Take the Pressure Off Pressure Measurement

Take the Pressure Off Pressure

Measurement

Pressure Measurement eHANDBOOK

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Safe and accurate level measurement in the chemical and petrochemical industries

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This pressure sensor uses a metallic measuring cell capable of measuring liquids and gases up

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TABLE OF CONTENTSUnderstand Instrument Capabilities 6

Proper data interpretation depends upon knowledge of a device’s limits

Think Straight About Orifice Plates 9

Insufficient flow conditioning often undermines measurement accuracy

Find the Real Maximum Pressure 12

Always consider static head when assessing pressure vessels

Additional Resources 15

Pressure Measurement eHANDBOOK: Take the Pressure Off Pressure Measurement 3

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The Optibar LC 1010 submersible level probe with ceramic

diaphragm is a simple and continuous hydrostatic level

measurement solution for water wells, tanks and rainwater

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For safe and easy cleaning on site, the diaphragm is flush-mounted. With a diameter of 22 mm/1

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Plant testing and troubleshooting

often require data gathering to

identify specific problems. Data

gathering needs vary with the process

and the equipment. Data must reflect

fundamentals of the equipment opera-

tion. Getting usable results, of course, is

crucial — but data gathering efforts must

contend with time and money limitations

as well as safety constraints. Wonderful

high-precision instruments are available.

However, they may be too difficult to use

or too expensive for many routine tasks.

Troubleshooters must know the limits of

conventional measurement instruments

and understand what data are necessary

to come to usable conclusions. An exam-

ination of centrifugal pump performance

illustrates the need to adequately appreci-

ate instrument capabilities.

Every centrifugal pump has a performance

curve. As flow rate increases, dynamic

head across the pump drops. So, knowing

dynamic head across a centrifugal pump

can give information about flow rates. The

key question then becomes how accurate

must the required pressure readings be to

estimate dynamic head.

Figure 1 highlights the challenges posed

by pressure gauge accuracy on estimat-

ing flow rates. This particular pump was

installed in a liquid recovery system that

collected slop streams. The nominal design

Understand Instrument CapabilitiesProper data interpretation depends upon knowledge of a device’s limits

By Andrew Sloley, Contributing Editor

How accurate must pressure readings be to estimate dynamic head?



rate was 50 gpm but normal “expected”

operation was 35 gpm. The system did not

meter flow. One part of the troubleshoot-

ing effort required getting an accurate

estimate of the liquid rate. This pump has

a constantly rising dynamic head down

to shutoff (zero flow). One characteristic

of a pump with a constantly rising head

is that the dynamic head changes only a

small amount with flow rate as that rate

approaches zero.

The suction pressure was 15 psig. To get

a dynamic head reading, the trouble-

shooters used two pressure gauges: one

on the suction and one on the discharge.

The maximum expected discharge pres-

sure was ~90 psig based on a 0.7 specific

gravity fluid. For safety, an initial check

was done with a 0–200-psig gauge on the

discharge, then a 0–100-psig gauge was

used for the final readings. Because no

lower-pressure-range gauges were avail-

able, a 0–100-psig gauge was used for the

suction pressure as well.

If the pump were operating at the normal

rate of 35 gpm and on the pump curve,

the gauge on the discharge would give a

reading of about 82 psig. The best case

for standard commercial pressure gauges

found at most plants is an accuracy of ±2%

of the range for readings. So, the 0–100-

psig gauges used for the suction and

discharge pressure readings each should

have an error of ±2 psig. Additionally, both

PUMP CURVEFigure 1. A modest inaccuracy in pressure measurement can translate into a far greater error in flow rate.

221 ft @ 35 gpm

215 ft @ 44 gpm

228 ft @22 gpm

240

200

160

120

8

Hea

d, f

t

Flow gpm

0 40 80 120

Effi

cien

cy, %

40

30

20

10

0

HeadEfficiency

Design: 50 gpmNormal: 35 gpm

3,550 rpm7.5-in. impeller



gauges will be taking readings close to

the ends of their range: the 15-psig suction

pressure is in the low end of the range while

the 82-psig discharge pressure is in the high

end. Readings like these towards the end

of the range likely will have higher errors

than readings more in the middle of the

gauge’s range.

To get a number, let’s examine the effect

of a 2-psi error in the dynamic head. If the

actual pressure is 82 psi but the gauge

reads 80 psi, the height is 215 ft and the

estimated flow rate is 44 gpm. If the gauge

instead reads 84 psi, the height is 228 ft and

the estimated flow rate is 22 gpm. The flow

error range is -36% to +25% for a reason-

able error in measuring dynamic pressure.

Lower flow rates shift all readings to the

right on the curve, making likely errors

greater. Higher flow rates shift the read-

ings to the left on the curve, making flow

errors smaller.

The accuracy needed on the flow measure-

ments depends upon the problem you are

trying to solve. However, troubleshooting

rarely benefits from measurements with a

-36% to +25% error range on flow rates. At

high enough rates, pump discharge pres-

sure might give a useful indication of flow

rate but it’s never the best method.

A better option for pumps with motors is

to get an estimate of pump power from

the electric load and then calculate a flow

rate. In this case, the flow-rate error drops

to -18% to +10%. While this certainly isn’t

great, it still is a 50% improvement.

Of course, these analyses all depend upon

knowing the pump curve and assuming

the pump accurately follows it. At low

flow rates, small deviations due to wear

or other factors will significantly change

any flow rate estimate from pump or

motor data.

The case discussed is a difficult one. The

best solution here, if possible, is to directly

measure flow with an ultrasonic flow

meter. Nevertheless, the case dramatically

illustrates the importance of understand-

ing the accuracy of measurements and

how they may affect troubleshooting.

ANDREW SLOLEY is a contributing editor for CP’s

Plant Insites column. Email him at [email protected].



Think Straight About Orifice PlatesInsufficient flow conditioning often undermines measurement accuracy


Plants frequently rely on differential

pressure created by an obstruction

in a line to measure flow. Accuracy

depends upon two factors: the correctness

of the differential pressure measurement

obtained via taps upstream and down-

stream, and the calculation for turning that

measurement into a flow rate.

The obstruction placed in the line most

often is an orifice plate — a flat plate with

a machined orifice. (For more on orifice

plates, see: “Remember the Old Reliable

Orifice Plate,” https://bit.ly/3cmBvZK; for

other differential-pressure flow metering

options, see: “Look Beyond Orifice Plates,”

https://bit.ly/3pFXAIx.) Orifice plates

are cheap and reliable. Moreover, orifice

plates manufactured to specific dimensions

and tolerances generate known pressure

drops for a given flow rate. The Interna-

tional Standards Organization (ISO) has

summarized the dimensional criteria; all

reputable orifice-plate manufacturers meet

these standards.

ISO standards also cover installation

requirements. Proper installation plays

a crucial role in achieving accurate ori-

fice-plate measurements. The major criteria

include a stable flow pattern, a fluid-filled

pipe and an unobstructed flow path (no

blockages). If these criteria are met, flow

meter calculations can be based on the

Many orifice meters inside process units don’t meet ISO standards.



physical dimensions of the system; no

in-place measurement or calibration is

required.

Let’s look in detail at the first requirement,

a stable flow pattern. An oft-repeated rule

of thumb states that a length of straight-run

pipe equal to 10–15 piping diameters creates

a sufficiently stable flow pattern. How does

this compare to the ISO standards?

The ISO standards include multiple

upstream piping configurations — from

fully open full-bore valves upstream and

downstream to multiple right-angle turns

to tee-branch connections. They also detail

for multiple values of β — the orifice diam-

eter/pipe diameter, in consistent units

— the length of straight-run piping required

(Table 1). In general, the lower the b, the

shorter the pipe run necessary. During

piping design, the final β ratio is unknown.

So, many engineering standards attempt to

reduce overall cost by specifying a maxi-

mum β of 0.55 to 0.63.

The best cases are fully open full-bore

valves with a straight run upstream of them,

and a single right-angle bend upstream.

The required piping runs for a 0.55 β are

13 diameters for the full-bore valves and 16

diameters for the single right-angle bend.

Every other configuration is worse — in

some cases, much worse. Higher β values

increase upstream requirements.

For two 90° bends in series, an orifice with

a 0.55 b requires 44 diameters of upstream

CAS

3”x4” 4”x3”

FIC

XXX

FE

XXX

FT

XXX

Upstream Configuration value

<0.32 0.45 0.55 0.63 0.70 0.77 0.84

Fully open, full-bore valve 12 12 13 16 20 27 38

Two right-angle bends in same plane, Two or three

bends at right angles with straightening vanes15 18 22 28 36 46 57

Two or three bends at right angles, Flow branch 35 38 44 52 63 76 89

Fully open globe valve 18 20 23 27 32 40 49

Single right-angle bend 10 13 16 22 29 44 56

BAD IDEAFigure 1. Installing a short section of larger diameter pipe would create flow pattern with unknown impact on meter.

ISO INSTALLATION REQUIREMENTSTable 1. Required number of pipe diameters in upstream straight run generally decreases with value.



piping to meet ISO standards. Even with

properly installed straightening vanes, this

layout needs 22 diameters. A β of 0.84

raises the requirement to 40+ diameters for

all types of installations.

What this all means is that if your plant

needs maximum accuracy, use lots of

pipe run upstream of orifice plates. In

some cases, 90 diameters are necessary.

Additionally, if you’re having flow meter

problems, check the installation. I’ve

observed many orifice meters inside pro-

cess units that don’t meet ISO standards.

The 10–15-diameters rule only applies

to a “best case” — i.e., everything else is

done correctly and a low-b orifice plate is

installed. Most industrial installations require

20+ diameters. Using straightening vanes

can help, but doesn’t completely solve the

problem. The toughest installations are

downstream of flow branches and where

multiple elbows in series are at right angles

to each other. To paraphrase a quote from

pump installation guidelines, the only thing

worse than one elbow upstream of a flow

orifice is two elbows.

While a plant may start with low-b orifice

plates, as hydraulics become tighter it may

put in new plates with lower pressure drops

(and higher β values). Installing a short run

of larger diameter pipe doesn’t solve the

problem (Figure 1). The upstream expansion

creates a flow pattern with unknown effect

on the orifice meter.

If the piping configuration doesn’t meet ISO

standards, accuracy will suffer. For monitor-

ing unit trends, reduced accuracy may be

an acceptable tradeoff for a cheaper meter

installation. For high and reliable accuracy,

always follow the ISO requirements.

ANDREW SLOLEY is contributing editor for CP’s Plant

Insites column. Email him at [email protected].



Find the Real Maximum PressureAlways consider static head when assessing pressure vessels


Chemical plant vessels serve many

purposes, including for storage

and surge control and as reactors,

fractionators, absorbers, strippers and

crystallizers. Pressure is a key parameter

for safe operation. Vessel operating and

design pressures may appear in piping

and instrumentation diagrams (P&IDs),

specification sheets, operating instruc-

tions and fabrication drawings.

Process safety analyses invariably address

over-pressure protection. Such analyses

generally rely on P&IDs for plant design

information. The P&IDs usually include

vessel maximum operating and working

pressures. But what do those pressures

mean? Are the P&ID pressures the ones

we really must worry about?

For vessels stamped as complying with

the American Society of Mechanical Engi-

neers (ASME) Boiler and Pressure Vessel

Code, the ASME U-1 form summarizes the

vessel’s design temperature and pressure.

It includes the design conditions and the

specific materials used, allowable materi-

als stresses, testing conditions and other

critical mechanical details. That form (and

any attachments stemming from vessel

modifications or repairs) defines the oper-

ating limits. The vessel should be code

stamped with the same values shown on

the U-1 form. While useful and convenient,

other paperwork doesn’t override the

U-1 form.

So, it’s important to understand how

P&ID design pressures compare to those



on ASME U-1 forms and to know some

common errors in P&ID pressures.

Section VIII Division 1 of the ASME Code

covers rules for construction of pressure

vessels. Subsection UG-21 defines design

pressure requirements: “Each element of

a pressure vessel shall be designed for at

least the most severe condition of coinci-

dent pressure (including coincident static

head in the normal operating position) and

temperature expected in normal operation.”

The design pressure of a vessel is the max-

imum pressure that any part of the vessel

can tolerate — it includes both system and

static pressures. Design pressures may

vary with temperature and vessels may be

stamped for multiple design temperature/

pressure combinations.

Common practice during hazard and oper-

ability (HAZOP) reviews is to use pressure

ratings on P&IDs rather than referring

back to U-1 forms. As long as the values

are correct and people properly under-

stand how to interpret them, this doesn’t

cause problems.

The U-1 form design values are for any

point on the vessel and must include static

head in the pressure evaluation. Too often,

HAZOP and other safety reviews look at

operating pressure from a pressure reading

point and fail to consider the implications

of static head. All vessels not under vacuum

have a static head component of pressure.

The static head may vary from insignificant

for a horizontal vessel under vacuum to

very high for a tall liquid-filled vessel. As

an example, let’s consider a vessel that has

a seam-to-seam height of 47 ft. 8 in. and

contains a mix of hydrocarbons and a liquid

ionic catalyst (hydrogen fluoride). The

average density of the liquid is 48.8 lb/ft3

at operating conditions. The U-1 form indi-

cates the vessel is designed to handle 165

psig at 250°F.

PRESSURE RELIEF VALVE OPTIONSFigure 1. Location of the valve can markedly affect the maximum release setting.

May 2011 cheMicalprocessing.coM 42

plant insites

Find the Real Maximum PressureAlways consider static head when assessing pressure vessels

CheMiCal Plant vessels serve many purposes, including for storage and surge control and as reac-tors, fractionators, absorbers, strippers and crystal-lizers. Pressure is a key parameter for safe operation. Vessel operating and design pressures may appear in piping and instrumentation diagrams (P&IDs), specification sheets, operating instructions and fabri-cation drawings.

Process safety analyses invariably address over-pressure protection. Such analyses generally rely on P&IDs for plant design information. The P&IDs usually include vessel maximum operating and working pressures. But what do those pressures mean? Are the P&ID pressures the ones we really must worry about?

For vessels stamped as complying with the American Society of Mechanical Engineers (ASME) Boiler and Pressure Vessel Code, the ASME U-1 form summarizes the vessel’s design temperature and

pressure. It includes the design conditions and the specific materials used, allowable materials stresses, testing conditions and other critical mechanical details. That form (and any attachments stemming from vessel modifications or repairs) defines the operating limits. The vessel should be code stamped with the same values shown on the U-1 form. While useful and convenient, other paperwork doesn’t over-ride the U-1 form.

So, it’s important to understand how P&ID design pressures compare to those on ASME U-1 forms and to know some common errors in P&ID pressures.

Section VIII Division 1 of the ASME Code covers rules for construction of pressure vessels. Sub-section UG-21 defines design pressure requirements: “Each element of a pressure vessel shall be designed for at least the most severe condition of coincident pressure (including coincident static head in the nor-mal operating position) and temperature expected in normal operation.”

The design pressure of a vessel is the maximum pressure that any part of the vessel can tolerate — it includes both system and static pressures. Design pressures may vary with temperature and vessels may be stamped for multiple design temperature/pressure combinations.

Common practice during hazard and operabil-ity (HAZOP) reviews is to use pressure ratings on P&IDs rather than referring back to U-1 forms. As long as the values are correct and people properly understand how to interpret them, this doesn’t cause problems.

The U-1 form design values are for any point on the vessel and must include static head in the pres-sure evaluation. Too often, HAZOP and other safety reviews look at operating pressure from a pressure reading point and fail to consider the implications of static head. All vessels not under vacuum have a static head component of pressure.

The static head may vary from insignificant for a horizontal vessel under vacuum to very high for a tall liquid-filled vessel. As an example, let’s con-sider a vessel that has a seam-to-seam height of 47 ft. 8 in. and contains a mix of hydrocarbons and a liquid ionic catalyst (hydrogen fluoride). The average density of the liquid is 48.8 lb/ft3 at operating condi-tions. The U-1 form indicates the vessel is designed to handle 165 psig at 250°F.

Other paperwork

doesn’t override

the U-1 form.

PRVOption 1

Piping Elevation67’-4” —

Hydrocarbon Out59’-4” —

Catalyst Out20’-11” —

Feed3’-3” —

Piping Elevation— 12’-0”

PRVOption 2

PRVOption 1

Pressure Relief Valve Options

Figure 1. Location of the valve can markedly affect the maximum release setting.

CP1105_42_43_InSites.indd 42 4/26/11 2:37 PM



Per the ASME code, the pressure relief

device must open at or before any part

of the vessel reaches the design pressure.

Figure 1 shows the vessel layout with the

current pressure relief valve (PRV) location

as Option 1 and a new proposed location

as Option 2.

The first common misconception often

encountered is that a vessel’s rating allows

operation at the design pressure at the

pressure measurement point. This plant

mythology is false. The ASME code doesn’t

specify the location of required pressure

measurement points. The code specifies

that the system pressure plus the coinci-

dent liquid head must be below the vessel

design limit at all points on the vessel.

For our example, this limit is the lower

edge of the feed nozzle at the bottom of

the vessel.

The feed nozzle centerline is at 3 ft. 3 in.

above grade. The lower edge of the nozzle

is 9 in. lower, at 2 ft. 6 in. The PRV cur-

rently is located in a pipe rack downstream

of the vessel (Option 1). The PRV inlet is

at 12 ft. 0 in. At an operating density of

48.8 lb/ ft3 this gives 5 psi of static head

(rounded up). The maximum PRV relief set-

ting is 161 psig (165 psig design limit minus

4 psi static head).

A proposal “to take full advantage of the

vessel design pressure” would move the

PRV to the Option 2 location. This idea

stems from the thought that the vessel’s

top seam defines the design pressure. This

idea is wrong. It’s also very curious. Even

if you believe that the top seam defines

the design pressure, why move the PRV?

Just reset it to account for the correct

static head.

What must happen if the PRV moves to the

Option 2 location? The elevation of the pipe

is 67 ft. 4 in. and the PRV inlet is at 69 ft. 0

in. At an operating density of 48.8 lbs/ft3

this gives 23 psi of liquid head (rounded up

again). The maximum PRV release setting

is 142 (165 psig design limit minus 23 psi

static head).

The key point is that pressure relief devices

must protect all points of the vessel from

exceeding design pressure. Moving PRVs

doesn’t change vessel design pressures.

Always go back to the U-1 forms when you

must verify design pressure. Other docu-

mentation, while convenient and helpful,

remains secondary to the U-1 forms and

related vessel code stamps.

ANDREW SLOLEY is contributing editor for CP’s Plant

InSites column. Email him at [email protected].

Other paperwork doesn’t override the U-1 form.




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Take the Pressure Off Pressure Measurement