Multiphase Pipeline Design Guide

PART I - MULTIPHASE PIPELINE & SLUG CATCHER DESIGN GUIDE

CPTC NOVEMBER 1994 1

Multiphase Pipeline Design Guide



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




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 PROBLEMS

6.1 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



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



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 report

briefly 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



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.



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:



Vsg = Superficial Gas Velocity, ft/sec

= (actual ft3/sec of gas) / (pipe cross-sectional area, ft2)

Vsl = Superficial Liquid Velocity, ft/sec

= (actual ft3/sec of liquid) / (pipe cross-sectional area, ft2)

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.



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



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 problems

are 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 these

prediction 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 Newton’s law of viscosity, which is defined as:

τ µyxxdv

dy=

where τyx = shear stress

µ = 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.



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.



Figure I:1-1 Flow Regimes in Horizontal Flow



Figure I:1-2 Flow Regimes in Vertical Flow



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

nsmax =

ρ(Eqn. 2.1)

where Vmax = Maximum mixture velocity, ft/sec

ρns = No-slip mixture density, lb/ft3

= ( ) ( )ρ ρg sg l sl

m

V V

V

+

ρg = Gas Density, lb/ft3



ρl = 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 to

increase 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 C value 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 Chevron’s 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 designer’s 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.



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 isn’t 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



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.



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 horizontal

flow 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.



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.



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 is

usually 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

dx elev

=

ρ αmix

c

g sin

144g

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

ρmix = Mixture Density, lb/ft3

= (ρl) (Hl) + (ρg) (1-Hl)

Hl = Liquid Holdup

g = Acceleration due to gravity, 32.2 ft/sec2

α = 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:



Slug Flow AnnularFlow

Stratified Flow DispersedBubble Flow

Superficial GasVelocity

Strong Strong Strong Strong

Superficial LiquidVelocity

Strong Strong Strong Strong

Gas Density Moderate Strong Strong None

Pipeline Diameter Moderate Weak Weak Weak

Angle ofInclination

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 Bernoulli’s 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

=

ρ144



where:ρ = Density, lbm/ft3

V = Velocity, ft/sec

For multiphase flow, the same type of relationship holds except that it refers to the flow

of the mixed phase fluid. Most methods assume a no-slip mixture and use the no-slip

mixture density (ρns) 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 <1% to more than 50% of the total pressure drop. Mechanistic models

include this loss, while most of the correlation based methods ignore this loss.

2.4.4 Allowable Pressure Drop

No clear-cut criteria exist for determining the amount of pressure drop to be allowed in a

pipeline design. Allowable pressure drop is a function of the parameters of the system

being designed. The following are some guidelines for specific systems:

a) For plant piping, rule of thumb values for pressure gradients, such as a frictional

gradient of 0.2-0.5 psi per 100 ft. of equivalent length, are generally used.

b) In the design of a gathering system, the ideal way to choose allowable pressure

drops is to simulate the system from the reservoir through the processing plant as a

function of time. This approach will account for the changes in reservoir pressure,



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

the field.

c) If it isn’t 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.



2.6 Section Highlights

Points to remember from Section 2.0 -

• No other equation has gained acceptance in the industry like the API equaton. The

recommended maximum velocity in the pipeline is the value calculated from Equation 2.1

with a C value of 100.

• The Taitel et al. Methods give reasonably good predictions of the various flow regime

transitions. The accuracy of the predictions has improved with each revision.

• The OLGAS method predicts flow regime transitions with similar accuracy to the Taitel et

al. methods.

• If the Taitel-Dukler map is used, the designer should be aware of the gross errors in the

slug 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 of

liquid 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 liquid

holdup.

• Kinetic energy acceleration losses are small for most oil industry applications. The main

exception is high velocity flow through low pressure piping.

• Pipephase 6.02 does not properly account for acceleration losses and is not suitable for

use in flare system design as a result.

• For plant piping, rule of thumb values for pressure gradients, such as a frictional gradient

of 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 initial

wellhead 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.



• The rule of thumb for estimating allowable pressure drop for long distance gas/condensate

pipelines is to allow 10-20 psi per mile for frictional pressure drop at design rates.



Figure I:2-1 Horizontal Flow Regime Map



Figure I:2-2 Vertical Flow Regime Map



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 transient

simulator, 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.



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 Chevron’s 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.



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 usually

characterized 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



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.)



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



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 pressure

drop 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 knows

water 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:



• 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 following

table. Note that the Chevron Fluid Flow manual also has an extensive list of thermal

conductivities for various types of materials.

Material ThermalConductivity,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 BondedEpoxy

0.15 0.32 75-90

Neoprene 0.12-0.15 0.50 90

Polyurethane Foam 0.011-0.022 0.38 2-12



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.



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 in

heat 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 ranges

in 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



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



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 can

be 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= −



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 doesn’t rely on reading the value for the liquid holdup is:

( )( )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:

( ) ( )ρ

ρ ρns

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".



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 the

user 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.





• 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 ratio

streams.

• 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 and

pressure 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 of

pressure 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 the

design 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.



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 be

very 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 are

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

operate.



SECTION 4.0 - TRANSIENT FLOW MODELING

4.1 Transient Flow Modeling (General)

Transient multiphase flow simulators have only been developed recently. The first widelyused commercial program, OLGA, began development in about 1983 and has beencommercially available since 1990. OLGA’s only current competitor, PLAC, wasintroduced to the market at about the same time. Chevron currently does not own eitherprogram but has used OLGA for specific projects through consultants. Chevron internallydeveloped a transient code, Transpire, in the same time frame as OLGA. This programhas 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 by

use 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 andenergy to calculate the liquid and gas flowrates, pressures, temperatures and liquidholdups. 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 andoutlet 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 beinaccurate if the line is in slug flow. Transient simulators can size pipelines moreaccurately, and they are valuable in several other areas such as the design of downstreamfacilities, 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



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 program

for slug length prediction.



• Transient simulators model the actual operation of pipelines much closer than steady

state simulators.

• General guideline for the use of steady state and transient modeling: 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.



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.



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.



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



length based on the Prudhoe Bay tests and small diameter laboratory data. Their

correlation is:

ln(Ls) = -2.663 + 5.44 (ln(D))0.5 + 0.059 ln(Vm)

where Ls = Average slug length, ft.

D = Pipeline inside diameter, inches

Vm = Mixture velocity, ft/sec

The Prudhoe Bay test data appeared to be a log normal distribution around the mean slug

length. Log normal distributions were fit to each of the Prudhoe Bay tests, and the mean

slug length and variance were calculated. Their method assumes that the variance for any

pipeline is the same as the average value of the variance from these tests. With this

assumption, the slug length distribution is the same for all pipelines. The distribution is:

Percent Probability Slug Length/Mean Slug Length

50.0000 1.00

84.1300 1.65

97.7200 2.72

99.8600 4.46

99.9900 6.42

99.9999 10.76

The Brill method is easy to use, and it has been used as the design basis for many

facilities. The Brill et al. model is the basis for the slug length predictions in Pipephase.

Unfortunately, the Brill method is very inaccurate. It gives poor predictions for almost

every data set that was not included in the original correlation. A comparison of the Brill

predictions against measured average slug lengths from a 16" laboratory line and two

field pipelines shows:



Actual Predicted

Data Source Avg. Slug Length, ft. Avg. Slug Length, ft.

BHRG 30 700

Field Data 1000 120

Field Data 400 1800

The inaccuracy of the Brill et al. method is due to the simplicity of their formula. The

slug length is predicted to be almost exclusively a function of the line diameter. As noted

previously, the slug length is a complex function of many variables, and a simple formula

like this cannot approximate reality.

There are two other methods based primarily on the Prudhoe Bay data, namely, the

Norris correlation and the Scott, et al. correlation. Their performance, while a bit better

than the Brill method for larger pipelines, is also weak. Both of these methods assume

that the slug length is only dependent on the line diameter.

In 1990, Hill and Wood of BP published a paper proposing an alternate way of modeling

slug frequency. Their work was based on both laboratory data and a large number of field

measurements. The model is more sophisticated than the Brill approach and attempts to

account for the many of the major variables. Their model correlates the slug frequency

with the diameter, gas-liquid slip velocity, and the equilibrium holdup at the beginning of

the pipeline. The model assumes a horizontal pipe and uses the Taitel-Dukler stratified

flow model to estimate the slip velocity and the equilibrium holdup at the beginning of the

pipeline. In order to calculate the Taitel-Dukler holdup, the value for the Lockhart &

Martinelli X factor is calculated from the following equation:

XV

Vsl

sg g g=

0 91

0 41

01. . .ρρ

µµ

where: µ1 = Liquid Viscosity, cp

µg = Gas Viscosity, cp



Once X is known, the value for the liquid holdup (Hle) for stratified flow can be read from

Figure I:5-1. The actual gas and liquid velocities can be calculated from the superficial

velocities and the stratified flow liquid holdup by the formulas:

UV

Hgsg

le=

−1

and UV

Hlsl

le=

Their equation for slug frequency is:

( )( )F

H U U

D Hs

le g l

le=

−

−

2 74

1

.

where Fs = Slug Frequency, slugs/hour

D = Pipe Inside Diameter, ft

Ug = Actual Gas Velocity, ft/sec

Ul = Actual Liquid Velocity, ft/sec

Hle = Liquid Holdup (based on stratified flow at the inlet of the pipe)

To calculate the slug length, the slug fraction, which is defined as the slug length divided

by the sum of the slug and bubble lengths, must be known. The slug fraction can be

calculated rigorously by use of mechanistic slug flow models, such as the Hubbard-Dukler

model or the Nicholson, Gregory, and Aziz model. A simplified alternative to the use of

these models is the following procedure.

First, the liquid holdup in the slug will be estimated. The Gregory, Nicholson, and Aziz

method is the easiest method to use, and it is as accurate as most of the methods. The

liquid holdup in the slug is given by:



HV

lsm

=

+

1

128 4

139

.

.

where Hls = Liquid Holdup in the slug

The liquid holdup in the bubble (Hlb) can be assumed to be ~0.20 for much of the slug

flow range. From a material balance, the slug fraction (SF) can be calculated from these

values and the overall liquid holdup prediction:

SFH H

H Hl lb

ls lb=

−−

The slug length can be calculated from:

LSF V

Fsm

s=

3600 ( )

where Ls = Mean Slug Length, ft

An example of the use of the Hill & Wood model is shown in Section 6.1.

To use the Hill & Wood model for slug catcher designs, the design slug length is needed

instead of the mean slug length. A rule of thumb, based on the limited amount of

experimental data available, is that the maximum slug length is approximately 6 times the

mean slug length.

The Hill & Wood model is more sophisticated than the Brill model and is based on a

broader data base. If a correlation based model is used for slug length prediction, the Hill

& Wood model is probably the best available method.

Another alternative approach to hydrodynamic slug length prediction is the method used

in the OLGA transient program. The slug tracking model in OLGA attempts to

dynamically track each slug in the system, from its inception to its dissipation or exit from

the pipeline. This method models the growth or dissipation of each slug and models

phenomena such as the merger of two slugs. The model accounts for terrain effects and



the effects of all of the major variables. This model, which has only been available to the

industry since early 1994, shows promise in prediction of slug lengths.

5.2.2 Terrain Slugging

Terrain slugging refers to slug flow resulting from a blockage of a low spot in the pipe

with liquid. When the liquid completely fills up a low spot, the liquid remains stationary

until gas pressure behind the slug builds to a suitably high level to push the liquid out as a

slug.

The most severe form of terrain slugging occurs in pipeline-riser systems, which have a

negatively sloping pipeline ahead of the riser. At low flowrates very large slugs may form

in this type of system. These slugs can be much larger than those associated with

hydrodynamic slugging. Figure I:5-2 illustrates the various stages in terrain slugging for

pipeline-riser systems.

Experimental work has shown that both hydrodynamic and terrain slugging may occur at

various points in the line, under certain conditions. This is one of the reasons that it is so

difficult to extrapolate laboratory slug length data to field conditions.

One of the problems with using a steady state program such as Pipephase is its inability to

identify terrain slugging. Pipelines in the terrain slugging regime will usually be identified

by the steady state simulator as being in stratified flow. If no further analysis is done, the

line may be designed in the belief that it is in a steady flow regime. On startup, the line

may experience slugs that may be thousands of feet long.

The best way to analyze the line for terrain slugging is to run a transient simulator. If

transient analysis is not possible, some simplified methods have been developed for

estimation of severe slug potential. These methods pertain only to the modeling of

pipeline-riser slugging. The method of Pots et al. Has been included in Pipephase, and it is

discussed here.

By equating pressure buildup rates, Pots developed a formula for a dimensionless severe

slugging factor πss:



( )πssg

l

c

f l

zRT

M

W

W

g

g L H=

−

144

1

where z = Gas compressibility

R = Gas Constant, 10.73 psia-ft3/(lb-mole-deg R)

T = Temperature, deg R

M = Gas Molecular Weight, lb/lb-mole

Wg = Gas mass flowrate, lb/sec

Wl = Liquid mass flowrate, lb/sec

gc = Conversion factor, 32.2 lb-ft/lbf-sec2)

g = Acceleration due to gravity, 32.2 ft/sec2

Lf = Length of the pipeline, ft

Hl = Liquid holdup in the pipeline

When the value for πss is less than 1, severe slugging could occur. Values less than 1 do

not ensure that severe slugging will occur. Pots indicates that the flowline must be in

stratified flow for severe slugging to take place. The Taitel-Dukler flow regime map can

be used to determine whether the flowline is in stratified flow at the pipeline-riser junction.

Pots also gives an expression for the slug length in severe slugging:

Ls = Lr / πss

where Ls = Slug length, ft

Lr = Riser height, ft

Pipephase calculates πss if the SLUG table option is turned on.

The equation for πss indicates that long pipelines, low pressure, and low gas-oil ratios

contribute to the likelihood and severity of terrain slugging.



The Pots method gives a reasonable approximation to terrain slugging behavior for a

simple constant slope pipeline-riser system, but it cannot rigorously model the

performance of a hilly terrain pipeline connected to a riser.

5.2.3 Pigging Slugs

Pigs are run through pipelines for a variety of reasons, including:

a) Liquid inventory control

b) Maintenance and data logging

c) Pipeline cleaning and dewaxing

d) Inhibitor application

The classic analysis of pigging was developed by Baker and McDonald in 1959. They

identified five zones in a pipeline undergoing pigging:

a) Undisturbed equilibrium flow

b) Liquid slug ahead of the pig

c) Pig itself

d) Dry gas region behind the pig

e) Reestablished equilibrium flow

Figure I:5-3 graphically illustrates the zones in a pigged pipeline.

Baker and McDonald developed a quasi-transient model of pigging in which the lengths

and pressure drops for the various zones were estimated and tracked with time. Although

some of the methods used in the Baker-McDonald model are antiquated, the model is

conceptually a very good representation of pigging.

Several investigators have modified the Baker-McDonald model to improve its

predictions. The model developed by Barua of Tulsa University has been included in

Pipephase.



The commercial transient simulators model the physics of pigging more rigorously than

the models of Baker & McDonald or Barua, and are more accurate. The simpler models,

however, give pretty good predictions. It is also possible to get a good estimate of the pig

volume and transit time by simple hand calculations.

To estimate the slug volume and transit time by hand, the designer can use the following

method. The pig and the slug in front of it move at a velocity that is equal to the gas

velocity behind the pig. For a first approximation, this velocity can be assumed to be

equal to the mixture velocity. The transit time of the pig would therefore be:

ttrans = Lf / Vm avg

where ttrans = Transit time for the pig, sec

Lf = Pipeline length, ft

Vm avg = Average mixture velocity, ft/sec

When the slug exits the line, the volume of liquid ahead of the pig is equal to the liquid

holdup in the line minus the amount of leakage past the pig minus the amount of liquid

produced while the pig is traversing the pipeline or:

Qslug = (Hl Ap Lf) (1 - fleak) - Vsl Ap ttrans

where Qslug = Volume of liquid ahead of the pig, ft3

Hl = Average liquid holdup in the pipe

Ap = Cross-sectional area of the pipe, ft2

fleak = Fraction of the liquid that leaks past the pig

The value for fleak is dependent on the type of pig and the pig velocity. In lieu of better

data, use fleak = 0.02.

The amount of time during which the liquid slug enters the slug catcher is:



tQ

A Vslugslug

p m avg=

where tslug = Time during which the slug enters the slug catcher, sec

The liquid rate while the slug enters the slug catcher is:

qslug = Vm avg Ap

where qslug = Liquid rate into slug catcher when slug is exiting, ft3/sec

The simplified analysis shown above doesn’t account for many effects that occur during

pigging, such as pig acceleration and deceleration, aeration of the liquid slug, etc., but it

gives a ball-park estimate of the flows and volumes needed for the slug catcher design.

5.2.4 Startup and Blowdown Slugs

When a pipeline is shut down, the liquid will drain to the low points in the line. When the

line is restarted, this liquid may exit the line in the form of slugs. To determine whether

slugging will occur, and to estimate the magnitude of the slugs, a transient simulator must

be used.

If the pipeline is depressured at shutdown, slugs may also form, due to high gas velocities

during the blowdown period. Transient simulation is needed to model these slugs.

5.2.5 Rate Change Slugs

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

holdup can either exit the line as a steady flow with increased liquid production, or it can

come out in the form of a slug, depending on the flowrate change.

Please note that the rate change slugs can occur in gas/condensate lines when the rates are

increased. The line may be in a steady flow pattern, such as stratified flow, at both the

initial and final flowrates but will slug during the transition period until the line re-

equilibrates at the higher rate.



As with startup slugs, it is impossible to predict whether slugs will occur when rates are

changed using steady state or hand methods. The line must be dynamically simulated

using a transient flow program.

5.2.6 Downstream Equipment Design for Slug Flow

The design of slug-catchers, separators, and control systems downstream of pipelines must

comprehend the presence and severity of slug flow. Estimates of parameters such as slug

volumes, liquid and gas rates exiting the pipeline as a function of time, etc. must be

factored into the design of this equipment. These variables should be calculated for the

design operation and for a series of off-design cases: turndown rates, pigging, shutdown

and startup, rate changes, etc.

As mentioned, transient modeling gives the best estimates of slug flow behavior. Some of

the transient simulators also allow the user to simulate separators and control systems as

part of the run, allowing the user to fine tune the design.

For approximate sizing of the equipment, the methods shown in Sections 5.2.1 to 5.2.3

can be used. Off-design cases should be estimated as well as the design point. In general,

the size of slug catching equipment for gas-condensate pipelines will be governed by

pigging considerations. For oil dominated systems, the size of the slug catcher is usually

governed by the maximum slug length due to either hydrodynamic or terrain slugs.

Because slug catching equipment can be a substantial cost item, it is possible to minimize

the cost of the equipment by considering alternative operating scenarios. Some of these

include:

a) Does the line need to be pigged routinely or is pigging only needed for maintenance?

If pigging is a maintenance item done once per year, it is possible to run the pig at a

rate low enough to keep a small slug catcher from overflowing. It is also possible to

send the liquid from the slug catcher to low pressure separation at high rates for

short periods, if the occurrence of pigging is infrequent. This, too, minimizes the

high pressure slug catcher sizing.



b) Pigging of the line after steady state conditions are reached may result in a very large

slug catcher. If a high pigging frequency is chosen, so that the line never achieves

steady state operation, the slug catcher size can be minimized.

c) It is possible to design the line for frequent pigging at low rates and steady state

operation at higher rates, thereby minimizing slug catcher size while eliminating the

need for frequent pigging when the rate is high.

d) If pigging isn’t necessary for a gas-condensate line, the sizing of the slug catcher

usually becomes dependent on rate changes. Limiting the turndown on the line and

limiting the amount that the rate is changed at one time can be beneficial in

minimizing slug catcher sizes.

e) The use of parallel pipelines should be considered as a way of limiting the turndown

for a line. When rates are low, all the production can be fed to one of the lines,

thereby keeping velocities high, minimizing liquid holdup.

f) If terrain slugging is shown to be a problem, there are several remedial steps which

have been employed to decrease the severity of the slugging or eliminate it

completely. One approach would be to reroute the pipeline to reduce or eliminate

dips in the pipeline profile. Two other methods are: choking of the flow at the tope

of the riser; and gas lifting the riser.

g) Because the slug length for hydrodynamic slug flow is a function of the diameter of

the line, the use of parallel pipelines instead of one pipeline can decrease the slug

length and volume going to the slug catcher/separator.

h) The use of multiphase pumps or subsea separators at well clusters can decrease slug

catcher sizes significantly.

i) Pigs have been developed with a variety of proprietary internals that limit the

velocity at which the pig moves through the pipeline. It may be possible to limit the

slug catcher size by restricting the pig velocity, which, in turn, limits the rate of

liquid exiting the line during the pigging operation.





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

1. Hydrodynamic Slugging

2. Terrain Slugging

3. Pigging

4. Startup

5. Rate Changes

• The Brill method is very inaccurate. It gives poor predictions for almost every data set

that was not included in the original correlation.

• Rule of thumb (based on limited amount of data available): the maximum slug length is

approximately 6 times the mean slug.

• If a correlation based model is used for slug length prediction, the Hill & Wood model

is probably the best available method.

• Pipelines, which are in the terrain slugging regime, will usually be identified by the

steady state simulator as being in stratified flow. If no further analysis is done, the line

may be designed in the belief that it is in a steady flow regime. On startup, the line

may experience slugs that may be thousands of feet long..

• In general, the size of slugcatching equipment for gas-condensate pipelines will be

governed by pigging considerations. For oil dominated systems, the size of the

slugcatcher is usually governed by maximum slug length due to either hydrodynamic

or terrain slugs.

• There are alternative operating scenarios which can be considered to minimize slug

catcher size.



Figure I: 5-1 Taital-Dukler Liquid Holdup Predictions for Simplified Flow in a Horizontal Pipeline



Figure I: 5-2 Stages in Terrain Slugging of Pipeline Riser Systems



Figure I: 5-3 Pipeline Pigging



SECTION 6.0 - EXAMPLE PROBLEMS

6.1 EXAMPLE1: Low Gas/Oil Line Between Platforms

This example is based on the Mejo to Tapa P/P line in Nigeria. The line runs 8 miles from

a Wellhead Platform to a Producing Platform. The gas-oil ratio for the production is about

345 SCF/bbl oil. Water production is initially zero, but builds rapidly with time. The

production profile is:

Year No. Oil Production,

bbl/day

Water Production,

bbl/day

Gas Production,

MMSCFD

1 15,478 0 5.32

4 7,954 8,865 2.74

8 2,057 7,584 0.70

12 765 4,438 0.26

The API gravity of the stock tank oil is 27.0. The specific gravity of the separator gas is

about 0.71. The dry wellstream composition, as measured by a separator test, is:

Component Mole %

CO2 0.90

N2 0.05

C1 31.94

C2 4.18

C3 1.55

iC4 0.45

nC4 0.49

iC5 0.48

nC5 0.37

C6’s 1.66

C7’s 2.04

C8’s 4.23

C9’s 3.55

C10’s 3.45



C11’s 3.07

C12's 2.57

Cl3’s 2.64

Cl4’s 2.22

C15's 2.06

C16's 1.82

Cl7's 1.48

C18's 1.41

Cl9's 1.09

C20+ 26.30

Total 100.00

The specific gravity of the C20+ fraction is 0.9499 at 60 degrees F. and the molecular

weight of the C20+ is 309.

The hydrocarbon wellstream composition is assumed to remain constant throughout the

life of the field.

The pipeline system consists of a downward riser at the wellhead platform, the 8 mile long

pipeline, and an upward riser at the production platform. The down riser is vertical and it

is 65 feet long. The up riser is vertical and 100 feet long. The pipeline elevation profile is

not known. The only information available on the pipeline elevation is that the water depth

at the wellhead platform is 15 feet, and the water depth at the production platform is 20

feet.

The pressure at the production platform is 80 psig. The maximum pressure at the wellhead

platform is 300 psig.

The line will be unburied, and it will not be coated with concrete. The wellstream

temperature at the wellhead platform is 130 degrees F, and the ambient air and water

temperatures are 80 degrees F.

What line size is required? Is slugging likely to be a problem? If slugging occurs, what

length and frequency are the slugs?



Solution

The operation of the line will be initially simulated by Pipephase. If it appears that slugging

is likely, additional simulations will be done using the OLGA transient model.

6.1.1 Line Size

In order to run Pipephase, an estimate of the line diameter is needed. The minimum line

size is given by the erosional velocity formula,

VC

nsmax =

ρ

where C = 100

To calculate ρns and Vmax rigorously, the phase equilibria and physical properties would

need to be found by flash calculations. Since the minimum diameter is only a starting place

for the pipe sizing calculations, it isn’t necessary to calculate these values with great

precision. When the Pipephase results are analyzed, the values of the erosional velocity

constant will be computed and compared against the maximum value of 100.

The maximum value of C should occur at the outlet of the up riser, since the mixture

velocity is the highest there. The pressure at the top of the riser should be almost equal to

the separator pressure, 80 psig. The temperature of the fluid at the top of the riser should

be close to sea temperature, 80 degrees F. At these conditions, almost all the gas should

be flashed out of solution, so that the amount of gas present should be equal to the gas/oil

ratio times the oil production rate. The compressibility factor of the gas at these conditions

would be about 0.98. Therefore the amount of gas and liquid present under the year 1

conditions should be approximately:

Gas ft3/sec:

cfsg = (345 SCF/bbl) (15478 bbL/D) (14.7 psia/94.7 psia) [(460+80) F/(460+60)

F] (1 D/86,400 sec) .98 = 9.76 ft3/sec

Liquid ft3/sec:



cfsl = (15478 bbl/D) (1 D/86,400 sec) (42 gal/bbl) (1 ft3/7.48 gal) = 1.01 ft3/sec

The densities of each phase are approximately:

Gas:

ρg = (.71) (28.97 lb/mole) (95 psia) / [(.98) (10.73 psia-ft3/mole-degF) (460+80

degF)] = 0.344 lb/ft3

Liquid:

ρl = (62.4 lb/ft3) [141.5/(131.5+27 deg API)] = 55.7 lb/ft3

The no-slip average density of the mixture, ρns, is given by:

( ) ( )

( )ρρ ρ

nsg g l l

g l

cfs cfs

cfs cfs=

+

+

= [(0.344 * 9.76) + (55.7 *1.01)] / (9.76 + 1.01) = 5.54 lb/ft3

The erosion velocity criteria gives a maximum velocity of:

Vmax = 100 / (5.54)0.5 = 42.5 ft/sec

The minimum diameter of the line, therefore, is:

Dmin = [( 4 * (9.76 + 1.01) ft3/sec) / (42.5 ft/sec) * 3.1416)]0.5 (12 in/ft)

= 6.82 inches

Rounding this result to the nearest standard line size indicates that the pipeline must be at

least an 8 inch line to meet the erosional velocity criterion.

Pipephase runs should be made for the production rates at the various years starting with a

line size of 8 inches. If the pressure drop becomes excessive, increase the line size, until



the inlet pressure for the limiting case remains below about 300 psig for all production

rates.

The approach discussed in the previous paragraph requires a trial and error solution for

the diameter. To cut down the number of trial diameters, it is possible to approximate the

frictional pressure drop by use of a "ball-park" method. The friction loss can be estimated

by assuming a homogenous mixture of the two phases. The assumption of homogenous

flow gives a friction loss that is low, typically being anywhere from 5% to 100% too low.

Adding 25% to the homogenous friction loss will bring the answer into a reasonable

range.

For this example, the elevational and accelerational pressure drops will be much less than

the frictional loss, and will be ignored in the rough line sizing calculations.

Using typical values for the viscosities of each phase (Gas Visc. = 0.015 cp, Liq. Visc. =

10 cp) gives a no-slip viscosity of:

[(0.015 x 9.76) + (10 x 1.01)] /(9.76 + 1.01) = 0.95 cp

The homogenous frictional pressure gradient is:

( )dp

dx

f V

g d

ns ns

c

m=4

2 12

2 ρ

Re

1488 Vd

12ns

ns m

ns=

ρ

µ

( )

Vm =+9.76 1.01 ft / sec x 4 x 144

3.1416 d

3

2

where d is in inches

Vm = 1975 / d2 (ft/sec)

ρns = 5.54 lb/ft3



µns = 0.95 cp

Substituting these values into the expression for Rens gives:

Rens = 1,426,000/d

For a smooth pipe, the Fanning friction factor can be approximated by:

fns = 0.046 / Rens0.2

Substituting,

fns = 0.046/(1426000/d)0.2 = .00270 d0.2

Combining this expression with the expressions above gives the following expression for

the homogenous frictional pressure gradient:

( ) ( )dp

dx=

4 0.00270d1975

d5.54

(64.4)(12d)

0.22

2

= 302.0/d4.8 psi/ft

The allowable pressure drop is 300-80=220 psi in the 8 mile long pipeline. Equating this

pressure drop to the homogenous pressure gradient, and adding 25% to the homogenous

pressure gradient to account for the two-phase behavior gives:

220 psi/(8 * 5280 ft) = 1.25 * 302.0/d4 8 psi/ft

d = 10.29 inches

These calculations indicate that an 8 inch line is probably too small, and that a 10 inch line

is probably the minimum line size that can be used to stay within the desired pressure

drop.



For the sake of this example, we will make Pipephase runs for 8, 10, and 12 inch pipelines

and risers.

Since we have both black oil properties and compositions available for this example, we

will run both the black oil model and the compositional model in Pipephase for the 10 inch

line as a comparison between the methods. In general, compositional modeling predicts

the phase equilibrium more accurately than black oil modeling, but the black oil models

may predict the physical properties of the liquid better than compositional modeling. For

low GOR systems, black oil modeling may give the best answers, while compositional

modeling is clearly superior for high GOR systems.

The black oil model only needs the gravities of each phase and the total gas-oil ratio of the

system.

For the compositional model, the program needs two of the following for each petroleum

cut: molecular weight; specific gravity; or normal boiling point. In the example

information, we are given the specific gravity and molecular weight for the C20 + cut, but

we have no information on the C6 to Cl9 cuts. Without this information, the only possible

characterization is to use the normal paraffin data for the cut. Use nCl0 as the basis for the

properties of the Cl0 cut, for instance. In reality, each of these cuts consists of hundreds of

different components boiling in the same boiling range as the normal paraffin, and the

phase behavior of this mixture of aromatics, naphthenes, and alkylated paraffins is not the

same as the behavior of the normal paraffin. The proper way to characterize the cuts is to

have laboratory analyses of the specific gravity and molecular weight of each of the cuts in

addition to their boiling ranges, but this information is rarely available.

Information is needed on the pipeline and risers. The diameters to be used in this exercise

are:

Nominal Diameter, inches Outside Diameter, inches Inside Diameter, inches

8 8.625 7.625

10 10.750 9.500

12 12.750 11.375



Since there is insufficient data to do rigorous heat transfer calculations, use an overall heat

transfer coefficient, U, of 5 BTU/hr/ft2/degF for the pipeline and a U of 20 for each of the

risers.

The pipeline elevation profile has not been given. The simplest way to simulate the

pipeline profile is to assume that it has a constant slope, with a 5 foot drop in elevation

over its 8 mile length. This assumption will give the most optimistic results, because it will

minimize the elevational pressure drop for hills. To better mirror reality, some terrain

effects could be added by adding hills and valleys to the profile. For high velocities, there

will be negligible difference in the holdup and pressure drop between the constant slope

profile and the profile with the hills added. For low velocities, however, the difference can

be significant. For the sake of simplicity in this example, we will use the straight-line

profile for the Pipephase runs, with the understanding that the pressure drop and liquid

holdup for low rates may be low.

The Pipephase runs will be made with a constant outlet ("Sink") pressure of 80 psig. The

inlet ("Source") pressure to the line will vary with the production rate.

Since this line is a low gas-oil ratio line, the Beggs and Brill correlation will be used to

calculate the pressure drop and liquid holdup for the line. If the OLGAS method had been

available, it would have been used for the calculations. Because of its mechanistic model,

OLGAS would show more of an impact of hills than Beggs and Brill, and would give

more accurate answers. The Beggs and Brill flow regime calculations are poor, and they

will be ignored. The Taitel-Dukler-Barnea flow regime predictions will be used instead.

A comparison of the predictions of the black oil model with three different liquid viscosity

methods and the compositional model was made for the 10-inch line at year 1 rates. The

predicted inlet pressures and liquid viscosity for the models were:



Model Inlet Pressure, psig Liquid Viscosity at

Pipeline Outlet, cp

Compositional 247 2

Black Oil, Glass Is. 301 18

Black Oil, Chew Vis. 320 29

Black Oil, Vazquez Vis. 335 62

There is a wide discrepancy between the model predictions, which can be explained by the

large differences in liquid viscosity. The viscosity of the reservoir fluid was measured

during the PVT analysis of the recombined wellstream. This viscosity indicated that the

Glaso correlation modeled the liquid viscosity best.

The poor viscosity prediction of the compositional model may be due in part to the

characterization of the heavy ends of the wellstream. Errors of this magnitude in physical

properties are disturbing, and merit further investigation.

It was decided to use the Black Oil model with the Glaso liquid viscosity correlation for all

the Pipephase runs for this example.

Pipephase runs were made for the three line sizes (8, 10, and 12 inch) for the 4 production

rates shown in the production forecast. The inlet pressures for the various cases were:

Year No. 8 inch 10 inch 12 inch

1 530 301 199

4 433 242 164

8 194 133 119

12 126 115 110

The 8-inch line is too small to handle the rates in the early years of production. The 10-

inch line is at the maximum inlet pressure for the Year 1 conditions. The 12-inch line is

comfortably within the maximum inlet pressure of the line.

The mixture velocities at the exit of the P/P riser for the various cases are:




1 33.21 21.06 14.58

4 18.60 12.03 8.37

8 5.88 3.75 2.60

12 2.51 1.59 1.12

Pipephase does not print the superficial gas and liquid velocities. The values for these

parameters can be calculated from the program output by the following relationships:

Vsg = Ug (I-HI)

Vsl = Ul HI

where the Ug is the actual gas velocity, Ul is the actual liquid velocity, and HI is the slip

liquid holdup. These three parameters are tabulated in the "Link" device detail report.

The superficial gas velocity at the exit of the P/P riser are:


1 30.01 19.00 13.14

4 15.13 9.80 6.81

8 3.90 2.47 1.71

12 1.44 0.90 0.64

The superficial liquid velocity at the exit of the P/P riser are:


1 3.20 2.06 1.44

4 3.47 2.23 1.56

8 1.98 1.28 0.89

12 1.07 0.69 0.48

The value of the constant C in the erosional velocity formula can be calculated by either of



the following equivalent relationships:

C = [Vm (ρl Vsl + pg Vsg)]0.5

= ρns0.5 Vm

The values of the liquid, gas, and no-slip densities can be read from the, "Property"

summary table in the Pipephase output. Substituting these values into the above equation

gives the following values for C at the exit of the P/P riser:


1 78.9 50.4 35.0

4 62.4 40.3 28.1

8 26.8 17.2 11.9

12 12.9 8.2 5.8

The values for C are within the API guidelines for all cases.

The Taitel-Dukler-Barnea flow regimes at the end of the pipeline for the various cases are:


1 Slug Slug Slug

4 Slug Slug Slug

8 Slug Slug Strat

12 Slug Strat Strat

The predicted liquid holdups (in barrels), using the Beggs and Brill correlation, are:


1 1064 1336 1664

4 1268 1683 2197

8 1367 2018 3607

12 1563 3093 5009



The conclusions regarding the line sizing are that the 10-inch line is the appropriate line

size. The pressure drop for the 10-inch during year 1 is right at the upper limit for inlet

pressure, so the production rate could be less than design if the Pipephase predictions are

optimistic. The 8-inch line is definitely undersized, and the 12-inch appears oversized. The

velocities in the later years of operation are a definite concern for all three-line sizes. Due

to these low velocities, severe slugging is a concern during the later years.

We need to quantify the occurrence and severity of slug flow. The Pipephase results

indicate that the pipeline will operate in the hydrodynamic slug flow regime for all rates

with the 8 inch line, years 1 to 8 for the 10 inch line, and years 1 to 4 for the 12 inch line.

Year 8 for the 12-inch and year 12 for both the 10 and 12 inch lines are shown to be in

stratified flow in the pipeline. This result is deceptive, since stratified flow in a

downwardly sloping pipeline connected to a vertical riser is the classical severe slugging

scenario. The key to estimating whether severe slugging will occur is to determine

whether the Pots πss factor is less than 1. Pipephase tabulates this value as part of its

"SLUG" report. The values for the various cases are:


1 0.098 0.084 0.078

4 0.052 0.044 0.041

8 0.024 0.023 0.032

12 0.020 0.042 0.121

All the values of πss are much less than 1, indicating that severe slugging is possible for this

pipeline. By Pots’ analysis, only those points, which are in stratified flow in the pipeline,

will be in the severe slugging regime.

Therefore, the predicted flow regimes for the system using the Pipephase results are:


1 Hydrodynamic Slug Hydrodynamic Slug Hydrodynamic Slug

4 Hydrodynamic Slug Hydrodynamic Slug Hydrodynamic Slug

8 Hydrodynamic Slug Hydrodynamic Slug Severe Slug

12 Hydrodynamic Slug Severe Slug Severe Slug



6.1.2 Slug Length Prediction

We need to estimate the slug lengths for the various cases. As mentioned in the text, the

use of correlations for slug length prediction is not very accurate. Unless a transient model

is available, however, these predictions are the best available method.

The output for the Pipephase "SLUG" option prints tables summarizing the predictions of

the Brill, et al. correlation. It also indicates an approximate slug length for severe slugging.

The Brill, et al. predictions will be used as the basis for predicting slug length in

hydrodynamic slug flow, and the Pots πss method will be used as the basis for predicting

slug length in severe slugging. These estimates will be checked against the predictions of

the Hill & Wood model for hydrodynamic slugging and the predictions of the OLGA

transient model.

In severe slugging, the operation is cyclic:

• The low spot fills with liquid, blocking the flow;

• Liquid builds in the pipeline and riser until the pressure behind the liquid slug builds to

a high enough level to push the slug out;

• The liquid exits the system as a slug, increasing in velocity as the gas behind the slug

expands;

• When the gas has exited the riser, liquid will fall back down the riser, and drain in from

the pipeline, starting the cycle again.

The slug length is estimated by:

Ls = Lr / πss

where Lr is the riser height in feet.

The slug length should not vary much for a constant rate, since the operation in cyclic.

The maximum slug length therefore is approximately equal to the average slug length.

The values for the slug length reported in Pipephase for the 3 severe slugging conditions

are:



a) For the 10" Line, Year 12 - Ls = 2392 ft.

b) For the 12" Line, Year 8 - Ls = 3140 ft

c) For the 12" Line, Year 12 - Ls = 829 ft

The large decrease in the slug length for the 12" line, year 12 is due to the large increase in

holdup in the line predicted by Beggs and Brill. The accuracy of this value is doubtful.

For the hydrodynamic slugging cases, the "SLUG" report in Pipephase contains

predictions of the average slug length and the slug length distribution using the Brill, et al.

model. As detailed in the text, the Brill method is based on a very small amount of data,

and it doesn’t account for most of the variables, which play a part in slug flow.

To use the Brill slug length model to estimate the maximum size of a slug for equipment

selection, the designer needs to estimate the slug length corresponding to a design

occurrence frequency. A reasonable philosophy would be to design the separator/slug

catcher for the once per year slug. Estimating the average slug frequency from the

Pipephase "SLUG" report, the number of slugs per year can be calculated. The slug

distribution table above can be used to estimate the slug length corresponding to one

occurrence in this number of slugs per year.

For example, the year 1 rate through the 10" line gives a slug frequency of 98.4 seconds

by the Brill model. This frequency corresponds to

(365 D/yr * 86,400 sec/D) / 98.4 sec = 320,500 slugs/yr

If one slug per year can be greater than the design slug, the design slug must be at the

(1 - 1/320500) * 100 = 99.9997 percentile

The average slug, according to Brill, is 293 ft long. The maximum allowable slug would

then be about 10.5 times as long as the average or 3080 ft long.

The design percent slug was calculated for all the other hydrodynamic slug cases. The

design slug was close to the 99.9999% slug length for all cases, using the one slug per



year criteria. Tabulating the 99.9999% data for the hydrodynamic slug cases gives design

slug lengths (in feet) for the various cases of:


1 2152 3153 4250

4 2079 3050 4113

8 1943 2847 3140*

12 1848 2392* 829*

* Terrain slugs

To calculate the liquid surge requirements for the separator based on these slug lengths,

additional information is needed. Since the slugs consist of aerated liquid, the liquid

holdup in the slug must be known in order to determine the volumetric rate entering the

separator. The surge volume required also depends in large measure on the capacity of the

liquid dump valves on the separator.

The Pipephase "SLUG" table includes a calculation of the liquid holdup in the slug, based

on the Brill, et al. method. The tabulated value for the "liquid from the slug" is equal to:

Volslug = Vm Hls Ap ts

where Hls is the liquid holdup in the slug, Ap is the pipe cross-sectional area, and ts

is the time that the liquid slug exits the pipeline.

The required surge volume for the separator is:

Volsurge = Volslug - ts Qdump

where Qdump is the volumetric capacity of the dump valve.

For this example, we will assume that the volumetric capacity of the dump valve is 1.33

times the design oil rate. This assumption gives a Qdump of 1.35 ft3/sec. The separator is

also assumed to be a 2-phase separator.



With these assumptions, the design surge volumes (ft3) for the hydrodynamic slugging

cases are:


1 145 561 1398

4 195 654 1506

8 101 294 —

12 0 — —

It is more difficult to estimate the slug volumes for the severe slugging cases using only

steady state analysis. For a crude approximation of the behavior, use the following:

a) Slug buildup phase

When the slug is building up, there is essentially no flow out of the pipeline. The time

required to build the slug is approximately:

tbu = Ls/vsl

The liquid rate out of the line is ~0 during this time.

b) Slow gas expansion phase

After the gas pressure has built to a high enough level to push the slug, the liquid rate

out of the line is approximately equal to the mixture velocity, Vm. The time required

for this phase is tsg.

c) Fast gas expansion

When the liquid slug is shorter than the riser pipe, the gas behind the slug expands

rapidly. The velocities in the riser may reach 60-80 ft/sec. This time period is so much

smaller than the others that it can be assumed to be virtually instantaneous. The liquid

rate is so high that essentially all the liquid produced during this time goes into the

separator surge.

With this simplified model, the flowrates out of the pipe can be estimated. The slug cycle

time (tc) is equal to tbu + tsg. The liquid produced during one cycle is equal to tc * V sl * Ap.

The liquid leaving the pipeline during the slug buildup phase is essentially 0, and the



volume of liquid exiting during the fast gas expansion is equal to the riser volume. The

remainder of the liquid must exit during the slow gas expansion phase. Therefore,

(tbu + tsg) Vsl Ap = (tsg Vm Ap) + (Lr Ap)

tsg = [(tbu Vsl) - Lr] / Vsg

tsg = (Ls - Lr) / Vsg

To illustrate the use of these equations, we will check the outlet rates for the 10" line at

year 12. The values for the key parameters are Ls = 2392 ft. Vsg = 0.90 ft/sec, Vsl = 0.69

ft/sec, Ap=0.492 ft2, and Lr = 100 ft. The value for tbu = 2392/0.69=3467 sec. The value

for tsg = (2392-100)/0.90 = 2547 sec.

The model indicates that the liquid flowrate out of the line is:

0 ft3/sec for 0<t<3467

(0.90+0.69) * 0.492 = 0.782 ft3/sec for 3467<t<6014

Infinity for t=6014, Volume=0.492 x 100-49.2 ft3

Since the 0.782 ft3/sec is less than the Qdump of 1.35 ft3/sec, the liquid will not accumulate

in the separator during the slow gas expansion phase. The required surge volume therefore

is 49.2 ft3.

Completing these calculations for the other severe slugging cases gives the following

tabulation for the required surge volumes (ft3) for all the cases:


1 145 561 1398

4 195 654 1506

8 101 294 935

12 0 49 71

Using the Brill analysis, the required surge volume for the downstream separator would be

654 ft3 for the 10" line.



6.1.3 Slug Frequency and Length by Hill & Wood Method

The Hill & Wood correlation for slug frequency will be used to check the slug length

predictions and surge volume predictions. To use the Hill & Wood method, a calculation

of the liquid holdup at the beginning of the pipeline must be made using the Taitel-Dukler

stratified flow model. The method assumes a horizontal pipeline. The value for the

Lockhart & Martinelli X factor can be calculated from the following equation:

XV

Vsl

sg g g

l l=

0 9 0 4 01. . .ρρ

µµ

Once X is known, the value for the liquid holdup for stratified flow can be read from

Figure I:5-1. The actual gas and liquid velocities can be calculated from the superficial

velocities and the stratified flow liquid holdup. Once these parameters are known, the slug

frequency can be calculated.

To illustrate the procedure, the slug frequency for the 10" line for the year 1 conditions

will be calculated. The values for the required parameters at the beginning of the pipeline

are:

Vsg = 5.40 ft/sec ρg = 1.10 lb/ft3 µg = 0.012 cp

Vsl = 2.10 ft/sec ρl = 53.97 lb/ft3 µl = 7.1 cp

X is then:

X= (2.1/5.4)°~9 * (53.97/1.1)° 4 * (7.1/.012)°~1 = 3.84

The value for the stratified flow liquid holdup, using the Taitel-Dukler model for

horizontal flow, is:

Hle=0.68

read from Figure I:5-1.

The actual gas velocity is:



Ug = Vsg/(1 HIe) = 5.40/(1-.68) = 16.88 ft/sec

The actual liquid velocity is:

Ul = Vsl/Hle = 2.1/.68 = 3.09 ft/sec

Putting these values into the Hill & Wood slug frequency equation gives:

Fs = 2.74 Hle (Ug - Ul) / (D (l -HIe))

= (2.74) (0.68) (16.88-3.09)/[(9.5/12)(1-.68)]

= 101 slugs/hour

The length of the average slug plus bubble unit would be:

Ls + Lb = (Vm) (3600)/Fs = (7.5 ft/sec) (3600 sec/hour) (1 hour)/101 slugs

= 267 ft

To estimate the slug length, the slug fraction must be calculated. There are several ways to

do this. This example will show the easiest way to estimate the slug fraction.

First, the liquid holdup in the slug will be estimated. The Gregory, Nicholson, and Aziz

method is the easiest method to use, and it is as accurate as most of the methods. The

liquid holdup in the slug is given by:

Hls = l [l +(V,,’/28.4)1.39]

= 1/[1+(7.5/28.4)1.39]

= 0.864

The liquid holdup in the bubble can be assumed to be ~0.20 for much of the slug flow

range. From a material balance, the slug fraction can be calculated from these values and

the overall liquid holdup prediction:

SF = (Hl - Hlb) /(Hls - Hlb)



Hl = 0.42 by the Beggs and Brill method at the beginning of the pipeline.

Therefore,

SF = (0.42 - 0.20)/(0.864 - 0.20)

= 0.33

Since the slug fraction is equal to the slug length divided by the slug length plus bubble

length, the slug length is equal to:

Ls = SF (Ls + Lb)

= 0.33 x 267

= 88ft.

This value for the average slug length is much less than the 293 ft average slug length

calculated by the Brill et al. method.

The average slug lengths for the 10" line are:

Year Number Average Slug Length, ft.

1 88

4 52

8 49

12 Terrain Slugging

The Hill and Wood method does not predict a slug length distribution. Experimental

evidence from the Prudhoe Bay field tests and BHRG laboratory studies indicate that the

maximum slug length is approximately 6 times the average slug length. Using this value,

the design slug lengths are:




1 528

4 312

8 294

12 2392

The design surge volume would be:

Volsurge = (Lsdes Hls Ap) - [(Lsdes Qdump) / Vm]

For the year 1 conditions, this volume is:

Volsurge = [(528 ft) x (.864) x (0.492 ft2 )] - [(528 ft) x (1.35 ft3/sec) / (7.50 ft/sec)]

= 129ft3

The values of the design surge volume for the 10" line during all the years are:


1 129

4 76

8 2

12 49

The design surge volume using the Hill & Wood method of 129 ft3 is considerably less

than the 654 ft3 calculated by use of the Brill et al. method. For comparison, the liquid

volume needed for separation is about 200 ft3, based on a 3 minute liquid residence time in

the vessel.

The transient OLGA program was run for the four design cases for the 10" line. The inlet

pressures using OLGA for the cases were:



Year Number OLGA Beggs and Brill

1 228 psig 301 psig

4 187 psig 242 psig

8 115 psig 133 psig

12 101-113 psig 115 psig

The pressure drops calculated by OLGA are much less than the Beggs and Brill pressure

drops. This is due in part to the physical properties used in the OLGA runs, since OLGA

uses a compositional model similar to Pipephase’s compositional model. Most of the

difference, however, is due to the different multiphase flow models. OLGA tends to

predict pressure drops that are too low in slug flow, and Beggs and Brill tends to

overpredict pressure drops. The actual inlet pressure is between the two answers. For year

1, for instance, the expected inlet pressure would probably be about 270-280 psig.

The flow regimes predicted by OLGA were:

Year Number Flow Regime

1 Hydrodynamic Slug Flow



12 Terrain Slug Flow

OLGA confirms the Pipephase flow regime predictions, which indicated the line should be

in slug flow for years 1 through 8, and then experience terrain slugging in year 12.

The liquid holdup predictions of OLGA were:

Year Number OLGA Beggs and Brill

1 1365 bbl 1336 bbl

4 1831 bbl 1683 bbl

8 2565 bbl 2018 bbl

12 2915 bbl 3093 bbl



The liquid holdup predictions of OLGA and Beggs and Brill are quite close. Beggs and

Brill’s holdup predictions in slug flow are pretty good, as seen above, but they can be very

poor for stratified flow, as shown in Example 2.

For the year 12 conditions, OLGA shows a cyclic terrain slugging pattern, as seen in

Figures 1:6-1 to I:6-4. The results shown for about the first 15,000 seconds reflect the

initial conditions selected for the problem, and can be ignored. After 15,000 seconds, a

cyclic pattern of terrain slugs is seen. Every 2400 seconds, a slug exits the pipeline with a

volume of about 11 cubic meters (-70 bbls). This slug volume corresponds to a slug

length of ~790 feet, which is considerably smaller than the 2392 feet predicted by the Pots

method. The OLGA results are much more believable than the Pots predictions. The

pressure oscillates between 101 psig and 113 psig. The liquid mass flow rate out of the

pipeline peaks at a rate of about 40 kg/sec, which corresponds to an instantaneous

flowrate of 21,700 BPSD, compared to the average flowrate of 5203 BPSD. The gas flow

rate peaks at about 0.28 kg/sec, which is equivalent to 1.0 mmSCFD, compared to its

average rate of 0.26 mmSCFD.

The slug-tracking model in OLGA V3.1 was run for the three years that were in

hydrodynamic slug flow. The slug-tracking model tracks each individual slug, and

determines whether it grows, dissipates, or merges with another slug as it flows down the

line.

The slug-tracking runs for these cases, however, were unable to converge on a value for

the average slug length. For year 1, for example, runs were made with three different

numbers of pipeline segments. If the model converges on a solution, the slug length would

be independent of the number of segments used. For this case, the average slug length vs.

number of pipeline segments was:

Number of Segments Average Slug Length, feet

79 197

154 82

294 51



The maximum number of pipeline segments in the commercial version of OLGA is 297.

Since the slug length was still decreasing using the maximum number of segments, the

program had not reached a value for the slug length that was independent of pipeline

discretization. As a result, the only conclusion that can be drawn from the slug-tracking

analysis is that the average slug length is probably less than 50 feet for the year 1

conditions.

The slug-tracking predictions are in line with the results of the Hill and Wood model.

Both predictions are much less than the Brill, et al. method.

In summary, the OLGA runs verified that the 10" pipeline should be capable of flowing all

rates within the allowable pressure drop. They confirmed that the pipeline would be in

hydrodynamic slug flow for the first 8 years, and then operate in the terrain slug flow

regime by year 12. The slug lengths predicted by Pipephase for the terrain slugging case

were found to be too large, and the slug lengths for the hydrodynamic slugging cases were

predicted to be of the same magnitude as the Hill and Wood predictions.

If this example was an actual design, checks of other transient phenomena such as startup,

shutdown, rate changes, etc. should be performed to determine whether they would cause

slugs which would govern the size of the slug catcher.

Based on this analysis, the slug catcher should be designed for a surge volume of about

150 ft3, based on the Hill and Wood results.

6.2 EXAMPLE 2 - Gas/Condensate Line

A pipeline is needed to flow the production from a subsea well development to shore. The

offshore field produces a mixture of gas, water, and condensate. Glycol is injected into

the line for hydrate suppression. The water quantity is about 4 bbl/mmSCF, and the

condensate production rate is also about 4 bbl/mmSCF. The peak production rate is 500

mmSCFD of gas and condensate. The minimum turndown is 25% of design rate.

The arrival pressure onshore is 900 psia. The inlet pressure to the line cannot exceed

1200 psia under normal operation.

The wellstream composition, including water and injected glycol is given in Table 1.



The pipeline elevation profile is shown in Table II.

The wellhead pressure will decrease with time as the reservoir becomes depleted. Initially,

the wellhead pressure will be over 4000 psig. For design, the wellhead temperature and

pressure are estimated to be 222 degrees F and 1215 psia, simulating conditions late in the

life of the field.

The pipeline is unburied, and is coated with 3 inches of concrete and 0.125 inches of

epoxy coating.

6.2.1 TABLE I, Wellstream Composition

Component Mole %

N2 3.68468

CO2 11.11650

C1 75.79904

C2 3.82831

C3 1.15655

iC4 0.16298

nC4 0.23880

iC5 0.08426

nC5 0.06506

C6’s 0.05431

C7’s 0.07390

C8’s 0.10359

C9’s 0.07833

C10’s 0.03915

C1l’s 0.02022

Cl2’s 0.01223



Cl3’s 0.00985

C14's 0.00794

C15's 0.00556

C16’s 0.00460

Cl7’s 0.00248

C18's 0.00157

C19's 0.00126

C20's 0.00074

C21's 0.00073

C22's 0.00054

C23's 0.00037

C24's 0.00031

C25's 0.00024

C26's 0.00019

C27’s 0.00016

C28+ 0.00021

Ethylene Glycol 0.12530

Water 3.32004

Totals 100.00000

6.2.2 TABLE 11, Pipeline Elevation Profile

Cumulative Distance, km Water Depth, m Segment Length, ft Elevation Change, ft

0.0 216

0.35 218 1148 -6.6

1.45 218 3608 0.0

1.75 219 984 -3.3

2.25 219 1640 0.0



2.55 218 984 3.3

3.15 218 1968 0.0

3.45 219 984 -3.3

3.95 219 1640 0.0

4.30 218 1148 3.3

6.35 218 6724 0.0

6.90 217 1804 3.3

7.30 217 1312 0.0

7.75 218 1476 -3.3

7.95 218 656 0.0

8.30 216 1148 6.6

8.90 214 1968 6.6

9.70 210 2624 13.1

10.00 208 984 6.6

10.40 208 1312 0.0

11.20 200 2624 26.2

12.40 190 3936 32.8

13.45 180 3444 32.8

14.45 170 3280 32.8

15.70 160 4100 32.8

16.95 150 4100 32.8

18.45 140 4920 32.8

20.20 130 5740 32.8

22.10 120 6232 32.8

24.10 110 6560 32.8

25.85 100 5740 32.8

30.25 90 14432 32.8

34.60 80 14268 32.8

42.10 70 24600 32.8

49.55 65 24436 16.4

56.15 65 21648 0.0

60.45 65 14104 0.0

61.75 60 4264 16.4

63.05 50 4264 32.8



74.25 40 36736 32.8

78.50 30 13940 32.8

79.45 25 3116 16.4

80.85 20 4592 16.4

81.95 15 3608 16.4

82.60 10 2132 16.4

83.40 5 2624 16.4

84.10 0 2296 16.4

6.2.3 Pipephase Simulation Comparisons

Pipephase simulations of the operation were made for 3 rates (125, 250, and 500

mmSCFD) for two line sizes (30" and 28"). Since this line is a gas-condensate line, the

Oliemans method was used to estimate the pressure drop in the system. The Taitel-

Dukler-Barnea method was used to predict the flow regime. As a check, the Beggs and

Brill method was run for the selected line size. A transient analysis of this line was also

performed using OLGA. It will therefore be possible to compare the predictions of

Oliemans and Beggs and Brill against OLGA.

The inside diameters used for the 30" line was 28.5", and the inside diameter for the 28"

line was 26.5".

The predicted inlet pressures using the Oliemans correlation were:

Production Rate, mmSCFD 30” Line 28” Line

125 932 psig 938 psig

250 970 psig 994 psig

500 1120 psig 1200 psig



The 30" line is comfortably within the wellhead pressure specification. The 28" line is right

on the specification at design rate.

A check of the pressure drop for the 28" line using the Beggs and Brill correlation gives:

Production Rate, mmSCFD Inlet Pressure, psig

125 947

250 1008

500 1226

Beggs and Brill indicates a higher pressure drop than Oliemans, making the 28" line more

questionable. Both line sizes will continue to be checked. The OLGA transient simulations

will be used as a third check of the pressure drop to determine whether the 28" line

appears acceptable.

The total liquid holdup for the line (barrels) using Oliemans and Beggs and Brill were:

Production Rate,

mmSCFD

Oliemans - 30” Oliemans - 28” Beggs-Brill - 28”

125 11,760 9,480 14,560

250 7,890 6,280 13,100

500 5,000 4,230 12,340

The superficial gas and liquid velocities are not printed in the Pipephase output. The

superficial velocities can be calculated from the actual gas or liquid velocities and the

liquid holdup. For gas/condensate lines, this is difficult because of the poor formatting of

the liquid holdup in the output table. Pipephase prints the value of the liquid holdup as a

two decimal place value (e.g. 0.02), which is not adequate to accurately determine the

superficial velocities. For holdup below 0.005, Pipephase prints 0.00. Fortunately, it is

possible to calculate accurate values for the superficial gas and liquid velocities and liquid



holdup from the information give in the output table. The superficial liquid velocity can be

calculated from the formula:

Vsl = [Ul(Ug - Vm)]/(Ug- Ul)

where Ug = Actual gas velocity, ft/sec

Ul = Actual liquid velocity, ft/sec

Vm = Mixture velocity, ft/sec

The superficial gas velocity can be calculated by:

Vsg = Vm - Vsl

The liquid holdup is given by:

HI = Vsl / Ul

The superficial gas velocities (ft/sec) for the various cases are:

Production Rate,

mmSCFD


125 4.66 5.39 5.38

250 9.32 10.77 10.77

500 18.54 21.40 21.38

The superficial liquid velocities (ft/sec) are:

Production Rate,

mmSCFD


125 0.0189 0.0224 0.0226



250 0.0381 0.0439 0.0471

500 0.0787 0.0884 0.0895

The values of C in the erosion velocity for the various cases are:

Production Rate,

mmSCFD


125 0.0189 0.0224 0.0226

250 0.0381 0.0439 0.0471

500 0.0787 0.0884 0.0895

All the C values are well below 100, so corrosion-erosion should not be a concern.

The dominant flow regimes for the various cases, using the Taitel-Dukler-Barnea method

are:

Production Rate,

mmSCFD


125 Slug Slug Slug

250 Strat./Slug Strat./Slug Strat./Slug

500 Strat./Annul. Strat./Annul. Strat./Annul.

The line is shown to be in slug flow at the 125 mmSCFD rate. Some of the segments of

the line are shown to be in slug flow at 250 mmSCFD, while others are shown to be in

stratified flow. For the 500 mmSCFD rate, all the segments are shown to be in either

stratified or annular flow.

Pipephase will not print a slug report for the runs in which the Oliemans correlation was

used. It did print slug lengths for the Beggs and Brill runs. The slug lengths predicted by



the Brill et al. method are 1460 to 1580 ft average slug length for the 28" line. These

numbers are very unrealistic for gas/condensate systems of this diameter.

The design criterion for the slug catcher surge volume is usually the liquid volume

associated with pigging the pipeline. This volume can be estimated from the Pipephase

output.

The pig travels at approximately the mixture velocity, Vm. The average values for Vm

(ft/sec) for the various cases is:

Production Rate,

mmSCFD


125 4.87 5.62 5.59

250 9.54 10.92 10.79

500 17.61 19.54 19.28

The pipeline is 84.1 km or 275,850 ft. long. The pig traverse times (seconds), therefore,

are:

Production Rate,

mmSCFD


125 56,640 49,080 49,350

250 28,920 25,260 25,570

500 15,660 14,120 14,310

An approximation of the volume of the slug ahead of the pig is that the slug volume is

equal to the liquid holdup minus the amount of liquid produced during the pig traverse

time. Numerically, this equates to:

Volslug = (Hl Ap Lp) -( tpig Vsl Ap)



The slug volumes (barrels) for the various cases are:

Production Rate,

mmSCFD


125 10,920 8,730 13,800

250 7,020 5,520 12,280

500 4,030 3,380 11,470

If transient analysis is not available, the table above would be used to design the surge

volume required for the slug catcher. For the 28" line, the slug catcher would be sized for

the 13,800 bbl surge volume, with the recognition that neither Oliemans nor Beggs-Brill

predicts liquid holdup well for gas/condensate pipelines.

We will use the OLGA transient program to check the design of the line to confirm the

line size, determine whether slugging is likely, and size the slug catcher.

OLGA simulations were run at the three design rates for both the 28 inch and 30 inch

lines. The predicted inlet pressures at the various rates were:

Rate, mmSCFD 30 Inch 28 Inch

125 1045 psig 1018 psig

250 980 psig 982 psig

500 1100 psig 1179 psig

A comparison of the OLGA, Oliemans, and Beggs and Brill predictions for the 28 inch

line at the various rates is:



Rate, mmSCFD OLGA Oliemans Beggs-Brill

125 1018 psig 938 psig 947 psig

250 982 psig 970 psig 1008 psig

500 1179 psig 1200 psig 1226 psig

The Oliemans predictions match the OLGA predictions quite closely, except at the lowest

rate. The difference between Oliemans and OLGA at the 125 mmSCFD rate is due to the

high liquid holdup predicted by OLGA, as shown below. The OLGA results confirm that

the pressure drop in the pipeline is less than the allowable drop for the 28" line, so the 28"

line can be used.

The predicted liquid holdups from OLGA are:


125 52,800 bbls 32,900 bbls

250 8,800 bbls 3,790 bbls

500 3,310 bbls 2,650 bbls

Comparing these results with Oliemans and Beggs and Brill for the 28 inch line shows the

wide disparity in results between the methods:

Rate, mmSCFD OLGA Oliemans Beggs-Brill

125 32,900 bbls 9,480 bbls 14,560 bbls

250 3,790 bbls 6,280 bbls 13,100 bbls

500 2,650 bbls 4,230 bbls 12,340 bbls

These results are plotted in Figure I:6-5. OLGA’s holdup predictions are much higher than

either Oliemans or Beggs and Brill at the lowest flowrate, and they are much lower at the

other rates. The OLGA holdup jumps when the gas velocity is too low to pull the liquid up

the upwardly inclined pipeline segments. OLGA indicates that, for these operating



conditions, this minimum superficial gas velocity is about 9 ft/sec, corresponding to a

throughput rate of about 220 mmSCFD. Below this rate, the liquid holdup builds up

rapidly. This behavior has been observed in the field and in experimental studies in inclined

flow. Neither Beggs and Brill nor Oliemans model this sensitivity to flowrate in inclined

pipes.

The predicted flow regimes from the OLGA simulations were:


125 Stratified and Slug Stratified and Slug

250 Stratified Stratified

500 Stratified Stratified

OLGA shows that some of the pipeline segments are in slug flow at the 125 mmSCFD

rate for both line sizes. For the higher rates, all of the segments are shown to be in

stratified flow. The comparison between the OLGA results and the Taitel-Dukler-Barnea

predictions is:

Rate, mmSCFD OLGA Taitel et al.

125 Strat./Slug Slug

250 Strat. Strat./Slug

500 Strat. Strat./Annul.

Although the comparison might not seem very good, the results of the two methods are

fairly close. OLGA predicts a somewhat lower transition velocity from Stratified to Slug

flow than Taitel-Dukler-Barnea, so that it shows slug flow only at the lowest rate. At the

highest rate, the predictions are really the same, since OLGA shows stratified flow with

considerable entrainment. Visually, this is indistinguishable from annular mist flow. The

only difference between the two methods is at the 250 mmSCFD. A good rule of thumb is

that slug flow does not occur for liquid holdups less than about 20%. The liquid holdup by

both the Oliemans and Beggs and Brill methods for the 250 mmSCFD rate is <7%, so the



prediction of slug flow by Taitel, et al. doesn’t appear likely. The OLGA prediction,

therefore, is more believable for this rate.

OLGA simulations for pigging were also performed. The liquid slug sizes ahead of the pig

for the various rates in the 28" line are:

Rate, mmSCFD Liquid Slug Size, bbls

125 33,900

250 3,770

500 1,920

If the line is run to equilibrium at the lowest rate, the slug catcher would have to be sized

for about 40,000 bbls. The capital expenditure to build a slug catcher that large would be

very significant. The best solution for the slug catcher design probably is to size the slug

catcher for the 250 mmSCFD rate, and pig the line frequently when the rate is less than

250 mmSCFD. To get an estimate of the pigging frequency that would be required,

calculate the time required to produce ~3770 bbls of liquid from the wellstream at the

desired rate. At 125 mmSCFD, the pigging frequency would be:

3770 bbls * 42 gal/bbl / 7.48 gal/ft / Vsl ft/sec / Ap ft2 / 3600 sec/hr / 24hr/D =

3770 * 42 / 7.48 / 0.0224 / 3.83 / 3600/ 24 = 2.86 days

If the line is pigged every 2-3 days at the 125 mmSCFD rate, the smaller slug catcher

could be used.

The design volume for the slug catcher should comprehend errors in the composition,

pipeline topography, and errors in the multiphase flow methods. The amount of the

increase should be based on sound engineering judgment. For our example, the 3770 bbl

maximum slug size calculated for the 250 mmSCFD rate should be increased to a design

slug catcher surge volume of at least 5000 bbls.



In summation, the best design for this pipeline would be:

a) install a 28" line;

b) install a slug catcher with a 5000 bbl surge capacity;

c) pig the line every 2 days at production rates less than 250 mmSCFD.



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



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



Figure I: 6-3 Liquid Flowrate Out of line, Example 1, Year 12



Figure I: 6-4 Gas Flowrate Out of line, Example 1, Year 12



Figure I: 6-5 Liquid Holdup Predictions for Example 2



SECTION 7.0 - REFERENCES

Barnea, D. “A Unified Model for Predicting Flow-Pattern Transitions for the Whole

Range of Pipe Inclinations,” Int. J. Multiphase Flow 13, No. 1, 1-12,, 1987.

Bendiksen, et al., “The Dynamic 2-Fluid Model OLGA: Theory and Application,” SPE

Production Engineering, pg. 171, May 1991.

Brill, J.P., et al., “Analysis of Two-Phase Tests in Large-Diameter Flow Lines in Prudhoe

Bay Field,” Society Petroleum Engineering Journal, pgs. 363-347, June 1981.

Hill, T.J. and Wood, D.G., “A New Approach to the Prediction of Slug Frequency,”

Society of Petroleum Engineers, pgs. 141-149, SPE 20629, 1990.

Taitel, Y. and Dukler, A.E., “A Model for Predicting Flow Regime Transitions in

Horizontal and Near Horizontal Gas-Liquid Flow,” AIChE Journal, 22, No. 1, pgs. 47-55,

1976.

Multiphase Pipeline Design Guide

Documents

pipeline design methods

design criteria

steady state design

pipeline elevation profile

section highlights

pressure gradient calculations

frictional losses

elevational losses