a partially processed material that requires further refining to make usable products), distillate (which includes diesel oil and furnace oil) and residual fuel (the heaviest and lowest-valued part of the product output, used to make heavy fuel oil and asphalt).

Refinery Configuration A refiner´s choice of crude oil will be influenced by the type of processing units at the refinery. Refineries fall into three broad categories. The simplest is a topping plant, which

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consists only of a distillation unit and probably a catalytic reformer to provide octane. Yields from this plant would most closely reflect the natural yields from the crude processed. Typically only condensates or light sweet crude would be processed at this type of facility unless markets for heavy fuel oil (HFO) are readily and economically available. Asphalt plants are topping refineries that run heavy crude oil because they are only interested in producing asphalt. The next level of refining is called a cracking refinery. This refinery takes the gas oil portion from the crude distillation unit (a stream heavier than diesel fuel, but lighter than HFO) and breaks it down further into gasoline and distillate components using catalysts, high temperature and/or pressure. The last level of refining is the coking refinery. This refinery processes residual fuel, the heaviest material from the crude unit and thermally cracks it into lighter product in a coker or a hydrocraker. The addition of a fluid catalytic cracking unit (FCCU) or a hydro cracker significantly increases the yield of higher-valued products like gasoline and diesel oil from a barrel of crude, allowing a refinery to process cheaper, heavier crude while producing an equivalent or greater volume of high-valued products. Hydrotreating is a process used to remove sulphur from finished products. As the requirement to produce ultra low sulphur products increases, additional hydrotreating capability is being added to refineries. Refineries that currently have large hydrotreating capability have the ability to process crude oil with a higher sulphur content. Figure 4 demonstrates that using the same crude input (heavy crude with a 27 API) yields a very different range of petroleum products depending on the refining units and processes used. In the case of the cracking refinery, the addition of other blending materials at various stages of production is required but the resulting volumetric output is greater than the volume of the crude oil input. Each refinery is unique due to age / technology and modifications over time, but generalizations are possible. The installation of additional conversion capability increases the yield of clean products and reduces the yield of heavy fuel oil. However, increased conversion capability would generally result in higher energy use and, therefore, higher operating costs. These higher operating and capital costs must be weighed against the lower cost of the heavier crude oil. Canada has primarily cracking refineries. These refineries run a mix of light and heavy crude oils to meet the product slate required by Canadian consumers. Historically, the abundance of domestically produced light sweet crude oils and a higher demand for distillate products, such as heating oil, than in some jurisdictions reduced the need for upgrading capacity in Canada. However, in more recent years, the supply of light sweet crude has declined and newer sources of crude oil tend to be heavier. Many of the Canadian refineries are now being equipped with upgraders to handle the heavier grades of crude oil currently being produced.

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[1] Source: NRCan surveys

Product Slate Refinery configuration is also influenced by the product demand in each region. Refineries produce a wide range of products including: propane, butane, petrochemical feedstock, gasolines (naphtha specialties, aviation gasoline, motor gasoline), distillates (jet fuels, diesel, stove oil, kerosene, furnace oil), heavy fuel oil, lubricating oils, waxes, asphalt and still gas. Nationally, gasoline accounts for about 40% of demand with distillate fuels representing about one third of product sales and heavy fuel oil accounting for only eight percent of sales. Total petroleum product demand is distributed almost equally across the regions, with Atlantic/Quebec, Ontario and the West each accounting for about one third of total sales. However, the mix of products varies quite significantly among the regions. [2] In the Atlantic provinces, where furnace oil (light heating oil) is the primary source of home heating, distillate fuels make up 40% of product demand, and heavy fuel oil, used to generate electricity, accounts for another 24%. Gasoline sales account for less than 30% of product demand. In Quebec, where natural gas and hydroelectricity are prevalent, distillate fuel has a 34% share of sales and gasoline is about 40%. Similarly, in Ontario, gasoline sales outpace distillate sales and account for more than 45% of total product demand, with distillates at less than 30%. In Western Canada, agricultural use is one of the primary drivers behind distillate demand and gasoline and distillate each account for about 40% of total petroleum product sales. These regional differences in product demand have influenced the configurations of the refineries in each area. By comparison, in the U.S., the demand for gasoline is much larger than distillate demand and, therefore, refiners configure their installations to maximize gasoline production.

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Gasoline sales account for nearly 50% of demand while distillate sales account for less than 30% of product demand. In several Western European countries, most notably Germany and France, policies exist that encourage the use of diesel engines creating a much stronger distillate component. Gasoline accounts for less than 20% of petroleum product sales in Europe. The US refineries are configured to process a large percentage of heavy, high sulphur crude and to produce large quantities of gasoline, and low amounts of heavy fuel oil. U.S. refiners have invested in more complex refinery configurations, which allow them to use cheaper feedstock and have a higher processing capability. Canada´s refineries do not have the high conversion capability of the US refineries, because, on average, they process a lighter, sweeter crude slate. Canadian refineries also face a higher distillate demand, as a percent of crude, than those found in the U.S. so gasoline yields are not as high as those in the US, but are still significantly higher than European yields. The relationship between gasoline and distillate sales can also create challenges for refiners. A refinery has a limited range of flexibility in setting the gasoline to distillate production ratio. Beyond a certain point, distillate production can only be increased by also increasing gasoline production. For this reason, Europe is a major gasoline exporter, primarily to the U.S.

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Mass-exchange processes, such as distillation, absorption, extraction, adsorption and drying are used in chemical technology for separation of substances into their components. These processes have common features. At least three substances are used in a mass-exchange process, namely, a distributive substance, which forms the first phase, the second distributive substance - the second phase, and a distributed substance, that migrates from one phase to another. The driving force of this process is determined by the difference between current and equilibrium concentrations of substances. The correlation between these concentrations could be linear (for absorption or extraction) or non-linear (for distillation). However, they (concentrations) strongly depend on the process parameters (temperature, pressure), and on the presence of various additives. Industrial apparatus is designed with respect to the certain values of these parameters and certain concentrations of initial products. In reality, disturbances could lead to the distortion of material and thermal balances in apparatus, to the deviation of pressure and temperatures from the desired values, and , finally, to the deviation of the composition (quality) of final products from the required ones. Therefore, the objective of control systems is to stabilise these process parameters in order to maintain the material and thermal balances by suppressing various disturbances. The majority of mass-exchange processes occur in columns, which have several meters in diameter and several dozens meters in height. Time delays in these apparatus could vary from several minutes to several dozens of minutes. Therefore, single-loop control systems have large offsets and long time transient processes. Employing cascade control systems one can improve the performance of these processes. Deficiency of instrumentation for continuous measuring of the composition of intermediate and final products creates difficulties for the accurate control. In such cases, control of the quality is performed indirectly, i.e. by controlling the boiling temperatures, densities or viscosities of mixtures.

22.1. Control of distillation columns. These columns are used for the separation of liquid homogeneous mixtures into its components or groups of components. Let's consider possible disturbances, manipulated and output variables. Since the initial product comes to the distillation column from the previous process units, therefore, variations of feed flowrate, its concentration and temperature are the major disturbances. Enthalpies of a heating vapor (steam) and a coolant, and heat losses are possible disturbances. Usually, only the feed temperature is stabilised (controlled), whereas the flowrate of feed is measured only. Flowrates of vapour for heating, heat-transfer agent, coolant, distillate product, bottoms product and reflux are manipulated variables. The concentration of distillate and bottom product, the levels of fluid in the column and the level of distillate in the tank, pressure in the column are the output variables.

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Example 1. Fig. 22.1 presents a schematic view of a simple control system, which comprise six singleloop control systems. This control system stabilises the composition of the distillate product and maintains the material and thermal balances in the distillation column. The major controller, which stabilises the composition of the distillate, is the temperature controller (pos. 5-2). It manipulates the flow rate of reflux (pos. 5-3). Temperature controller (pos. 1-2) controls feed temperature by manipulating the flowrate of heat-transfer agent. Level controllers (pos. 7-2) and (pos. 8-2) maintain the material balance of liquid phase in the column, and pressure controller (pos. 4-2) maintains the material balance of vapor phase. The flowrate controller (pos. 3-3) stabilises the flowrate of heating vapour into the re-boiler. If our task is to control the composition of the bottom product, then the flowrate of steam for heating is manipulated by the control signal from the temperature controller (pos. 2-2), and the flowrate of reflux is manipulated by the flowrate controller (pos. 6-3). Simultaneous control of compositions of the distillate and bottom products or temperatures at the top and at the bottom of the column usually is not used because these process variables are interdependent. An application of feedback control systems may reduce the stability of these control systems. This control system has several disadvantages: • stabilisation of the vapour flowrate without a respect to other process parameters causes an excessive consumption of the vapour. This happens because the set point value supplied to this controller is slightly higher in order to take into account possible variations in the enthalpy of steam, supercooling of reflux, etc.; • since the effect of disturbances, such as flowrate or temperature of the feed, is not suppressed, this can lead to significant deviations in the composition of the final products from their desired values. This happens because the temperature controller at the top of the column receives the signal about the deviation in temperature (composition) of the product only after the composition of the fluid mixture has been changed along the height of the column.

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PT TT 1-1 4-1

Coolant out

PC TC 1-2 4-2

LT TT 1-1 7-1

Coolant in TT 1-1 5-1

TC 1-2

4-3

TC 1-2 5-2

5-3

TT 1-1

Distillate

FE TC 1-2 6-1 7-3

Reflux FT TT 1-2 1-1 6-2

Heat-transfer agent

FC FT TT 1-2 1-1 6-3

1-3

TC 1-2 2-2

Feed mixture

TT 1-1 2-1

3-4

LT TT 1-1 8-1

LC TC 1-2 8-2

FE TC 1-2 3-1

Steam for heating

FT TT 1-2 1-1 3-2 8-3 FC FT TT 1-2 1-1 3-3

LC TC 1-2 7-2

Bottom product

Figure 66 Distillation column with six single-loop control systems.

155

PT TT 1-1 3-1

Coolant out

PC TC 1-2 3-2

LT TT 1-1 5-1

Coolant in

LC TC 1-2 5-2

3-3

TC 1-2

TT 1-1

4-5

Reflux

5-3

TT 1-1 4-1

Heat-transfer agent

TC 4-2 1-3

Feed mixture

Distillate

FT TT 1-2 1-1 4-3

TC 4-4

FE TC 1-2 2-3 FT TT 1-2 1-1 2-4

Steam for heating

2-6

LT TT 1-1 6-1

LC TC 1-2 6-2

FE TC 1-2 2-1 FT TT 1-2 1-1 2-2

Bottom product

FFC FT TT 1-2 1-1 2-5 6-3

Figure 67 Distillation column with single-loop and cascade control systems

156

Another significant disadvantage of using the temperature of the product to control its composition is as follows: variations of this temperature due to changes in the composition of the product are comparable with its variations caused by pressure changes in the column. These temperature variations may be comparable with the accuracy of a temperature sensor. Let variations in the composition of the product can not exceed of ±1%. The difference in boiling temperatures of components is 20 °C. Then corresponding variations of temperature are equal to ±0.2 °C. For a potentiometer with a temperature range from 0 to 150 °C and an accuracy of 0.5% the error in temperature measurements is 0.75 °C. One should take this into account when select a temperature sensing device.

Example 2. In Fig. 22.2 the controller (pos. 2-5) controls the flowrate ratio between a feed mixture and steam for heating in a re-boiler, thus reducing the consumption of energy for the separation of mixture into its components. A cascade control system is used to control temperature at the top of the distillation column by introduction of a correction signal (pos. 4-3) from the loop for measuring temperature on the selected tray of the column. These are only two simple examples, whereas in reality more complex control systems are used. 22.2. Control of absorption columns Absorption columns (or absorbers) are used as intermediate units in chemical processes. The objective of absorption processes is to maximise the degree of absorption or to minimise the consumption of energy for the separation of the mixture. The major sources of disturbances are the flowrate, composition and temperature of a gas stream entering for absorption, and, sometimes, the temperature and composition of the liquid absorbing stream. The major manipulated variables are the flowrate of the liquid absorbing stream and flowrate of the bottom product. When one controls pressure and level in the column this maintains the material balance between gaseous and liquid phases. A control system with several single control loops (see Fig. 22.3a) keeps the material and thermal balances by using the level controller (pos.2-2) and pressure controller (pos. 1-2); and keeps the composition of the bottoms product at the desired value by using the composition controller (pos. 3-2). An introduction of a control signal using a flow ratio controller (pos. 3-5 in Fig. 22.3b) suppresses the effect of the variation of the gaseous mixture flowrate (this is the disturbance) and improves the performance of this cascade control system. Cascade control system (see Fig. 22.3c) uses the composition of the gaseous-liquid mixture on the certain tray of the column as an auxiliary controlled variable. In this case the composition controller (pos. 3-2) is the primary, or master, controller, whereas the composition controller (pos. 3-5) is the secondary, or slave controller.

4

4

1-3

4

1-3

PT 1-1

1-3

PC 1-2

PT 1-1

2

FT 3-3

PC 1-2

PT 1-1

PC 1-2

2 2

3-3

3-6

3-5 CT 3-3

FFC 3-5

1

FT 3-4

LT 2-1

1

1 LT 2-1

LT 2-1

LC 2-2

LC 2-2

LC 2-2

2-3

2-3 CT 3-1

2-3

CC 3-2

3

CT 3-1

3

a

CC 3-4

CC 3-2

CT 3-1

3

b

c

Figure 22.3. Control of absorbers. 1 - gas mixture; 2 - liquid absorbent; 3 - bottom product; 4 - end gases.

158

CC 3-2

Fundamentals of Process Plant Equipment Control Manual

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