multiphase design guide part1

8/22/2019 Multiphase Design Guide Part1

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PART I - MULTIPHASE PIPELINE & SLUG CATCHER DESIGN GUIDE

CPTC NOVEMBER 1994 1

Multiphase Pipeline Design Guide


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PART I

TABLE OF CONTENTS

SECTION 1.0 - INTRODUCTION

1.1 Objective and Scope.....................................................................................................................................................1

1.2 Definition of Terms........................................................................................................................................................1

SECTION 2.0 OVERVIEW OF MULTIPHASE FLOW FUNDAMENTALS

2.1 Design Criteria..............................................................................................................................................................11

2.2 Velocity Guidelines .......................................................................................................................................................11

2.3 Flow Regimes...............................................................................................................................................................13

2.4 Pressure Gradient.........................................................................................................................................................16

2.4.1 Frictional Losses..........................................................................................................................................16

2.4.2 Elevational Losses........................................................................................................................................17

2.4.3 Acceleration Losses......................................................................................................................................18

2.4.4 Allowable Pressure Drop...............................................................................................................................20

2.5 Pressure Gradient Calculations......................................................................................................................................20

2.6 Section Highlights.........................................................................................................................................................21

SECTION 3.0 STEADY STATE DESIGN METHODS

3.1 Pipeline Design Methods...............................................................................................................................................25

3.2 Steady State Simulators................................................................................................................................................26

3.2.1 Phase Equilibrium and Physical Properties.................................................................................................... 26

3.2.2 Pipeline Elevation Profile .............................................................................................................................. 28

3.2.3 Heat Transfer ............................................................................................................................................... 30

3.2.4 Recommended Methods for Pressure Drop, Liquid Holdup, and

Flow Regime Prediction................................................................................................................................ 33

3.2.5 Interpretation of Results................................................................................................................................35


SECTION 4.0 TRANSIENT FLOW MODELING

4.1 Transient Flow Modeling (General) ................................................................................................................................41

4.2 Use of Transient Simulators........................................................................................................................................... 42


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SECTION 5.0 SLUG FLOW ANALYSIS

5.1 Slug Flow (General)......................................................................................................................................................45

5.2 Slug Length and Frequency Predictions.........................................................................................................................46

5.2.1 Hydrodynamic Slugging................................................................................................................................46

5.2.2 Terrain Slugging...........................................................................................................................................51

5.2.3 Pigging Slugs............................................................................................................................................... 53

5.2.4 Startup and Blowdown Slugs........................................................................................................................ 55

5.2.5 Rate Change Slugs ......................................................................................................................................56

5.2.6 Downstream Equipment Design for Slug Flow............................................................................................... 56


SECTION 6 EXAMPLE PROBLEMS6.1 Example Problem 1 Low Gas/Oil Line Between Platforms ...... ...... ....... ....... ...... ....... ....... ...... ....... ....... ...... ....... ...... ..... 63

6.1.1 Line Size...................................................................................................................................................... 65

6.1.2 Slug Length Prediction .................................................................................................................................75

6.1.3 Slug Frequency and Length by Hill & Wood Method ......................................................................................80

6.2 Example Problem 2 Gas Condensate Line..................................................................................................................88

6.2.1 Table 1, Wellstream Composition.................................................................................................................89

6.2.2 Table 2, Pipeline Evaluation Profile ...............................................................................................................90

6.2.3 Pipeline Simulation Comparison ...................................................................................................................92

SECTION 7.0 REFERENCES .................................................................................................................................................... 106

FIGURES

I: 1-1 Flow Regimes in Horizontal Flow...................................................................................................................................8

I: 1-2 Flow Regimes in Vertical Flow ......................................................................................................................................9

I: 2-1 Horizontal Flow Regime Map.........................................................................................................................................23

I: 2-2 Vertical Flow Regime Map.............................................................................................................................................24

I: 5-1 Taitel-Dukler Liquid Holdup Predictions..........................................................................................................................60

I: 5-2 Stages in Terrain Slugging............................................................................................................................................ 61

I: 5-3 Pipeline Slugging..........................................................................................................................................................62

I: 6-1 Liquid Holdup for Example 1, Year 12............................................................................................................................101

I: 6-2 Inlet Pressure for Example 1, Year 12............................................................................................................................102

I: 6-3 Liquid Flowrate Out of Line, Example 1, Year 12............................................................................................................ 103

I: 6-4 Gas Flowrate Out of Line, Example 1, Year 12............................................................................................................... 104


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I: 6-5 Liquid Holdup Predictions for Example 2........................................................................................................................ 105


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SECTION 1.0 - INTRODUCTION

1.1 Objective and Scope

The simultaneous flow of gas and liquid through pipes, often referred to as multiphase

flow, occurs in almost every aspect of the oil industry. Multiphase flow is present in well

tubing, gathering system pipelines, and processing equipment. The use of multiphase

pipelines has become increasingly important in recent years due to the development of

marginal fields and deep water prospects. In many cases, the feasibility of a design

scenario hinges on cost and operation of the pipeline and its associated equipment.

Multiphase flow in pipes has been studied for more than 50 years, with significant

improvements in the state of the art during the past 15 years. The best available methods

can predict the operation of the pipelines much more accurately than those available only

a few years ago. The designer, however, has to know which methods to use in order to

get the best results.

Part I of this guide consists of seven sections. The fundamentals of multiphase flow in

pipelines are discussed in Section 2.0. The third section describes the use of steady state

simulation methods. This section of the guide helps the designer choose the best methods

for the project, and it gives guidelines to use in designs. The fourth section of the reportbriefly describes transient flow modeling. The fifth section describes slug flow modeling,

giving suggestions on the best methods to use in slug flow simulation. The sixth section

includes two sample problems, based on actual designs, which illustrate the design steps

used in analyzing the pipeline designs.

1.2 Definition of Terms

In discussing the design of multiphase pipelines, it is necessary to define several terms

used repeatedly throughout this text.

Near Horizontal and Near Vertical Angles

The term "near horizontal" is used in this guide to denote angles of -10 degrees to +10


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degrees from horizontal. The term "near vertical" denotes upward inclined pipes with

angles from 75 to 90 degrees from horizontal.

Flow Regimes

In multiphase flow, the gas and liquid within the pipe are distributed in several

fundamentally different flow patterns or flow regimes, depending primarily on the gas and

liquid velocities and the angle of inclination. Observers have labeled these flow regimes

with a variety of names. Over 100 different names for the various regimes and sub-

regimes have been used in the literature. In this guide, only four flow regime names will

be used: slug flow, stratified flow, annular flow, and dispersed bubble flow.

Figure I:1-1 shows the flow regimes for near horizontal flow, and Figure I:1-2 shows the

flow regimes for vertical upward flow. Descriptions of the flow regimes

1. Stratified Flow - at low flowrates in near horizontal pipes, the liquid and gas separate

by gravity, causing the liquid to flow on the bottom of the pipe while the gas flows

above it. At low gas velocities, the liquid surface is smooth. At higher gas velocities,

the liquid surface becomes wavy. Some liquid may flow in the form of liquid droplets

suspended in the gas phase. Stratified flow only exists for certain angles of inclination.

It does not exist in pipes that have upward inclinations of greater than about 1 degree.Most downwardly inclined pipes are in stratified flow, and many large diameter

horizontal pipes are in stratified flow. This flow regime is also referred to as stratified

smooth, stratified wavy, and wavy flow by various investigators.

2. Annular Flow - at high rates in gas dominated systems, part of the liquid flows as a

film around the circumference of the pipe. The gas and remainder of the liquid (in the

form of entrained droplets) flow in the center of the pipe. The liquid film thickness is

fairly constant for vertical flow, but it is usually asymmetric for horizontal flow due to

gravity. As velocities increase, the fraction of liquid entrained increases and the liquid

film thickness decreases. Annular flow exists for all angles of inclinations. Most gas

dominated pipes in high pressure vertical flow are in annular flow. This flow regime is

referred to as annular-mist or mist flow by many investigators.


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3. Dispersed Bubble Flow - at high rates in liquid dominated systems, the flow is a frothy

mixture of liquid and small entrained gas bubbles. For near vertical flow, dispersed

bubble flow can also occur at more moderate liquid rates when the gas rate is very

low. The flow is steady with few oscillations. It occurs at all angles of inclination.Dispersed bubble flow frequently occurs in oil wells. Various investigators have

referred to this flow regime as froth or bubble flow.

4. Slug Flow - for near horizontal flow, at moderate gas and liquid velocities, waves on

the surface of the liquid may grow to sufficient height to completely bridge the pipe.

When this happens, alternating slugs of liquid and gas bubbles will flow through the

pipeline. This flow regime can be thought of as an unsteady, alternating combination

of dispersed bubble flow (liquid slug) and stratified flow (gas bubble). The slugs can

cause vibration problems, increased corrosion, and downstream equipment problems

due to its unsteady behavior.

Slug flow also occurs in near vertical flow, but the mechanism for slug initiation is

different. The flow consists of a string of slugs and bullet-shaped bubbles (called

Taylor bubbles) flowing through the pipe alternately. The flow can be thought of as a

combination of dispersed bubble flow (slug) and annular flow (Taylor bubble). The

slugs in vertical flow are generally much smaller than those in near horizontal flow.

Slug flow is the most prevalent flow regime in low pressure, small diameter systems.

In field scale pipelines, slug flow usually occurs in upwardly inclined sections of the

line. It occurs for all angles of inclination. Investigators have used many terms to

describe parts of this flow regime. Among them are: intermittent flow; plug flow;

pseudo-slug flow, and churn flow.

Superficial Velocities

The velocities of the gas and liquid in the pipe are prime variables in the prediction of the

behavior of the multiphase mixture. Most multiphase flow prediction methods use the

superficial gas and liquid velocities as correlating parameters. The superficial velocities

are defined as the in situ volumetric flowrate of that phase divided by the total pipe cross-

sectional area. In oil field units, this corresponds to:


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Vsg = Superficial Gas Velocity, ft/sec

= (actual ft3/sec of gas) / (pipe cross-sectional area, ft

2)

Vsl = Superficial Liquid Velocity, ft/sec

= (actual ft3/sec of liquid) / (pipe cross-sectional area, ft

2)

Mixture Velocity

The mixture velocity (Vm) is the volumetric average velocity of the gas-liquid mixture. It

is equal to the sum of the gas and liquid superficial velocities.

V V Vm sg sl= +

Slip and Liquid Holdup

Liquid holdup is defined as the volume fraction of the pipe that is filled with liquid. It is

the most important parameter in estimating the pressure drop in inclined or vertical flow.

It is also of prime importance in sizing downstream equipment, which must be able to

operate properly when the liquid holdup in the line changes because of pigging or rate

changes.

If there was no slip between the gas and liquid phases, both phases would move through

the pipe at the mixture velocity. The liquid would occupy the volume fraction equivalent

to the ratio of the liquid volumetric flowrate to the total volumetric flowrate. In

multiphase flow terminology, this equates to the liquid holdup being equal to the ratio

between the superficial liquid velocity and the mixture velocity:

Hlns = No-slip liquid holdup

= Vsl / Vm

Under most conditions, however, the liquid phase, which is more dense and viscous,

moves more slowly than the gas. When this occurs, the liquid holdup (Hl) is greater than

the no-slip holdup.


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H Hl s> ln

Under these conditions, the actual gas velocity is greater than the mixture velocity, and

the actual liquid velocity is smaller than the mixture velocity. The expressions for the

actual gas velocity (Ug) and actual liquid velocity (Ul) are:

UV

g

sg

l

=1 H

UV

Hlsl

l

=

For small diameter, low pressure piping, there is frequently a vast difference between Ug

and Ul. For field piping, there is generally less slip between the phases, and the flow may

approximate no-slip flow in dispersed bubble and annular flows.

It is possible to get conditions where the liquid holdup is less than no-slip, but this only

occurs over a small range of flowrates in downwardly inclined pipes.

Pressure Gradient

Two definitions of the term "pressure gradient" are used in the literature. In this guide, the

term "pressure gradient" will be used to describe the pressure drop per unit length of pipe,

(Pin - Pout)/L. In many papers, the term "pressure gradient" is used to denote the pressure

change per unit length (dp/dx = (Pout - Pin)/L). The magnitude of the pressure gradient is

the same in either definition, but the sign of the pressure drop per unit length is usually

positive, while the sign of dp/dx is usually negative. Most people prefer to work with

positive numbers, so the majority of people use the pressure drop per unit length

definition. The choice of the definition is somewhat arbitrary, but it should be noted when

reading the multiphase flow literature, and working with some of the available software.

3-Phase Flow vs. 2-Phase Flow

In most of this guide, the discussion will consider 2-phase flow, or gas-liquid flow. In the

majority of oil field applications, there will actually be 3 phases present (gas, oil, and

water). The rigorous prediction of 3-phase flow is in its infancy. 3-Phase flow methods


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are not generally available, so most simulators use 2-phase models with a mixed liquid

stream using averaged properties for the oil and water. The use of 2-phase models with

averaged properties generally gives acceptable results unless either: emulsions are present;

or the flowrates are low enough to cause stratification of all three phases. These problemsare discussed in more depth in Section 3.2.1

Mechanistic Models vs. Correlations

The prediction of multiphase flow behavior has improved considerably during the 50+

years that the subject has been studied. For many years, multiphase flow prediction

methods were correlations, based on curve fits of experimental data. The correlations

frequently used arbitrarily selected variables and were based on limited databases,

consisting almost entirely of low pressure, small diameter data. Extrapolations of theseprediction methods to field conditions frequently proved to be seriously in error. In the

1960s and 1970s, several investigators undertook experimental studies to try to

understand the fundamental mechanisms of the various flow regimes. In the past 15 years

models have been developed, which are based on simulation of these mechanisms. These

models, referred to as mechanistic models, have proven to extrapolate best to field

conditions.

Newtonian vs. Non-Newtonian Fluids

Most condensates and crude oils follow Newtons law of viscosity, which is defined as:

yxxdv

dy=

where yx = shear stress

= viscosity

vx = velocity

y = distance

Some liquids, however, exhibit behavior that is very different from Newton's law. These

fluids are referred to as non-Newtonian. In the oil field, examples of non-Newtonian fluids

are drilling muds, polymeric additives, and crude oils below their cloud point.


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Flowline simulators are based on Newtonian fluids. If a non-Newtonian liquid is present,

the simulator must be tricked into giving a Newtonian viscosity equivalent to the actual

viscosity at the given temperature and shear stress. The methods of doing this are beyond

the scope of this guide.


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Figure I:1-1 Flow Regimes in Horizontal Flow


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Figure I:1-2 Flow Regimes in Vertical Flow


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SECTION 2.0 - OVERVIEW OF MULTIPHASE FLOW FUNDAMENTALS

2.1 Design Criteria

The majority of lines are sized by use of three primary design criteria: available pressure

drop; allowable velocities; and flow regime. In some cases, a more optimal line size may

be found that better suits the overall design of the pipeline system. These considerations

will be discussed later in the transient modeling section of the guide. A description of each

of the primary design criteria follows in Sections 2.2, 2.3, and 2.4.

2.2 Velocity Guidelines

The velocity in multiphase flow pipelines should be kept within certain limits in order to

ensure proper operation. Operating problems can occur if the velocity is either too high

or too low, as described in the following sections.

It is difficult to accurately define the point at which velocities are "too high" or "too low".

This section of the guide will try to quantify limits, but these limits should be considered

as guidelines and not absolute values.

Maximum Velocity

For the maximum design velocity in a pipeline, API RP-14E recommends the following

formula:

VC

ns

max = (Eqn. 2.1)

where Vmax = Maximum mixture velocity, ft/sec

ns = No-slip mixture density, lb/ft3

=( ) ( ) g sg l sl

m

V V

V

+

g = Gas Density, lb/ft3


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l = Liquid Density, lb/ft3

C = Constant, 100 for continuous service, 125 for intermittent service.

This equation attempts to indicate the velocity at which erosion-corrosion begins toincrease rapidly. Many people think this equation is an oversimplification of a highly

complex subject, and as a result, there has been considerable controversy over its use.

For wells with no sand present, values of C have been reported to be as high as 300

without significant erosion/corrosion. For flowlines with significant amounts of sand

present, there has been considerable erosion-corrosion for lines operating below C= 100.

The use of the API equation has been the subject of several research projects. It has been

generally agreed that the form of the equation is not sophisticated enough, and should

include additional parameters. Unfortunately, no other equation has been proposed which

has gained acceptance in the industry as an alternative to the API equation. As a result,

the recommended maximum velocity in the pipeline is the value calculated from Equation

2.1 with a Cvalue of 100.

It should be noted that Equation 2.1 is also used by many people as an estimate of the

maximum velocity for noise control.

For additional guidance on the use of the API equation, refer to Chevrons Piping Manual.

Minimum Velocity

The concept of a minimum velocity for the pipeline is an important one and should be

considered in the design of the line. Turndown conditions frequently govern the design of

the downstream equipment. Velocities that are too low are frequently a greater problem

than excessive velocities, so that the designers natural tendency to add "a bit of fat" to the

design by increasing pipe diameter can cause severe problems in the operation of the line

and the downstream facilities.

At low velocities, several operating problems may occur:

a) Water may accumulate at low spots in the line. If there is an appreciable amount of

CO2 or H2S in the well stream, this water may be very corrosive.


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b) Liquid holdup may increase rapidly at low mixture velocities. A large accumulation of

liquid may cause problems in downstream separators or slug catchers if the line is

pigged or the rate is changed rapidly.

c) Low velocities may cause terrain induced slugging in hilly terrain pipelines and

pipeline-riser systems.

It isnt possible to give a simple formula quantifying the velocity when the phenomena

discussed above will occur. The minimum velocity depends on many variables, including:

topography; pipeline diameter; gas-liquid ratio; and operating conditions of the line. A

ball-park value for the minimum velocity would be a mixture velocity of 5-8 ft/sec. The

actual value of the minimum velocity can only be quantified by simulation of the system

using the methods discussed in Section 5.2.2.

2.3 Flow Regimes

As discussed in Section 1, the gas and liquid in the pipe are distributed differently in each

of the four major flow regimes (stratified, annular, slug, and dispersed bubble flows). The

prediction of the correct flow regime is important for several reasons. The flow regime

prediction can show whether the line will operate in a stable flow regime or an unstable

regime. The prediction of liquid holdup and pressure drop are highly dependent on the

flow regime, with each regime exhibiting different behavior when the design variables are

changed.

The transitions between the flow regimes are frequently depicted in a flow regime map,

such as that shown in Figure I:2-1. The flow regime map typically has the superficial gas

velocity (Vsg) on the X-axis and the superficial liquid velocity (Vsl) on the Y-axis. As

discussed later in this section, the flow regime map is only valid for a single point in the

pipeline. As the angle of inclination, pressure and temperature change with position in the

pipeline, the flow regime map also changes.

Some general comments, however, can be made about the flow regime transitions.

Stratified flow occurs at low superficial gas and liquid velocities. Dispersed bubble flow

occurs at high superficial liquid velocities. Annular flow occurs at high superficial gas


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velocities. Slug flow occurs at moderate superficial gas and liquid velocities. Figure I:2-2

shows a typical flow regime map for vertical flow.

Many experimental studies of the transitions between the flow regimes for various systems

have been made, and many flow regime transition prediction methods have been published.

Some of these methods work fairly well, but most are poor. The designer needs to

carefully choose the method that will work best for the set of conditions. The best

methods are discussed in the remainder of this section.

Experimental studies of flow regime transitions have shown that each of the flow regime

boundaries reacts differently to changes in the system variables. The following table shows

the sensitivity of the transitions to changes in the major system variables:

Transition

Variable

Slug to

Dispersed

Bubble

Slug to

Annular

Slug to

Stratified

Stratified to

Annular

Angle of

Inclination

Small Effect Moderate

Effect

Strong Effect Strong Effect

Gas Density Small Effect Strong Effect Strong Effect Strong Effect

Pipeline

Diameter

Small Effect Small Effect Strong Effect Moderate

Effect

Liquid Physical

Properties

Moderate

Effect

Small Effect Moderate

Effect

Moderate

Effect

Many people have attempted to develop simple flow regime maps, usually using some

arbitrary dimensionless parameter on each axis (e.g. Baker, Beggs & Brill). These

methods are inherently inaccurate since no single parameter can model the sensitivity

effects shown in the previous table. The only flow regime map prediction methods that

have been effective for a wide range of conditions are those using mechanistic models to

estimate the flow regime transitions.


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In 1976, Taitel and Dukler published a landmark article describing a method of predicting

flow regime transitions by modeling the mechanism of each transition. By modeling each

transition, this method can show the same type of behavior observed in the experimental

work. The original Taitel-Dukler paper covered flow regime transitions in near horizontalflow only, and one of the transitions (slug-dispersed bubble) is very much in error. Taitel

and his co-workers at the University of Tel Aviv have subsequently published several

articles that expand the range of angles of inclination and correct the errors in the original

paper. The Taitel-Dukler paper and the latest paper from Tel Aviv model flow regime

transitions for all angles of inclination.

The Taitel, et al. methods give reasonably good predictions of the various flow regime

transitions, and the accuracy of the predictions has improved with each revision.

Another approach to the modeling of flow regime transitions is the method used in the

OLGAS method. It employs mechanistic models of each flow regime and links the models

by the assumption that the flow regime giving the lowest liquid holdup is the correct one.

This assumption holds up well in practice. The OLGAS method predicts flow regime

transitions with similar accuracy to the Taitel, et al. models.

Within Chevron, there are several programs available for flow pattern prediction.

Pipephase will print a flow regime map based on the Taitel-Dukler method for near

horizontal flow and the Taitel-Dukler-Barnea model for near vertical flow. Unfortunately,

these methods are the oldest and weakest of this family of methods. Two programs are

available within CPTC that incorporate the latest versions of the Taitel, et al. models.

These programs are FLOPAT, developed by Tulsa University, and FLEX, developed by

Advance Multiphase Technology. CPTC should be consulted if it is desired to use these

programs.

As in many aspects of multiphase flow, the flow regime prediction methods are not exact.

Errors of +/- 25% for the transition velocities are typical, even for the best prediction

methods. If the Taitel-Dukler map is used, the designer should be aware of the gross

errors in the slug to dispersed bubble transition. The errors for this transition can be

1000%. The dispersed bubble to slug transition typically occurs at a superficial liquid

velocity of about 10 ft/sec. Taitel-Dukler frequently predicts this transition velocity to be

50-100 ft/sec.


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2.4 Pressure Gradient

In most pipelines, the pipeline diameter is determined by the allowable pressure drop in the

line. The overall pressure gradient is composed of three additive elements:

a) pressure drop due to friction;

b) pressure changes due to elevational effects;

c) accelerational losses.

The calculation of the constituent parts of the pressure gradient will be discussed in the

next three sections.

The Chevron Fluid Flow Manual contains a good discussion of these pressure loss terms

for single phase flow and can be consulted as a reference.

2.4.1 Frictional Losses

In multiphase flow, frictional losses occur by two mechanisms: friction between the gas or

liquid and the pipe wall; and frictional losses at the interface between the gas and liquid.

The friction calculations, therefore, are highly dependent on the flow regime, since the

distribution of liquid and gas in the pipe changes markedly for each regime.

In stratified flow, there is wall friction between the gas and the pipe wall at the top of the

pipe, and wall friction between the liquid and the wall at the bottom of the pipe. There is

also friction between the gas and liquid at the gas-liquid interface. The interfacial friction

can be similar in magnitude to the wall friction if the interface is smooth, or it can be

considerably higher if waves are present.

In annular flow, there is friction between the liquid film and the wall. There is also

considerable interfacial friction between the gas in the core of the pipe and the liquid film.

The interfacial friction is usually the larger component.

In dispersed bubble flow, friction occurs between the liquid and the wall. There is

negligible interfacial friction between the gas and liquid.


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Slug flow has several frictional components. In the slug, the friction losses are caused by

the friction between the liquid and the pipe wall. In the gas bubble, the frictional

components are the same as in stratified flow, namely gas and liquid friction with the pipe

walls and interfacial friction between the gas and liquid. The friction loss in the slug isusually much higher than the losses in the bubble.

2.4.2 Elevational Losses

Elevational losses may be the major pressure loss component in vertical flow and flow

through hilly terrain. The calculation of elevational losses is governed by the following

equation:

dp

dxelev

=

mix

c

g sin

144g

where: (dp/dx)elev = Pressure gradient due to elevation, psi/ft

mix = Mixture Density, lb/ft3

= (l) (Hl) + (g) (1-Hl)

Hl = Liquid Holdup

g = Acceleration due to gravity, 32.2 ft/sec2

= Angle of inclination

gc = Gravitational conversion factor, 32.2 lb-ft/(lbf-sec2)

In order to calculate the elevational gradient, the liquid holdup must be determined. The

holdup in each flow regime has its own sensitivity to the important operating variables. A

summary of the effect of the major operating variables on the liquid holdup is:


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Slug Flow Annular

Flow

Stratified Flow Dispersed

Bubble Flow

Superficial Gas

Velocity

Strong Strong Strong Strong

Superficial Liquid

Velocity

Strong Strong Strong Strong

Gas Density Moderate Strong Strong None

Pipeline Diameter Moderate Weak Weak Weak

Angle of

Inclination

Moderate Weak Very Strong None

Liquid Properties Moderate Moderate Moderate Weak

As seen in the previous table, the influence of the major variables on the holdup is very

different for each of the flow regimes. As a result, it is impossible to develop a general

holdup correlation that will apply to all the flow regimes. Unfortunately, almost all of the

commonly used holdup methods available in commercial software try to do this. They

work poorly over much of the operating range as a result. The only way to accurately

predict liquid holdup is to use mechanistic models for each of the flow regimes. The

accuracy of available holdup methods is discussed further in Section 3.2.4.

2.4.3 Acceleration Losses

Although acceleration losses are present for all flow regimes, they are only significant for

two flow regimes: annular flow and slug flow. The mechanisms for the losses in these two

flow regimes are very different and will be discussed separately.

In single phase flow, acceleration losses can be calculated from Bernoullis equation.

Acceleration losses represent the change in kinetic energy as the fluid flows down the

pipe. The expression for acceleration gradient is:

dp

dx

V

g

dV

dxaccel c

=

144


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where: = Density, lbm/ft3

V = Velocity, ft/sec

For multiphase flow, the same type of relationship holds except that it refers to the flowof the mixed phase fluid. Most methods assume a no-slip mixture and use the no-slip

mixture density (ns) and the mixture velocity (Vm) in the calculation of acceleration

losses.

The kinetic energy acceleration losses are small for most oil industry applications. The

main exception is high velocity flow through low pressure piping. Flare systems would be

an example of piping that has high acceleration losses. Acceleration may account for 30-

50% of the overall pressure loss in such lines. For a typical high pressure gathering system

line, acceleration is usually less than 1% of the total drop and is frequently ignored.

Please note that the present version of Pipephase, 6.02, does not properly account for

acceleration losses, and, as a result, is not suitable for use in flare system design.

In slug flow, another source of acceleration contributes significantly to the total pressure

drop. As a slug propagates down the pipeline, it overruns and entrains the slower moving

liquid from the film ahead of the slug front. Accelerating the liquid from the film velocity

to the slug velocity can produce significant pressure losses. The acceleration loss may be

anywhere from


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flowrate, and composition that the gathering system must handle during the life of

the field.

c) If it isnt feasible to do the rigorous simulations for a gathering system, the allowable

pressure drop can be estimated from the initial wellhead pressure and the processing

plant inlet separator pressure. A rule of thumb to use for this method is to take 1/3

of the difference between the wellhead pressure and the separator pressure as the

allowable pressure drop in the pipeline. The remainder of the difference would equal

the initial choke pressure drop. This approach would allow for future operation at

reduced reservoir pressures.

d) A rule of thumb estimate of allowable pressure drop for long distance

gas/condensate pipelines is to allow 10-20 psi per mile for frictional pressure drop at

design rates.

2.5 Pressure Gradient Calculations

As indicated in sections 2.4.1 to 2.4.3, the calculation of the pressure gradient for

multiphase flow is very complicated. Hundreds of methods have been proposed to predict

pressure drops, but only a few methods work well over a wide range of conditions. The

best available methods are discussed in Section 3.2.4.


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2.6 Section Highlights

Points to remember from Section 2.0 -

No other equation has gained acceptance in the industry like the API equaton. Therecommended maximum velocity in the pipeline is the value calculated from Equation 2.1

with a Cvalue of 100.

The Taitel et al. Methods give reasonably good predictions of the various flow regimetransitions. The accuracy of the predictions has improved with each revision.

The OLGAS method predicts flow regime transitions with similar accuracy to the Taitel etal. methods.

If the Taitel-Dukler map is used, the designer should be aware of the gross errors in theslug to dispersed bubble transiton.

Overall pressure gradient is composed of three additive elements: pressure drop due to friction pressure changes due to elevational effects accelerational losses

Frictional calculations are highly dependent on the flow regime, since the distribution ofliquid and gas in the pipe changes markedly for each regime.

Elevational losses may be the major pressure loss component in vertical flow and flow

through hilly terrain.

Using mechanistic models for each flow regime is the only way to accurately predict liquidholdup.

Kinetic energy acceleration losses are small for most oil industry applications. The mainexception is high velocity flow through low pressure piping.

Pipephase 6.02 does not properly account for acceleration losses and is not suitable foruse in flare system design as a result.

For plant piping, rule of thumb values for pressure gradients, such as a frictional gradientof 0.2-0.5 psi per 100 ft. of equivalent length, are generally used.

The allowable pressure drop for a gathering system can be estimated from the initialwellhead pressure and the processing plant inlet separator pressure. The rule of thumb for

this method is to take 1/3 of the difference between the wellhead pressure and the

separator pressure as the allowable pressure drop in the pipeline.


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Figure I:2-1 Horizontal Flow Regime Map


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Figure I:2-2 Vertical Flow Regime Map


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SECTION 3.0 - STEADY STATE DESIGN METHODS

3.1 Pipeline Design Methods

As stated in the previous sections, the pipeline designer needs to estimate the pressure

drop, flow regime, and velocities in the line in order to select the proper line size. The

calculation of these parameters is laborious and is usually done by computer simulation.

Line sizing is usually performed by use of steady state simulators, which assume that the

pressures, flowrates, temperatures, and liquid holdup in the pipeline are constant with

time. This assumption is rarely true in practice, but line sizes calculated from the steady

state models are usually adequate.

Within Chevron, Pipephase and PIPEFLOW-2 are available for steady state pipeline

simulation.

For a more rigorous pipeline sizing, the simulations could be done using transient

simulators. Transient simulators allow changes in parameters such as inlet flowrate and

outlet pressure as a function of time, and calculate values for the outlet flowrates,

temperatures, liquid holdup, etc. as a function of time. If the line is operating in slug flow,

the line size calculated from the transient model may be different from that calculated

from a steady state simulator.

The principal uses of transient simulators are in the design of downstream equipment and

the development of operating guidelines. Transient simulators can model transient

behavior such as slug flow, pigging, rate changes, etc.

Transient simulators are quite new, developed in the last 10 years, and are not in general

use. Chevron has used the OLGA program for transient flowline analysis on several

projects, utilizing outside consulting services. CPTC developed an in-house transientsimulator, but it currently does not have as many features as the commercially available

codes.

The use of steady state models will be further discussed in Section 3.2, and transient

modeling will be briefly discussed in Section 4.


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3.2 Steady State Simulators

This section contains some general guidelines on the use of steady state simulators.

Although there are several steady state programs available, the discussion will center on

the use of Pipephase, which is Chevrons currently recommended simulator. The topics

covered include:

a) Phase Equilibrium and Physical Properties

b) Pipeline Elevation Profile

c) Heat Transfer

d) Recommended Methods for Pressure Drop, Liquid Holdup, and Flow Regime Prediction

e) Interpretation of Results

3.2.1 Phase Equilibrium and Physical Properties

Accurate prediction of the phase behavior and physical properties for the fluid flowing

through the pipeline is essential to a good simulation of the pipeline operation. The

estimates of these parameters depend in large part on the quality of the input data

available.

During conceptual design work, the only data available may be an estimate of the oil rate

and gas-oil ratio. After well tests have been performed, compositions of the wellstream

and PVT data may be available as well as projections of the flowrates of oil, gas and water

as a function of time. Obviously, as the accuracy of the input data improves, the quality of

the pipeline simulation improves.

Pipephase has two fundamentally different models available within it for estimation of

phase behavior and physical properties. The black oil model estimates the phase behavior

and physical properties by use of a series of correlations that are based on operating

temperature, pressure and some global parameters such as specific gravity of the oil and

gas. Compositional models use an equation of state to estimate the quantity of liquid and

gas at the operating conditions; then, correlations are used to estimate the physical

properties.

The decision on whether to use the black oil model or compositional modeling depends on

the available information and the type of system that is being modeled.


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The choice of models for gas-condensate and volatile oil systems is clear. Compositional

models should be used for any gas-condensate or volatile oil system. This recommendation

covers gas-oil ratios above about 3500 SCF/bbl.

For lower gas-oil ratios, the choice of models is more difficult. Compositional models

should give more accurate phase equilibrium results, but the physical property estimates

from the compositional models may not be as good as the black oil model. (Section 6-1

illustrates this point.) As a result, it cannot be stated categorically that either the black oil

model or the compositional model is superior for low gas-oil ratio systems. General

practice with Pipephase has been to use the black oil model for lower gas-oil ratio streams.

The accuracy of compositional modeling depends, in a large part, on the characterization

of the heavy ends of the well stream. The materials heavier than hexane (C6+) are usuallycharacterized by use of pseudo-components or cuts. The heavy ends could be

characterized by one C6+ cut, or by a series of cuts corresponding to various boiling

ranges. In general, the accuracy of the predictions increases when more cuts are used.

Pipephase requires two of the following parameters in order to characterize a cut: specific

gravity; molecular weight; or normal boiling point. In many cases, the mole fractions for

cuts heavier than C6 may have been measured in the PVT analysis, but cut properties

were not noted. In cases like this, the customary assumption is to use the properties of the

corresponding normal paraffin as the cut properties. This adds some error to the analysis,

but it is unavoidable in many circumstances.

If tests of the phase equilibrium and physical properties have been done as part of the

wellstream analysis, Pipephase allows the users of the black oil model to adjust the model

predictions for solution GOR, densities, and liquid viscosity to match experimental

values. The pipeline predictions after PVT matching should be considerably better than

those obtained with use of the standard correlations.

If the compositional model is used in Pipephase, the only variable that can be easily

manipulated to match experimental data is the liquid viscosity. Pipephase does not have an

option that will automatically adjust the phase equilibrium calculations to match

experimental data. It is possible to manually modify the phase equilibrium calculations, but


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it requires considerable effort, and the methods to do this are beyond the scope of this

guide.

Although it is possible to get good estimates of the phase equilibrium for 3-phase (gas-oil-

water) systems, the available software does not allow rigorous simulation of threephase

flow. The models present in Pipephase can only do two-phase (gas-liquid) flow

calculations. Pipephase averages the properties of the liquid hydrocarbon and liquid

water, and uses those average in the two-phase flow methods. Volumetric averaging,

however, may not give good values for the viscosity and surface tension of the mixture. If

the oil and water form an emulsion, the viscosity estimate may be off considerably using

simple volumetric averaging, because the viscosity of an emulsion can be as much as 50

times as high as the viscosity of the oil or water. If it is likely that an emulsion will form,

the Woeflin method, which is available in Pipephase, should be used to estimate the

viscosity of the emulsion.

3.2.2 Pipeline Elevation Profile

The pipeline elevation profile used in the simulation can have a significant impact on the

calculated pressure drop. Because the liquid holdup in upwardly inclined flow is greater

than the holdup in downward flow, the elevational pressure drop in uphill legs is greater

than the pressure recovery in downhill legs. As a result, elevational losses can account for

much of the pressure drop in hilly terrain pipelines, even if the inlet and outlet of the line

are at the same relative elevation.

If the velocities in the line are high, the uphill and downhill holdups may be close. As the

mixture velocity decreases, there will be an increasing difference between uphill and

downhill holdups.

The following table illustrates how sensitive the liquid holdup is to mixture velocity at

various angles of inclination from horizontal. The feed stream is a gas-condensate with

about 4 bbl/mm SCF of liquid present. (The values shown are predictions of the OLGAS

model.)


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MIXTURE VELOCITY FT/SEC

ANGLE,

DEGREES

2.7 4.1 5.4 8.1 16.2

-2.0 0.0041 0.0053 0.0064 0.0091 0.0115

-1.0 0.0052 0.0068 0.0085 0.0108 0.0122

-0.5 0.0068 0.0087 0.0108 0.0124 0.0126

0.0 0.0224 0.0218 0.0198 0.0156 0.0131

0.2 0.5797 0.4134 0.2249 0.0179 0.0134

0.5 0.5961 0.4988 0.3846 0.0317 0.0135

1.0 0.5997 0.5000 0.4314 0.3023 0.0144

2.0 0.6009 0.5024 0.4337 0.3428 0.0158

Using the values in the above table, a comparison of two models for a given section of a

pipeline has been made. In the first model, the pipeline segment consists of two equal

length sections of -0.5 degree and +0.5 degree each. The second model consists of a

single horizontal pipeline segment. The liquid holdups for the two models are:

MIXTURE VELOCITY,

FT/SEC

HOLDUP FOR -0.5

DEGREE AND +0.5

MODEL

HOLDUP FOR

HORIZONTAL MODEL

2.7 0.3015 0.0224

4.1 0.2538 0.0218

5.4 0.1977 0.0198

8.1 0.0221 0.0156

16.2 0.0131 0.0131


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The liquid holdups are far apart at low velocities and are the same at higher velocities.

This comparison makes two points:

The pipeline profile must be realistic if the calculations of liquid holdup and pressuredrop are to be accurate.

Low velocities cause severe problems in prediction of the pipeline performance.

For very low velocities, it would be necessary to know the pipeline elevation profile

within an accuracy of about one pipe diameter in order to get accurate holdup predictions.

This is generally not practical.

In many cases, the pipeline topography is not known when the preliminary pipeline sizing

calculations are run. Frequently, in offshore pipeline design, the designer only knowswater depths at subsea wells or platforms. Instead of assuming a straight line pipeline

profile, it is recommended that the designer add some terrain features to the pipeline

profile to simulate hills and valleys that are inevitably present in the actual profile.

To improve the accuracy of the simulation, many calculation segments should be used in

simulating the pipeline. Increasing the number of calculation segments always improves

the accuracy of the simulation, but it increases the computer simulation time. The number

of segments required depends on how rapidly the temperature, pressure and holdup are

changing in the pipeline. For a system with rapid changes in pressure, e.g. flare systems,

the number of calculation segments should be greater. If the temperature and pressure are

changing slowly, a more coarse grid can be used to simulate the pipeline.

3.2.3 Heat Transfer

The temperature profile along the pipeline is important in many circumstances. The

amount of condensation of liquids along a gas-condensate line, for instance, has a large

impact on the pressure drop and liquid holdup in the line. Hydrate and wax deposition may

occur in the line, requiring accurate estimates of temperatures. Corrosion is a strong

function of temperature, so good heat transfer estimates are vital to corrosion prediction.

To properly model the heat transfer between the pipeline and the surroundings, it is

necessary to have information on the following:


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thicknesses of the pipewall, pipeline coatings and insulation

whether the pipe is buried or exposed

the burial depth of the line

type of surroundings

ambient temperatures

thermal conductivities of the pipe, coatings and insulation.

With this information the programs can calculate heat transfer coefficients, which are then

used to calculate the temperature profile in the pipeline.

Values of the thermal properties for various materials can be read from the followingtable. Note that the Chevron Fluid Flow manual also has an extensive list of thermal

conductivities for various types of materials.

Material Thermal

Conductivity,

Btu/hr-ft-degF

Specific Heat,

Btu/lb-degF

Density, lb/ft3

Carbon Steel 26 0.11 490

Stainless Steel 8-13 0.11 488

Concrete

(Saturated)

0.75-1.2 0.10 147-200

Onshore Soil (Wet) 1.35 0.20 90-110

Subsea Sandy Soil 1.25-1.50 0.30 105-115

Coal Tar Epoxy 0.20 0.35 92

Fusion Bonded

Epoxy

0.15 0.32 75-90

Neoprene 0.12-0.15 0.50 90

Polyurethane Foam 0.011-0.022 0.38 2-12


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At the early stages of a project, there may not be enough information to enable rigorous

calculation of the heat transfer coefficient.. The following are rule of thumb values for heat

transfer coefficients for subsea flowlines, which can be used in these instances:

Applications U Value, BTU/hr/ft2/degF

Wells 2

Risers 20-40

Buried Pipelines 1-3

Concrete Coated Nonburied Pipelines 3-5

Nonburied Pipelines without Concrete 5-10

For gas/condensate pipelines, temperature loss by the Joule-Thomson expansion (J-T)

effect can be significant. In many gas pipelines, the temperature of the gas leaving the

pipeline is less than ambient because of the J-T effect.

Several concerns arise when using Pipephase for heat transfer calculations:

a) Pipephase only estimates temperature loss by the Joule-Thomson expansion cooling

effect if the compositional model is used. The J-T effect is ignored in black oil

simulations.

b) The default velocity of water flowing past a pipeline is 10 miles per hour in

Pipephase. This velocity is generally too high. More typical values are 1 to 3 ft/sec

(0.7-2 mph).

c) The Pipephase viscosity routine does not estimate viscosities at temperatures below

60 degrees F. At lower temperatures, it uses the viscosity at 60 degrees F. This can

lead to errors for pipelines in deep water or cold climates.

d) The thermal conductivity for saturated concrete is much higher than that for dry

concrete. The saturated concrete value should be used for subsea pipelines with

concrete coating.


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e) Unless a value is entered for Hrad, radiation is ignored in the heat transfer calculations.

For subsea or buried pipelines, radiation is negligible, but it can be a significant effect

for surface flowlines.

f) The convective heat transfer routines in Pipephase are not very rigorous. Errors inheat transfer calculations can occur for systems in which convection is the prime

source of heat transfer.

3.2.4 Recommended Methods for Pressure Drop, Liquid Holdup, and Flow

Regime Prediction

There have been hundreds of multiphase flow design methods developed in the past 50

years. Most computer programs contain dozens of options to select for pressure drop,

liquid holdup, and flow regime predictions. Most of these methods only have small rangesin which their predictions are accurate. This section of the guide discusses this problem

and gives some recommendations on which methods to use for certain applications.

Most of the methods available in Pipephase are correlations based on data taken in small

diameter (0.5-2 inch) test loops having an air-water flow operating at nearly atmospheric

pressure. The correlations developed from these data sets frequently do not include the

effects of all the key variables, such as pressure, because changes in these variables were

not studied in the experimental work. These correlations extrapolate poorly from field

conditions.

In the past 10 years, the development of mechanistic modeling has created a marked

improvement in prediction capabilities. As noted in Section 1.2, mechanistic models

attempt to model the physical phenomena associated with each flow regime. Mechanistic

models solve a set of simultaneous equations developed for a specific flow regime.

Correlations for a few key parameters are required in order to solve the equation set.

Mechanistic models extrapolate to field conditions much better than correlations because

the mechanistic models account for the effects of all the major variables.

Several mechanistic models have been developed in the past few years. Tulsa University

has developed models for near vertical flow (Ansari) and a general model covering all

inclinations (Xiao). The physics in these models are good, but the correlations built into

them are based solely on small diameter, low pressure data. The OLGAS model is


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currently the best available method for general multiphase flow calculations. OLGAS is

based on a wide range of data (diameter from 1 to 8 inches, pressures from atmospheric

to 1400 psi), and it extrapolates best to field conditions.

OLGAS is a proprietary program that has not been available within Chevron. As this

guide is being written, however, negotiations are underway to add OLGAS to Pipephase

and several other programs as options. If OLGAS becomes available, it is the

recommended method for prediction of pressure drop, liquid holdup and flow regime.

Methods are available that are as good or slightly better than OLGAS in certain ranges,

but they are not as good overall.

The following methods can be used in Pipephase as a check of OLGAS or as the design

method if OLGAS is not available:

a) Pressure Drop

1) Near Horizontal Low Gas-Oil Ratio - Beggs and Bril1 (Moody) is good.

2) Near Horizontal Gas/Condensate - Eaton-Oliemans is good for relatively high

velocities. All of the methods are poor for low velocities.

3) Near Vertical Gas/Condensate - Both Gray and Hagedorn-Brown are good.

4) Near Vertical Gas/Oil - Hagedorn and Brown is good.

5) Inclined Up - Nothing is good; Beggs and Brill (Moody) is fair.

6) Inclined Down and Vertical Down - Everything is poor. Use Beggs and Brill

(Moody), but answers may be suspect at times.

b) Liquid Holdup

1) Near Horizontal Low Gas-Oil Ratio - Beggs and Brill (Moody) is O.K.

2) Near Horizontal Gas/Condensate Lines - All available methods are poor. The

Eaton holdup correlation is better than the other methods.

3) Near Vertical Gas/Condensate - The most accurate method is no-slip.

4) Near Vertical Gas/Oil - Hagedorn and Brown is pretty good.

5) Inclined Up - Beggs and Brill (Moody) is usable for low GOR lines, nothing is

accurate for gas/condensate. If gas velocities are high, use no-slip; otherwise use


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Beggs and Brill (Moody). The user must be careful because the holdups can be a

factor of 10 in error in some cases.

6) Inclined Down and Vertical Down - Everything is poor. Use Beggs and Brill

(Moody), but answers may be suspect.

c) Flow Regimes

1) The Taitel-Dukler flow regime map is as good as OLGAS for near horizontal

flow with the exception of the slug-dispersed bubble boundary. This boundary is

very poorly predicted. If this method is used, it is recommended that a value of

~10 ft/sec be used as the superficial liquid velocity for the slug-dispersed bubble

transition rather than the Taitel-Dukler prediction.

2) The Taitel-Dukler-Barnea map for near vertical flow is also as accurate as

OLGAS.

On occasion, the conditions for a simulation may cause otherwise good multiphase flow

methods to give erroneous results. It is usually a good idea to spot-check the results by

use of another method to ensure that the answers are reasonable. If there is a wide

variance in results, CPTC should be contacted for guidance.

3.2.5 Interpretation of Results

When a multiphase simulator such as Pipephase is run, the interpretation of the results canbe difficult. The following section provides assistance in understanding Pipephase output,

and ensuring that the design criteria for the line (velocities, flow regime, and allowable

pressure drop) are met.

As discussed in Section 2.2, the velocity in the pipeline should be kept within a limited

range. Calculation of the velocities from a Pipephase output is not straightforward. The

designers of Pipephase chose to include the actual gas and liquid velocities in their output

table rather than the superficial gas and liquid velocities which are needed in the erosional

velocity calculations. As discussed in Section 1.2, the superficial and actual velocities are

related by simple formulas:

( )V Ug 1 Hsg l=


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and V U Hsl l= 1

The liquid holdup is read from the "slip holdup" column. This calculation is made more

difficult by the poor formatting of the liquid holdup in the Pipephase output table. (The

liquid holdup is shown to only two decimal places in the table. For gas-condensate lines, if

the liquid holdup is below 0.5 percent, the printout will show 0.00 for the holdup.)

A more accurate way of calculating the superficial velocities from the Pipephase output

tables which doesnt rely on reading the value for the liquid holdupis:( )( )

HU Vm

U Ul

g

g l

=

V U Hsl l l=

V V Vsg m sl=

To calculate the C value in the API-RP14E equation, the value of the no-slip mixture

density must be known. Pipephase apparently only calculates and tabulates this value in

the output table if the Beggs and Brill (Moody) method is used. If other methods are used,

a value of 0.00 is given in the output table for the no-slip mixture density. The no-slip

mixture density can be calculated, however, from the phase densities shown on the output

table and the superficial velocities calculated above:

( ) ( )

ns

g sg l sl

m

V V

V=

+

Pipephase allows the user to print a flow regime map based on either the Taitel-Dukler

map for near horizontal flow or the Taitel-Dukler-Barnea map for near vertical flow. The

flow regime map is printed only for the last "device" in a "link". If the "link" contains

several pipes with different inclinations, the flow regime map for some of these sections

may be quite different from the map at the last "device". The only way to print the flow

regime map at specific points along the line is to make these points ends of Pipephase

"links".


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The "link" summary tables print the flow regime predictions for each pipeline segment.

The printout shows both the predictions of the multiphase flow design method (e.g. Beggs

and Brill) and the Taitel-Dukler method. If OLGAS is available, the flow regime

predictions of OLGAS can be compared directly with the Taitel-Dukler prediction, and theuser can feel confident that the predicted flow regime is valid if the two methods match. If

methods other than OLGAS are used, disregard their flow regime predictions and only

consider the Taitel-Dukler predictions as reasonable.

Once the flow regime is determined, the designer needs to decide if this flow regime is

acceptable. This decision is more difficult than it may appear. Ideally, the flow line should

not be in the slug flow regime. In practice, it may be very difficult to design a line to avoid

slug flow under all anticipated flow conditions. The only variables the designer can change

are diameter and operating pressure; the changes in these variables required to avoid slug

flow may be impractical. It should be pointed out that many pipelines operate successfully

in slug flow. As long as the pipeline and downstream equipment are designed with proper

consideration of slug flow effects, they can be successfully operated.

The flow regime analysis may show that the line is in stratified flow. In many instances,

this is an excellent flow regime in which to operate. At low flowrates, however, slugging

may occur in lines predicted to be in stratified flow, induced by the terrain. Terrain

induced slugs are generally much longer than the slugs in normal slug flow and can cause

severe operating problems. Terrain slugging is discussed in more detail in Section 5.2.2.

If the pressure drop and velocities for lines in dispersed bubble or annular flow are within

acceptable limits, these flow regimes are usually good regimes in which to operate.

The pressure drop in the line should be compared with the allowable pressure drop. The

pressure drop in the line can be read from the Pipephase "link summary" table. It should be

pointed out that pressure drop is not always a maximum at the highest flowrate. If the

pipeline contains inclined or vertical elements, it is possible that the highest pressure drop

may occur at a low flow condition due to high elevational losses at low flows.

It is worthwhile to emphasize the point that the pipeline design should be checked at off-

design points as well as the nominal design point. For most pipelines, worst case

conditions for liquid holdup and flow regime occur at turn-down conditions.


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Compositional models should be used for any gas-condensate or volatile oil system.This recommendation covers gas-oil ratios above 3500 SCF/bbl.

General practice with Pipephase: use the black oil model for lower gas-oil ratiostreams.

If it is likely an emulsion will form, the Woeflin method (available in Pipephase)should be used to estimate the viscosity of the emulsion.

The pipeline profile must be realistic if the calculations of liquid holdup andpressure drop to be accurate.

Low velocities cause severe problems in prediction of the pipeline performance.

If OLGAS becomes available, it is the recommended method for prediction ofpressure drop, liquid holdup, and flow regime.

Mechanistic models extrapolate to field conditions much better than correlations,since the mechanistic models account for the effects of all the major variables.

The following methods can be used in Pipephase as a check of OLGAS or as thedesign method if OLGAS is not available:

1. Pressure Drop

a) Near Horizontal Low Gas-Oil Ratio - Beggs and Brill (Moody) is good.

b) Near Horizontal Gas/Condensate - Eaton-Oliemans is good for relatively

high velocities. All of the models are poor for low velocities.

c) Near Vertical Gas/Condensate - Both Gray and Hagedorn-Brown are

good.

d) Near Vertical Gas/Oil - Hagedorn and Brown is good.

e) Inclined Up - Nothing is good; Beggs and Brill (Moody) is fair.

f) Inclined Down and Vertical Down - Everything is poor. Use Beggs and

Brill (Moody), but answers may be suspect at times.


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Liquid Holdup

g) Near Horizontal Low Gas-Oil Ratio - Beggs and Brill (Moody) is O.K.

h) Near Horizontal Gas/Condensate Lines - Nothing is accurate. The Eaton

holdup correlation is poor, but better than the other methods.

i) Near Vertical Gas/Condensate - The most accurate method is no-slip.

j) Near Vertical Gas/Oil - Hagedorn and Brown is pretty good.

k) Inclined Up - Beggs and Brill (Moody) issuables for low GOR lines,

nothing is accurate for gas/condensate. If gas velocities are high, use no-

slip; otherwise use Beggs and Brill (Moody). Be careful because the

holdups can be a factor of 10 in error in some cases.

l) Inclined Down and Vertical Down - Everything is poor. Use Beggs and

Brill (Moody), but answers may be suspect.

Flow Regimes

a) The Taitel-Dukler flow regime map is as good as OLGAS for near

horizontal flow with the exception of the slug-dispersed bubble boundary.

This boundary is very poorly predicted. If this method is used, it is

recommended that a value of ~10 ft/sec be used as the superficial liquid

velocity for the slug-dispersed bubble transition rather than the Taitel-

Dukler prediction.

b) The Taitel-Dukler-Barnea map for near vertical flow is also as accurate

as OLGAS.

The flow line should, ideally, not be in the slug flow regime. In practice, it may bevery difficult to design a line to avoid slug flow under all anticipated flow conditions.

At low flow rates slugging may occur in lines predicted to be in stratified flow,induced by the terrain.

If the pressure drop and velocity for lines in dispersed bubble or annular flow arewithin acceptable limits, these flow regimes are usually good regimes in which to

operate.


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SECTION 4.0 - TRANSIENT FLOW MODELING

4.1 Transient Flow Modeling (General)

Transient multiphase flow simulators have only been developed recently. The first widely

used commercial program, OLGA, began development in about 1983 and has been

commercially available since 1990. OLGAs only current competitor, PLAC, was

introduced to the market at about the same time. Chevron currently does not own either

program but has used OLGA for specific projects through consultants. Chevron internally

developed a transient code, Transpire, in the same time frame as OLGA. This program

has not been widely used, and it does not have as many features as the commercial codes.

Steady state simulators assume that all flowrates, pressures, temperatures, etc. are

constant through time. Inherently transient phenomena, such as slug flow, are modeled byuse of their average holdups and pressure drops. Transient models allow all the input

variables to change with time. Transient programs can model phenomena such as slug

flow and can show the variations in parameters such as outlet gas and liquid flowrates as a

function of time. Transient simulators, therefore, model the actual operation of pipelines

closer and with more detail than steady state simulators.

Transient simulators solve a set of equations for conservation of mass, momentum and

energy to calculate the liquid and gas flowrates, pressures, temperatures and liquid

holdups. These calculations are done at each time step. The programs utilize an iterativeprocedure, which ensures that a set of boundary conditions (such as inlet flowrates and

outlet pressures as a function of time) are met while solving the conservation equations.

Steady state modeling can be used to size pipelines, but the predicted size may be

inaccurate if the line is in slug flow. Transient simulators can size pipelines more

accurately, and they are valuable in several other areas such as the design of downstream

facilities, development of operating guidelines, and the diagnosis of operating problems.

Steady state simulators cannot properly address any of these other concerns.

4.2 Use of Transient Simulators

Because of their power, transient simulators have been used for a variety of purposes.

These uses include:

a) Slug flow modeling


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b) Estimates of the potential for terrain slugging

c) Pigging simulation

d) Estimation of corrosion potential in low spots in the line

e) Startup, shutdown and pipeline depressuring simulations

f) Development of operating guidelines

g) Real time modeling, including leak detection

h) Operator training

i) Design of control systems for downstream equipment

j) Slug catcher design

A general guideline for the use of steady state and transient modeling would be to use

steady state modeling during the feasibility level design of a system but use transient

modeling in the final design of the pipeline and its associated equipment.

As transient simulators improve and computer power increases, it is likely that transient

simulators will eventually supplant steady state simulators.

Because Chevron does not own a transient simulator at this time, this guide does not

contain any guidelines for their use. Section 5.1 discusses the use of the OLGA programfor slug length prediction.



Transient simulators model the actual operation of pipelines much closer than steadystate simulators.

General guideline for the use of steady state and transient modeling: use steady statemodeling during the feasibility level design of a system, but use transient modeling in

the final design of the pipeline and its associated equipment.


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SECTION 5.0 - SLUG FLOW ANALYSIS

5.1 Slug Flow - General

The formation of slugs of liquid can be caused by a variety of mechanisms:

a) Hydrodynamic Slugging

b) Terrain Slugging

c) Pigging

d) Startup and Blowdown

e) Flowrate Changes

Each of the mechanisms will be briefly discussed here, and will be further discussed in

Section 5.2.

Hydrodynamic slugging refers to operating in the slug flow regime. In near horizontal

flow, slugs are formed by waves growing on the liquid surface to a height sufficient to

completely fill the pipe. When this happens, alternating slugs of liquid and bubbles of gas

flow through the pipe, as illustrated in Figure I:1-1.

Terrain slugging occurs when a low point in the line fills with liquid. The liquid does not

move until gas pressure behind the blockage builds high enough to push the liquid out of

the low spot as a slug. Terrain slugging can produce very long slugs in pipeline-riser

systems. Although terrain slugging occurs at low superficial gas and liquid velocities, the

actual velocities during slug release can be very high.

When a pipeline is pigged, most of the liquid inventory is pushed from the line as a liquid

slug ahead of the pig.

When a line is shut down, liquid that is left in the line will drain to the low points in the

line. When the flow is restarted, the accumulated liquid may exit the pipeline as a slug.


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When the flowrate is increased, the liquid holdup in the line decreases. This change in

holdup can either exit the line as a gradual increase in liquid flow, or it can come out in

the form of a slug, depending on the flowrate.

Each of the slug flow mechanisms is highly transient in nature. Steady state models cannot

properly simulate slug flow behavior and are very limited in their ability to predict slug

characteristics such as slug length and frequency. The next two sections of the guide

discuss the slug flow mechanisms in more detail, discuss available methods of predicting

slug flow behavior and give some recommendations on sizing of slug catchers and

separators.

5.2 Slug Length and Frequency Predictions

Although estimates of slug length and frequency are of prime importance in design of

pipeline system facilities, most of the prediction methods available are poor. Development

of prediction methods has been hampered by the difficulty of the problem and the meager

amount of available test data. This section discusses each of the mechanisms for slug flow,

discusses the available test data, and give recommendations on the best available

prediction methods.

5.2.1 Hydrodynamic Slugging

Experimental measurements of the slug length in hydrodynamic slug flow show several

interesting results:

a) The slug length is not constant. At a given point in the line, the slug length varies

greatly around an average value. Different investigators have characterized the slug

length distribution as log normal, truncated Gaussian, inverse Gaussian, or fractal

distributions. The maximum slug length may be several times greater than the

average.

b) The average slug length and the slug length distribution change with the position

down the pipe. Slugs may grow, dissipate, or merge as the flow continues down the

pipe. As a result, the average slug length usually increases with the position in the

pipe, while the standard deviation of the slug length distribution decreases.


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c) Slugs in vertical pipes are much smaller than slugs in horizontal pipes.

d) The slug length in laboratory experiments can be fairly well correlated. These tests

show that the average slug length (in feet) is approximately 32 times the Pipe

Diameter (in feet) for horizontal pipes.

e) In the data base of published pipeline field test results, the average slug length is

much higher than the results observed in the laboratory. The field tests results show

average slug lengths of 300-2000 times the Pipe Diameter, with some slugs as long

as 10,000 times the Pipe Diameter.

The differences between laboratory and field data shown in points d) and e) above are due

to factors such as:

- terrain features have a large effect on the slug length and frequency;

- slug flow in the field can be combination of mechanisms such as hydrodynamic

slugging causing terrain slugging;

- field pipelines are much longer, allowing more time for slug growth.

Average slug length is a complex function of many variables: the diameter and length of

the pipeline; the topography of the line; the gas and liquid superficial velocities; the liquid

physical properties; and the gas density.

Several correlations have been presented for the prediction of slug length and slug

frequency for horizontal piping and pipelines. Most of these correlations are based solely

on laboratory data, which means they are of limited use in the design of pipelines in the

field.

A few correlation methods have been presented based on field data. Two of these

methods, the Brill, et al. Correlation and the Hill & Wood method, have been widely used

for slug length prediction. Both methods will be discussed in detail.

Brill, et al. took several sets of data on 12 and 16 inch pipelines at Prudhoe Bay in about

1978. They were the first experimentalists to report the wide disparity between the

extrapolation of lab results and field data. They developed a simple correlation for slug


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CPTC NOVEMBER 1

multiphase design guide part1

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