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GAS FACILITIES

“RULE OF THUMB”

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INDEX Page Treating 6 to 12 Sulfur Recovery Units 6 Glycol Dehydration 6 Amine Treating 9 Mol Sieve Treating 9 Corrosion Information 9 Copper Strip Testing 9 Conversion Factors 10 Caustic Washer Design 10 Metallurgy for Amine Treaters 10 H2S Toxicity Data 12 Iron Sponge 12 Sulfur Compounds in NGL 12 Separation 13 to 16 Vertical Knockout Drum Sizing 13 Crude Oil Separator Sizing 13 Vertical Separator Design 13 Horizontal Separator Design 14 Brown-Souders Equation for Vessel Sizing 15 Mist Extractor Selection 16 Refrigeration 17 Condensers 17 Propane Refrigeration Systems 17 Condensing Temperature Effects 17 Heat Transfer 18 to 27 Overall Heat Transfer Coefficients 18 Heat Exchanger Velocities 19 Allowable Pressure Losses 19 Cooling Water Temperatures 19 Mean Temperature Differences 19 Fouling Factors 20 Materials of Construction 20 Heat Exchangers - General 21 Condensers 21 Reboilers and Chillers 22 Plate Heat Exchangers - Calculation Method 22 Brazed Aluminum Exchangers 23 Air Fin Exchangers 23 Fired Heaters 24

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Organic Heat Transfer Fluids 24 Cooling Towers 25 Insulation 25 NGL Expander Plants 26 Plant Systems 26 Sizing Shell/Tube Exchangers 26 Physical Fan Laws 27 Fractionation 28 Fluid Flow 30-35 Miscellaneous 30 Expander Plants 30 Piping Sizing 31 Piping 31 Maximum Operating Velocities 32 Fricition Factor 33 Piping Noise 33 Control Valves 33 Two Phase Flow 34 Compressors, Expander and Pumps 36 to 40 Reciprocating Compressors 36 Compressor Quickies 36 LNG 37 Energy Conservation 37 Fuel Consumption 38 Expander Plants 38 Simulation Guidelines 38 Pump Sizing 38 Pumps 39 General 40 Combustion 41 Flares 41 Fired Heater 41 Fuel Requirements 41 Miscellaneous 42 to 44 Water / Steam Systems 42 Economics 42 Hydrates 42 Expander Plants 43 Misc. Plant Systems 43 Wind Loading 43

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Steam Leaks 43 Composition of Air 43 Storage Vessel Capacity 43 Pipeline Volumes 43 Pressure Vessels 43 NACE Requirements 44 Pressure Waves 44 Absolute Pressure of Height 44 Boiler HP 44 Solar Radiation 44 Combustion - Air Requirements 44

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Treating

1. Sulfur Recovery Units a. Thermal zone will produce 55-65% of the sulfur and is a function of the H2S content

of the feed. The catalytic region makes the rest. b. If the acid gas feed is less than 30% H2S, then flame stability in the reaction furnace is

a potential problem. Minimum temperature for effective operation is 1700 F. c. Temperature in catalyst beds should be kept below 800 F. d. SRU steam production will be approximately 6700 lbs of steam per long ton of sulfur

produced. e. Glossy carbon deposits on catalyst indicates amine carryover. f. Sulfur fog is caused by too much cooling capacity. Sulfur mist can be caused by

excessive velocity in the condenser. g. Ferrules should extend at least 6" inside the tubesheet. Refractory lining is usually 2½-3" thick on the tubesheet. h. Mass velocity in waste heat exchanger and sulfur condenser tubes should be 2-6 lbs/sec-ft2. i. Space velocity through catalyst beds should be 700-1000 standard cubic feet per hour

(SCFH) of gas per cubic foot of catalyst. Lean streams require the lower value and rich streams the higher value.

j. Sulfation of catalyst is caused by SO3; oxygen combines with SO2 to form SO3 which is chemisorbed on alumina surface.

k. Velocity in process piping should not exceed 100 ft/sec. l. Liquid sulfur solidifies at 246 F and becomes very viscous above 320-350 F. m. Approximate Stack Gas Flow, scfm: SGF = (Sulfur Production, LT/D) x (100)

2. Glycol Dehydration

a. TEG - Dew point depression temperature ranges from -80 to -140 F. The degree of dehydration which can be obtained depends on amount of water removed from glycol in the reboiler and the circulation rate. The minimum circulation rate to assure good glycol/gas contact is approximately 2 gallons of glycol for each pound of water to be removed. Seven gallons is maximum and standard is three gallons.

b. Stripping Gas - Approximately 3-8 scf/gal of glycol circulation. c. Glycol will absorb approximately one (1) scf of gas/gallon of glycol. Glycol Contractor

- For best scrubbing of overhead gas install "Mist Pad" on the face of "Vane Type" mist extractor.

d. Estimate total reboiler duty from 2000 Btu/US gallons TEG circulation rate. The use of glycol/glycol heat exchangers will reduce the total reboiler duty.

e. Glycol loss should be approximately 0.1 gallon TEG/MMscf. f. Packing - Minimum of 4' in any gas-glycol contactor. g. Triethylene Glycol Dehydration Unit - Maximum recommended heat flux for a direct

fired TEG regenerator is 8000 Btu/square foot of fire tube surface area. The recommended heat flux of maximum fire tube life is 6000 Btu/ft2.

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h. Troubleshooting - A black viscous solution indicates that heavy hydrocarbons have been carried over with the gas. A sweet, burnt sugar smell accompanied by low pH and a dark is a clear signal that thermal degradation is occurring.

3. Amine Treating a. Amine Circulation: 3 cu.ft acid gas/gal amine

MEA gpm = 41.0 * Q*X/Z DEA gpm = 45.0 * Q*X/Z (conventional) DEA gpm = 32.0 * Q*X/Z (high load)

where Q = Gas, MMscfd X = Acid Gas, volume percent Z

= Amine Concentration, wt.%

b. Maximum acid gas pickup is not more than 0.35 mols/mol of MEA. Normal value is 0.30.

c. Amine treating processes tend to be troubled by the same problems regardless of the type amine used.

d. Typical MEA losses due to entrainment:

Absorber: 1.0 Lbs/MMscf Still:

2.5 Lbs/MMscf

e. Flow Velocity - Rich Stream: Not to Exceed 5 ft/sec Lean Stream: Not to Exceed 7 ft/sec f. Filter Beds: Recommended flow rate through a carbon bed is 4 gpm/ft2 (cross

sectional area) which is approximately equal to 20 minutes superficial contact time g. Loadings: 0.36 mols CO2/mol MEA [absorber RICH] 0.12 mols CO2/mol MEA [still LEAN] h. Reflux Ratio: MEA and DEA 1.5 to 3.0 [mols H20/mols acid gas leaving reflux drum] i. Equivalent Steam Rate: MEA 0.9 to 1.2 lbs steam/gal amine DEA 0.8 to 1.1 lbs steam/gal amine j. Lean amine can contain 0.05 to 0.08 mols total acid gas and still meet specs. k. CO2 and H2S gases appreciably increase total water content and dehydration

requirements of gas streams. l. Recommended maximum ranges for amine strength and acid gas loadings that have

proven historically to adequately address corrosion concerns are:

Amine

Wt% Rich Loading

Moles Acid Gas / Mole Amine MEA 15 – 20 0.30 - 0.35 DEA 25 – 30 0.35 - 0.40

MDEA 50 – 55 0.45 - 0.50

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m. Recommended loading in the lean circuit to minimize acid gas flashing are: Total Lean Loading

Amine MEA

Moles Acid Gas / Mole Amine* 0.10 - 0.15

DEA 0.05 - 0.07 MDEA 0.004 - 0.010

* These loadings should be easily achieved with a 1.0-2.0 mole/mole stripper reflex ratio.

n. Recommended Minimum Water Quality Standards for make-up water for amine plants:

Total Dissolved Solids

<100 ppm

Total Hardness <3 grains/gal Chlorides <2 ppm Sodium <3 ppm Potassium <3 ppm Iron <10 ppm

o. Liquid/Liquid Contactors: (Feasibility Sizing Data) For rough diameter sizing of liquid/liquid contactors for amine treating of light

hydrocarbon liquids, use 12-15 gallons per minute (gpm) of amine. This should correspond to approximately 10% of flooding velocity. For hydrocarbon distributor nozzles for liquid/liquid contactors, use an orifice velocity of approximately 1 ft/sec. Higher velocities than this can lead to emulsion problems. Velocities lower than 0.5 ft/sec can result in NGL being entrained in the sour amine stream.

To estimate height of packing required, assume 6-8 ft of packing for each theoretical

separation stage. p. Mercaptan Removal from Gas MEA and DEA will remove approx.: 40-55 mol% methylmercaptan 20-25 mol% ethylmercaptan 0 -10 mol% propylmercaptan Regenerative caustic process will remove mercaptans down to <10 ppm. Activated Carbon, Calgon FCA, will remove 4-5 wt% mercaptans.

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4. Mol Sieve Treating a. Bed Design (Length/Diameter)

Minimum Maximum Liquid L/D 3:1 5:1 Gas L/D 2:1 4:1

b. Maximum gas velocity: 0.33 to 0.75 ft/sec (superficial linear velocity) c. Maximum liquid velocity: 50 bbls/hr/ft2 (bed area) Dehy-ALCOA: 30 gpm/ft2 (=43 bbl/hr/ft2) d. Minimum velocity: Liquid: 60 second contact time, or 0.01 psi/ft (liquid). Gas: 3.5 second contact time e. Mol Sieve: Draining bed leaves approximately 25 vol% of total bed volume on bed as

sponged liquid. f. Alumina: Draining bed leaves approximately 0.048 gals. per pound of alumina on bed

as sponged liquid. g. Molecular Sieve Dehydrators: As strictly a rule of thumb based on many Phillips

designs, when the pressure drop through a mol sieve bed reaches 20 psi, the bed support is nearing its maximum load capacity and action should be taken to reduce the pressure drop.

5. Corrosion Information

a. CO2 Corrosion: Low Corrosion - Pco2 7 psia High Corrosion - Pco2 = 7-15 psia High Corrosion - Pco2 > 15 psia Where: Pco2 = Partial Pressure of CO2

b. Corrosion rate is also directly related to temperature. 6. Copper Strip Testing Copper Strip Test ASTM 5.05 D1838 No. 1A copper strip normally < 1-2 ppm H2S. H2S corrosive to copper strip 1 ppm or .16 gr/100 sc. Copper Strip will not detect mercaptan or other organic sulfides.

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7. H2S and CO2 Conversion Factors 1 mol% H2S = 630 grains/100 scf 1 mol% CO2 = 813 grains/100 scf Grains of H2S/100 scf x 15.9 = mol% H2S 1 MMscf H2S = 37.6 long tons sulfur 1 grain H2S/100 scf = 17.1 ppmw = 22.8 mg/m grain H2S/100 scf = 15.9 ppmv grain CO2/100 scf = 12.3 ppmv 2S ppmw = [gr. H2S/100 scf] x [542/(mol wt gas)] 8. Caustic Washer Design Vertical washers are sized by using a factor of 400-500 gal/hr/sq.ft. of cross-sectional area

of empty tower. In the tower, 5-7 ft. of raschig rings equal one stage. 9. Metallurgy Requirements For Amine Treaters a. Vessels 1. Amine contactor, flash tank, stripper, surge tank, accumulator, inlet scrubber, and

outlet scrubber should be manufactured using carbon steel and stress relieved with corrosion allowances as shown in 9.a.3.

2. Trays for the contractor and stripper should be constructed using 304 stainless. 3. Corrosion allowances for MEA and DEA Systems.

CO2/H2S<20

Inches CO2/H2S>20

Inches Inlet Scrubber 1/8 1/8 Amine Contactor 1/8 1/8 Outlet Scrubber 1/16 1/16 Flash Tank 1/8 1/8 Cross Exchanger 1/8 1/8 Amine Stripper 3/16 3/16 Reflux Accumulator 1/8 1/4 Reboiler 1/8 1/8 Reclaimer ¼ 1/4 Surge Tank 0 0 Piping 1/16 1/16 Amine Cooler* 1/8 1/8 Stripper Overhead Condenser* 1/8 1/4

* Corrosion allowance applies to shell side exchangers with water cooling in the tubes.

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b. Heat Exchangers 1. Shells - carbon steel and stress relieved 2. Tubes - 12 gauge minimum, carbon steel, seamless

3. Reclaimer element - carbon steel, 2-inch schedule 80 tubes. Amine temperature in reclaimer should not exceed 310 F.

4. Temperature of amine in reboiler should not exceed 250°F. 5. U-bends of U-tube carbon steel bundles should be stress relieved.

c. Pumps

1. The amine circulation and stripper reflux pumps should be constructed of carbon steel with 316 stainless trim.

d. Piping 1. Piping should be carbon steel. The weld and heat affected zone of piping

containing H2S and H2O, with or without amine, should have a hardness no greater than Brinnell hardness number 235.

2. Velocities in the rich and lean solution piping should be limited to 2-3 ft/sec. for MEA solution, and to 7 ft/sec. for DEA solution.

3. Piping from the letdown valve to the flash tank and from the letdown valve to the stripper should be 304 stainless with the letdown valves of 316 stainless.

e. An inert gas blanket should be maintained on the fresh amine storage and amine surge

to prevent oxygen contact with the amine. f. No copper-bearing materials such as Admiralty, Monel, etc., should be used anywhere

in amine units. g. When the CO2/H2S ratio is greater than 20, the following exceptions to the above

requirements should be used.

• The exchanger tubes in the reboiler, stripper overhead condenser, lean-rich cross exchanger, and amine cooler should be 304 stainless. Those exchangers with water in the 304 stainless tubes should have a minimum water flowrate of 5ft/sec. Limit amine velocity in tubes to 5ft/sec.

• The stripper vessel head and shell down through the top three trays should be 304 stainless clad or solid 304 stainless.

• The stripper overhead piping from the stripper to the accumulator should be 304 stainless. The reflux piping from the accumulator back to the stripper can be either thin wall 304 stainless or carbon steel with 1/16" corrosion allowance.

The vapor line from the reboiler to the stripper should be 304 stainless.

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10. H2S Gas Toxicity Data

ppm gr./100 scf 10 0.65 Can smell. Safe for 12 hr. exposure.

100 6.48 Kills sense of smell in 2-15 minutes. 200 12.96 Kills sense of smell quickly. Stings eyes and throat. 500 32.96 Loses sense of reasoning and balance. Respiratory

paralysis in 30-40 minutes. 700 45.36 Breathing will stop and death result if not rescued

promptly. Immediate artificial resuscitation. 1000 64.80 Unconscious at once. Permanent brain damage or death

may result unless rescued promptly.

11. Iron Sponge H2S Removal: W = 1.43*GQ = .09*PQ Where: W = H2S removal lbs/day G = H2S gr/100 scf inlet P = H2S ppmv Q = MMscfd

12. Distribution of Sulfur Compounds in NGL Product • H2S will be concentrated mainly in the C3 and lighter. • COS will be concentrated mainly in the C3 streams. • CH3SH will primarily be split between the C3 and C4 streams. • CH3CH2SH will be concentrated mainly in the C4 and heavier streams.

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Separation 1. Vertical Knockout Drum Preliminary Sizing

a. Size for vapor

W=1100 [Density Vapor (Density Liquid-Density Vapor)] 1/2

Where: W = maximum allowable mass velocity in pounds/hour/ft2

1100 = empirically determined constant Densities = pounds/cubic foot at process T & P

b. Size for liquid

Should be able to contain maximum slug expected depending on pipe configuration. Never size for less than one minute liquid holdup. Size 8-10 ft tall.

2. Crude Oil Service Separator Sizing

a. Use vertical separator for high vapor to liquid ratios and for two phase separation.

b. Use horizontal separator for high liquid to vapor ratios and for three phase separation. Vessel L/D 3 to 5.

c. Check size for both gas and liquid handling (i.e., gas superficial velocity and liquid

residence time).

d. Use 3 minute liquid residence time for the hydrocarbon phase in a crude oil system. e. Use 3-6 minutes residence time for the water phase in a crude oil system.

f. Estimate 15-30 minutes water residence time for electrostatic coalescers (100% filled).

Vessel L/D 4 to 6. 3. Vertical Separator Design

a. The disengaging space - the distance between the bottom of the mist elimination pad and the inlet nozzle, should be equal to the vessel internal diameter or a minimum of 3'-0".

b. The distance between the inlet nozzle and the maximum liquid level should be equal to

one-half the vessel diameter, or a minimum of 2'-0".

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c. A mist eliminator pad should be installed. Otherwise the separator should be designed so that actual gas velocity should be no greater than 15% of the maximum allowable gas velocity as calculated by the following equation:

Vg = (11.574)*(MMcfd) /A

Where: Vg = Vertical velocity of gas, ft/sec MMcfd = Actual gas volume at operating conditions, MMcfd

A = Cross Sectional Area of vessel, ft2

d. The dimension between the top tangent line of the separator and the bottom of the mist eliminator pad should be a minimum of 1'-0".

e. Inlets should have an internal arrangement to divert flow downward.

f. Liquid outlets should have antivortex baffles.

g. Mist eliminator pads should be specified as a minimum of 4 inches thick, nominal 9

lb/ft3 density and stainless steel.

h. Normal practice for calculating liquid retention time is to allow for the volume contained in the shell portion of the vessel only. No credit is taken for any liquid retention time attributable to the volume contained in the vessel head. Sump height should be a minimum of 1'-6' to allow for liquid level control.

4. Horizontal Separator Design

The following are commonly used rules of thumb for sizing horizontal separators:

a. Oil level is usually controlled by a weir, which is commonly placed at a point corresponding to 15% of the tangent-to-tangent length of the vessel. This results in 85% of the vessel being available for separation. Height of the weir is commonly set at 50% of the internal diameter of the vessel.

b. The maximum liquid level should provide a minimum vapor space height of 1'-3" but

not be substantially below the center line of the vessel.

c. Separators designed for gas-oil-water separation should provide residence time and separation facilities for removal of the water.

d. For separators handling fluids where foaming is considered a possibility, additional

foam disengagement space and foam control baffling should be provided. Mist eliminator devices should be located external to the vessel to maximize foam disengagement potential within the vessel.

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e. The volume of dished heads should not be taken into account in vessel sizing calculations.

f. Inlet and outlet nozzles should be located as closely as possible to vessel tangent lines.

g. Liquid outlets should have antivortex baffles.

5. Brown - Souders Equation For Vessel Sizing

W = C [Dv (D1 - Dv)]½ W = Vapor Loading - lb/hr/sq. ft.

Dv = Vapor Density - lb/cu ft. at operating condition D1 = Liquid Density - lb/cu ft.

C = Constant (a) for absorbers use 600 (b) for scrubbers use 1100 (c) for still use 500

Example: Size scrubber for field engine discharge

Dv = 29.423 lb/mol x 63 psia x 520 = 0.333 lb/cu. ft.

380.6 cu. ft./mol 14.65 550

D1 = 0.82 Sp. Gr. x 62.3 lb/cu. ft. (H20) = 51.0 lb/cu. ft. W = 1100 [.333(51.0 - .333)]½

= 1100 [16.9]½

= 4520 lb/hr./sq. ft.

Gas Flow = 158,311 Mpd x 29.423 lb/mol = 194,000 lb/hr.

Cross Section Area Required = 194,000 lb/hr. = 43 sq. ft. 4520 lb/hr/sq. ft.

Dia. = Q(43/.7854)½ = 7.4 ft. Use 8 ft. diameter scrubber

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6. Mist Extractor Selection a. The stainless steel mesh pad type mist extractor is generally less expensive than the

vane type and is adequate for most clean service applications. Similar liquid removal efficiencies can be achieved (within certain velocity constraints) with mist particle sizes of 10 microns and larger.

b. The pad type usually has less clean pressure drop than the vane type.

c. The vane type usually performs better than the pad type where tacky solids such as iron sulfide are present in the flowing gas stream. The liquid flow from the mist extractor is at right angles to the gas flow in vane type and it tends to wash solids away better.

d. If the van type is used in corrosive service (hydrogen sulfide, carbon dioxide, or

oxygen with water wet gas), the vanes should be 316 stainless steel. Experience has shown that a small amount of corrosion with carbon steel vanes roughs the surface and solids tend to accumulate and plug the vanes rapidly.

e. For retrofit or sometimes new applications, it’s possible to use a smaller diameter

vessel for the vane type as it may be fitted in different orientations to limit the velocity to acceptable ranges. The pad type is usually installed horizontally.

f. It is usually cheaper to retrofit vessels with the pad type as both would have to be cut

and match marked to fit through an 18" or smaller manway and reinstalled inside the vessel. The vane type usually has boxing that must be welded together inside the vessel while the pad type can usually be bolted.

g. The vane type may be used for small in-line applications where the pad type usually

can not. h. If the pad type plugs with solids or hydrates, the pressure drop will likely dislodge the

mist extractor and plug downstream piping or equipment. i. For tough separation applications where it is necessary to remove mist particles

smaller than 10 microns (such as inlet to glycol or amine systems where the foreign liquid may cause foaming or chemical contamination), often a combination of pad type (for coalescing) and vane type (for mist removal) is used.

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Refrigeration 1. Condensers

For water cooling of propane condensers, cool no further than a 10 F approach to the warm cooling water leaving the condenser. Use 4 to 8 ft/sec velocity for water through the tubes. Restrict cooling water return temperature to a maximum of 125 F.

2. Propane Refrigeration Systems

a. Use a composition of 3.0 mol% C2, 95 mol% C3 and 2 mol% C4. Experience has indicated it is difficult and expensive (propane losses) to maintain 99%+ propane content. Continuous purge systems usually result in high losses.

b. Air (pulled in around piston rod packing with low suction pressure of 1 to 2 psig and high valve losses) tends to accumulate in the propane accumulator vessel vapor space after the condensers. It causes high compressor discharge pressure and a potentially hazardous situation. A manual purge of the accumulator vapor space weekly will keep air concentration down.

c. For water cooled propane condenser design, generally use 10 F temperature rise on cooling water through exchanger and a 10 F approach of condensed propane out to the warm water from the exchanger.

d. For reciprocating propane compressor calculations, add 10% to the final horsepower calculated by conventional means and 10 F to the final stage discharge temperature for preliminary design.

e. Gas Processing - Simulation Guidelines Refrigerant chiller temperature approach to gas is normally 5-10°F.

f. Troubleshooting: Refrigerant composition can be checked by operating conditions. Light hydrocarbons are found in the surge tank where they accumulate. This usually shows up as higher than expected discharge pressures. The quick check for heavies is temperature of the low stage chiller at operating pressure being higher than expected. If light hydrocarbons are a problem, venting off the surge tank can help. If no vent system is in place, the surge tank can be vented manually for short periods. Let the system come back to equilibrium and vent a little more. This should eliminate the lights and minimize propane losses.

3. Condensing Temperature Effects

For gas turbine driven propane refrigeration systems, there will be approximately ½-1% increase in gas load horsepower for every degree Fahrenheit increase in refrigerant condensing temperature.

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Heat Transfer 1. Overall Heat Transfer Coefficients Overall heat transfer coefficients suitable for feasibility design estimates are provided in Table 1 below:

TABLE 1 Hot Fluid Cold Fluid Overall U

Btu / HR-FT2-F Water Water 250-500

Ammonia Water 250-500

MEA or DEA Water 140-200

Fuel Oil Water 15-25

Fuel Oil Oil 10-15

Gasoline Water 60-100

Heavy Oil Water 15-50

Heavy Oil Heavy Oil 10-40

Reformer Stream Reformer Stream 50-120

Light Organics Water 75-120

Medium Organics Water 50-125

Heavy Organics Water 5-75

Gas Oil Water 25-70

Gases Water 2-50

Gases Gases 2-25

Condensing Steam Water 200-700

Condensing Steam Light Organics 100-200

Condensing Steam Medium Organics 50-100

Condensing Steam Heavy Organics 6-60

Condensing Steam Propane (Boiling) 200-300

Steam Gases 5-50

Light Organics Light Organics 40-75

Medium Organics Medium Organics 20-60

Heavy Organics Heavy Organics 10-40

Crude Oil Gas Oil 80-90

Crude Oil Gasoline (Condensing) 20-30

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2. Heat Exchanger Velocities Recommended shellside and tubeside liquid velocities for various tube materials are summarized as follows:

Tube Material

Velocity (Ft/Sec)

Admiralty, Carbon Steel 4 to 8 Copper, Brass (85 - 15) 2 to 4 Nickel, Copper-Nickel 5 to 10 Stainless Steel, Monel 6 to 12 Titanium 6 to 15

Permissible tubeside velocities for dry gases range from 50 to 150 feet/sec. The

recommended minimum shellside liquid velocities is 1.5 feet/sec. 3. Allowable Pressure Losses

Recommended maximum allowable shellside and tubeside pressure losses are 10-15 psi for plate-baffle exchangers. Allowable shellside pressure losses for RODbaffle should range from 4-8 psi.

4. Cooling Water Temperatures

Maximum cooling water and tube wall temperatures to minimize fouling deposition are 125 F and 145 F, respectively.

5. Mean Temperature Differences

Log Mean Temperature Difference (LMTD) correction factors (F) for single shellpass, multiple tubepass exchangers should be greater than 0.75 to avoid temperature approach problems.

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6. Recommended Fouling Factors Recommended Tubular Exchanger Manufacturers Association (TEMA) fouling factors are provided in Table 2 below.

TABLE 2 Fouling (HR-FT2-F/BTU)

Exchanger Service Less 125F Greater 125F

Cooling Tower Water 0.001 0.002

Brackish Water 0.002 0.003

Sea Water 0.0005 0.001

Boiler Feedwater 0.001

Condensate 0.0005

Steam 0.0005

Compressed Air 0.001

Natural Gas & LPG Gas 0.001 - 0.002

Acid Gases 0.002 - 0.003

Reformer Feed-Effluent Gas 0.0015

Hydrocracker Feed-Effluent Gas 0.002

HDS Feed-Effluent Gas 0.002

MEA and DEA Solutions 0.002

DEG and TEG Solutions 0.002

Heat Transfer Fluids 0.002

Propane and Butane 0.001

Gasoline 0.002

Kerosene, Naptha, & Light Distillates 0.002 - 0.003

Light Gas Oil 0.002 - 0.003

Heavy Gas Oil 0.003 - 0.005

Heavy Fuel Oil 0.005 - 0.007

Vacuum Tower Bottoms 0.010

Natural Gas Combustion Products 0.005 7. Recommended Materials of Construction

Tubes: Inhibited Admiralty tubes are strongly recommended for non-chromate containing, cooling water services where tubewall temperatures range from 145-450 F. Inhibited Admiralty tubes are also recommended for conventionally treated cooling water service for tubewall temperature between 165-450 F. Do not use admiralty or other copper bearing alloys when cooling tower water may become contaminated with ammonia or where cooper is incompatible with the process fluid. Carbon steel tubes are recommended for cooling water services where tubewall temperature is below 165 F. Low-chrome steel tubes are recommended for high-temperature, sulfur-bearing streams. Austenitic stainless steel alloys are recommended for low temperature services (below -150 F).

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Monel tubes are recommended for HF acid-containing streams above 160 F, while titanium tubes are recommended for brackish and sea water services. Welded, fully killed carbon steel (ASTM A-214) should be avoided in low pH water soluble hydrocarbons, furfural, phenol, sulfuric acid, amine service, HF alkylation, and in final overhead crude tower coolers. Seamless carbon steel tubes (A-179 or A-83) should be used where welded tubes are not permitted.

Baffles, Tie Rods, & Spacers should be constructed of minimum quality material compatible with tube and tubesheet material.

Tube sheets must be compatible with channel materials.

Shell and channels must be compatible with service conditions.

Direct question about material suitability should be directed to Engineering Materials and Services.

8. Heat Exchangers (General)

a. Heat Exchanger Area, A (ft2) = Q / (U x LMTD) Where: U = heat transfer coefficient

Q = Heat Duty, Btu/Hr LMTD = Log Mean Temperature Difference b. Log Mean Temperature Difference, LMTD = TL - TS 1n (TL / TS) Where: TL = Largest temperature difference (F) TS = Smallest temperature difference (F) c. Limit temperature approach in gas to gas exchanger to 20 F. d. For preliminary design for cooling water systems, use a cooling water temperature rise

of 15-20 F through the heat exchangers. In most cases, a process stream temperature approach of 10 F to the cold water to the exchanger is reasonable.

e. If flow through the exchanger is not countercurrent, hot fluid outlet temperature

should be greater than cold fluid outlet temperature. f. For exchangers of 200 ft2 and less, use modular fintubes. 9. Condensers

a. Be aware when condensing pure components such as propane that the limiting temperature occurs when the desuperheating stops and condensing starts.

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b. For water cooling, try to cool no further than a 10 F approach to the warm cooling water leaving the condenser. Use 4-8 ft/sec velocity for water through the tubes. Restrict cooling water return temperature to maximum of 125 F.

c. For water-cooled propane condenser design, generally use 10 F temperature rise on

cooling water through the exchanger.

10. Reboilers and Chillers a. Many failures occur because the pressure on the condensate return header is higher

than the low pressure steam at the reboiler. b. Limit the approach temperature of the gas to the refrigerant in gas chillers to 10 F.

Less than 10 F delta T requires excess exchanger surface area. c. Usually design reboilers for a conservative heat flux of 8,000-12,000 Btu/ft2 and

reduce pressure of steam to prevent film boiling. d. Submergence of Bundle - Level generally controlled at top of bundle. e. For thermosiphon and side reboiler designs for demethanizer columns, limit the

vaporization of the reboiler liquid stream to a maximum of about 35% by volume. Attempting to vaporize more fluid may result in problems with the thermosiphon flow.

11. Rough Calculation Methods For Plate Heat Exchangers

A rough method for calculating plate heat exchangers is presented below. With this method, the required heating surface can be estimated for water-to-water duties. Corrections can be made for liquids with physical properties other than those of water.

Note: Calculations must be performed with SI-Units

a. Determine the inlet and outlet temperatures for both fluids. b. Calculate the LMTD for full counter-current flow. c. Calculate the -value for the primary side:

1 = temperature change of primary fluid/LMTD d. Enter diagram at 1 and find the specific area required for the relevant pressure drop,

at a primary fluid flow of 1 kg/s. e. Calculate the area required for the actual flowrate A=(A per kg/s)*m1, where

m1=kg/s. f. Calculate secondary side pressure drop

dP2 = dP1 (m2/m)1.9

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Sample Calculation:

Hot side 140 kg/s water 80 - 45 C, 50 KPa Cold side 170 kg/s water 20-?

Outlet temperature on cold side=20+(14/17)*(80-45)=49 C

31 {80 45} 25

{4920} LMTD = (31-25)/[ln(31/25)]=27.8 C 1 = 35/27.8=1.26 HTU A per kg/s=1.4 m2 A for 140 kg/s=196 m2 dP2=50*(170/140)1.9=72 kPa

12. Brazed Aluminum Plate Heat Exchangers

a. For aluminum plate fin reboilers, methanol may tend to accumulate in the reboiler and eventually log off the exchanger limiting thermosiphon flow. It can usually be cleared if a drain is provided on the lower header of the exchanger.

b. Mercury occurring naturally in some natural gas steams is extremely corrosive to

aluminum heat exchangers used extensively in LNG plant processes. Plan to check for mercury in feed gas up front in any project.

c. For Aluminum Plate Fin Core-in-Shell evaporator heat exchange design use a 3-5

degree temperature approach to shell side evaporating fluid temperature. Use a maximum evaporation of 30% of the thermosiphon circulated fluid in the evaporator when preparing preliminary core specifications.

13. Air Fin Heat Exchangers

a. Design for 150 F outlet temperature if cooling water follows. b. Normally design for 40 F approach to inlet air temperature if no cooling water used

(20 F minimum). c. For preliminary estimates, assume four rows of tubes. Estimate power requirement at

3 HP/MMbtu/hr for fans for face velocity of air to the coil of 450-550 ft/minute. d. Limit tip speed of fans: < 9 feet in diameter to 12,000 ft/minute (FPM) > 9 feet in diameter to 11,000 ft/minute

e. Hot air recirculation can be a problem, especially in hot weather. Consider air

recirculation when locating air cooled exchanagers.

23

Locate coolers away from taller buildings or structures, especially downwind of the cooler.

Do not locate coolers downwind of other heat generating equipment, i.e., furnaces, boilers, etc.

Mount coolers high enough from the ground to avoid high inlet air approach velocities. Consider mounting them on pipe lanes or provide at least ½ fan diameter clearance between the ground and the plenum.

Locate large banks of coolers with the banks long axis perpendicular to the prevailing summer wind direction.

Do not mix forced and induced draft coolers in close proximity and do not locate coolers of different heights in close proximity.

14. Fired Heaters

a. Maximum recommended heat flux for a direct fired triethylene glycol regenerator in a TEG dehydration unit is 8000 Btu/ft2 of fire tube surface area. The recommended heat flux for maximum fire tube life is 6000 Btu/ft2.

b. For most process heaters, assume a thermal efficiency of 75-80% when calculating fuel

requirements. Where: % Thermal Efficiency = (Heat Transferred Heat Released)*100.

15. Organic Heat Transfer Fluids a. Fired heaters for organic heat transfer fluids are usually designed with average radiant

heat fluxes ranging from 5000-12,000 Btu/hr-sq ft. Actual allowable heat flux is usually limited by fluid maximum allowable film temperature. Film temperature is dependent on:

- Maximum fluid bulk temperature - Velocity of the fluid across the heat transfer surface - Uniformity of heat distribution in the furnace - Heat transfer properties of the heat transfer fluid.

b. If too high film temperature results, too much fluid is vaporized and the heat transfer

surface is blanketed with vapors. The heat transfer coefficient is rapidly reduced and dangerously high surface temperatures can develop resulting in severe fluid degradation and mechanical failure.

c. High surface temperatures may also cause the fluid to carbonize forming carbon scale on the heat transfer surface which may lead to over heating and tube metal failure.

24

d. All other things being equal, any organic heat transfer fluid degrades in proportion to its temperature. Operation at approximately 100 F below vendors maximum recommended bulk fluid operation temperature may extend the life of the fluid by ten times.

16. Cooling Towers

a. The evaporation rate on a cooling tower is dependent on the amount of water being cooled and the temperature differential. For each 10 F temperature drop across the tower, 1% of the recirculation rate is evaporated. In other words, 0.001 times the circulation rate in gpm times the temperature drop equals the evaporation rate in gpm.

b. Windage Losses for Cooling Towers:

Spray ponds 1.0-5.0% of circulation Atmospheric cooling towers 0.3-1.0% of circulation

Forced draft cooling towers 0.1-0.3% of circulation

Evaporation losses for cooling towers: Evaporation losses are usually 0.85-1.25% of the tower circulation rate. An evaporation loss of 1% of tower circulation per each 10 F temperature drop across the tower can be assumed for estimating purposes.

c. Cooling Water System Feasibility design:

Feasibility designs for cooling water systems may be completed by setting the water temperature rise across all exchangers, usually 15-20 F rise, (or at a 10 F approach to the process outlet temperature if the assumed rise results in a temperature cross for some exchanger), and setting the inlet water temperature to the exchangers to the site wet bulb temperature plus 8 F.

d. Cooling Water System Fluid Flow and Piping:

For preliminary sizing branch offs with different flowrates from the main header, the following rule of thumb equations may be used.

qi/di2 = Q/D2 Where: Q and qi are volumetric flowrates through the header and branch i; and D and di are the diameters of the header and branch i.

D = Summation di2 round to nearest standard size.

17. Insulation

a. For estimating insulation thickness: Thickness = {3 +[(T - 100)/100)]} / 2

Where Thickness = inches T = Process Temperature, F 18. NGL Expander Plants

25

a. For thermosiphon reboiler and side reboiler designs for demethanizer columns, limit the vaporization of the reboiled liquid stream to a maximum of about 35% by volume. Attempting to vaporize more fluid may result in problems with thermosiphon flow.

b. For aluminum plate fin reboilers, methanol may tend to accumulate in the reboiler and

eventually log off the exchanger limiting thermosiphon flow. It can usually be cleared if a drain is provided on the lower header of the exchanger.

19. Miscellaneous Plant Systems

a. Cooling Water Systems - For preliminary design for cooling water systems, use a cooling water temperature rise of 15-20 F through the heat exchangers. In most cases a process stream temperature approach of 10 F to the cold water to the exchanger is reasonable.

b. For Aluminum Plate Fin Core in Shell evaporator heat exchanger design use a 3-5

degree temperature approach to shell side evaporating fluid temperature. Use a maximum evaporation of 30% of the thermosiphon circulated fluid in the evaporator when preparing preliminary core specifications.

c. Wind Chill and Tw = 33-[(10.45+10V)(33-T)]/32 Heat Loss H = (10.45+10V - V)(33-T) Where: Tw = Wind chill temp., C T = actual temp., C V = wind speed, meters/sec. H = heat loss, kcal/m2-hr.

20. Method For Feasibility Study Sizing of Gas Plant Gas/Gas Shell and Tube Heat Exchanger a. From the process simulator output for the process, determine the required UA rate for

the gas/gas exchanger. A = UA = UA Assume U = 60 BTU/Hr Ft2 F U 60 Assume a 20-foot exchanger with 3/4" OD tubes on a 15/16" triangular pitch.

Go to a Tube Count Table and read the number of tubes required for the Area A and unit diameter and/or number of units.

You now have a feasibility estimate which includes: - Exchanger Area (Ft2) - Number of 3/4" tubes - Unit length, diameter, and number of units.

26

21. Physical Fan Laws a. The following relations are characteristic of fans operating in a given system with

constant air density: With constant fan size and varying fan speed: (1) Volume (CFM) varies directly as the fan speed. CFM2 / CFM1 = RPM2 / RPM1 (2) Static Pressure Varies Directly as the square of the RPM. SP2 / SP1 = (RPM2 / RPM1)2 (3) Horsepower absorbed by the fan varies directly as the fan speed cubed. HP2 / HP1 = (RPM2 / RPM1)3 b. With varying fan sizes at the same speed: (1) Volume (CFM) varies directly as the fan size cubed. CFM2 / CFM1 = (DIA2 / DIA1)3 (2) Static pressure varies directly as the fan size squared. SP2 / SP1 = (DIA2 / DIA1)2 (3) Horsepower absorbed by the fan varies directly as the fifth power of size. HP2 / HP1 = (DIA2 / DIA1)5 c. The following relations are characteristic when a fan or a given size delivers a constant

mass of air of varying density. (Density varies directly as absolute temperature and inversely as the atmospheric pressure.):

(1) Volume, fan speed, and total pressure vary inversely as the density. (2) Horsepower absorbed by the fan varies inversely as the square of the density. d. Fan horsepower varies directly as the product of the volume (ACFM) times the total

pressure (inches W.G.) divided by the constant 6370 times the total aerodynamic efficiency.

Actual Fan Horsepower = ACFM x Total Pressure 6370 x Total Aerodynamic Efficiency

27

FRACTIONATION

1. For optimum economy, in terms of minimum column diameter and maximum tray efficiency, distillation columns should be designed with 5-10 percent entertainment.

2. Practical limits for packed columns are:

a. 30 gpm/ft2 liquid rate maximum b. 3 gpm/ft2 liquid rate maximum

(These can be extended by proper distributor design.)

3. For a quick estimate of the number of trays for a given separation, double the minimum number of stages obtained from the Fenske equation:

Dmin = log[(D1/B1)/(Dh/Bh1)] log

Where: D1 = Mole fraction of light key in distillate Dh = Mole fraction of heavy key in distillate B1 = Mole fraction of light key in bottoms Bh = Mole fraction heavy key in bottoms = Relative volatility, light key/heavy key

The minimum reflux ratio can be calculated by the following equation, if there are no divided keys, i.e., no components whose volatilizes lie between the light and heavy keys. Multiply the minimum reflux ratio by 1.3 to correspond to double the minimum stages.

Rmin = D1 (1 - D1 ) ( - 1) F1 (1 - F1) Where: F1 = Mole fraction light key in feed D1= Mole fraction light key in distillate

If divided keys are present it is probably easier to run a rigorous simulation than to solve the short cut equations by hand.

4. Sizing reflux accumulators:

Generally accumulators are sized based on liquid holdup time from normal liquid level to low liquid level, where the Normal Liquid Level is approximately the center of the vessel. Holdup time of 8-10 minutes is generally acceptable, based on total overhead condensed. Low liquid level should be 8-12 inches above the bottom of the vessel or the height of the vortex breaker, if one is present. Generally reflux accumulators should be two feet or greater in diameter.

28

For a column with a partial condenser sufficient vapor space must be left above high level for vapor/liquid separation. This can be roughly estimated for light hydrocarbon systems as:

Area vapor, required (ft2) = (Vapor flow rate, ft3/sec)/0.7

This vapor cross sectional area should be added to the total cross sectional area of the vessel to obtain the diameter. Even if the condenser is total condenser, the high liquid level should be 6-8 inches below the top of the vessel.

5. Estimating column diameter (low pressure columns only): The active area of a trayed

column can be fairly accurately estimated. The downcomer areas can also be reasonably estimated. (Downcomer area should never be less than 5% of the total area of the column.) Therefore a rough estimate of the diameter can be readily obtained.

A = W/ (1.6 * Pv)

DC = Liquid rate (gpm) / 175

Tower area = A + 2 * DC

Where W = Vapor rate, lb/sec A = Active area of column, ft2

Pv = Vapor density, lb/ft3 DC = Downcomer area, ft2

6. Fractionation: Reflux to Feed Ratio 40 Tray = 0.55 mol/mol 30 Tray = 0.7 mol/mol

29

FLUID FLOW

1. Miscellaneous a. Absolute pressure of atmosphere at height 'H' above sea level:

P, psia = P1(1-0.00000687H) n = 5.256 where: P1 = pressure at sea level, psia

Density: W = W1(1-0.00000687H) n = 4.256

where: W1 = density of air at sea level

P, psia = exp [2.6876 - 0.00000368(H)] Where H is height, feet above sea level

b. Acoustic Velocity Va = 80.53(P/)

where: P = psia = density lb/ft3 Va = ft/sec

For perfect gas: Va = (gc. k x r x t/m)½ where: gc. = 32.17 k = Cp/Cv r = 1546 t = R m = mol wt.

c. Vortex Breaker

Vortex breaker is needed if flow is greater than 1.9 ft./sec.

2. NGL Expander Plants Flowrate through an expander can generally be controlled from 0-150 % of the design

value. Expanders adiabatic efficiency will generally be within four percentage points of the design between 75-125% of design flowrate (lb/hr).

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3. Preliminary Piping Sizing

For preliminary liquid piping sizing, the following table may be used as a guide:

Liquid Service

Pressure Drop (Psi / 100 FT)

Velocity (FT / Sec)

Pump Suction 0.4 2.0 – 4.0 Pump Discharge 1.5 – 3.0 7.0 – 10.0 To Reboiler ---- 3.0 – 5.0 Gravity Flow 0.4 ---- Water ---- 2.5 – 9.0** High Viscosity to 200 CP Pump Suction 0.5 – 1.0 0.25 – 0.5 Pump Discharge 10.0 – 1.0 1.0 – 1.5

** Water with high CO2, seawater, etc., requires lower maximum velocities, linings, or special material.

For preliminary vapor piping sizing, the following table may be used as a guide.

Vapor Service

Pressure Drop (PSI / 100 FT)

Velocity (FT / Sec)

Total Allowable Pressure Drop (PSI)

Tower Overhead 0.5 ---- 1.0 On Plot Gas 0.5 ---- ---- Comp. Suct 0.3 5 – 150 1.5 Comp. Disc 0.5 100 4.0 Steam ---- 100 4.0

4. Piping

a. Initial maximum fluid velocities for line sizing: Most liquids: 10 ft/s All vapors: 50 ft/s Raw sea water: 11.5 ft/s (CuNi piping) Gravity drains: 1.5 ft/s

Steam condensate: 1.0 ft/s

b. Fluid velocity for vapor and two-phase flows should not exceed the erosional velocity. Estimate erosional velocity from Ve 100/ where Ve = erosional velocity in ft/see and = fluid density in lb/ft³.

Limiting Velocities - Liquids: (Another Source)

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Normal limiting velocities (highest normal design velocities) in process lines are given by the following formula:

V = 100 ½

Where: v is limiting design velocity in ft/sec is density in lb/ft3 at system T&P

is denisty in lb/ft3 system T&P

Erosion velocity (velocity at which erosion of process line is expected) for clear liquids are given by the following formula:

ve = 150 ½

Where: ve is the erosion velocity in ft/sec and is density in lb/ft3 at system T&P

c. Compressible gases (i.e., hydrocarbons, air, steam) can be treated as incompressible when the pressure loss for the segment in question is less than 10% of the inlet pressure.

5. Maximum Operable Velocities When rating existing liquid and/or gas piping systems, it is sometimes desirable to

determine the limiting fluid velocity for the system. This may be estimated as follows: a. For liquids: 48

Limiting velocity vm = ½ For erosive or corrosive liquids = 0.5 x vm Where: vm = limiting velocity, ft/sec and = density, lb/ft3 at system T&P b. For gases:

Turbulent Flow Average limiting velocity vm = 148.7 (kZT/m)½ ( sonic velocity)

For erosive or corrosive gases = 0.5 x vm Where: vm = Average limiting velocity, ft/sec

k = Specific heat ratio Z = Compressibility factor T = Temperature, R m = Molecular weight lb/lb mol

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Note: Economic factors such as pressure loss usually necessitates operating at lower than limiting velocity.

6. Friction Factor - Project Life Piping friction factor increases with operating time. As a result, the piping head loss

nearly in: 1. 25 to 30 years for clean gases and light hydrocarbons. 2. 15 to 25 years for most middle distillates. 3. 10 to 15 years for residues.

While sizing pipelines and pumping facilities, this aspect should be duly considered.

7. Piping Noise Liquid velocities above 20-30 ft/sec can cause noise. As a rule, a velocity head less than 1.3 psi avoids excessive noise.

8. Allocating Pressure Drops to Control Valves

a. In a pumped circuit, the pressure drop allocated to the control valve should be 33% of all other friction losses in the system at pump rated flow (exclusive of the valve pressure drop itself) or 15 psi whichever is greater.

Valid for <750 gpm and <150 psi pump delta p

Valid for >300 gpm and 150-275 psi pump delta p

If outside these ranges, pressure drop allocated may be 25% of system dynamic losses at pump rated head, or 15 psi whichever is greater.

In both cases above, use no more than 90% of valves Cv .

b. Compressor discharge and suction lines The pressure drop allocated to a control valve in the suction or discharge line of a

centrifugal compressor should be 5% of the suction absolute pressure, or 50% of the system dynamic losses (exclusive of the control valve) at the compressor rated point, whichever is larger. Also, no more than 90% of the valve Cv should be used.

c. Pressure motivated systems In a system where tank pressure moves liquid from one vessel to another, the pressure

drop should be 10% of the lower terminal vessel pressure, or 50% of the system dynamic losses, whichever is greater.

The above rule also applies to vapor, but in addition, the assigned Delta P should not

exceed 42% of the upstream pressure to avoid critical flow problems through the valve.

d Steam and flashing water

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Valves in steam lines to turbines, reboilers, and process vessels, should be allocated 10% of the design absolute pressure of the system or 5 psi, whichever is greater. The valve should be sized for twice the normal flow rate since steam usage rates can vary widely, especially during start-up.

For valves handling a flashing mixture, the allocated pressure drop should be equal to

0.9 times the difference in between the absolute inlet pressure and the absolute saturation pressure if flowing temperature is more than 5 F below the saturation temperature. If less than 5 F below the saturation temperature, the pressure drop should not be greater than 0.06 times the absolute inlet pressure.

9. Two Phase Flow For the seven types of two phase flow patterns in pipes, some guidelines on liquid and

vapor superficial velocities (LSV and GSV respectively) which can be used to make initial predictions on the type of flow pattern are given below:

Horizontal Pipes: a. In DISPERSED FLOW PATTERN, nearly all the liquid is entrained as spray by

the gas. This occurs at GSV >200 ft/sec. b. In ANNULAR FLOW PATTERN, liquid forms a film around the inside wall of

pipe and gas flows at a high velocity as a central core. This occurs at GSV >20 ft/sec.

c. In BUBBLE FLOW PATTERN, bubbles of gas move along at about the same velocity as the liquid. This occurs at LSV of 5-15 ft/sec, and GSV of 1-10 ft/sec.

d. In STRATIFIED FLOW PATTERN, liquid flows along the bottom of the pipe and gas flows over the smooth liquid gas interface. This normally occurs for LSV <0.5 ft/sec, and GSV of 2-10 ft/sec.

e. In WAVE FLOW PATTERN, the interface is disturbed by waves moving in the direction of flow; otherwise it is similar to stratified flow pattern. This occurs for LSV <1 ft/sec and GSV of about 15 ft/sec.

f. In SLUG FLOW PATTERN, waves are picked up periodically in the gas stream and form a slug that moves at much greater velocity than average liquid velocity. Slugs can cause severe vibration due to impact on fitting such as return bends.

g. In PLUG FLOW PATTERN, alternate plugs of liquid and gas move along the pipe. This occurs at LSV <2 ft/sec and GSV <3 ft/sec.

34

Upflow Vertical Pipes: a. For dispersed flow, GSV >70 ft/sec. b. For annular flow, LSV <2 ft/sec. and GSV >30 ft/sec. c. For bubble flow, GSV <2 ft/sec.

d. Stratified flow does not occur. e. For wave flow, the SV's are unpredictable. f. For slug flow, GSV = 2-30 ft/sec. g. For plug flow, GSV = 2-30 ft/sec.

35

Compressors, Expanders & Pumps

1. Reciprocating Compressors

a. Limit compression ratio to a maximum discharge temperature of 300 F. This is generally less than a compression ratio of 3.7 per stage.

b. For reciprocating propane compressor calculations, add 10% to the final horsepower

calculated by conventional means and 10 F to the final discharge temperature for preliminary design.

c. A significant consumer of propane refrigeration horsepower can be leaking recycle

valves.

d. To avoid excessive vibration, the mass of the foundation must be approximately five times the mass of the unit.

2. Compressor Quickies

a. 1 lb-mole (ideal) gas occupies 379 scf. Thus mass flow in lb/min =

MMscfd*106*MW/(1440*379). b. The lowest discharge volume flow through a centrifugal compressor is 175 ACFM

(300 m3/hr). ACFM = MMscfd*106*14.7*T*z/(1400*P*520)(T, in R) To achieve reasonable compressor efficiency with a centrifugal compressor, the

suction afm needs to be above 1500-2000. Avoid applications at lower acfm.

36

c. Centrifugal compressor head Ha = (z*1545*Ts/MW)*(Y/Y-1)*(Rc

(Y-1/Y)-1) Where: Y = Cp/Cv Rc = Pd/Ps Ts = Suction T (in R) Maximum head per impeller is 10,000 ft. d. Discharge temperature Reciprocating Td = Ts*(Rc

(Y-1/Y)) Centrifugal Td = Ts*(Rc

(Y-1/Y*np)

np = Polytropic efficiency e. Maximum allowable discharge temperature for associated gas (i.e., gas from crude oil

wellhead separation) compression is 300 F (150 C). Recommended maximum discharge temperature for centrifugal compressor is 350 F.

Absolute maximum discharge temperature for centrifugal compressors is 400 F. High temperature seals are absolutely necessary to operate at the maximum temperature.

f. "Head" for a centrifugal compressor is really energy imported to the gas. "Feet" head

is in fact ft-lb force/ft-lb mass. SI expresses it as kiloJoule/kilogram. Conversion: 1000 ft-lbf/ft-lbm = 2.989 kJ/kg Other expressions such as "meters" head (converting feet into meters) are meaningless

as they do not take the gravitational constant into account. g. For centrifugal compressors, a 20% surge margin from the design operating point is

recommended, 30% is preferred. Absolute minimum acceptable is 10%.

3. Liquefied Natural Gas (LNG) Plants For centrifugal compressor calculations, assume a polytropic efficiency of 77 to 80% for LNG plant preliminary design.

4. Energy Conservation Natural Gas Engines

If compression ratio is increased from 8:1 to 10:1 a 5% reduction in fuel rate and a 9% increase in brake horsepower results (assumes 1000 Btu fuel & minimum octane of 115).

Octane Number: C1-120 iC4-97.6 C5-80.2

C2-100.7 nC4-89.1 C6-26.0 C3-98.1 iC5-61.9 C7-0.0

5. Fuel Consumption Reciprocating Engines - 1000 hp 7200 Btu/hr/bhp Turbines - Solar T4500 9800 Btu/hr/bhp

37

6. NGL Expander Plants

a. For expander-compressor preliminary designs, limit the adiabatic efficiency to 72% for the expander end and 65% for the compressor end. Vendors will claim higher efficiencies, but actual tests indicate the above efficiencies to be more representative in service. Use a 95% mechanical efficiency for transferring horsepower from the expander to the compressor.

b. For expanders, limit liquefaction of feed stream through expander to 15 to 18 weight % maximum. In excess of this, mechanical problems can be expected with expander.

7. Gas Processing - Simulation Guidelines a. Turboexpander maximum size approximately 8,000 hp (6 MW).

b. Turboexpander maximum pressure ratio 2.5:1. c. Temperature reduction on expansion (F/psi):

Turboexpander JT-valve Lean gas 0.06 0.03 Rich gas 0.1 0.05

8. Pump Sizing

Example: Size Lean Oil Pump

Suction Conditions = 145 psig @ 90 F Discharge = 485 psig Capacity = 700,000 GPD @ 60 F SP. GR. @ Pump Temp. = 0.805 GPM @ Pump Temp. = 700,000 0.815 = 492 GPM 1440 min/day 0.805

Head = P x 2.31 = (485-145)2.31 = 976 Ft. of Liquid SP. GR. 0.805

B.H.P. = (GPM)(Ft of Hd.)(SP. GR.)

(3960) x Eff. B.H.P. = (492)(976)(0.805) = 139.5 H.P.

(3960) (0.70)

Use 150 H.P. Electric Motor Drive. Check Curve of Pump Purchased for End of Curve H.P. Requirement.

Specify Pump for 492 GPM @ 976 Ft. Head.

Case to be good for 485# Disch. and shut in head of pump.

38

NPSH Not Critical.

H.P. x 33,000 = Ft. - LB (min.)

LB. H20 @ 62 F x 0.12 = Gal. H.P. x 33,000 x 0.12 = 3960 x H.P. = Ft. - Gal.(min) 9. Pumps

a. Suction Specific Speed:

When selecting centrifugal pumps, the suction specific speed (Nss) for the pump should be less than 11,000. Experience has indicated that pumps operating with suction specific speeds above 11,000 have a much higher failure frequency.

RPM x (GPM) 0.5 Nss = (NPSHR)0.75

Where: Nss = Suction Specific Speed GPM = Flow

Rate in gallons per minute NPSHR = Net positive suction head required for the pump in feet of fluid.

If the pump is double suctioned, divide flow rate by two.

b. Net Positive Suction Head: Be sure to check the size and configuration of the pump suction and piping before

increasing the pump speed as it affects the pump required Net Positive Suction Head (NPSHR) considerably. See the relation:

NPSHR2 / NPSHR1 = (n2 / n1)2

Where: n is pump speed in rpm and subscripts 1 & 2 indicate initial and final conditions.

c. Axial Compressors 1. Axial compressors cannot be used with side loads. 2. Maximum discharge pressure is around 300 psia. 3. The minimum suction actual cubic feet per minute is 60,000.

4. Axial compressors cannot operate over as large a range of flow as centrifugal compressors

5. The maximum recommended discharge temperature is similar to a centrifugal or about 350 F.

10. General

a. Maintenance costs for gas engine driven reciprocating compressor units is about 6 times those of gas turbine driven centrifugal compressor units. For 1993: Reciprocating = $30/HP Yr Centrifugal = $ 5/HP Yr

b. Gas Turbine Wasteheat Availability

39

About 1/3 of fuel gas BTU requirements can be considered recoverable as wasteheat from a gas turbine for quickly preliminary estimates.

40

Combustion

1. Flare

a. Flare tip velocities Open pipe flare: 0.5 MACH emergency flaring

0.2 MACH continuous flaring Proprietary flare tips: 0.5 - 0.7 MACH (max)

b. Limit flare piping velocities to 0.7 MACH in laterals and to 0.5 MACH in headers.

CHECK BACK PRESSURES!

c. Normally limit PSV back pressure to 10%.

d. Estimate final temperature of gas remaining inside vessel following depressurization to flare from mid-point between resulting temperatures calculated by isenthalphic ("JT") and isentropic ("turboexpander") expansion.

e. Flare, reference API RP 521.

2. Fired Heaters

For most process heaters, assume a thermal efficiency of 75 to 80% when calculating fuel requirements.

Where: % Thermal Efficiency = Heat Transferred x 100

Heat Released 3. Fuel Requirements

For determining fuel requirements for process equipment, always use the net heating value of the fuel rather than the gross heating value.

41

MISCELLANEOUS

1. Water and Steam Systems a. Approximate break point for steam pressure at which silica becomes a problem with

vaporization and deposition on turbine blades is at 500 psig. b. The evaporation rate on a cooling tower is dependent on the amount of water being

cooled and temperature differential. For each 10 F temperature drop across the tower, 1% of the recirculation rate is evaporated. In other words, 0.001 times the circulation rate in gpm times the temperature drop equals the evaporation rate in gpm.

2. Economics

a. Capex Ratio Exponents For Processing Plants and Ancillaries Same No. of Units COST 2=(SIZE OR CAPACITY 2/SIZE OR CAPACITY 1)0.5(COST 1)

Unit number change required for new capacity. USE 0.6 exponent Infrastructure (Camps, Warehouses, Maintenance facilities) USE 0.3 exponent

b. Capex Factors From Major Equipment Cost

Installed Cost Onshore 2.5 x (Major equipment cost) Offshore 5.0 x (Major equipment cost)

[excludes deck and jacket costs]

c. Annual Operating Costs [excludes fuel and depreciation] Onshore; 3% of Capital Cost Offshore; 5% of Capital Cost

d. Capex (Total) Remote Area LNG Plants

[One Train] $ 2x109 per 2x106 MTY [1990 BASIS] Note: Cost Reductions via technology offset regulatory increases. Use 5%/yr

esc. in general costs. MTY = Metric Tonnes/Year

3. Hyates a. Hydrates generally form at 50 to 60 F.

b. Expect 1 degree F depression for each % methanol in liquid drainage. (Methanol content may be estimated with a hydrometer.)

c. Hydrate Control Add a margin of 50% to calculate hydrate inhibitor injection rates. d. Typical hydrate inhibitor concentrations: MEG: 70 - 80 wt% at inlet and 60 wt% in solution outlet MeOH: 98 wt% at inlet and 90 wt% in solution outlet e. Glycol inhibitor loss estimate is 1 lb/MMscf plus 200 ppm(v) in liquid hydrocarbon.

42

f. MeOH will melt hydrates, already formed. MEG will not. 4. NGL Expander Plants

If the CO2 in the feed gas to the cryogenic plant is in excess of 0.25 mol %, be sure to check for CO2 solidification in both the liquid and vapor phases immediately downstream of the expander and in the top four stages of the demethanizer.

5. Miscellaneous Plant Systems

Instrument Air - As a preliminary estimate for instrument air requirements for feasibility study design, use 0.5 to 0.75 scfm per control instrument.

6. Wind Loadings Flat Surfaces P= 0.004 x V2 Cylindrical Members P = 0.6 x 0.004 x V2 Where: P = Pressure in PSF V = Velocity of wind MPH 7. Steam Leaks @ 100 psi

Orifice Steam Energy Loss / Month Size Wasted / Mo. Gas (scf) #2 Fuel Oil

1/2 835000 878947 7383 3/8 470000 494737 4155 1/4 210000 221053 1857 1/8 52500 552263 464 1/16 13200 13895 117 1/32 3400 3579 30

8. Composition of Air

Air composition N2 - 78.07 mol% 02 - 20.99 mol% Ar - 0.94 mol%

9. Storage Vessel Capacity

a. Vessel Capacity: Capacity (gallons)=(Diameter, ft)2÷2) x Length, inches.

10. Pipeline Volume: (Diameter of pipe, inches)2 = Barrels/1000 feet (3% high)

11. Pressure Vessels In general, the maximum operating pressure (MOP) of a process pressure vessel is established from the maximum internal or external pressure at which the vessel operates while fulfilling it's normal function.

In general, for vessels subject to internal pressure only, the maximum allowable working

pressure (MAWP) of a process vessel can be arrived at by adding the greatest of 10% or 10 to 25 psi to the maximum operating pressure.

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Economic L/D ratios for pressure vessels generally fall in the 2 to 5 range where L = shell seam length and D = inside diameter, both in feet.

12. NACE Requirements

"Material shall be selected to be resistant to SSC or the environment should be controlled if the gas being handled is at a total pressure of 65 psia or greater and if the partial pressure of H2S in the gas is greater than 0.05 psia."

13. Pressure Waves (e.g. water hammer) Magnitude of P.W. in lbs f/in2 psi PW = a.d.vd/144-g) a = velocity of sound in the fluid ft/sec d = density of fluid, lb mass/ft3

vd = velocity decrease, i.e. velocity before change less the velocity after change, Ft/sec g = conversion factor = standard acceleration due to gravity, 32.174 ft/sec2

14. Absolute Pressure of Atmosphere at Height 'H' feet above Sea Level P1 = P1 (1-0.00000687H)5.256

P

1 = pressure at sea level - psia

Density: w = we (1-0.00000687H)4.256

W1 = density of air at sea level

15. Boiler Horsepower: 1 hp = 33,480 Btu/hr.

16. Solar Radiation: approximately 300 Btu/hr-ft2

17. Combustion Air Required, Stoichiometric: 10 ft3 air/ ft3 natural gas 1300 ft3 air / gal of #2 fuel oil o/bustech/landon/wkfiles/rulebook.4th