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
3.3 Section Highlights... 38
SECTION 4.0 – TRANSIENT FLOW MODELING 4.1 Transient Flow Modeling (General) ... 41
4.3 Section Highlights... 43
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
5.3 Section Highlights... 59
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
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
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.
Vm =Vsg + Vsl
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
Hl >Hlns
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:
Ug V sg l = − 1 H U V H l sl 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:
τyx µ
x
dv 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.
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:
V C
ns
max = ρ (Eqn. 2.1)
where Vmax = Maximum mixture velocity, ft/sec
ρns = No-slip mixture density, lb/ft3
=
(
ρg sg)
(
ρl sl)
m V V V + ρg = Gas Density, lb/ft 3ρl = 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
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 = ρmix c α 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 Annular Flow
Stratified Flow Dispersed
Bubble Flow Superficial Gas
Velocity
Strong Strong Strong Strong
Superficial Liquid Velocity
Strong Strong Strong Strong
Gas Density Moderate Strong Strong None
Pipeline Diameter Moderate Weak Weak Weak
Angle of Inclination
Moderate Weak Very Strong None
Liquid Properties Moderate Moderate Moderate Weak
As seen in the previous table, the influence of the major variables on the holdup is very different for each of the flow regimes. As a result, it is impossible to develop a general holdup correlation that will apply to all the flow regimes. Unfortunately, almost all of the commonly used holdup methods available in commercial software try to do this. They work poorly over much of the operating range as a result. The only way to accurately predict liquid holdup is to use mechanistic models for each of the flow regimes. The accuracy of available holdup methods is discussed further in Section 3.2.4.
2.4.3 Acceleration Losses
Although acceleration losses are present for all flow regimes, they are only significant for two flow regimes: annular flow and slug flow. The mechanisms for the losses in these two flow regimes are very different and will be discussed separately.
In single phase flow, acceleration losses can be calculated from 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 dx accel c = ρ 144
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.
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 Thermal Conductivity, Btu/hr-ft-degF Specific Heat, Btu/lb-degF Density, lb/ft3 Carbon Steel 26 0.11 490 Stainless Steel 8-13 0.11 488 Concrete (Saturated) 0.75-1.2 0.10 147-200
Onshore Soil (Wet) 1.35 0.20 90-110
Subsea Sandy Soil 1.25-1.50 0.30 105-115
Coal Tar Epoxy 0.20 0.35 92
Fusion Bonded Epoxy
0.15 0.32 75-90
Neoprene 0.12-0.15 0.50 90
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
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:
(
)
and Vsl =U Hl 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:
(
)
(
)
H U Vm U U l g g l = − − Vsl =U Hl l Vsg =Vm−VslTo 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:
(
)
(
)
ρns ρ ρ 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 off-design point. For most pipelines, worst case conditions for liquid holdup and flow regime occur at turn-down conditions.
3.3 Section Highlights
Points to remember from Section 3.0
-• 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
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 widely used commercial program, OLGA, began development in about 1983 and has been commercially available since 1990. OLGA’s only current competitor, PLAC, was introduced to the market at about the same time. Chevron currently does not own either program but has used OLGA for specific projects through consultants. Chevron internally developed a transient code, Transpire, in the same time frame as OLGA. This program has not been widely used, and it does not have as many features as the commercial codes.
Steady state simulators assume that all flowrates, pressures, temperatures, etc. are constant through time. Inherently transient phenomena, such as slug flow, are modeled 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 and energy to calculate the liquid and gas flowrates, pressures, temperatures and liquid holdups. These calculations are done at each time step. The programs utilize an iterative procedure, which ensures that a set of boundary conditions (such as inlet flowrates and outlet pressures as a function of time) are met while solving the conservation equations.
Steady state modeling can be used to size pipelines, but the predicted size may be inaccurate if the line is in slug flow. Transient simulators can size pipelines more accurately, and they are valuable in several other areas such as the design of downstream facilities, development of operating guidelines, and the diagnosis of operating problems. Steady state simulators cannot properly address any of these other concerns.
4.2 Use of Transient Simulators
Because of their power, transient simulators have been used for a variety of purposes. These uses include:
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.
4.3 Section Highlights
Points to remember from Section 4.0
-• 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